RU2393146C2 - Optimised liquid-phase oxidation - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: invention relates to a method for liquid-phase catalytic oxidation of an aromatic compound and can be used for oxidation of aldehydes to acids, alkylaromatic compounds to alcohols, acids or diacids. The method involves feeding starting material to be oxidised into the reaction zone of an initial reactor, feeding the gaseous phase stream of the oxidising agent containing molecular oxygen, oxidation of the compound to be oxidised into the liquid phase of the reaction medium contained in an agitated reactor, where the said reaction medium has gas content of at least approximately 0.6 averaged on time and averaged on volume.

EFFECT: technology provides a highly efficient reaction at relatively low temperatures, where if the oxidised compound is para-xylene and the oxidation reaction product is crude terephthalic acid, such a product can be purified and extracted using technologies cheaper than those that could have been used if the product was obtained using a traditional high-temperature oxidation method.

30 cl, 58 dwg, 10 ex

Description

FIELD OF THE INVENTION

The present invention relates generally to a process for the liquid phase catalytic oxidation of an aromatic compound. One aspect of the present invention relates to the partial oxidation of a dialkyl aromatic compound (e.g., paraxylene) to produce a crude aromatic dicarboxylic acid (e.g., crude terephthalic acid), which can then be purified and separated. Another aspect of the present invention relates to an improved reactor-type bubble column which provides a more efficient and economical liquid phase oxidation process.

State of the art

Liquid phase oxidation reactions are used in many existing industrial processes. For example, liquid phase oxidation is currently used to oxidize aldehydes to acids (e.g. propionaldehyde to propionic acid), to oxidize cyclohexane to adipic acid and to oxidize alkyl aromatic compounds to alcohols, acids or diacids. A particularly important industrial method of oxidation in the latter category (oxidation of alkyl aromatic compounds) is a liquid-phase catalytic partial oxidation of paraxylene to terephthalic acid. Terephthalic acid is an important compound with many uses. The primary use of terephthalic acid is to use polyethylene terephthalate (PET) as starting materials. PET is a well-known plastic used in large quantities in the world for the manufacture of products such as bottles, fibers and packaging.

In a typical liquid phase oxidation process, including partial oxidation of paraxylene to terephthalic acid, a liquid phase feed stream and a gas phase oxidizer stream are introduced into the reactor and form a multiphase reaction medium in the reactor. The liquid phase feed stream introduced into the reactor contains at least one oxidizable organic compound (e.g., paraxylene), while the gas phase oxidizer stream contains molecular oxygen. At least a portion of the molecular oxygen introduced into the reactor as a gas is dissolved in the liquid phase of the reaction medium, making oxygen available for the liquid phase reaction. If the liquid phase of the multiphase reaction medium contains an insufficient concentration of molecular oxygen (for example, if certain parts of the reaction medium are "oxygen depleted"), undesirable side reactions can cause the formation of impurities and / or target reactions can slow down in speed. If the liquid phase of the reaction medium contains too little oxidizable compound, the reaction rate may be undesirably slow. In addition, if the liquid phase of the reaction medium contains an excess concentration of the oxidizable compound, additional undesirable side reactions can cause the formation of impurities.

Conventional liquid phase oxidation reactors are equipped with means for mixing the multiphase reaction medium contained therein. Stirring of the reaction medium is done to facilitate the dissolution of molecular oxygen in the liquid phase of the reaction medium, to maintain relatively uniform concentrations of dissolved oxygen in the liquid phase of the reaction medium, and to maintain relatively uniform concentrations of oxidizable organic compound in the liquid phase of the reaction medium.

Mixing of the reaction medium subjected to liquid phase oxidation is often achieved by mechanical means of mixing in containers, such as, for example, flow mixing reactors (CSTR). Although CSTRs can provide thorough mixing of the reaction medium, CSTRs have several disadvantages. For example, CSTRs have a relatively high capital cost due to their need for expensive engines, fluid tight suspensions and drive shafts and / or complex mixing mechanisms. In addition, the rotating and / or oscillating mechanical components of conventional CSTRs require regular maintenance. The labor and shutdown times associated with this service add to the cost of the CSTR. However, even with regular maintenance, the mechanical agitation systems used in the CSTR are prone to mechanical failure and may require replacement after relatively short periods of time.

Reactor bubble columns provide an attractive alternative to CSTR and other mechanically stirred oxidation reactors. Reactor bubble columns provide mixing of the reaction medium without requiring expensive and unreliable mechanical equipment. Reactor bubbler columns typically comprise a reaction zone elongated from bottom to top in which the reaction medium is contained. Stirring of the reaction medium in the reaction zone is ensured, first of all, by the natural emergence of gas bubbles ascending through the liquid phase of the reaction medium. This natural buoyancy mixing provided in bubble column reactors reduces capital and maintenance costs compared to mechanically stirred reactors. In addition, the absence of essentially movable mechanical parts, associated with bubble column reactors, provides an oxidation system that is less prone to mechanical failure than mechanically stirred reactors.

When the liquid phase partial oxidation of paraxylene is carried out in a conventional oxidation reactor (CSTR or bubble column), the product recovered from the reactor is typically a suspension containing crude terephthalic acid (CTA) and mother liquor. CTA contains relatively high levels of impurities (for example, 4-carboxybenzaldehyde, paratoluylic acid, fluorenones and other colored particles), which render it unsuitable as starting materials for making PET. Thus, the CTA obtained in conventional oxidation reactors is typically subjected to a purification process that converts the CTA to purified terephthalic acid (PTA) suitable for producing PET.

One typical purification method for converting CTA to PTA involves the following steps: (1) replacing the mother liquor of the CTA-containing suspension with water, (2) heating the CTA / water suspension to dissolve the CTA in water, (3) catalytic hydrogenation of the CTA / solution water to convert impurities into more desirable and / or readily separable compounds, (4) precipitating the resulting PTA from the hydrogenated solution through multiple crystallization steps, and (5) recovering crystallized PTA from the remaining liquids. Although effective, this type of conventional cleaning method can be very expensive. Individual factors contributing to the high cost of conventional CTA purification methods include, for example, the thermal energy needed to facilitate the dissolution of CTA in water, the catalyst required for hydrogenation, the hydrogen stream required for hydrogenation, yield losses caused by the hydrogenation of a portion of the terephthalic acid, and many containers required for multi-stage crystallization. Thus, it would be desirable to obtain a CTA product that could be purified without the need for dissolution, which is facilitated by heating in water, hydrogenation and / or multi-stage crystallization.

OBJECTS OF THE INVENTION

Therefore, the aim of the present invention is to provide a more efficient and economical reactor and method of liquid phase oxidation.

Another objective of the present invention is to provide a more efficient and economical reactor and process for the liquid phase catalytic partial oxidation of paraxylene to terephthalic acid.

Another objective of the present invention is the creation of a bubble column reactor type, which facilitates improved liquid-phase oxidation reactions while reducing the formation of impurities.

Another objective of the present invention is to provide a more efficient and economical system for producing pure terephthalic acid (PTA) by liquid phase oxidation of paraxylene, to obtain crude terephthalic acid (CTA), and subsequently purifying CTA to PTA.

In addition, it is an object of the present invention to provide a reactor type bubble column to oxidize paraxylene and produce a CTA product that can be purified without the need for dissolution, which is facilitated by heating the CTA in water, hydrogenating the dissolved CTA and / or multi-stage crystallization of hydrogenated PTA.

It should be noted that the scope of the present invention, as defined in the attached claims, is not limited to methods or devices that can realize all the goals listed above. Rather, the scope of the claimed invention may cover a variety of systems that do not achieve all of the above objectives or any of them. Additional objectives and advantages of the present invention will be readily apparent to those skilled in the art upon viewing the following detailed description and the annexed drawings.

SUMMARY OF THE INVENTION

One embodiment of the present invention relates to a method comprising oxidizing an oxidizable compound in a liquid phase of a reaction medium contained in a stirred reactor, wherein the reaction medium has a gas content of at least about 0.6, averaged over time and volume .

Another embodiment of the present invention relates to a method for producing terephthalic acid, comprising the following steps: (a) oxidizing paraxylene in the liquid phase of a three-phase reaction medium contained within a bubble column reactor type, with the formation of crude terephthalic acid, where the reaction medium contains a gas of at least about 0.6, averaged over time and averaged over volume; and (b) oxidizing at least a portion of the crude terephthalic acid in the secondary oxidation reactor, thereby producing cleaner terephthalic acid.

Brief Description of the Drawings

Preferred embodiments of the present invention are described in detail below with reference to the accompanying drawings.

Figure 1 is a side view of an oxidation reactor constructed in accordance with one embodiment of the present invention, in particular, illustrating the introduction of feed streams, an oxidizing agent and reflux into the reactor, the presence of a multiphase reaction medium in the reactor, and the extraction of gas and slurry from the top and lower parts of the reactor, respectively;

FIG. 2 is an enlarged cross-sectional side view of the bottom of the reactor type bubble column along line 2-2 of FIG. 3, in particular illustrating the location and configuration of the oxidizing agent distribution device used to introduce the oxidizing agent stream into the reactor;

FIG. 3 is a top view of the oxidizing agent distribution device of FIG. 2, in particular illustrating the holes for the oxidizing agent in the upper part of the oxidizing agent distribution;

FIG. 4 is a bottom view of the oxidizer distribution device of FIG. 2, in particular illustrating the holes for the oxidizer in the lower part of the oxidizer distribution device;

FIG. 5 is a side sectional view of the oxidizer distribution device along line 5-5 of FIG. 3, in particular illustrating the orientation of the oxidizer holes in the upper and lower parts of the oxidizer distribution device;

6 is an enlarged side view of the lower part of the bubble column reactor type, in particular, illustrating a system for introducing a stream of raw materials into the reactor in a variety of positions, separated from each other by a certain vertical distance;

FIG. 7 is a top sectional view taken along line 7-7 of FIG. 6, particularly illustrating how the feed system shown in FIG. 6 distributes the flow of feed materials within, in a preferred radial inlet (FZ) zone, and in more than one azimuthal quadrant (Q ₁ , Q ₂ , Q ₃ , Q ₄ );

Fig. 8 is a sectional top view similar to Fig. 7, but illustrating alternative means for releasing a feed stream in a reactor using bayonet (quick-detachable) tubes, each of which has a plurality of small inlets;

Fig. 9 is an isometric view of an alternative system for introducing a feed stream into the reaction zone in a plurality of positions separated by a certain vertical distance, without the need for multiple passes through the vessel wall, in particular illustrating that a feed distribution system can at least partially supported on an oxidizer distribution device;

Figure 10 is a side view of the distribution system of raw materials with one passage through the wall and the distribution device of the oxidizer, illustrated in figure 9;

11 is a top sectional view taken along line 11-11 of FIG. 10, further illustrating a one-pass feed distribution system supported on an oxidizer distribution device;

12 is an isometric view of an alternative oxidizing agent distribution device having all oxidizing holes located at the bottom of the annular element;

Fig.13 is a top view of an alternative oxidizing agent distribution device of Fig.12;

Fig. 14 is a bottom view of the alternative oxidizing agent distribution device of Fig. 12, in particular illustrating the arrangement of openings in the lower portion for introducing an oxidizing agent stream into the reaction zone;

Fig. 15 is a side sectional view of the oxidizer distribution device along line 15-15 of Fig. 13, in particular illustrating the orientation of the lower oxidant openings;

Fig is a side view of a bubble column reactor type equipped with an internal deaeration tank near the lower outlet of the reactor;

FIG. 17 is an enlarged cross-sectional side view of the bottom of the reactor type bubble column of FIG. 16 taken along line 17-17 of FIG. 18, in particular illustrating the configuration of an internal deaeration vessel located at the bottom outlet of the reactor type bubble column;

Fig. 18 is a sectional top view taken along line 18-18 of Fig. 16, in particular illustrating a vortex quencher located in a deaeration vessel;

Fig is a side view of a bubble column reactor equipped with an external deaeration tank, illustrating the method in which part of the deaerated suspension leaving the lower part of the deaeration tank can be used to clean the discharge line connected to the lower part of the reactor;

FIG. 20 is a side view of a reactor type bubble column equipped with a hybrid internal / external deaeration vessel for separating a gas phase of a reaction medium extracted from an elevated lateral position in a reactor; FIG.

Fig is a side view of a bubble column reactor type equipped with an alternative hybrid deaeration tank near the bottom of the reactor;

Fig is an enlarged side view in section of the lower part of the bubble column reactor type in Fig.21, in particular, illustrating the use of an alternative oxidizer distribution device using inlet passages that receive the flow of oxidant through the lower head of the reactor;

FIG. 23 is an enlarged cross-sectional side view similar to FIG. 22, in particular, illustrating alternative means for introducing an oxidizer stream into the reactor through a plurality of openings at the bottom of the reactor, optionally using fenders plates to more uniformly distribute the oxidizer stream in the reactor;

Fig is a side view of a bubble column reactor using an internal flow passage to help improve dispersion of the oxidizable compound by recycling a portion of the reaction medium from the top of the reactor to the bottom of the reactor;

Fig is a side view of a bubble column reactor using an external flow passage to help improve dispersion of the oxidizable compound by recycling a portion of the reaction medium from the top of the reactor to the bottom of the reactor;

FIG. 26 is a side cross-sectional view of a horizontal ejector that can be used to improve dispersion of an oxidizable compound in an oxidation reactor, in particular, illustrating an ejector that uses incoming liquid feed to pump the reaction medium into the ejector and releases a mixture of the starting materials and the reaction medium in the reaction zone at high speed;

FIG. 27 is a side cross-sectional view of a vertical ejector that can be used to improve dispersion of an oxidizable compound in an oxidation reactor, in particular illustrating an ejector that combines liquid feed materials and an inlet gas and uses a combined two-phase fluid to pump the reaction medium into an ejector and releasing a mixture of liquid starting materials, a feed gas and a reaction medium in the reaction zone at high speed;

Fig. 28 is a side view of a bubble column reactor containing a multiphase reaction medium, in particular illustrating a reaction medium that is theoretically distributed over 30 horizontal sections of equal volume to quantify the gradients of certain parameters in the reaction medium;

Fig is a side view of a bubble column reactor containing a multiphase reaction medium, in particular, illustrating the first and second discrete 20 percent continuous volumes of the reaction medium, which have essentially different oxygen concentrations and / or rates of its consumption;

FIG. 30 is a side view of two reaction tanks arranged one above the other, with or without mechanical stirring, containing a multiphase reaction medium, in particular illustrating that the containers contain discrete 20% continuous volumes of the reaction medium having substantially different oxygen concentration and / or its consumption rate;

FIG. 31 is a side view of three reaction containers located side by side, with or without mechanical stirring, containing a multiphase reaction medium, in particular illustrating that the containers contain discrete 20% continuous volumes of the reaction medium having substantially different oxygen concentration and / or its consumption rate;

Fig is a side view of a bubble column reactor type with a step change in speed, having a wide lower reaction zone and a narrow upper reaction zone;

Fig. 33 is a side view of a bubble column reactor equipped with a vertically spaced separation wall to add a vertically spaced surface area that comes into contact with the reaction medium;

Fig. 34 is a sectional view taken along line 34-34 in Fig. 33, in particular illustrating that the dividing wall is a planar element dividing the reaction zone into two substantially equal sections;

Fig. 35 is a side view of a bubble column reactor equipped with a shortened vertically arranged dividing wall to add a vertically located surface area that comes into contact with the reaction medium;

Fig. 36 is a side view of a bubble column reactor equipped with a shortened and curved vertically arranged separation wall to add a vertically located surface area that comes into contact with the reaction medium;

Fig. 37 is a sectional view taken along line 37-37 of Fig. 36, in particular, illustrating that the curved vertically dividing wall as a whole is an S-shaped element dividing a portion of the reaction zone into two substantially equal sections ;

Fig. 38 is a side view of a bubble column reactor equipped with a shortened vertically disposed inner member to add a vertically disposed surface area that comes into contact with the reaction medium;

Fig. 39 is a sectional view taken along line 39-39 of Fig. 38, in particular, illustrating that the vertically disposed inner element is in the form of an "X" and that the edges of the inner element do not extend everywhere to the side wall of the reactor;

Fig. 40 is a side view of a bubble column reactor equipped with intermittent, variously configured, vertically disposed internal elements to add a vertically disposed surface area that comes into contact with the reaction medium;

Fig. 41 is a sectional view taken along line 41-41 in Fig. 40, in particular illustrating one of the configurations of vertically arranged elements that are in the form of an "X" and divide a portion of the reaction zone into four substantially equal quadrants;

Fig. 42 is a sectional view taken along line 42-42 of Fig. 40, in particular illustrating another configuration of vertically arranged elements that divide a portion of the reaction zone into eight substantially equal wedge-shaped sections;

Fig. 43 is a side view of a bubble column reactor equipped with a plurality of screw-shaped internal elements to add a vertically arranged surface area that comes into contact with the reaction medium;

Fig. 44 is a sectional view taken along line 44-44 in Fig. 43, in particular illustrating the shape of one of the screw-shaped internal elements;

Fig. 45 is a side view of a bubble column reactor equipped with a plurality of deflectors, each of which comprises a plurality of cylindrical rods for coming into contact with the reaction medium;

Fig. 46 is an enlarged isometric view of the deflectors in Fig. 45, in particular illustrating the manner in which the cylindrical rods of adjacent deflectors are rotated 90 ° with respect to each other;

Fig. 47 is a sectional view taken along line 47-47 of Fig. 45, in particular illustrating one of the deflectors individually;

Fig. 48 is a side view of a bubble column reactor equipped with a plurality of deflectors, each of which contains a plurality of L-shaped elements for coming into contact with the reaction medium;

Fig. 49 is an enlarged side view of the deflectors in Fig. 48, in particular illustrating the way in which the L-shaped elements of adjacent deflectors are rotated 90 ° with respect to each other;

Fig. 50 is a sectional view taken along line 50-50 of Fig. 48, in particular illustrating one of the deflectors individually;

Fig is a side view of a bubble column reactor type, equipped with a separate monolithic cylindrical baffle rhombic shape for contact with the reaction medium;

Fig.52 is an enlarged side view of the monolithic baffle in Fig.51;

Fig. 53 is a sectional view taken along line 53-53 of Fig. 51, illustrating the cylindrical shape of a monolithic deflector;

FIG. 54A and 54B are enlarged views of crude terephthalic acid (CTA) particles obtained in accordance with one embodiment of the present invention, in particular illustrating that each CTA particle is a low density, large surface area particle consisting of a plurality of fragile related subparticles of CTA;

55A and 55B are enlarged views of a typically produced CTA, in particular illustrating that a conventional CTA particle has a larger particle size, lower density and lower surface area than the CTA particles of the present invention in FIG. 54A and 54B;

56 is a simplified flowchart of a method known from the literature for producing purified terephthalic acid (PTA);

Fig. 57 is a simplified flowchart of a method for producing a PTA in accordance with one embodiment of the present invention; and

Fig. 58 is a table summarizing various operating parameters of an oxidation reactor type bubble column, in which some operating parameters are set in accordance with the description given in the Examples section.

Detailed description

One embodiment of the present invention relates to a liquid phase partial oxidation of an oxidizable compound. Such oxidation is preferably carried out in the liquid phase of a multiphase reaction medium contained in one or more stirred reactors. Suitable stirred reactors include, for example, bubbling stirred reactors (e.g. reactor type bubble columns), mechanically stirred reactors (e.g. flow mixing reactors) and stream stirred reactors (e.g. jet reactors). In one of the embodiments of the present invention, the liquid phase oxidation is carried out in a single bubble column reactor type.

As used here, the term "reactor-type bubbler column" is intended to mean a reactor to facilitate chemical reactions in a multiphase reaction medium, where mixing of the reaction medium is ensured primarily by moving gas bubbles from the bottom up through the reaction medium. As used here, the term "mixing" should mean work dissipated in the reaction medium, causing a flow of fluid and / or mixing. As used here, the terms "majority", "mainly" and "predominantly" should mean more than 50%. As used here, the term "mechanical mixing" should mean the mixing of the reaction medium caused by the physical movement of a rigid or flexible element (s) with respect to the reaction medium or inside it. For example, mechanical mixing can be achieved by rotating, oscillating and / or vibrating internal mixers, vanes, vibrators or acoustic diaphragms located in the reaction medium. As used here, the term "flow mixing" should mean the mixing of the reaction medium caused by high-speed injection and / or recirculation of one or more fluids into the reaction medium. For example, flow mixing may be provided by nozzles, ejectors, and / or ejection devices.

In a preferred embodiment of the present invention, less than about 40% of the stirring of the reaction medium in a bubble column reactor during oxidation is provided by mechanical stirring and / or flow stirring, more preferably less than about 20% of the stirring is provided by mechanical stirring and / or mixing with a stream, and most preferably, less than 5% of the mixing is provided by mechanical mixing and / or Ivanov stream. Preferably, the amount of mechanical stirring and / or stream stirring imparted to the multiphase reaction medium during oxidation is less than about 3 kW / m ^{3 of the} reaction medium, more preferably less than about 2 kW / m ³ , and most preferably less than 1 kW / m ³ .

Turning now to FIG. 1, a preferred reactor type bubble column 20 is illustrated as comprising a housing 22 having a reaction section 24 and a separation section 26. The reaction section 24 defines the internal reaction zone 28, while the separation section 26 defines the internal zone of the separation 30. Basically, a liquid phase feed stream is introduced into the reaction zone 28 through feed inlets 32a, b, c, d. Basically, a gas phase oxidant stream is introduced into the reaction zone 28 through an oxidizer distribution device 34 located at the bottom of the reaction zone 28. The liquid phase feed stream and the gas phase oxidizer stream together form a multiphase reaction medium 36 within the reaction zone 28. The multiphase reaction medium 36 contains a liquid phase and gas phase. More preferably, the multiphase reaction medium 36 comprises a three-phase medium having solid phase, liquid phase and gas phase components. The solid-phase component of the reaction medium 36 is preferably deposited within the reaction zone 28 as a result of the oxidation reaction carried out in the liquid phase of the reaction medium 36. The bubble column reactor 20 contains a slurry outlet 38 located near the bottom of the reaction zone 28 and a gas outlet 40 located near the upper part of the separation zone 30. The effluent stream of the suspension containing the liquid-phase and solid-phase components of the reaction medium 36 is removed from the reaction zone 28 through the outlet for the suspension 38, While predominantly gaseous effluent is withdrawn from separation zone 30 via gas outlet 40.

The liquid phase feed stream introduced into the reactor bubbler 20 through feed ports 32a, b, c, d preferably contains an oxidizable compound, a solvent, and a catalyst system.

The oxidizable compound present in the liquid phase feed stream preferably contains at least one hydrocarbon group. More preferably, the oxidizable compound is an aromatic compound. Even more preferably, the oxidizable compound is an aromatic compound with at least one attached hydrocarbon group, or at least one attached substituted hydrocarbon group, or at least one attached heteroatom, or at least one attached carboxylic acid functional group (-COOH). Even more preferably, the oxidizable compound is an aromatic compound with at least one attached hydrocarbon group or at least one attached substituted hydrocarbon group, with each attached group containing from 1 to 5 carbon atoms. Even more preferably, the oxidizable compound is an aromatic compound having exactly two attached groups, with each attached group containing exactly one carbon atom and consisting of methyl groups and / or substituted methyl groups, and / or at most one carboxylic acid group. Even more preferably, the oxidizable compound is paraxylene, metaxylene, paratolualdehyde, metatolualdehyde, paratoluylic acid, metatoluyl acid and / or acetaldehyde. Most preferably, the oxidizable compound is paraxylene.

A “hydrocarbon group” as defined herein is at least one carbon atom that is bonded only to hydrogen atoms or to other carbon atoms. A “substituted hydrocarbon group,” as defined herein, is at least one carbon atom bonded to at least one heteroatom and at least one hydrogen atom. “Heteroatoms,” as defined herein, are all atoms other than carbon and hydrogen atoms. Aromatic compounds, as defined here, contain an aromatic ring, preferably having at least 6 carbon atoms, even more preferably having only carbon atoms as part of the ring. Suitable examples of such aromatic rings include, but are not limited to, benzene, biphenyl, terphenyl, naphthalene, and other carbon-based fused aromatic rings.

Suitable examples of an oxidizable compound include aliphatic hydrocarbons (eg, alkanes, branched alkanes, cyclic alkanes, aliphatic alkenes, branched alkenes, and cyclic alkenes); aliphatic aldehydes (e.g. acetaldehyde, propionaldehyde, isobutyraldehyde and n-butyraldehyde); aliphatic alcohols (e.g. ethanol, isopropanol, n-propanol, n-butanol and isobutanol); aliphatic ketones (e.g. dimethyl ketone, ethyl methyl ketone, diethyl ketone and isopropyl methyl ketone); aliphatic esters (e.g. methyl formate, methyl acetate, ethyl acetate); aliphatic peroxides, peracids and hydroperoxides (for example, tert-butyl hydroperoxide, peracetic acid and di-tert-butyl hydroperoxide); aliphatic compounds with groups that are combinations of the above aliphatic species, plus other heteroatoms (for example, aliphatic compounds containing one or more molecular segments of hydrocarbons, aldehydes, alcohols, ketones, esters, peroxides, peracids and / or hydroperoxides in combination with sodium, bromine, cobalt, manganese and zirconium); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups with one or more attached hydrocarbon groups (for example, toluene, ethylbenzene, isopropylbenzene, n-propylbenzene, neopentylbenzene, paraxylene, methaxylene, orthoxylene, all isomers of trimethylbenzenes, all pentamethylbenzene, hexamethylbenzene, all isomers of ethyl methylbenzenes, all isomers of diethylbenzenes, all isomers of ethyl dimethylbenzenes, all isomers of dimethylnaphthalenes, all isomers of ethylmethylnaphthalenes, all isomers of diethylnaphthalenes, all isomers of dimethylbiphenyls, all isomers of ethylmethylbiphenyls and all isomers of diethylbiphenyls, stilbene, and it is with one or more attached hydrocarbon groups, fluorene, and it is with one or more attached hydrocarbon groups, anthracene, and it is also with one or several attached hydrocarbon groups, and diphenylethane, and it is with one or more attached hydrocarbon groups); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups with one or more attached hydrocarbon groups and / or with one or more attached heteroatoms that can combine with other atoms or groups of atoms (e.g. phenol, all methyl phenol isomers, all isomers of dimethylphenols, all isomers of naphthols, simple benzyl methyl ether, all isomers of bromophenols, bromobenzene, all isomers of bromotoluene, including alpha-bromtoluene, dibrombe nzol, cobalt naphthenate, and all isomers of bromobiphenyls); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups with one or more attached hydrocarbon groups, and / or one or more attached heteroatoms, and / or with one or more attached substituted hydrocarbon groups (for example, benzaldehyde, all isomers of bromobenzaldehyde , all isomers of brominated tolualdehydes, including all isomers of alpha-bromotolualdehydes, all isomers of hydroxybenzaldehydes, all isomers of bromohydroxy nzaldegidov, all isomers benzenedicarboxaldehyde, all isomers benzoltrikarboksoaldegidov, paratolualdegidom, metatolualdegidom, ortotolualdegidom, all isomers toluoldikarboksaldegidov, all isomers toluoltrikarboksaldegidov, all isomers toluoltetrakarboksaldegidov, all isomers dimetilbenzoldikarboksaldegidov, all isomers of dimethylbenzene trikarboksaldegidov, all isomers of dimethylbenzene tetrakarboksaldegidov, all isomers of trimethylbenzene trikarboksaldegidov, all isomers of ethyl olualdegidov, all isomers of trimethylbenzene dicarboxaldehyde, tetramethylbenzene dicarboxaldehyde, hydroxymethylbenzene, gidroksimetiltoluolov all isomers, all isomers gidroksimetilbromtoluolov, gidroksimetiltolualdegidov all isomers, all isomers gidroksimetilbromtolualdegidov, benzyl hydroperoxide, benzoyl hydroperoxide, all isomers tolilmetilgidroperoksidov and all isomers of methylphenol methyl hydroperoxide); various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups with one or more attached selected groups, selected groups mean hydrocarbon groups and / or attached heteroatoms, and / or substituted hydrocarbon groups, and / or carboxylic acid groups, and / or groups of peroxyacids (e.g., benzoic acid, paratoluic acid, metatoluilic acid, orthotoluic acid, all isomers of ethylbenzoic acids, all isomers of propylbenzoic acids, all isomers of butyl benzoic acids, all isomers of pentylbenzoic acids, all isomers of dimethylbenzoic acids, all isomers of ethylmethylbenzoic acids, all isomers of trimethylbenzoic acids, all isomers of tetramethylbenzoic acids, all isomers of pentamethylbenzoic acids, all isomers of benzene isomers methylbenzene dicarboxylic acids, all isomers of dimethylbenzene dicarboxylic acids, all isomers of methyl benzene tricarboxylic acids, all isome of bromobenzoic acid, all isomers of dibromobenzoic acids, all isomers of bromotoluic acids, including alpha bromotoluic acids, tolyl acetic acid, all isomers of hydroxybenzoic acids, all isomers of hydroxymethylbenzoic acids, all isomers of hydroxymethylbenzoic acids, all hydroxybromobenzoic acids, all isomers of hydroxybromotoluenoic acids, all isomers of hydroxymethyl bromobenzoic acids, all and Zomer carboxybenzaldehyde, all isomers dikarboksibenzaldegidov, perbenzoic acid, all isomers gidroperoksimetilbenzoynyh acids, all isomers gidroperoksimetilgidroksibenzoynyh acids, all isomers gidroperoksikarbonilbenzoynyh acids, all isomers gidroperoksikarboniltoluolov, all isomers metilbifenilkarbonovyh acids, all isomers dimetilbifenilkarbonovyh acids, all isomers metilbifenildikarbonovyh acids, all isomers bifeniltrikarbonovyh acids, all stilbene isomers with one or more and connected selected groups, all isomers of fluorenone with one or more attached selected groups, all isomers of naphthalene with one or more attached selected groups, benzyl, all isomers of benzyl with one or more attached selected groups, benzophenone, all isomers of benzophenone with one or more attached selected groups, anthraquinone, all isomers of anthraquinone with one or more attached selected groups, all isomers of diphenylethane with one sludge and several attached selected groups, benzocoumarin, and all benzocoumarin isomers with one or more attached selected groups).

If the oxidizable compound present in the liquid phase feed stream is a solid compound under normal conditions (i.e., solid at standard temperature and pressure), it is preferable that the oxidizable compound substantially dissolves in the solvent when introduced into the reaction zone 28 It is preferred that the boiling point of the oxidizable compound at atmospheric pressure is at least about 50 ° C. More preferably, the boiling point of the oxidizable compound is in the range of about 80 to about 400 ° C, and most preferably in the range of 125 to 155 ° C. The amount of oxidizable compound present in the liquid phase starting materials is preferably in the range of about 2 to about 40% by weight, more preferably in the range of about 4 to about 20% by weight, and most preferably in the range of 6 to 15% by weight.

Now note that the oxidizable compound present in the liquid-phase starting materials may contain a combination of two or more different oxidizable chemical compounds. These two or more different chemical materials may be introduced mixed in a liquid phase feed stream, or may be introduced separately in a plurality of feed streams. For example, an oxidizable compound containing paraxylene, metaxylene, paratolualdehyde, paratoluylic acid and acetaldehyde can be introduced into the reactor through one inlet or multiple separate inlets.

The solvent present in the liquid phase feed stream preferably contains an acid component and a water component. The solvent is preferably present in the liquid phase feed stream at a concentration in the range of about 60 to about 98% by weight, more preferably in the range of about 80 to about 96% by weight, and most preferably in the range of 85 to 94% by weight. The solvent acid component is preferably essentially an organic low molecular weight monocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the solvent acid component is mainly acetic acid. Preferably, the acid component is at least about 75% wt. solvent, more preferably at least about 80% wt. solvent, and most preferably from 85 to 98% wt. solvent, the remainder is mainly water. The solvent introduced into the bubbler column reactor type 20 may contain small amounts of impurities, such as, for example, paratolualdehyde, terephthaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, paratoluylic acid, paratoluyl aldehyde, alpha-isofromobaltol , phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended particles. It is preferred that the total amount of impurities in the solvent introduced into the bubble column reactor 20 is less than about 3% by weight.

The catalyst system present in the liquid phase feed stream is preferably a homogeneous liquid phase catalyst system that can facilitate the oxidation (including partial oxidation) of the oxidizable compound. More preferably, the catalyst system contains at least one transition metal of variable valency. Even more preferably, the transition metal of variable valency includes cobalt. Even more preferably, the catalyst system includes cobalt and bromine. Most preferably, the catalyst system includes cobalt, bromine and manganese.

When cobalt is present in the catalyst system, it is preferred that the amount of cobalt present in the liquid phase feed stream is such that the cobalt concentration in the liquid phase of reaction medium 36 is maintained in the range of about 300 to about 6,000 ppm. (ppm wt.), more preferably in the range of from about 700 to about 4200 ppm. wt., and most preferably in the range from 1200 to 3000 ppm. wt. When bromine is present in the catalyst system, it is preferred that the amount of bromine present in the liquid phase feed stream is such that the bromine concentration in the liquid phase of reaction medium 36 is maintained in the range of about 300 to about 5000 ppm. wt., more preferably in the range of from about 600 to about 4000 ppm. wt., and most preferably in the range from 900 to 3000 ppm. wt. When manganese is present in the catalyst system, it is preferred that the amount of manganese present in the liquid phase feed stream is such that the concentration of manganese in the liquid phase of reaction medium 36 is maintained in the range of about 20 to about 1000 ppm. wt., more preferably in the range of from about 40 to about 500 ppm. wt., most preferably in the range from 50 to 200 ppm wt.

The concentrations of cobalt, bromine and / or manganese in the liquid phase of the reaction medium 36 above are expressed as time averaged and volume averaged. As used here, the term "time averaged" should mean averaging of at least 10 measurements taken at regular intervals over a continuous period of operation of at least 100 seconds. As used here, the term "volume averaged" should mean averaging over at least 10 measurements carried out at uniform 3-dimensional gaps in a certain volume.

The mass ratio of cobalt to bromine (Co: Br) in the system of catalysts introduced into reaction zone 28 is preferably in the range of about 0.25: 1 to about 4: 1, more preferably in the range of about 0.5: 1 to about 3: 1, and most preferably in the range of 0.75: 1 to 2: 1. The mass ratio of cobalt to manganese (Co: Mn) in the system of catalysts introduced into reaction zone 28 is preferably in the range of about 0.3: 1 to about 40: 1, more preferably in the range of about 5: 1 to about 30: 1, and most preferably in the range from 10: 1 to 25: 1.

The liquid phase feed stream introduced into the bubble column reactor 20 may contain small amounts of impurities, such as, for example, toluene, ethylbenzene, paratolualdehyde, terephthaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, paratoluylic acid, paratoluidyl, alpha-bromo-paratoluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended particles. When a bubble column reactor type 20 is used to produce terephthalic acid, methaxylene and orthoxylene are also considered impurities. It is preferred that the total amount of impurities in the liquid phase feed stream introduced into the bubble column reactor type 20 is less than about 3% by weight.

Although FIG. 1 illustrates an embodiment where the oxidizable compound, solvent, and catalyst system are mixed together and introduced into the reactor type bubbler 20 as a single feed stream, in an alternative embodiment of the present invention, the oxidizable compound, solvent and catalyst may be in the reactor type bubbler. 20 entered separately. For example, it is possible to introduce a clean stream of paraxylene into the bubble column reactor 20 through an inlet separate from the inlet (s) for the solvent and catalysts.

Basically, the gas-phase oxidant stream introduced into the bubble column reactor 20 through the oxidizer distribution device 34 contains molecular oxygen (O ₂ ). Preferably, the oxidizing agent stream contains from about 5 to about 40 mol% of molecular oxygen, more preferably from about 15 to about 30 mol% of molecular oxygen, and most preferably from 18 to 24 mol% of molecular oxygen. It is preferred that the remainder of the oxidizing stream consist mainly of gas or gases, such as nitrogen, which are inert to oxidation. More preferably, the oxidizing agent stream consists essentially of molecular oxygen and nitrogen. Most preferably, the oxidizing agent stream is dry air that contains about 21 mol% of molecular oxygen and from about 78 to about 81 mol% of nitrogen. In an alternative embodiment of the present invention, the oxidizing agent stream may comprise substantially pure oxygen.

Referring again to FIG. 1, the reactor type bubbler column 20 is preferably provided with a reflux dispenser 42 located above the top surface 44 of the reaction medium 36. The reflux sparger 42 operates by introducing droplets of a substantially liquid phase reflux stream into the separation zone 30 using any means of droplet formation known in this area. More preferably, the reflux dispenser 42 produces a droplet spray directed downward towards the upper surface 44 of the reaction medium 36. Preferably, this downward directed spray of droplets has an effect (i.e., overlaps and affects) of at least about 50% of the maximum the horizontal cross-sectional area of the separation zone 30. More preferably, the droplet spray has an effect of at least about 75% of the maximum horizontal cross-sectional area of the separation zone 30. The most more preferably, the droplet spray has an effect of at least 90% of the maximum horizontal cross-sectional area of the separation zone 30. This downward-directed liquid phlegm spray can help prevent foaming on or above the top surface 44 of reaction medium 36 and can also help with separating any droplets of liquid or suspension trapped in the gas moving from the bottom up, which flows towards the outlet for the gas 40. In addition, liquid reflux can serve to reduce the number of particles and potentially precipitating compounds (for example, dissolved benzoic acid, paratoluylic acid, 4-CBA, terephthalic acid and metal catalytic salts) leaving the gas effluent recovered from the separation zone 30 through the gas outlet 40. In addition, the introduction of reflux drops into the separation zone 30 can, through distillation, be used to adjust the composition of the gas effluent recovered through the gas outlet 40.

The liquid reflux stream introduced into the bubble column reactor 20 through the reflux dispenser 42 preferably has approximately the same composition as the solvent component of the liquid phase feed stream introduced into the bubble column reactor 20 through the feed inlets 32a, b, c, d. Thus, it is preferred that the liquid reflux stream contains an acid component and water. The acid component of the reflux stream is preferably a low molecular weight organic monocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the reflux stream is acetic acid. Preferably, the acid component is at least about 75% wt. a reflux stream, more preferably at least about 80% wt. flow of reflux, and most preferably from 85 to 98% wt. reflux stream, with the remainder being water. Since the reflux stream generally has substantially the same composition as the solvent in the liquid phase feed stream, when this description refers to the “total amount of solvent” introduced into the reactor, such a “total amount of solvent” should include as part of the feed stream in the reflux stream, and part of them in the solvent.

During liquid phase oxidation in a bubble column reactor 20, it is preferable that the feed streams, oxidizing agent and reflux are substantially continuously introduced into the reaction zone 28, while the effluent gas and slurry streams are substantially continuously withdrawn from the reaction zone 28. As used here, the term "essentially continuous" should mean for a period of at least 10 hours with interruptions of less than 10 minutes. During oxidation, it is preferred that the oxidizable compound (e.g., paraxylene) is substantially continuously introduced into reaction zone 28 at a rate of at least about 8000 kg / h, more preferably at a rate in the range of from about 13,000 to about 80,000 kg / h, even more preferably in the range of about 18,000 to about 50,000 kg / h, and most preferably in the range of 22,000 to 30,000 kg / h. Although it is generally preferred that the flow rates for the feed streams of the starting materials, oxidizing agent and reflux are substantially stationary, it has now been noted that one embodiment of the present invention involves pulsating the feed stream of the starting materials, oxidizing agent and / or reflux for improved mixing and mass transfer. When the feed streams of the starting materials, oxidizing agent and / or reflux are introduced in a pulsed manner, it is preferred that their flow rates range from about 0 to about 500% of the steady-state flow rates mentioned here, more preferably from about 30 to about 200% of the stationary flow rates mentioned here, and most preferably in the range of 80 to 120% of the stationary flow rates mentioned here.

The space and time-averaged reaction rate (STR) in a bubble column of an oxidation type reactor 20 is defined as the mass of an oxidizable compound introduced per unit volume of the reaction medium 36 per unit time (e.g., kilograms of paraxylene introduced per cubic meter per hour). In normal use, the amount of oxidizable compound not converted to the product should, as a rule, be subtracted from the amount of oxidizable compound in the feed stream before calculating STR. However, the degrees of conversion and yields, as a rule, are high for many of the oxidizable compounds preferred here (for example, paraxylene), and the definition of the term as set forth above is common here. For reasons of capital and work costs, among other things, it is generally preferred that the reaction be carried out with high STR. However, the reaction at ever higher STRs may affect the quality or yield of partial oxidation. The reactor type bubbler column 20 is particularly suitable when the STR of an oxidizable compound (e.g., paraxylene) is in the range of about 25 kg / m ³ × h to about 400 kg / m ³ × h, more preferably in the range of about 30 kg / m ³ × h up to about 250 kg / m ³ × h, even more preferably from about 35 kg / m ³ × h to about 150 kg / m ³ × h, and most preferably in the range from 40 kg / m ³ × h up to 100 kg / m ³ × h

The oxygen STR in the bubble column of the oxidation type reactor 20 is defined as the mass of molecular oxygen consumed per unit volume of the reaction medium 36 per unit time (for example, kilograms of molecular oxygen consumed per cubic meter per hour). For reasons of capital cost and solvent consumption for oxidation, among other things, it is generally preferable that the reaction is carried out at large oxygen STRs. However, carrying out the reaction at increasingly high STRs in oxygen may reduce the quality or yield of the partial oxidation products. Without being limited by theory, it can be seen that this is possibly related to the rate of transfer of molecular oxygen from the gas phase to the liquid over the surface area of the interface, and, therefore, to the volume of the liquid. Too high STR in oxygen, possibly leads to too low dissolved oxygen in the volume of the liquid phase of the reaction medium.

The total oxygen averaged STR is defined here as the mass of all oxygen consumed in the entire volume of the reaction medium 36 per unit time (for example, kilograms of molecular oxygen consumed per cubic meter per hour). The reactor type bubbler column 20 is particularly suitable when the total oxygen averaged STR is in the range of about 25 kg / m ³ × h to about 400 kg / m ³ × h, more preferably in the range of about 30 kg / m ³ × h up to about 250 kg / m ³ × h, even more preferably from about 35 kg / m ³ × h to about 150 kg / m ³ × h, and most preferably in the range from 40 kg / m ³ × h to 100 kg / m ³ × h

During oxidation in a bubble column reactor 20, it is preferable that the ratio of the mass flow rate of the total amount of solvent (both from the feed stream and from the reflux stream) to the mass flow rate of the oxidizable compound entering reaction zone 28 is maintained within about from 2: 1 to about 50: 1, more preferably in the range of about 5: 1 to about 40: 1, and most preferably in the range of 7.5: 1 to 25: 1. Preferably, the ratio of the mass flow rate of the solvent introduced as part of the feed stream to the mass flow rate of the solvent introduced as part of the reflux stream is maintained in the range of about 0.5: 1 to the complete absence of the reflux stream, more preferably in the range of about 0 , 5: 1 to about 4: 1, even more preferably in the range of about 1: 1 to about 2: 1, and most preferably in the range of 1.25: 1 to 1.5: 1.

During liquid phase oxidation in a reactor type 20 bubbler column, it is preferable that the oxidizing stream is introduced into the reactor type 20 bubbler column in an amount that provides molecular oxygen slightly exceeding the stoichiometric oxygen demand. The amount of excess molecular oxygen needed for best results with a particular oxidizable compound affects the overall economic performance of liquid phase oxidation. During liquid phase oxidation in a bubble column reactor 20, it is preferable that the ratio of the mass flow rate for the oxidizing stream to the mass flow rate of the oxidizable organic compound (e.g., paraxylene) entering the reactor 20 is maintained in the range of about 0.5: 1 to about up to 20: 1, more preferably in the range of about 1: 1 to about 10: 1, and most preferably in the range of 2: 1 to 6: 1.

Referring again to FIG. 1, the feed streams, oxidizing agent and reflux introduced into the bubble column reactor 20 together form at least a portion of the multiphase reaction medium 36. The reaction medium 36 is preferably a three-phase medium containing a solid phase, a liquid phase and gas phase. As discussed above, the oxidation of an oxidizable compound (eg, paraxylene) takes place mainly in the liquid phase of the reaction medium 36. Thus, the liquid phase of the reaction medium 36 contains dissolved oxygen and an oxidizable compound. The exothermic nature of the oxidation reaction, which takes place in a bubble column reactor type 20, causes the boiling / evaporation of a portion of the solvent (e.g., acetic acid and water) introduced through the feed inlets 32a, b, c, d. Thus, the gas phase of the reaction medium 36 in the reactor 20 is formed mainly from the evaporated solvent and the undissolved, unreacted part of the oxidizing stream. Certain oxidation reactors known from the literature use heat exchange tubes / fins to heat or cool the reaction medium. However, such heat transfer structures may not be desirable in the reactor of the present invention and in the method described herein. Thus, it is preferable that the bubble column reactor type 20 essentially does not contain any surfaces in contact with the reaction medium 36, and a time-averaged heat flux greater than 30,000 W / m ² .

The concentration of dissolved oxygen in the liquid phase of the reaction medium 36 is the result of a dynamic equilibrium between the mass transfer rate from the gas phase and the reaction rate in the liquid phase (that is, it is not controlled simply by the partial pressure of molecular oxygen in the incoming gas phase, although it is one of the factors for the feed rate of dissolved oxygen, and it does not affect the maximum upper concentration of dissolved oxygen). The amount of dissolved oxygen varies locally, being higher near the bubble interface. In general, the amount of dissolved oxygen depends on the equilibrium of accumulation and flow factors in various areas of the reaction medium 36. Instantly, the amount of dissolved oxygen depends on the uniformity of gas and liquid mixing with respect to chemical flow rates. In the design, in order to suitably match the supply and flow rate in dissolved oxygen in the liquid phase of the reaction medium 36, it is preferable that the time-averaged and volume-averaged oxygen concentration in the liquid phase of the reaction medium 36 is maintained above about 1 ppm. mol., more preferably in the range of from about 4 to about 1000 ppm. mol., even more preferably in the range of from about 8 to about 500 ppm. mol., and most preferably in the range from 12 to 120 ppm. pier

The liquid phase oxidation reaction carried out in a bubble column reactor type 20 is preferably a precipitation reaction that forms solid products. More preferably, liquid phase oxidation carried out in a bubble column reactor type 20 causes at least about 10% wt. an oxidizable compound (eg, paraxylene) introduced into the reaction zone 28 to form a solid compound (eg, crude terephthalic acid particles) in the reaction medium 36. Even more preferably, the liquid phase oxidation causes at least about 50 wt. oxidizable compounds to form solid compounds in the reaction medium 36. Most preferably, liquid phase oxidation causes at least 90% wt. the oxidizable compound to form solid compounds in the reaction medium 36. It is preferred that the total amount of solid products in the reaction medium 36 is greater than about 3% by weight, time averaged and volume averaged. More preferably, the total amount of solid products in reaction medium 36 is maintained in the range of about 5 to about 40% by weight, even more preferably in the range of about 10 to about 35% by weight, and most preferably in the range of 15 to 30% by weight. . It is preferred that a substantial portion of the oxidation product (eg, terephthalic acid) obtained in the bubble column reactor type 20 is present in the reaction medium 36 as solid products, rather than remaining dissolved in the liquid phase of the reaction medium 36. The amount of solid phase oxidation product present in the reaction medium 36, preferably is at least about 25% wt. from the total amount of the oxidation product (solid and liquid phase) in the reaction medium 36, more preferably at least about 75% wt. from the total amount of oxidation product in the reaction medium 36, and most preferably at least 95% wt. of the total amount of oxidation product in the reaction medium 36. The numerical ranges given above for the amount of solid products in the reaction medium 36 relate to stationary, essentially continuous operation of the bubble column 20 for a substantially continuous period of time, but not to start-up, the shutdown or suboptimal mode of operation of the bubble column reactor type 20. The amount of solid products in the reaction medium 36 is determined gravimetrically. In this gravimetric method, a representative sample of the suspension is removed from the reaction medium and weighed. Under conditions that effectively support the overall distribution of solid products — the liquid present in the reaction medium, free liquid is removed from part of the solid products by sedimentation or filtration, effectively without loss of precipitated solid products and with less than about 10% of the initial liquid mass, remaining as solid products. The remaining liquid on solid products is evaporated to dryness, efficiently, without sublimation of solid products. The remainder of the solid products are weighed. The ratio of the mass of a part of the solid products to the mass of the initial part of the suspension is the proportion of solid products, usually expressed as a percentage.

The deposition reaction carried out in a bubble column reactor type 20 can cause the formation of a precipitate (that is, the deposition of solid products) on the surface of certain rigid structures in contact with the reaction medium 36. Thus, in one embodiment of the present invention, it is preferable that the bubble column reactor type 20, essentially, did not contain internal structures of heat transfer, mixing or baffles in the reaction zone 28, since such structures are prone to the formation of adka. If internal structures are present in the reaction zone 28, it is desirable to exclude internal structures having external surfaces that contain a significant area of the flat surface facing up, since such looking up flat surfaces would be highly prone to sedimentation. Thus, if any internal structures are present in the reaction zone 28, it is preferable that less than about 20% of the total surface area of the open outer surface facing upwards, such internal structures would be formed by essentially flat surfaces, tilted at an angle less than about 15 ° from the horizontal.

Referring again to FIG. 1, the physical configuration of the bubble column reactor 20 helps to provide optimized oxidation of the oxidizable compound (eg, paraxylene) with minimal formation of impurities. It is preferred that the elongated reaction section 24 of the housing 22 comprises a substantially cylindrical main body 46 and a bottom 48. The upper surface level of the reaction zone 28 is defined by a horizontal plane 50 extending along the upper part of the cylindrical main body 46. The lower surface level 52 of the reaction zone 28 determined by the lowest inner surface of the bottom 48. Typically, the lower level of the surface 52 of the reaction zone 28 is located near the opening 38 for the exit of the suspension. Thus, the elongated reaction zone 28 defined inside the bubble column reactor 20 has a maximum length "L", measured from the upper level of surface 50 to the lower level of surface 52 of reaction zone 28 along the longitudinal axis of the cylindrical main body 46. The length "L" of the reaction zone 28 is preferably in the range of about 10 to about 100 m, more preferably in the range of about 20 to about 75 m, and most preferably in the range of 25 to 50 m. The reaction zone 28 has a maximum diameter ( irin) "D", which is usually equal to the maximum inner diameter of the cylindrical main body 46. The maximum diameter "D" of the reaction zone 28 is preferably in the range of about 1 to about 12 m, more preferably in the range of about 2 to about 10 m, even more preferably in the range of about 3.1 to about 9 m, and most preferably in the range of 4 to 8 m. In a preferred embodiment of the present invention, the reaction zone 28 has a length to diameter "L: D" in the range of about from 6 : 1 to about 30: 1. Even more preferably, reaction zone 28 has an L: D ratio in the range of about 8: 1 to about 20: 1. Most preferably, reaction zone 28 has an L: D ratio of from 9: 1 to 15: 1.

As discussed above, the reaction zone 28 of the bubble column reactor 20 accepts a multiphase reaction medium 36. The reaction medium 36 has a lower surface level coinciding with the lower edge 52 of the reaction zone 28 and an upper surface level located on the upper surface 44. The upper surface 44 of the reaction medium 36 is determined along a horizontal plane that intersects the reaction zone 28 in a vertical position, where the contents of the reaction zone 28 passes from a continuous gaseous state to a continuous liquid-phase state. The upper surface 44 is preferably located in a vertical position, where the local time-averaged gas content of the thin horizontal portion of the contents of the reaction zone 28 is 0.9.

The reaction medium 36 has a maximum height "H", measured between its upper and lower surface levels. The maximum width “W” of the reaction medium 36 is typically equal to the maximum diameter “D” of the cylindrical main body 46. During liquid phase oxidation in a bubble column reactor type 20, it is preferred that H is maintained from about 60 to about 120% of L, more preferably from about 80 to about 110% of L, and most preferably from 85 to 100% of L. In a preferred embodiment of the present invention, the reaction medium 36 has a height to width "H: W" ratio greater than about 3: 1. More preferably, reaction medium 36 has an H: W ratio in the range of about 7: 1 to about 25: 1. Even more preferably, reaction medium 36 has an H: W ratio in the range of about 8: 1 to about 20: 1. Most preferably, reaction medium 36 has an H: W ratio in the range of 9: 1 to 15: 1. In one embodiment of the present invention, L = H and D = W, so that the various sizes or ratios given here for L and D apply also to H and W, and vice versa.

The relatively high L: D and H: W ratios provided in accordance with one embodiment of the present invention can contribute to several important advantages of the system of the present invention. As discussed in more detail below, it has been found that higher L: D and H: W ratios, as well as some other features discussed below, can contribute to creating advantageous gradients in the elevation of molecular oxygen and / or oxidizable concentrations (e.g., paraxylene) in reaction medium 36. In contrast to the generally accepted concept, which tends to a well-mixed reaction medium with relatively uniform concentrations in it, it was found that a vertical stepwise change in the concentration of acid ode and / or the oxidizable compound facilitates a more effective and economical oxidation reaction. Minimizing the concentrations of oxygen and the oxidizable compound near the top of the reaction medium 36 can help prevent the loss of unreacted oxygen and unreacted oxidizable compounds through the upper gas outlet 40. However, if the concentrations of the oxidizable compound and unreacted oxygen are low throughout the reaction medium 36, then the rate and / or oxidation selectivity is reduced. Thus, it is preferable that the concentrations of molecular oxygen and / or the oxidizable compound are significantly higher near the lower part of the reaction medium 36 than near the upper part of the reaction medium 36.

In addition, the high L: D and H: W ratios make the pressure in the lower part of the reaction medium 36 substantially higher than the pressure in the upper part of the reaction medium 36. This pressure gradient in height is the result of the high height and density of the reaction medium 36. One of the advantages of this gradient in pressure height is that the increased pressure in the lower part of the tank provides greater oxygen solubility and greater mass transfer than those that would otherwise be achieved at comparable temperatures x and pressures in the upper part of the column in smaller reactors. Thus, the oxidation reaction can be carried out at lower temperatures than would be required in smaller reactors. When a bubble column reactor type 20 is used to partially oxidize paraxylene to crude terephthalic acid (CTA), the ability to operate at lower reaction temperatures, at the same or better oxygen mass transfer rates, has several advantages. For example, the low temperature oxidation of paraxylene reduces the amount of solvent burned during the reaction. As discussed in more detail below, low temperature oxidation also favors the formation of small, high surface area, loosely bonded, readily soluble CTA particles that can be exposed to more economical cleaning technologies than large, low surface area, dense CTA particles produced by conventional high temperature oxidation processes.

During oxidation in the reactor 20, it is preferable that the time-averaged and volume-averaged temperature of the reaction medium 36 is maintained in the range of about 125 to about 200 ° C, more preferably in the range of about 140 to about 180 ° C, and most preferably ranges from 150 to 170 ° C. The pressure at the top of the column above the reaction medium 36 is preferably maintained in the range of about 1 to about 20 bar in the sensor (bar), more preferably in the range of about 2 to about 12 bar, and most preferably in the range of 4 to 8 bar. Preferably, the pressure difference between the upper part of the reaction medium 36 and the lower part of the reaction medium 36 is in the range of about 0.4 to about 5 bar, more preferably the pressure difference is in the range of about 0.7 to about 3 bar, and most preferably the difference pressure ranges from 1 to 2 bar. Although it is generally preferred that the pressure in the top of the column above the reaction medium 36 is maintained at a relatively constant value, one embodiment of the present invention provides a pulsed pressure mode in the upper part of the column to facilitate improved mixing and / or mass transfer in the reaction medium 36 When the pressure at the top of the column is pulsating, it is preferable that the pulsed pressure mode be between about 60 and about 140% of the steady state. the pressure at the top of the column mentioned here, more preferably between about 85 and about 115% of the stationary pressure at the top of the column mentioned here, and most preferably between 95 and 105% of the stationary pressure at the top of the column mentioned here.

An additional advantage of the high L: D ratio of reaction zone 28 is that this can contribute to an increase in the average surface velocity of the reaction medium 36. The terms “surface velocity” and “surface velocity of gas”, as used herein with respect to reaction medium 36, should denote the volumetric flow rate of the gas phase of the reaction medium 36 at a certain height in the reactor, divided by the horizontal cross-sectional area of the reactor at this height. The increased surface velocity provided by the high L: D ratio of reaction zone 28 can accelerate local mixing and increase the gas content in reaction medium 36. The time-averaged surface velocities of reaction medium 36 are at one fourth height, half height and / or three quarters of height reaction medium 36 is preferably greater than about 0.3 m / s, more preferably in the range of about 0.8 to about 5 m / s, even more preferably in the range of about 0.9 to about 4 m / s, and nai olee preferably in the range of 1 to 3 m / s.

Referring again to FIG. 1, the reactor-type bubble column separation section 26 is simply an expanded portion of the housing 22 located directly above the reaction section 24. The separation section 26 section reduces the speed of the bottom-up gas phase in the bubble-type column of the reactor type 20 when the gas phase rises above the upper level of surface 44 of reaction medium 36 and reaches an outlet for gas 40. This decrease in the bottom-up velocity of the gas phase helps to facilitate the removal of trapped liquid and / or solid 's products in the current bottom-up gas phase and thereby reduces undesirable loss of certain components present in the liquid phase of reaction medium 36.

The separation section 26 preferably comprises a generally diverging conical transition wall 54, a generally cylindrical wide side wall 56 and an upper head 58. A narrow lower surface level of the transition wall 54 is connected to the upper part of the cylindrical main body 46 of the reaction section 24. A wide upper level of the surface of the transition wall 54 is connected to the bottom of the wide side wall 56. It is preferable that the transition wall 54 extends up and out from its narrow lower edge at an angle within from about 10 to about 70 ° from the vertical, more preferably from about 15 to about 50 ° from the vertical, and most preferably from about 15 to 45 ° from the vertical. The wide side wall 56 has a maximum diameter "X", which is usually larger than the maximum diameter "D" of the reaction section 24, although when the upper part of the reaction section 24 has a diameter smaller than the maximum diameter of the reaction section 24 as a whole, then X may actually be less than D. In a preferred embodiment of the present invention, the ratio of the diameter of the wide side wall 56 to the maximum diameter of the reaction section 24 "X: D" is in the range from about 0.8: 1 to about 4: 1, most preferably in pre elah from 1.1: 1 to 2: 1. The top 58 is connected to the top of the wide side wall 56. The top 58 is preferably a generally elliptical cap element defining a central opening that allows the gas to escape from the separation zone 30 through the gas outlet 40. Alternatively, the top 58 may be of any shape including conical. The separation zone 30 has a maximum height "Y", measured from the upper part 50 of the reaction zone 28 to the uppermost part of the separation zone 30. The ratio of the length of the reaction zone 28 to the height of the separation zone 30 "L: Y" is preferably in the range from about 2: 1 up to about 24: 1, more preferably in the range of about 3: 1 to about 20: 1, and most preferably in the range of 4: 1 to 16: 1.

Turning to FIG. 1-5, the position and configuration of the oxidizer distribution device 34 will now be discussed in more detail. FIG. 2 and 3 show that the oxidizer distribution device 34 may comprise an annular element 60, a diametrical element 62, and a pair of passages 64a, b for accumulating the oxidizing agent. It is convenient that these passages 64a, b for accumulating the oxidizing agent can enter the vessel slightly higher than the annular element 60, and then turn down, as shown in FIG. 2 and 3. Alternatively, the oxidizer accumulation passage 64a, b may extend into the vessel below the annular element 60 or in approximately the same horizontal plane as the annular element 60. Each oxidizer accumulation passage 64a, b contains a first surface level connected to a corresponding oxidant inlet 66a, b, formed in the housing 22, and a second surface level connected to the fluid communication with the annular element 60. The annular element 60 is preferably formed from passages, more preferably from many straight x passes sections, and most preferably a plurality of straight pipe sections, rigidly connected to each other to form a tubular polygonal ring. Preferably, the annular element 60 is formed of at least 3 straight pipe sections, more preferably from 6-10 pipe sections, and most preferably from 8 pipe sections. Accordingly, when the annular element 60 is formed of 8 tube sections, it has a generally octagonal configuration. The diametral element 62 is preferably formed essentially from a straight pipe section, which is connected to the fluid communication with the opposite pipe sections of the annular element 60 and extends diagonally between them. The tube section used for the diametric element 62 preferably has substantially the same diameter as the tube sections used to form the ring element 60. It is preferred that the tube sections, which comprise the passages 64a, b, for the accumulation of oxidizing agent, ring element 60 and diametral element 62 had a nominal diameter greater than about 0.1 m, more preferably in the range of about 0.2 to about 2 m, and most preferably in the range of 0.25 to 1 m. As best ver clearly, illustrated in FIG. 3, the annular element 60 and the diametrical element 62 each provide a plurality of upper oxidant openings 68 for releasing the oxidant stream from bottom to top in the reaction zone 28. As best illustrated in FIG. 4, the annular element 60 and / or diametral element 62 may provide one or more lower oxidant openings 70 for releasing the oxidant stream down into the reaction zone 28. The lower oxidant openings 70 may also be used to release liquid and / or solid products that could enter ring element 60 and / or diametral element 62. To prevent solid products from precipitating inside oxidizer distribution device 34, fluid flow can continuously or periodically pass through distribution devices 34, washing off any accumulated solid products.

Referring again to FIG. 1-4, during oxidation in a bubble column reactor 20, oxidizer streams are pumped through oxidizer inlets 66a, b and into passageways 64a, b for oxidizer accumulation, respectively. Then, the oxidizer streams are transferred through passages 64a, b for accumulating the oxidizing agent in the ring element 60. When the oxidant stream enters the ring element 60, the oxidizer stream is distributed among the internal volumes of the ring element 60 and the diametrical element 62. Then, the oxidizer stream is forcibly removed from the oxidizer distribution device 34 and into the reaction zone 28 through the upper and lower holes 68, 70 for the oxidizing agent of the ring element 60 and the diametrical element 62.

The outputs of the upper oxidant openings 68 are laterally separated from each other and are located at substantially the same height in the reaction zone 28. Thus, the outputs of the upper oxidant openings 68 are generally located along a substantially horizontal plane defined by the upper part of the oxidizing agent distribution device 34. The outputs of the lower oxidizing holes 70 are laterally separated from each other and are located at substantially the same height in the reaction zone 28. Thus, the outlets of the bottom openings 70 for the oxidizing agent are generally arranged along a substantially horizontal plane defined by the lower part of the oxidizing agent distribution device 34.

In one embodiment of the present invention, the oxidizing agent distribution device 34 has at least about 20 upper oxidizing holes 68 formed therein. More preferably, the oxidizer distribution device 34 has in the range of about 40 to about 800 upper oxidant openings formed therein. Most preferably, the oxidizer distribution device 34 has from 60 to 400 upper oxidant openings 68 formed therein. The oxidizer distribution device 34 preferably has at least about 1 lower oxidant hole 70 formed therein. More preferably, the oxidizer distribution device 34 has in the range of about 2 to about 40 lower oxidant openings 70 formed therein. Most preferably, the oxidizer distribution device 34 has from 8 to 20 lower oxidant holes 70 formed therein. The ratio of the number of upper oxidant openings 68 to the lower oxidant openings 70 in the oxidizing agent distribution device 34 is preferably in the range of about 2: 1 to about 100: 1, more preferably in the range of about 5: 1 to about 25: 1, and most preferably in the range from 8: 1 to 15: 1. The diameters of essentially all of the upper and lower oxidizing holes 68, 70 are preferably substantially the same, so that the ratio of the volumetric flow rate of the oxidizing agent from the upper and lower holes 68, 70 is substantially the same as the ratio above, for the relative number of upper and lower holes 68, 70 for the oxidizing agent.

Figure 5 illustrates the direction of release of the oxidizing agent from the upper and lower oxidizing holes 68, 70. When mentioning the upper oxidizing holes 68, it is preferred that at least a portion of them are arranged to release the oxidizing stream at an angle “A”, that is, at an angle to the vertical. It is preferable that the percentage of upper oxidant openings 68 that deviate from the vertical by an angle “A” is in the range of about 30 to about 90%, more preferably in the range of about 50 to about 80%, even more preferably in the range from 60 to 75%, and most preferably, was about 67%. The angle “A” is preferably in the range of about 5 to about 60 °, more preferably in the range of about 10 to about 45 °, and most preferably in the range of 15 to 30 °. With respect to the lower oxidizing holes 70, it is preferable that substantially all of the lower oxidizing holes 70 are located near the lowermost part of the annular element 60 and / or the diametrical element 62. Thus, any liquid and / or solid products that may inadvertently enter into the oxidizer distribution device 34, can be easily released from the oxidizer distribution device 34 through the lower oxidizer openings 70. Preferably, the lower openings for the oxidizing agent 70 release the oxidizing stream downward at a substantially vertical angle. For the purposes of the present description, the upper hole for the oxidizing agent may be any hole that releases the oxidizing stream in the direction generally from the bottom up (i.e., at an angle above the horizontal), and the lower opening for the oxidizing agent may be any hole that releases the oxidizing stream in the direction in general, from top to bottom (that is, at an angle below the horizontal).

In many conventional reactor-type bubblers containing a multiphase reaction medium, substantially all of the reaction medium below the oxidizer distribution device (or other mechanism for introducing the oxidizer stream into the reaction zone) has a very low gas content. As is known in the art, a “gas content” is simply the volume fraction of a multiphase medium that is in a gaseous state. Low gas zones in the medium may also be referred to as “unaerated” zones. In many conventional reactor type bubbler columns, a significant portion of the total volume of the reaction medium is located below the oxidizing agent distribution device (or other mechanism for introducing the oxidizing agent stream into the reaction zone). Thus, a significant portion of the reaction medium present in the lower portion of conventional bubble column columns of the reactor type is unaerated.

It has been found that minimizing the number of unaerated zones in a reaction medium that undergoes oxidation in a bubble column reactor type reactor can minimize the formation of certain types of undesirable impurities. Unaerated zones of the reaction medium contain relatively few oxidizing bubbles. This small volume of oxidizer bubbles reduces the amount of molecular oxygen available for dissolution in the liquid phase of the reaction medium. Thus, the liquid phase in the non-aerated zone of the reaction medium has a relatively low concentration of molecular oxygen. These oxygen-depleted, unaerated zones of the reaction medium tend to accelerate undesirable side reactions, instead of the desired oxidation reaction. For example, when paraxylene is partially oxidized to form terephthalic acid, the insufficient availability of oxygen in the liquid phase of the reaction medium can cause the formation of undesirably high amounts of benzoic acid and associated aromatic rings, in appreciable quantities containing highly undesirable colored molecules known as fluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid phase oxidation is carried out in a bubble column reactor configured and operating in such a way that the volume fraction of the reaction medium with low gas contents is minimized. This minimization of unaerated zones can be quantified by theoretically dividing the entire volume of the reaction medium into 2000 discrete horizontal sections of the same volume. With the exception of the uppermost and lowest horizontal sections, each horizontal section represents a discrete volume bounded on its sides by the side wall of the reactor and bounded above and below by imaginary horizontal planes. The highest horizontal section is bounded below by an imaginary horizontal plane, and above by the upper surface of the reaction medium. The lowest horizontal section is bounded above by an imaginary horizontal plane, and below by the lower edge of the tank. When the reaction medium is theoretically divided into 2,000 discrete horizontal sections of equal volume, the time-averaged and volume-averaged gas contents of each horizontal section can be determined. When using this method of quantifying the number of unaerated zones, it is preferable that the number of horizontal sections having a time-averaged and volume-averaged gas content less than 0.1 is less than 30, more preferably less than 15, even more preferably less than 6, even more preferably less than 4, and most preferably less than 2. It is preferred that the number of horizontal sections having a gas content less than 0.2, was less than 80, more preferably less than 40, even more preferably less than 20, even more preferably less than 12, and most preferably less than 5. It is preferred that the number of horizontal sections having a gas content of less than 0.3 was less than 120, more preferably less than 80, even more preferably less than 40, even more preferably less than 20, and most preferably less than 15 .

Referring again to FIG. 1 and 2, it was found that the arrangement of the oxidizing agent distribution device 34 below in the reaction zone 28 provides several advantages, including reducing the number of unaerated zones in the reaction medium 36. At a given height “H” of the reaction medium 36, the length “L” of the reaction zone 28, and the maximum the diameter "D" of the reaction zone 28, it is preferable that most (i.e.> 50% by weight) of the oxidizing agent stream is introduced into the reaction zone 28 within about 0.025H, 0.022L and / or 0.25D from the lower surface level 52 of the reaction zone 28. More preferably, the bol most of the oxidant stream is introduced into reaction zone 28 within about 0.02H, 0.018L and / or 0.2D from the lower surface level 52 of reaction zone 28. Most preferably, most of the oxidant stream is introduced into reaction zone 28 within 0.015H, 0.013L and / or 0.15D from the lower level of surface 52 of reaction zone 28.

In the embodiment illustrated in FIG. 2, the vertical distance “Y ₁ ” between the lower edge 52 of the reaction zone 28 and the outlet of the upper oxidizing holes 68 of the oxidizing agent distribution device 34 is less than about 0.25H, 0.022L and / or 0 25D, so that essentially the entire oxidant stream enters the reaction zone 28 within about 0.25H, 0.022L and / or 0.25D from the lower level of the surface 52 of the reaction zone 28. More preferably, Y _{1 is} less than than 0.02H, 0.018L and / or 0.2D. Most preferably, Y _{1 is} less than 0.015H, 0.013L and / or 0.15D, but more than 0.005H, 0.004L and / or 0.06D. Figure 2 illustrates the start line of the bend 72 in a position where the lower surface level of the cylindrical main body 46 of the housing 22 is connected to the upper edge of the elliptical bottom 48 of the housing 22. Alternatively, the bottom 48 can be of any shape, including a conical one, and the start line of the bend is still defined as the bottom of the cylindrical main body 46. The vertical distance "Y ₂ " between the start line of the bend 72 and the upper part of the oxidizer distribution device 34 is preferably equal to at least about 0.0012 H, 0.001 L and / or 0.01 D; more preferably at least about 0.005H, 0.004L and / or 0.05D; and most preferably at least 0.01H, 0.008L and / or 0.1D. The vertical distance "Y ₃ " between the lower edge 52 of the reaction zone 28 and the outlet of the lower oxidizer openings 70 for the oxidizer distribution device 34 is preferably less than about 0.015H, 0.013L and / or 0.15D; more preferably less than about 0.012H, 0.01L and / or 0.1D; and most preferably, less than 0.01H, 0.008L and / or 0.075D, but more than 0.003H, 0.002L and / or 0.025D.

In a preferred embodiment of the present invention, the holes that release the oxidizing stream and the feed stream in the reaction zone are configured so that the amount (by weight) of the oxidizing agent or feed stream released from the hole is directly proportional to the open area of the hole. Thus, for example, if 50% of the total open area defined by all openings for the oxidizing agent is within 0.15D from the bottom of the reaction zone, then 50% wt. the oxidant stream enters the reaction zone within 0.15D from the bottom of the reaction zone, and vice versa.

In addition to the advantages provided by minimizing unaerated zones (ie, low gas zones) in the reaction medium 36, it has been found that oxidation can be improved by maximizing the gas content in the entire reaction medium 36. The reaction medium 36 preferably has a time-averaged and volume-averaged gas content of at least about 0.4, more preferably in the range of about 0.6 to about 0.9, and most preferably in the range of 0.65 to 0.85. Several physical and operational attributes of the reactor type bubble column 20 contribute to the high gas content discussed above. For example, for a given reactor size and flow rate for an oxidizer stream, a high L: D ratio of reaction zone 28 gives a smaller diameter that increases the surface velocity in reaction medium 36, which in turn increases the gas content. In addition, the actual bubble column diameter and the L: D ratio are known to affect the average gas content even for a given constant surface velocity. In addition, minimizing unaerated zones, especially in the lower part of reaction zone 28, contributes to an increase in gas content. In addition, the pressure at the top of the column and the mechanical configuration of the bubble column reactor may affect stability at the high surface speeds and gas contents described herein.

Moreover, the authors found the importance of working with optimized pressure in the upper part of the column to obtain an increased gas content and increased mass transfer. It might seem that working at a lower overpressure, which reduces the solubility of molecular oxygen in accordance with Henry's law, would reduce the rate of mass transfer of molecular oxygen from gas to liquid. In a mechanically stirred tank, this usually happens because the aeration levels and mass transfer rates are determined by the design of the mixer and the pressure in the upper part of the column. However, in a reactor bubbler column in accordance with a preferred embodiment of the present invention, it has been found how to use a lower pressure in the upper part of the column to create a given mass of a gas-phase oxidizer stream, occupying a larger volume, increasing the surface velocity in the reaction medium 36 and, in turn, increasing gas content and molecular oxygen transfer rate.

The balance between coalescence and bubble destruction is an extremely complex phenomenon, leading, on the one hand, to a tendency to foam, which reduces the internal circulation rates of the liquid phase and which may require very large separation zones, and, on the other hand, to the tendency to form several very large bubbles, which give a lower gas content and a lower rate of mass transfer from the oxidant stream to the liquid phase. Regarding the liquid phase, its composition, density, viscosity and surface tension, among other factors, are known to interact in a very complex way, with very complex results, even in the absence of a solid phase. For example, laboratory researchers find it useful to qualify whether tap water, distilled water, or deionized water is “water” when they describe and evaluate observations even for simple bubble columns in a water-air system. For complex mixtures in the liquid phase and with the addition of a solid phase, the degree of complexity increases further. The roughness of the surface of individual particles of solid products, the average size of the solid products, the distribution of particle sizes, the amount of solid products relative to the liquid phase and the ability of a liquid to wet the surface of solid products, among other things, are all important in their interaction with the liquid phase and the oxidizing stream when establishing what behavior of the bubbles and what structures of the flow of natural convection will result.

Thus, the ability of a bubbler column reactor to function efficiently at the high surface speeds and high gas contents described herein depends, for example, on the right choice: (1) the composition of the liquid phase of the reaction medium; (2) the amount and type of precipitated solid products, both of which can be controlled by reaction conditions; (3) the magnitude of the flow of oxidizing agent introduced into the reactor; (4) the pressure at the top of the column, which affects the volumetric flow rate, for the oxidizer flow, bubble stability, and, through an energy balance, the reaction temperature; (5) the reaction temperature itself, which affects the properties of the fluid, the properties of the precipitated solid products, and the specific volume of the oxidizing stream; and (6) the geometry and mechanical details of the reaction vessel, including the L: D ratio.

Turning again to FIG. 1, it has been found that an improvement in the distribution of an oxidizable compound (eg, paraxylene) in the reaction medium 36 can be achieved by introducing a liquid phase feed stream into the reaction zone 28 in a plurality of positions separated by a certain vertical distance. Preferably, a liquid phase feed stream is introduced into the reaction zone 28 through at least 3 inlets, more preferably at least 4 inlets. As used herein, the term “inlet openings” shall mean openings where a liquid-phase feed stream is released into the reaction zone 28 for mixing with the reaction medium 36. It is preferred that at least 2 of the inlet openings are separated by a certain distance. in height of at least about 0.5D, more preferably at least about 1.5D, and most preferably at least 3D. However, it is preferable that the tallest inlet is separated by a certain distance in height from the bottommost opening for the oxidizing agent, not greater than about 0.75H, 0.65L and / or 8D; more preferably not greater than about 0.5H, 0.4L and / or 5D; and most preferably, not greater than 0.4H, 0.35L, and / or 4D.

Although it is desirable to introduce a liquid phase stream of starting materials in a plurality of vertically arranged positions, it has also been found that an improved distribution of the oxidizable compound in the reaction medium 36 is provided if most of the liquid phase stream of the starting materials is introduced in the lower half of the reaction medium 36 and / or reaction zone 28. Preferably, at least about 75% wt. a liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or reaction zone 28. Most preferably, at least 90% by weight. a liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or reaction zone 28. In addition, it is preferable that at least about 30% wt. a liquid phase feed stream was introduced into the reaction zone 28 within about 1.5 D from the lowest vertical position, where the oxidizing stream is introduced into the reaction zone 28. This is the lowest vertical position where the oxidizing stream is introduced into the reaction zone 28, as a rule the bottom of the oxidizer distribution device; however, a preferred embodiment of the present invention contemplates many alternative configurations for introducing an oxidizing agent stream into reaction zone 28. Preferably, at least about 50 wt. liquid-phase starting materials are introduced within about 2.5 D from the lowest vertical position, where the oxidant stream is introduced into the reaction zone 28. Preferably, at least about 75% wt. the liquid phase feed stream is introduced within about 5D from the lowest vertical position, where the oxidant stream is introduced into the reaction zone 28.

Each inlet defines an open area through which starting materials are released. It is preferred that at least about 30% of the total open area of all raw material inlets is within about 1.5 D from the lowest vertical position where the oxidant stream is introduced into reaction zone 28. Preferably, at least about 50 % of the total open area of all inputs for the starting materials is within about 2.5 D from the lowest vertical position, where the oxidant stream is introduced into the reaction zone 28. Preferably, at least about 75% of the total Access the square of all inputs for the starting material located within about 5D of the lowest vertical position, where the oxidant stream is introduced into reaction zone 28.

Referring again to FIG. 1, in one embodiment of the present invention, the inputs for raw materials 32a, b, c, d are simply a series of vertically arranged openings along one side of the housing 22. These inlet openings preferably have substantially similar diameters smaller than about 7 cm, more preferably in the range of about 0.25 to about 5 cm, and most preferably in the range of 0.4 to 2 cm. The bubble column reactor 20 is preferably provided with a system for controlling the flow rate,A liquid-phase feed stream out of each inlet. Such a flow control system preferably comprises an individual flow control valve 74a, b, c, d for each respective inlet for feed 32a, b, c, d. In addition, it is preferable that the bubble column reactor 20 is equipped with a flow control system that enables the introduction of at least a portion of the liquid-phase feed stream into the reaction zone 28 at an increased inlet surface velocity of at least about 2 m / s, more preferably at least about 5 m / s, even more preferably at least about 6 m / s, and most preferably in the range from 8 to 20 m / s. As used here, the term "inlet surface velocity" refers to the time-averaged volumetric flow rate for the flow of raw materials from the inlet divided by the area of the inlet. Preferably, at least about 50% wt. a feed stream is introduced into the reaction zone 28 at an increased input surface velocity. Most preferably, substantially all of the feed stream is introduced into reaction zone 28 at an increased inlet surface velocity.

Turning now to FIG. 6-7, an alternative system for introducing a liquid phase feed stream into the reaction zone 28 is illustrated. In this embodiment, the feed stream is introduced into the reaction zone 28 at four different levels in height. Each level in height is equipped with an appropriate source distribution system 76a, b, c, d. Each source distribution system 76 comprises a main source passage 78 and a collector 80. Each collector 80 is provided with at least two outlets 82, 84 connected to respective insertion passages 86, 88, which extend into the reaction zone 28 of the housing 22. Each passage 86, 88 provides a corresponding inlet 87, 89 for releasing a feed stream into the reaction zone 28. The inlets 87, 89 preferably have substantially similar diameters of less than about 7 cm or more than preferably in the range of about 0.25 to about 5 cm, and most preferably in the range of 0.4 to 2 cm. It is preferred that the inlets 87, 89 of each feed distribution system 76a, b, c, d are diametrically opposed so as to introduce a feed stream into the reaction zone 28 in opposite directions. In addition, it is preferable that the diametrically opposite inlets 86, 88 of the adjacent feed distribution systems 76 are oriented at a 90 ° rotation angle with respect to each other. In operation, a liquid phase feed stream is charged through a feed passage 78 and subsequently flows into a manifold 80. A collector 80 distributes the feed stream evenly for simultaneous injection on opposite sides of the reactor 20 through inlets 87, 89.

FIG. 8 illustrates an alternative configuration where each feed distribution system 76 is provided with bayonet (quick-release) tubes 90, 92 instead of passages 86, 88 (shown in FIG. 7). The bayonet tubes 90, 92 protrude into the reaction zone 28 and comprise a plurality of small inlets 94, 96 for releasing liquid phase feed materials into the reaction zone 28. It is preferred that the small inlets 94, 96 of the bayonet tubes 90, 92 have substantially the same diameters smaller than about 50 mm, more preferably from about 2 to about 25 mm, and most preferably from 4 to 15 mm.

FIG. 9-11 illustrate an alternative source material distribution system 100. The source material distribution system 100 introduces a liquid phase stream of source materials in a plurality of positions, separated from each other by a certain vertical distance and separated from each other by a certain lateral distance, without the need for multiple passes through the side wall bubble column reactor type 20. The input system 100, as a rule, contains one inlet passage 102, a collector 104, many located vertically distributed distribution pipes 106, a lateral support mechanism 108 and a vertical support mechanism 110. The inlet passage 102 passes through the side wall 46 of the housing 22. The inlet passage 102 is connected to the fluid communication with the manifold 104. The manifold 104 evenly distributes the feed stream received from the input passage 102, among vertically arranged distribution pipes 106. Each distribution pipe 106 has a plurality of inlet openings 112a, b, c, d d separated from each other by a vertical distance I release the feed stream to reaction zone 28. The lateral support mechanism 108 is connected to each distribution pipe 106 and impedes the relative lateral movement of the distribution pipes 106. The vertical support mechanism 110 is preferably connected to the lateral support mechanism 108 and to the top of the oxidizer distribution device 34. The mechanism the vertical support 110, essentially makes it difficult to vertically move the distribution pipes 106 in the reaction zone 28. It is preferred th to the inlet openings 112 have substantially the same diameters of less than about 50 mm, more preferably from about 2 to about 25 mm, and most preferably from 4 to 15 mm. The vertical distance between adjacent inlets 112 of the feed distribution system 100 illustrated in FIG. 9-11 may be substantially the same as described above with reference to the feed distribution system of FIG. 1.

It has been found that the flow patterns of the reaction medium in a plurality of bubble column columns of the reactor type can make possible an uneven azimuthal distribution of the oxidizable compound in the reaction medium, especially when the oxidizable compound is mainly introduced along one side of the reaction medium. As used here, the term "azimuthal" should mean the angle or distance around the vertically aligned longitudinal axis of the reaction zone. As used here, “vertically positioned” shall mean: within 45 ° of the vertical. In one embodiment of the present invention, a feed stream containing an oxidizable compound (e.g., paraxylene) is introduced into the reaction zone through a plurality of inlet openings separated by a certain azimuthal distance. These inlet openings separated by a certain azimuthal distance can help prevent the occurrence of regions of excessively high and excessively low concentrations of the oxidizable compound in the reaction medium. The various feed systems illustrated in FIG. 6-11 are examples of systems that provide the correct azimuthal distance between inlets.

Referring again to FIG. 7, to quantitatively determine the azimuthally distributed introduction of a liquid-phase feed stream into the reaction medium, the reaction medium can theoretically be divided into four vertically arranged azimuthal quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ " of approximately equal volume. These azimuthal quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ " are determined by a pair of imaginary intersecting perpendicular vertical planes "P1, P2", extending through the maximum vertical size and maximum radial size of the reaction medium. When the reaction medium is contained in a cylindrical container, the intersection line of the imaginary intersecting vertical planes P1, P2 will approximately coincide with the vertical center line of the cylinder, and each azimuthal quadrant Q ₁ , Q ₂ , Q ₃ , Q ₄ will be a generally wedge-shaped vertical a volume having a height equal to the height of the reaction medium. It is preferred that a substantial portion of the oxidizable compound is released into the reaction medium through inlets located in at least two different azimuthal quadrants.

In a preferred embodiment of the present invention, not more than about 80% wt. the oxidizable compound is released into the reaction medium through inlet openings, which may be located in one azimuthal quadrant. More preferably, not more than about 60% wt. the oxidizable compound is released into the reaction medium through inlet openings, which may be located in one azimuthal quadrant. Most preferably, not more than 40% wt. the oxidizable compound is released into the reaction medium through inlet openings, which may be located in one azimuthal quadrant. These parameters for the azimuthal distribution of the oxidizable compound are measured when the azimuthal quadrants are azimuthally oriented so that the maximum possible amount of oxidizable compound is released into one of the azimuthal quadrants. For example, if the entire feed stream is released into the reaction medium through two inlet openings, which are 89 ° azimuthally separated from each other, for the purpose of determining the azimuthal distribution in four azimuthal quadrants, 100% wt. the feed stream is released into the reaction medium in one azimuthal quadrant, since the azimuthal quadrants can be azimuthally oriented so that both inlets are located in the same azimuthal quadrant.

In addition to the advantages associated with the correct azimuthal distance between the inlet openings, it has also been found that the correct radial distance between the inlet openings in the bubble column reactor type can also be important. It is preferred that a substantial portion of the oxidizable compound introduced into the reaction medium is released through inlets that are radially distributed inward from the side wall of the vessel. Thus, in one embodiment of the present invention, a substantial portion of the oxidizable compound enters the reaction zone through inlets located in the “preferred radial inlet zone”, which is distributed inward from the vertically arranged side walls defining the reaction zone.

Referring again to FIG. 7, the preferred radial inlet zone “FZ” may take the form of a theoretical upright cylinder centered in the reaction zone 28 and having an outer diameter “D ₀ ” of 0.9D, where “D” is the diameter of the reaction of zone 28. Thus, an outer OA gap is defined having a thickness of 0.05 D between the preferred radial inlet zone FZ and the inner side of the side walls defining reaction zone 28. It is preferable that either a small amount of oxidizable compound was introduced into the reaction zone 28 through inlets located in this outer annular gap OA, or it was not introduced at all.

In another embodiment, it is preferable that a small amount of the oxidizable compound is introduced into the center of the reaction zone 28 or not introduced at all. Thus, as illustrated in FIG. 8, the preferred radial inlet zone FZ may take the form of a theoretical vertically annular gap, centered in the reaction zone 28 having an outer diameter D _{0 of} 0.9 D and an inner diameter of D _{i of} 0.2 D. Thus, in the present embodiment, the inner cylinder IC having a diameter of 0.2D is “cut out” from the center of the preferred radial inlet zone FZ. It is preferable that a small amount of the oxidizable compound is introduced into the reaction zone 28 through the inlets located in this inner cylinder IC, or it is not introduced at all.

In a preferred embodiment of the present invention, a substantial portion of the oxidizable compound is introduced into the reaction medium 36 through inlets located in the preferred radial inlet zone, regardless of whether the preferred radial inlet zone has the cylindrical or ring shape described above. More preferably at least about 25% wt. the oxidizable compound is released into the reaction medium 36 through inlets located in the preferred radial inlet zone. Even more preferably, at least about 50% wt. the oxidizable compound is released into the reaction medium 36 through inlets located in the preferred radial inlet zone. Most preferably, at least 75% wt. the oxidizable compound is released into the reaction medium 36 through inlets located in the preferred radial inlet zone.

Although the theoretical azimuthal quadrants and the theoretical preferred radial inlet zone illustrated in FIG. 7 and 8 are described with reference to the distribution of the liquid-phase flow of the starting materials, it is found that the correct azimuthal and radial distribution of the gas-phase flow of the oxidizing agent can also provide certain advantages. Thus, in one embodiment of the present invention, the description of the azimuthal and radial distribution of the liquid-phase feed stream given above is also applicable to the method by which a gas-phase oxidizer stream is introduced into the reaction medium 36.

Turning now to FIG. 12-15, an alternative oxidizer distribution device 200 is illustrated as containing a generally annular element 202 and a pair of passages 204, 206 for accumulating the oxidizing agent. The oxidizer distribution device 200 in FIG. 12-15 is similar to the oxidizer distribution device 34 in FIG. 1-11, with the following three main differences: (1) the oxidizer distribution device 200 does not contain a diagonal diametrical element; (2) the upper part of the annular element 202 does not have openings for releasing the oxidizing agent in the upward direction; and (3) the oxidizing agent distribution device 200 has much more openings in the lower part of the annular element 202.

As probably best illustrated in FIG. 14 and 15, the lower portion of the oxidizer distribution device ring 202 provides a plurality of oxidizer holes 208. The holes for the oxidizer 208 are preferably configured so that at least about 1% of the total open area defined by the holes 208 for the oxidizer is located below the center line 210 (FIG. 15) of the ring element 202, where the center line 210 is located at the height of the volumetric the centroid of the annular element 202. More preferably, at least about 5% of the total open area defined by all of the oxidizing holes 208 is located below the center line 210, with at least about 2% of the total its open area is determined by openings 208, which release the flow of oxidizing agent in the direction, generally, downward, within about 30 ° from the vertical. Even more preferably, at least about 20% of the total open area defined by all openings for the oxidizing agent 208 is located below the center line 210, with at least about 10% of the total open area being determined by openings 208 that release the oxidant stream in the direction generally down, within 30 ° from the vertical. Most preferably, at least about 75% of the total open area defined by all of the oxidant openings 208 is located below the center line 210, with at least about 40% of the total open area defined by openings 208 that release the oxidant flow in the direction generally down, within 30 ° of the vertical. The fraction of the total open area defined by all oxidizer openings 208 that are located above the center line 210 is preferably less than about 75%, more preferably less than about 50%, even more preferably less than about 25 %, and most preferably less than 5%.

As illustrated in FIG. 14 and 15, the oxidizer openings 208 comprise downwardly directed openings 208a and inclined openings 208b. The downwardly directed openings 208a are configured to release the oxidizing stream, generally downward, at an angle within about 30 ° from the vertical, more preferably within about 15 ° from the vertical, and most preferably within 5 ° from the vertical. The inclined openings 208b are configured to release the oxidizing stream, generally outward and downward, at an angle “A” that is in the range of about 15 to about 75 ° from the vertical, more preferably the angle A is in the range of about 30 to about 60 ° from the vertical, and most preferably, the angle A is in the range of 40 to 50 ° from the vertical.

It is preferred that substantially all of the oxidizing holes 208 have approximately the same diameter. The diameter of the oxidizing holes 208 is preferably in the range of about 2 to about 300 mm, more preferably in the range of about 4 to about 120 mm, and most preferably in the range of 8 to 60 mm. The total number of oxidizer holes 208 in the ring member 202 is selected to satisfy the criteria for low pressure drop, described in detail below. Preferably, the total number of oxidizer holes 208 formed in the ring member 202 is at least about 10, more preferably the total number of oxidizer holes 208 is in the range of about 20 to about 200, and most preferably the total number of holes 208 for oxidizing agent ranges from 40 to 100.

Although FIGS. 12-15 illustrate a very specific configuration of the oxidizer distribution device 200, it is now noted that various configurations of the oxidizer distribution device can be used to achieve the advantages described herein. For example, the oxidizer distribution device does not need to have the configuration of an octagonal ring element illustrated in FIGS. 12-13. Rather, it is possible for the oxidizer distribution device to be formed from any flow path (s) that use a plurality of openings separated by a certain distance to release the flow of oxidant. The size, quantity, and direction of release of the oxidant openings in the flow passage are preferably within the ranges set forth above. In addition, the oxidizer distribution device is preferably configured to provide the azimuthal and radial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidizer distribution device, it is preferable that the oxidizer distribution device be physically configured and operate in a manner that minimizes the pressure drop associated with the release of the oxidizer stream from the passage (s) for the stream, through the openings for the oxidizer, and into the reaction zone. This pressure drop is calculated as the time-averaged static pressure of the oxidizer stream inside the flow passage at the inlet to the oxidizer 66a, b of the oxidizer distribution device minus the time-averaged static pressure in the reaction zone at a height where half of the oxidizer flow is introduced above this vertical position and half an oxidizing agent stream is introduced below this vertical position. In a preferred embodiment of the present invention, the time-averaged pressure drop associated with the release of the oxidizing agent stream from the oxidizing agent distribution device is less than about 0.3 MPa, more preferably less than about 0.2 MPa, even more preferably less than about 0.1 MPa, and most preferably less than 0.05 MPa. Under the preferred operating conditions of the reactor bubbler column described herein, the pressure of the oxidizer stream inside the passage (s) for the flow in the oxidizer distribution device is preferably in the range of about 0.35 to about 1 MPa, more preferably in the range of about 0.45 to up to 0.85 MPa, and most preferably in the range from 0.5 to 0.7 MPa.

As indicated previously with reference to the configuration of the oxidizing agent distribution device illustrated in FIG. 2-5, it may be desirable to continuously or periodically clean the oxidizing agent distribution device with a liquid (e.g., acetic acid, water and / or paraxylene) to prevent the formation of a precipitate in the oxidizing agent distribution device from solid products. When such a liquid purge is used, it is preferable that an effective amount of liquid (i.e., not only a small amount of liquid droplets that might naturally be present in the oxidizer stream) pass through the oxidizer distribution device and from the oxidizer openings for at least one period more than one minute daily. When a liquid is continuously or periodically released from an oxidizer distribution device, it is preferable that the time-averaged ratio of the mass flow rate of the liquid through the oxidizer distribution device to the mass flow rate of molecular oxygen through the oxidizer distribution device is in the range of about 0.05: 1 to about 30 : 1, or in the range of about 0.1: 1 to about 2: 1, or even in the range of 0.2: 1 to 1: 1.

In one embodiment of the present invention, a significant portion of the oxidizable compound (e.g., paraxylene) can be introduced into the reaction zone through an oxidizer distribution device. In such a configuration, it is preferable that the oxidizable compound and molecular oxygen are released from the oxidizer distribution device through the same openings in the oxidizer distribution device. As noted above, the oxidizable compound is typically a liquid at STP (at standard temperature and pressure). Therefore, in the present embodiment, the two-phase stream can be released from the oxidizer distribution device, wherein the liquid phase contains an oxidizable compound and the gas phase contains molecular oxygen. It should be noted, however, that at least a portion of the oxidizable compound may be in a gaseous state when released from the oxidizer distribution device. In one embodiment, the liquid phase released from the oxidizer distribution device is formed primarily from the oxidizable compound. In another embodiment, the liquid phase released from the oxidizer distribution device has substantially the same composition as the feed stream described above. When the liquid phase released from the oxidizer distribution device has substantially the same composition as the feed stream, such a liquid phase may contain a solvent and / or catalyst system in the amounts and ratios described above with reference to the composition of the feed stream.

In one of the embodiments of the present invention, it is preferred that at least about 10% wt. the entire oxidizable compound introduced into the reaction zone was introduced through an oxidizer distribution device, more preferably at least about 40 wt%. the oxidizable compound was introduced into the reaction zone through an oxidizing agent distribution device, and most preferably at least 80% by weight. the oxidizable compound was introduced into the reaction zone through an oxidizer distribution device. When all or part of the oxidizable compound is introduced into the reaction zone through an oxidizing agent distribution device, it is preferred that at least about 10% by weight. all of the molecular oxygen introduced into the reaction zone was introduced through the same oxidizer distribution device, more preferably at least about 40 wt%. the oxidizable compound was introduced into the reaction zone through the same oxidizer distribution device, and most preferably at least 80% by weight. the oxidizable compound was introduced into the reaction zone through the same oxidizer distribution device. When a significant portion of the oxidizable compound is introduced into the reaction zone through an oxidizer distribution device, it is preferable that one or more temperature monitoring devices (eg, thermocouples) be located in the oxidizer distribution device. These temperature sensors can be used to help make sure that the temperature in the oxidizer distribution device does not become dangerously high.

Turning now to FIG. 16-18, the reactor type bubbler column 20 is illustrated as containing an internal deaeration vessel 300 located in the lower part of the reaction zone 28 near the outlet for the slurry 38. Side reactions forming impurities were found to occur at a relatively high speed during deaeration of the reaction medium 36 As used here, "deaeration" should mean the separation of the gas phase from the multiphase reaction medium. When the reaction medium 36 is highly aerated (gas content> 0.3), the formation of impurities is minimal. When the reaction medium 36 is highly unaerated (gas content <0.01), the formation of impurities is also minimal. However, when the reaction medium is partially aerated (gas content 0.01-0.3), undesirable side reactions are facilitated and more impurities are formed. The deaeration vessel 300 solves this and other problems by minimizing the volume of the reaction medium 36 in a partially aerated state and by minimizing the time that deaeration of the reaction medium 36 takes. In essence, the deaerated suspension is obtained from the bottom of the deaeration vessel 300 and leaves the reactor 20 through outlet 38 for the suspension. Essentially, the deaerated suspension preferably contains less than about 5% by volume of the gas phase, more preferably less than about 2% by volume of the gas phase, and most preferably less than 1% by volume of the gas phase.

In FIG. 16, a reactor type bubble column 20 is illustrated as comprising a level controller 302 and a flow control valve 304. The level controller 302 and the flow control valve 304 together support the reaction medium 36 at substantially constant height in the reaction zone 28. The level controller 302 operates by tracking (for example, by monitoring the level by pressure difference or by internal level monitoring) the height of the upper level of the surface 44 of the reaction medium 36 and generating a control signal 306 that is sensitive to height the reaction medium 36. The flow control valve 304 receives a control signal 306 and controls the flow rate of the suspension through the passage to exit the suspension 308. Thus, the flow rate of the suspension from the outlet 38 for the suspension can vary between the maximum volumetric flow rate of the suspension (F _max ) when the height of the reaction medium 36 is too large, and the minimum volumetric flow rate of the suspension (F _min ) when the height of the reaction medium 36 is too low.

To remove the solid-state oxidation product from the reaction zone 28, the batch must first pass through the deaeration vessel 300. The deaeration vessel 300 provides an internal volume with low turbulence that allows the gas phase of the reaction medium 36 to naturally rise from the liquid and solid phases of the reaction medium 36 when the liquid and solid products flow down towards the outlet for slurry 38. The rise of the gas phase from the liquid and solid phases is caused by the natural positive rise of the gas phase s in liquid and solid phases. When a deaeration vessel 300 is used, the transition of the reaction medium 36 from a fully aerated three-phase medium to a fully deaerated two-phase suspension is quick and efficient.

Turning now to FIG. 17 and 18, the deaeration vessel 300 comprises a generally vertically positioned side wall 308 defining a deaeration zone 312 within itself. Preferably, the side wall 308 extends from bottom to top within about 30 ° from the vertical, more preferably within about 10 ° from the vertical . Most preferably, the side wall 308 is substantially vertical. The deaeration zone 312 is separate from the reaction zone 28 and has a height "h" and a diameter of "d." The upper level of the surface 310 of the side wall 308 is open, so as to receive the reaction medium from the reaction zone 28 into the internal volume 312. The lower level of the surface of the side wall 308 is connected to the fluid communication with the outlet for the suspension 38 through the transition section 314. In certain cases, for example, when the opening of the outlet 38 for the suspension is larger or when the diameter "d" of the side wall 308 is small, the transition section 314 can be eliminated. As probably best illustrated in FIG. 18, the deaeration vessel 300 may also include a vortex quencher 316 located in the deaeration zone 312. The vortex quencher 316 may have any structure that works to slow the formation of vortices when the solid and liquid phases flow down towards exit 38 for the suspension.

To enable appropriate separation of the gas phase from the solid and liquid phases in the deaeration vessel 300, the height "h" and the horizontal cross-sectional area of the internal deaeration zone 312 are carefully selected. The height "h" and the horizontal cross-sectional area of the inner deaeration zone 312 should provide sufficient distance and time, so that even when the maximum amount of suspension is removed (that is, when the suspension is removed at F _max ), essentially the entire volume of gas bubbles can rise from solid and liquid phases before gas bubbles reach the lower outlet of the deaeration vessel 300. Thus, it is preferable that the cross-sectional area of the deaeration zone 312 be such that the maximum I down speed (V _dmax) of the liquid and solid phases through deaeration zone 312 was substantially less than the natural rise velocity (V _u) of the gas phase bubbles through the liquid and solid phases. The maximum downward velocity (V _dmax ) of the liquid and solid phases through the deaeration zone 312 is achieved at the maximum volumetric flow rate of the suspension (F _max ), discussed above. The natural rate of rise (V _u ) of gas bubbles through the liquid and solid phase varies depending on the size of the bubbles; however, the natural rate of rise (V _u0.5 ) of gas bubbles with a diameter of 0.5 cm through the liquid and solid phases can be used as a threshold value, since essentially the entire volume of the bubbles, initially in the reaction medium 36, will be greater than 0, 5 cm. Preferably, the cross-sectional area of the deaeration zone 312 is such that V _dmax is less than about 75% of V _u0.5 , more preferably V _dmax is less than about 40% of V _u0.5 , most preferably V _dmax is less than 20% of V _u0.5 .

The downward velocity of the liquid and solid phases in the deaeration zone 312 of the deaeration tank 300 is calculated as the volumetric flow rate of the deaerated suspension through the outlet for the slurry 38 divided by the minimum cross-sectional area of the deaeration zone 312. The downward velocity of the liquid and solid phases in the deaeration zone 312 300 is preferably less than about 50 cm / s, more preferably less than about 30 cm / s, and most preferably less than 10 cm / s.

We note now that, although the vertically arranged side wall 308 of the deaeration vessel 300 is illustrated as having a cylindrical configuration, the side wall 308 may include a plurality of side walls that form a plurality of configurations (e.g., triangular, square, or oval), insofar as the walls define an internal volume having an appropriate volume, cross-sectional area, width "d" and height "h". In a preferred embodiment of the present invention, “d” is in the range of about 0.2 to about 2 m, more preferably in the range of about 0.3 to about 1.5 m, and most preferably in the range of 0.4 to 1, 2 m. In a preferred embodiment of the present invention, “h” is in the range of about 0.3 m to about 5 m, more preferably in the range of about 0.5 to about 3 m, and most preferably in the range of 0.75 to 2 m

In a preferred embodiment of the present invention, the side wall 308 is substantially vertical, so that the horizontal cross-sectional area of the deaeration zone 312 is substantially constant over the entire height “h” of the deaeration zone 312. Preferably, the maximum horizontal cross-sectional area of the deaeration zone 312 is less than about 25% of the maximum horizontal cross-sectional area of the reaction zone 28. More preferably, the maximum horizontal cross-sectional area the cross section of the deaeration zone 312 is in the range of about 0.1 to about 10% of the maximum horizontal cross-sectional area of the reaction zone 28. Most preferably, the maximum horizontal cross-sectional area of the deaeration zone 312 is in the range of 0.25 to 4% of the maximum horizontal the cross-sectional area of the reaction zone 28. Preferably, the maximum horizontal cross-sectional area of the deaeration zone 312 is in the range from about 0.02 to about 3 m ² , more preferred particularly in the range of about 0.05 to about 2 m ² , and most preferably in the range of 0.1 to 1.2 m ² . The volume of the deaeration zone 312 is preferably less than about 5% of the total volume of the reaction medium 36 or reaction zone 28. More preferably, the volume of the deaeration zone 312 is in the range from about 0.01 to about 2% of the total volume of the reaction medium 36 or reaction zone 28. Most preferably, the volume of deaeration zone 312 is in the range from 0.05 to about 1% of the total volume of reaction medium 36 or reaction zone 28. The volume of deaeration zone 312 is preferably less than about 2 m ³ , more preferably ranges from about 0.01 to about 1 m ³ , and most preferably ranges from 0.05 to 0.5 m ³ .

Turning now to FIG. 19, the reactor type bubbler column 20 is illustrated as having an external deaeration vessel 400. In this configuration, the aerated reaction medium 36 is removed from the reaction zone 28 through an elevated opening on the side of the housing 22. The recovered aerated medium is transferred to the external deaeration vessel 400 through an outlet passage 402 to separate the gas phase from the solid and liquid phases. The separated gas phase leaves the deaeration vessel 400 through passage 404, while the substantially deaerated suspension leaves the deaeration vessel 400 through passage 406.

In FIG. 19, the exit passage 402 is shown as being approximately straight, horizontal, and orthogonal to the housing 22. This is simply one of the convenient configurations; and the outlet passage 402 can be excellent in any respect, provided that it is suitable for use to connect a bubble column reactor 20 to an external deaeration tank 400. Turning to the passage 404, it is useful that this passage is connected on top of or near the deaeration tank 400 its upper part for reliable control of problems associated with a stagnant gas pocket containing an oxidizable compound and an oxidizing agent. Moreover, passages 402 and 404 may be suitable for containing flow isolation means, such as valves.

When the reaction medium 36 is withdrawn from the reactor 20 through a heightened outlet, as shown in FIG. 19, it is preferable that the bubble column reactor 20 is provided with a lower outlet 408 near the lower part 52 of the reaction zone 28. The lower outlet 408 and the lower passage 410 connected to it can be used for a spare (i.e. empty) reactor 20 during shutdowns. Preferably, one or more lower exits 408 are provided in the lower third of the height of the reaction medium 36, more preferably in the lower quarter of the reaction medium 36, and most preferably at the lowest point of the reaction zone 28.

For the elevated suspension and deaeration recovery system shown in FIG. 19, the lower passage 410 and outlet 408 are not used to extract the suspension from reaction zone 28 during oxidation. It is known in the art that solid products tend to settle by gravity in non-aerated and otherwise unmixed portions of the suspension, including stagnant flow zones. In addition, precipitated solid products (e.g. terephthalic acid) may tend to cure in the form of large agglomerates by continuing precipitation and / or crystalline reorganization. Thus, to prevent clogging of the lower passage for stream 410, a portion of the deaerated slurry from the lower part of the deaeration vessel 400 can be used to continuously or periodically clean the lower passage 410 during normal operation of the reactor 20. Preferred means for such a cleaning of the suspension passage 410 is to periodically open valve 412 in passage 410 and allowing part of the deaerated slurry to flow through passage 410 and into reaction zone 28 through the bottom its opening 408. Even when the valve 412 is fully or partially open, only part of the deaerated suspension flows through the lower passage 410 and back into the reaction zone 28. The remaining part of the deaerated suspension, not used to clean the lower passage 410, is transferred through the passage 414 from the reactor 20 for additional subsequent processing (e.g. cleaning).

During normal operation of the bubbler column reactor 20 for a sufficient length of time (e.g.> 100 hours), it is preferable that the amount of deaerated slurry used to clean the lower passage 410 be less than 50% by weight of the total deaerated suspension obtained from the lower part of the deaeration tank 400, more preferably less than about 20 wt.%, and most preferably less than 5 wt.%. In addition, it is preferable that, for a sufficient length of time, the average mass flow rate of the deaerated slurry used to clean the lower passage 410 is less than approximately 4 times the average mass flow rate of the oxidizable compound in reaction zone 28, more preferably less approximately more than 2 times the average mass flow rate of the oxidizable compound in the reaction zone 28, even more preferably less than the average mass flow rate of the oxidizable compound Ia in reaction zone 28, and most preferably less than 0.5 times the average mass flow rate of the oxidizable compound into reaction zone 28.

Referring again to FIG. 19, the deaeration vessel 400 comprises a substantially vertically preferably cylindrical side wall 416 defining the deaeration zone 418. The deaeration zone 418 has a diameter “d” and a height “h.” The height "h" is measured as the vertical distance between the position where the aerated reaction medium enters the deaeration vessel 400 and the lower part of the side wall 416. The height "h", the diameter "d", the area and volume of the deaeration zone 418 are preferably essentially the same as described above, with reference to the deaeration zone 312 of the deaeration vessel 300 illustrated in FIG. 16-18. In addition, the deaeration vessel 400 includes an upper section 420 formed by a protruding side wall 416 above the deaeration zone 418. The upper section 420 of the deaeration vessel 400 can be of any height, although it preferably extends up to the level of the reaction medium 36 in the reaction zone 28 or higher . The upper section 420 ensures that the gas phase has room for appropriate separation from the liquid and solid phases before exiting the deaeration vessel 400 through passage 404. Now, note that although passage 404 is illustrated as returning the separated gas phase to the separation zone of reactor 20, passage 404 alternatively, it can be connected to the housing 22 at any height above the exit passage 402. Optionally, the passage 404 can be connected to the exit passage of the gas 40, so that the separated gas phase from the deaeration vessel 400 is combined with Dalen from column top vapor flow in the passage 40, and sent on for further processing.

Turning now to FIG. 20, a reactor type bubbler column 20 is illustrated as containing a hybrid internal-external deaeration vessel 500. In this configuration, a portion of the reaction medium 36 is removed from the reaction zone 28 through a relatively elevated opening 502 in the side wall of the housing 22. Then, the extracted reaction medium 36 is transferred through a relatively large diameter crank passage 504 and enters the upper part of the deaeration vessel 500. In FIG. 20, the crank passage 504 is shown as connecting orthogonally to the wall of the housing 22 and containing a smooth rotation at an angle of about 90 °. This is just one of the convenient configurations; and the crank passage 504 may be excellent in any respect, provided that it connects a usable reactor type bubble column 20 with an external deaeration tank 500 as described. In addition, the crank passage 504 may include suitable means for isolating the flow, such as valves.

In the deaeration vessel 500, the gas phase moves from bottom to top, while the solid and liquid phases move from top to bottom. Moving from the bottom up, the gas phase can re-enter the cranked passage 504, and then go through the opening 502 back into the reaction zone 28. Thus, a counterflow of the incoming reaction medium 36 and the outgoing separated gas can take place in the opening 502. The deaerated slurry leaves the deaeration tank 500 through the passage 506. The deaeration tank 500 comprises a substantially vertically preferably cylindrical side wall 508 defining the deaeration zone 510. The deaeration zone 510 has a height “h” and a diameter “d”. It is preferable that the raised hole 502 and the elbow bore 504 have a diameter equal to or greater than the diameter "d" of the deaeration zone 510. The height "h", the diameter "d", the area and volume of the deaeration zone 510 are preferably essentially the same as described above with reference to the deaeration zone 312 of the deaeration vessel 300 illustrated in FIG. 16-18.

FIG. 19 and 20 illustrate an embodiment of a reactor type bubbler column 20, wherein a solid product (e.g., crude terephthalic acid) obtained in reaction zone 28 is recovered from reaction zone 28 through an elevated outlet. Removing the aerated reaction medium 36 from a raised position above the bottom of the bubble column reactor 20 can help prevent the accumulation and stagnation of the poorly aerated reaction medium 36 in the lower part 52 of reaction zone 28. In accordance with other aspects of the present invention, the concentration of oxygen and oxidizable compound (for example, paraxylene) in the reaction medium 36 near the top of the reaction medium 36 is preferably lower than near the bottom. Thus, recovering the reaction medium 36 in a raised position can increase the yield by lowering the amount of unreacted reagents recovered from the reactor 20. Also, the temperature of the reaction medium 36 changes significantly in the vertical direction when the bubbler column of reactor type 20 operates at high STR and chemical gradients composition as described here. Under such conditions, the temperature of the reaction medium 36, as a rule, will have local minima near the lower surface level and the upper surface level of the reaction zone 28. Near the lower surface level, the minimum is associated with the evaporation of the solvent near the region where the oxidizing agent or part thereof is taken. Near the upper level of the surface, the minimum is again caused by evaporation of the solvent, although the decrease in pressure in the reaction medium caused here is. In addition, other local minima may occur between the upper and lower surface levels when additional starting materials or an oxidizing agent are received in the reaction medium. Thus, there is one or more temperature maxima arising from the exothermic heat of the oxidation reactions between the lower edge and the upper edge of the reaction zone 28. Removing the reaction medium 36 in a raised position with elevated temperature can be especially advantageous when subsequent processing is carried out at higher temperatures, since the cost of energy associated with heating the extracted medium for subsequent processing is reduced.

Thus, in a preferred embodiment of the present invention, and especially when the subsequent processing is carried out at higher temperatures, the reaction medium 36 is removed from the bubble column of the reactor type 20 through an elevated outlet (s) located above the position (s) where reaction zone 28, at least 50% wt. a liquid phase stream of starting materials and / or a gas phase stream of an oxidizing agent. More preferably, the reaction medium 36 is removed from the bubble column reactor 20 through a heightened outlet (s) located above the position (s) where substantially all of the liquid phase feed stream and / or gas phase oxidant stream enter reaction zone 28 . Preferably, at least 50% wt. solid-phase and liquid-phase components, extracted from the bubble column reactor type 20, are extracted through a heightened outlet (s). More preferably, substantially all of the solid phase and liquid phase components recovered from the bubble column reactor type 20 are recovered via a heightened outlet (s). Preferably, the elevated output (s) is at least about 1D higher than the lower surface level 52 of the reaction zone 28. More preferably, the elevated output (s) is at least about 2D higher than the lower surface level. 52 of reaction zone 28. Most preferably, the elevated output (s) are at least 3D above the lower surface level 52 of reaction zone 28. At a given height “H” of reaction medium 36, it is preferable that the solder the elevated yield (s) were vertically in the range between about 0.2H and about 0.8H, more preferably between about 0.3H and about 0.7H, and most preferably between 0.4H and 0.6H. In addition, it is preferable that the temperature of the reaction medium 36 at an elevated outlet from the reaction zone 28 be at least 1 ° C higher than the temperature of the reaction medium 36 at a lower level of the surface 52 of the reaction zone 28. More preferably, the temperature the reaction medium 36 at a heightened outlet from the reaction zone 28 is higher, in the range of about 1.5 to about 16 ° C, than the temperature of the reaction medium 36 at a lower level of the surface 52 of the reaction zone 28. Most preferably, the reaction temperature hydrochloric medium 36 at an elevated height of the outlet 28 above the reaction zone, in the range of 2 to 12 ° C, than the temperature of reaction medium 36 on the lower level surface 52 of reaction zone 28.

Referring now to FIG. 21, the reactor type bubbler column 20 is illustrated as containing an alternative hybrid deaeration vessel 600 located at the bottom of the reactor 20. With this configuration, the aerated reaction medium 36 is removed from the reaction zone 28 through a relatively large opening 602 at a lower surface level 52 casing 22. Hole 602 defines the open upper surface level of deaeration vessel 600. In deaeration vessel 600, the gas phase moves from bottom to top, while the solid and liquid phases moving from top to bottom. Moving from the bottom up, the gas phase can re-enter the reaction zone 28 through the opening 602. Thus, the counterflow of the incoming reaction medium 36 and the leaving separated gas can be carried out at the opening 602. The deaerated slurry leaves the deaeration vessel 600 through passage 604. The deaeration vessel 600 comprises a substantially vertically preferably cylindrical side wall 606 defining the deaeration zone 608. The deaeration zone 608 has a height “h” and a diameter “d”. It is preferable that the aperture 602 has a diameter the same as the diameter "d" of the deaeration zone 608, or larger. The height “h”, the diameter “d”, the area and volume of the deaeration zone 608 are preferably substantially the same as described above with reference to the deaeration zone 312 of the deaeration vessel 300 illustrated in FIG. 16-18.

Turning now to FIG. 22, the reactor type bubbler column 20 of FIG. 21 is illustrated as containing an alternative oxidizer distribution device 620. The oxidizer distribution device 620 comprises an annular element 622 and a pair of inlet passages 624, 626. The annular element 622 preferably has substantially such the same configuration as the ring member 202 described above with reference to FIG. 12-15. The inlet passages 624, 626 extend from the bottom up through the holes in the bottom 48 of the housing 22 and provide an oxidizer flow to the ring element 622.

Turning now to FIG. 23, the reactor type bubbler column 20 of FIG. 21 is illustrated as containing means for introducing an oxidizing agent stream into the reaction zone 28, not containing distribution devices. In the configuration of FIG. 23, the oxidizer stream is delivered to the reactor 20 through the passages for the oxidizer 630, 632. The passages for the oxidizer 630, 632 are connected to the corresponding holes for the oxidizer 634, 636 in the bottom 48 of the housing 22. The oxidant stream is introduced directly into the reaction zone 28 through the holes for oxidizer 634, 636. Optional deflector plates 638, 640 may be provided to divert the flow direction, for the oxidizer flow, after it initially enters the reaction zone 28.

As discussed above, it is preferable that the oxidation reactor is configured and operated in a manner that eliminates zones with a high concentration of oxidizable compound in the reaction medium, since such zones can lead to the formation of impurities. One way to improve the initial dispersion of an oxidizable compound (e.g., paraxylene) in the reaction medium is to dilute the oxidizable compound with a liquid. The fluid used to dilute the oxidizable compound may occur from a portion of the reaction medium located at a considerable distance from the position (s) where the oxidizable compound is introduced into the reaction zone. This liquid from a remote part of the reaction medium can circulate to a position close to the accumulation position of the oxidizable compound through a passage that is located inside and / or outside the main reaction vessel.

Figs. 24 and 25 illustrate two preferred methods of circulating liquid from a remote portion of the reaction medium to a position near the inlet of the oxidizable compound using an internal (Fig. 24) or external (Fig. 25) passage. Preferably, the length of the passage for the flow from its inlet (i.e., the hole (s) where the liquid enters the passage) to its exit (i.e., the hole (s) where the liquid is released from the passage) is greater than about 1 m, more preferably , greater than about 3 m, even more preferably greater than about 6 m, and most preferably more than 9 m. However, the actual length of the passage becomes less important if the liquid is obtained from a separate container, probably located directly above the capacity to which initially Sources of released oxidizable compound, or next to it. A fluid from any single container containing at least a portion of the reaction medium is the preferred source for the initial dilution of the oxidizable compound.

It is preferred that the fluid flowing through the passage, whatever its source, has a lower constant concentration of the oxidizable compound than the reaction medium in the immediate vicinity with at least one exit from the passage. In addition, it is preferable that the liquid flowing through the passage has a concentration of the oxidizable compound in the liquid phase of below about 100,000 ppm. wt., more preferably below about 10,000 ppm. wt., even more preferably below about 1000 ppm. wt., and most preferably below 100 ppm. wt., where the concentration is measured before adding to the passage of a small additive of the starting materials of the oxidizable compound and any optional, separate source materials of the solvent. When the measurements are carried out after adding a small additive of the starting materials of the oxidizable compound and the optional starting materials of the solvent, it is preferable that the combined liquid stream entering the reaction medium has a concentration of the oxidizable compound in the liquid phase below about 300,000 ppm. wt., more preferably below about 50,000 ppm. wt., and most preferably below 10000 ppm. wt.

It is desirable to maintain the flow through the passage at a sufficiently low speed so that the circulating liquid does not suppress the desired overall gradient of the oxidizable compound in the reaction medium. In this regard, it is preferable that the ratio of the mass of the liquid phase in the reaction zone to which a small additive of the oxidizable compound is first released to the mass flow rate of the liquid flowing through the passage be greater than about 0.3 minutes, more preferably large, about than 1 minute, even more preferably, was between about 2 minutes and about 120 minutes, and most preferably between 3 minutes and 60 minutes.

There are many means available to cause fluid to flow through a passage. Preferred means include gravity, all types of ejectors using either gas or liquid, or both, as a motive fluid, and mechanical pumps of all types. When an ejector is used, one embodiment of the present invention uses at least one fluid selected from the group consisting of starting materials of an oxidizable compound (liquid or gas), starting materials of an oxidizing agent (gas), and starting materials as a driving fluid solvent (liquid) and pumped source of the reaction medium (suspension). Another embodiment uses at least two fluids selected from the group consisting of starting materials of an oxidizable compound, starting materials of an oxidizing agent and starting materials of a solvent as a driving fluid. Another embodiment utilizes a combination of starting materials of an oxidizable compound, starting materials of an oxidizing agent and starting materials of a solvent as a driving fluid.

The corresponding diameter or diameters of the circulation passage can vary in accordance with the quantity and properties of the material that is being transported, the energy available for forced flow, and capital investment considerations. It is preferred that the minimum diameter for such a passage is greater than about 0.02 m, more preferably between about 0.06 m and about 2 m, and most preferably between 0.12 and 0.8 m.

As noted above, it is desirable to control the flow through the passage within certain preferred ranges. There are many tools known in the art for implementing this control by selecting the appropriate fixed geometry during the construction of the flow passage. Another preferred embodiment is the use of geometry that changes during operation, explicitly including valves of all sorts and descriptions, including both manually actuated and actuated by the actuator, by any means, including loop control feedback from the measuring element or without them. Other preferred means for controlling the dilution fluid flow are to vary the energy storage between the inlet and the outlet of the passage. Preferred means include changing the flow rate of one or more motive fluids for the ejector, changing the energy storage for driving the pump, and the density difference or height difference when gravity is used. These preferred agents may also be used in any combination.

The passage used to circulate liquid from the reaction medium may be of any type known in the art. One embodiment uses a passage constructed in whole or in part using conventional pipe materials. Another embodiment uses a passage constructed in whole or in part using the wall of the reaction vessel as one part of the passage. The passage may be designed as completely enclosed within the boundaries of the reaction vessel (Fig. 24), or it may be constructed completely outside the reaction vessel (Fig. 25), or it may contain sections both inside and outside the reaction vessel.

The authors believe that, in particular, in large reactors, it may be desirable to have multiple passages and their various designs for moving fluid through the passage. In addition, it may be desirable to create multiple outputs in multiple positions on one or all passes. The design features will provide the desired overall gradient of constant concentrations of the oxidizable compound with the desired initial dilution of the starting materials of the oxidizable compound in accordance with other aspects of the present invention.

Both FIG. 24 and FIG. 25 illustrate structures that use a deaeration vessel connected to the passage. This deaeration vessel ensures that part of the reaction medium used to dilute the incoming oxidizable compound is a substantially deaerated suspension. Now note, however, that the liquid or suspension used to dilute the incoming oxidizable compound may be in an aerated form as well as in a deaerated form.

The use of a fluid flowing through the passage to ensure dilution of the starting materials of the oxidizable compound is particularly useful in bubble column reactors. In addition, in reactor-type bubbler columns, good gains can be achieved for the initial dilution of the starting materials of the oxidizable compound, even without adding the starting materials of the oxidizable compound directly to the passage, provided that the exit from the passage lies close enough to the position of adding the oxidizable compound, in such an embodiment, it is preferable that the exit from the passage is within about 27 diameters of the exit from the passage from the nearest addition position for approx the compound to be used, more preferably within about 9 diameters of the exit of the passage, even more preferably within about 3 diameters of the exit of the passage, and most preferably within 1 diameter of the exit of the passage.

It has also been found that flow ejectors may be suitable for the initial dilution of an oxidizable compound in oxidation bubble columns in accordance with one embodiment of the present invention, even without using passages to obtain a dilution fluid from a remote portion of the reaction medium. In such cases, the ejector is located in the reaction medium and has an open path from the reaction medium to the neck, where low pressure draws into the adjacent reaction medium. Examples of two possible ejector configurations are illustrated in FIG. 26 and 27. In a preferred embodiment of these ejectors, the closest position of the incoming oxidizable compound is within about 4 m, more preferably within about 1 m, and most preferably 0.3 m from the neck of the ejector. In another embodiment, the oxidizable compound is introduced under pressure as a driving fluid. In yet another embodiment, either a solvent or an oxidizing agent is introduced under pressure as an additional driving fluid along with the oxidizable compound. In yet another embodiment, both the solvent and the oxidizing agent are introduced under pressure as an additional driving fluid along with the oxidizable compound.

The authors suggest that, especially in large reactors, it may be desirable to have many ejectors and their various designs located at different positions in the reaction medium. Structural features will balance the desired overall gradient of constant concentrations of the oxidizable compound with the desired initial dilution of the starting materials of the oxidizable compound, in accordance with other aspects of the present invention. In addition to this, the authors suggest that the jets of the output stream from the ejector can be oriented in any direction. When multiple ejectors are used, each ejector can be oriented individually, again, in any direction.

As discussed above, certain physical and operational features of the bubble column reactor 20 described above with reference to FIG. 1-27 provide gradients in elevation of pressure, temperature and concentrations of reactants (ie, oxygen and an oxidizable compound) in the reaction medium 36. As described above, these elevation gradients can provide a more efficient and economical oxidation method than conventional oxidation methods which well perceive a well-mixed reaction medium with relatively uniform pressure, temperature, and reagent concentrations in it. The height gradients for oxygen, an oxidizable compound (e.g., paraxylene) and temperatures that are made possible by using the oxidation system in accordance with an embodiment of the present invention will now be discussed in more detail.

Turning now to FIG. 28, to quantify the concentration gradients of reactants existing in reaction medium 36 during oxidation in a bubble column reactor type 20, the entire volume of reaction medium 36 can theoretically be divided into 30 discrete horizontal sections of equal volume. 28 illustrates the concept of dividing the reaction medium 36 into 30 discrete horizontal sections of equal volume. With the exception of the highest and lowest horizontal sections, each horizontal section is a discrete volume bounded above and below by imaginary horizontal planes and bounded on its sides by the wall of the reactor 20. The highest horizontal section is bounded on its lower part by an imaginary horizontal plane, and on its the upper part of the upper surface of the reaction medium 36. The lowest horizontal portion is bounded on its upper part imagined horizontally th plane, and on its lower part the lower part of the body. After the reaction medium 36 is theoretically divided into 30 discrete horizontal sections of equal volume, then the time-averaged and volume-averaged concentration of each horizontal section can be determined. A single horizontal section having a maximum concentration of all 30 horizontal sections can be identified as a “C-max horizontal section”. A single horizontal section located above the horizontal C-max section and having a minimum concentration of all horizontal sections located above the C-max horizontal section can be identified as a "C-min horizontal section". The concentration gradient over height can then be calculated as the ratio of the concentration in the horizontal section C-max to the concentration in the horizontal section C-min.

Regarding the quantitative determination of the oxygen concentration gradient, when the reaction medium 36 is theoretically divided into 30 discrete horizontal sections of equal volume, the horizontal O ₂ -max section is identified as having the maximum oxygen concentration of all 30 horizontal sections, and the horizontal O ₂ -min section is identified as having the minimum the oxygen concentration from the horizontal sections located above the horizontal sections O ₂ -max. The oxygen concentrations of the horizontal sections are measured in the gas phase of the reaction medium 36 as a time-averaged and volume-averaged mole concentration in the wet state. It is preferable that the ratio of the oxygen concentrations of the horizontal O ₂ -max portion to the oxygen concentration in the horizontal O ₂ -min portion be in the range of about 2: 1 to about 25: 1, more preferably in the range of about 3: 1 to about 15: 1, and most preferably in the range from 4: 1 to 10: 1.

Typically, the horizontal portion of O ₂ -max will be located near the bottom of the reaction medium 36, while the horizontal portion of O ₂ -min will be located near the top of the reaction medium 36. Preferably, the horizontal portion of O ₂ -min is one of 5 the highest horizontal sections, of 30 discrete horizontal sections. Most preferably, the horizontal portion O ₂ -min is the uppermost portion of 30 discrete horizontal portions, as illustrated in FIG. 28. Preferably, the horizontal portion O ₂ -max is one of the 10 lowest horizontal portions of the 30 discrete horizontal portions. Most preferably, the horizontal portion O ₂ -max is one of the 5 lowest horizontal portions of the 30 discrete horizontal portions. For example, FIG. 28 illustrates the horizontal portion of O ₂ -max as the third horizontal portion from the bottom of reactor 20. It is preferred that the vertical distance between the horizontal portions of O ₂ -min and O ₂ -max be at least about 2W, more preferably at least about 4W, and most preferably at least 6W. It is preferred that the vertical distance between the horizontal portions of O ₂ -min and O ₂ -max is at least about 0.2H, more preferably at least about 0.4H, and most preferably at least 0 , 6H.

The time averaged and volume averaged wet oxygen concentration for the horizontal O ₂ -min portion is preferably in the range of about 0.1 to about 3 mol%, more preferably in the range of about 0.3 to about 2 mol%. and most preferably in the range from 0.5 to 1.5 mol%. The time-averaged and volume-averaged oxygen concentration of the horizontal O ₂ -max portion is preferably in the range of about 4 to about 20 mol%, more preferably in the range of about 5 to about 15 mol%, and most preferably in the range of 6 up to 12 mol% The time-averaged oxygen concentration, in the dry state, in the gaseous effluent released from the reactor 20 through the gas outlet 40, is preferably in the range of about 0.5 to about 9 mol%, more preferably in the range of about 1 to about 7 mol%, and most preferably in the range from 1.5 to 5 mol%.

Since the oxygen concentration attenuates strongly towards the upper part of the reaction medium 36, it is desirable that the oxygen demand in the upper part of the reaction medium 36 is reduced. This reduced oxygen demand near the top of the reaction medium 36 can be achieved by creating a height gradient of the concentration of the oxidizable compound (for example, paraxylene), where the minimum concentration of the oxidizable compound is near the top of the reaction medium 36.

Regarding the quantitative determination of the concentration gradient of an oxidizable compound (e.g., paraxylene), when the reaction medium 36 is theoretically divided into 30 discrete horizontal sections of equal volume, the horizontal section OC-max is identified as having the maximum concentration of the oxidizable compound from all 30 horizontal sections and the horizontal section OC-min identified as having a minimum concentration of oxidizable compound, from horizontal sections located above the horizontal stka OC-max. The concentrations of the oxidizable compound of the horizontal sections are measured in the liquid phase with respect to the time-averaged and volume-averaged mass fraction. It is preferable that the ratio of the concentration of the oxidizable compound of the horizontal portion OC-max to the concentration of the oxidizable compound of the horizontal portion OC-min is greater than about 5: 1, more preferably large, about 10: 1, even more preferably large more than about 20: 1, and most preferably, ranged from 40: 1 to 1000: 1.

Typically, the horizontal portion of OC-max will be located near the bottom of the reaction medium 36, while the horizontal portion of OC-min will be located near the top of the reaction medium 36. Preferably, the horizontal portion of OC-min is one of the 5 topmost horizontal plots of 30 discrete horizontal plots. Most preferably, the horizontal portion OC-min is the uppermost of the 30 discrete horizontal portions, as illustrated in FIG. 28. Preferably, the horizontal portion of the OC-max is one of the 10 lowest horizontal sections of the 30 discrete horizontal sections. Most preferably, the horizontal portion of OC-max is one of the 5 lowest horizontal sections of the 30 discrete horizontal sections. For example, FIG. 28 illustrates the horizontal portion of OC-max as the fifth horizontal portion from the bottom of reactor 20. It is preferred that the vertical distance between the horizontal portions of OC-min and OC-max be at least about 2W, where “W” represents the maximum width of the reaction medium 36. More preferably, the vertical distance between the horizontal sections OC-min and OC-max is at least about 4W, and most preferably at least 6W. For a given height “H” of the reaction medium 36, it is preferable that the vertical distance between the horizontal sections OC-min and OC-max be at least about 0.2H, more preferably at least about 0.4H, and most preferably at least 0.6H.

The time-averaged and volume-averaged concentration of an oxidizable compound (e.g., paraxylene) in the liquid phase of the horizontal section OC-min is preferably less than about 5000 ppm. wt., more preferably, less than about 2000 ppm. wt., even more preferably, less than about 400 ppm. wt., and most preferably is in the range from 1 ppm. wt. up to 100 ppm wt. The time-averaged and volume-averaged concentration of the oxidizable compound in the liquid phase of the horizontal section OC-max is preferably in the range of about 100 ppm. wt. up to about 10,000 ppm wt., more preferably, in the range of about 200 ppm. wt. up to about 5000 ppm wt., and most preferably in the range of 500 ppm wt. up to 3000 ppm wt.

Although it is preferable that the bubble column reactor 20 provides gradients in height of the concentration of the oxidizable compound, it is also preferred that the volume percent of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase is higher than 1000 ppm. wt., was minimized. Preferably, a time-averaged volume percent of reaction medium 36 having a concentration of an oxidizable compound in the liquid phase above 1000 ppm. wt., less than about 9%, more preferably less than about 6%, and most preferably less than 3%. Preferably, a time-averaged volume percent of reaction medium 36 having a concentration of an oxidizable compound in a liquid phase above 2,500 ppm. wt., less than about 1.5%, more preferably less than about 1%, and most preferably less than 0.5%. Preferably, a time-averaged volume percent of reaction medium 36 having a concentration of an oxidizable compound in the liquid phase above 10,000 ppm. wt., less than about 0.3%, more preferably less than about 0.1%, and most preferably less than 0.03%. Preferably, a time-averaged volume percent of reaction medium 36 having a concentration of an oxidizable compound in a liquid phase above 25,000 ppm. wt., less than about 0.03%, more preferably less than about 0.015%, and most preferably less than 0.007%. The authors note that the volume of the reaction medium 36, having elevated levels of the oxidizable compound, does not have to lie in a single continuous volume. In many cases, the chaotic structure of the stream in the vessel of the bubble column reactor type simultaneously form two or more solid, but separated from each other parts of the reaction medium 36 having elevated levels of oxidizable compounds. In each case, the use of time averaging, all such continuous, but separated from each other volumes, greater than 0.0001% vol. from the reaction medium as a whole, are added together to determine the total volume having elevated levels of the concentration of the oxidizable compound in the liquid phase.

In addition to the oxygen and oxidation concentration gradients discussed above, it is preferred that a temperature gradient exist in reaction medium 36. Referring again to FIG. 28, this temperature gradient can be quantified in a manner similar to concentration gradients, by theoretically dividing the reaction medium 36 into 30 discrete horizontal sections of equal volume and measuring the temperature averaged over time and volume averaged over each section. The horizontal section with the lowest temperature of the bottom 15 horizontal sections can then be identified as a horizontal section of T-min, and a horizontal section located above the horizontal section of T-min, and having a temperature maximum of all sections above the horizontal section of T-min, can then be identified as a "T-max horizontal section". It is preferred that the temperature of the horizontal T-max portion is at least about 1 ° C higher than the temperature of the horizontal T-min portion. More preferably, the temperature of the horizontal T-max portion is higher, in the range of about 1.25 to about 12 ° C, than the temperature of the horizontal T-min portion. Most preferably, the temperature of the horizontal portion of T-max is higher, in the range of 2 to 8 ° C, than the temperature of the horizontal portion of T-min. The temperature of the T-max horizontal portion is preferably in the range of about 125 to about 200 ° C, more preferably in the range of about 140 to about 180 ° C, and most preferably in the range of 150 to 170 ° C.

Typically, the horizontal T-max section will be located near the center of the reaction medium 36, while the horizontal T-min section will be located near the bottom of the reaction medium 36. Preferably, the horizontal T-min section is one of the 10 lowest horizontal sections , of the 15 lowest horizontal sections. Most preferably, the T-min horizontal portion is one of the 5 lowest horizontal sections, of the 15 lowest horizontal sections. For example, FIG. 28 illustrates a T-min horizontal portion as a second horizontal bottom portion in the reactor 20. Preferably, the T-max horizontal portion is one of 20 middle horizontal portions of 30 discrete horizontal portions. Most preferably, the horizontal T-min section is one of 14 average horizontal sections, out of 30 discrete horizontal sections. For example, FIG. 28 illustrates the T-max horizontal portion as the twentieth horizontal portion from the bottom of the reactor 20 (i.e., one of the middle 10 horizontal sections). It is preferred that the vertical distance between the horizontal sections of T-min and T-max is at least about 2W, more preferably at least about 4W, and most preferably at least 6W. It is preferred that the vertical distance between the horizontal sections of T-min and T-max is at least about 0.2H, more preferably at least about 0.4H, and most preferably at least 0.6H .

As discussed above, when there is a gradient in height of temperature in the reaction medium 36, it may be advantageous to extract the reaction medium 36 in an elevated position where the temperature of the reaction medium is the highest, especially when the recovered product is subjected to further post-processing at a higher temperature. Thus, when the reaction medium 36 is removed from the reaction zone 28 through one or more elevated outlets, as illustrated in FIG. 19 and 20, it is preferable that the elevated outlet (s) be located near the horizontal T-max portion. Preferably, the elevated outlet is located within 10 horizontal sections of the T-max horizontal section, more preferably within 5 horizontal sections of the T-max horizontal section, and most preferably within 2 horizontal sections of the T-max horizontal section.

Now, we note that many of the features of the present invention described herein can be used in systems with multiple oxidation reactors, and not only in systems using a single oxidation reactor. In addition, certain features of the present invention described herein can be used in agitated reactors and / or flow stirred oxidation reactors, and not just in mixing bubble reactors (i.e. reactor type bubblers). For example, the authors found certain advantages associated with a stepwise change / smooth change in oxygen concentration and / or oxygen consumption rate in the reaction medium. The advantages realized by stepwise changing the concentration / flow rate of oxygen in the reaction medium can be realized regardless of whether the entire volume of the reaction medium is contained in one vessel or in multiple vessels. In addition, the advantages realized by stepwise changing the concentration / flow rate of oxygen in the reaction medium can be realized regardless of whether the reactor is (capacitive) with mechanical stirring, with stirring with a stream and / or bubbled mixing.

One way to quantify the level of a stepwise change in the concentration and / or rate of oxygen consumption in the reaction medium is to compare two or more separate 20% of the continuous volumes of the reaction medium. These 20% continuous volumes need not be determined by any particular form.

However, each 20% continuous volume must be formed from a continuous separate volume of the reaction medium (that is, each volume is “continuous”), and 20% continuous volumes must not overlap with each other (that is, the volumes are “separate”). FIG. 29-31 illustrate that these individual 20% continuous volumes can be located in the same reactor (FIG. 29) or in multiple reactors (FIGS. 30 and 31). Note that the reactors illustrated in FIG. 29-31 may be mechanically agitated, flow agitated, and / or bubble agitated reactors. In one embodiment, it is preferred that the reactors illustrated in FIG. 29-31, were bubbler reactors (i.e., bubble column reactors).

Turning now to FIG. 29, reactor 20 is illustrated as containing a reaction medium 36. The reaction medium 36 includes a first separate 20% continuous volume 37 and a second separate 20% continuous volume 39.

Turning now to FIG. 30, a plurality of reactors is illustrated as comprising a first reactor 720a and a second reactor 720b. The reactors 720a, b together contain the total volume of the reaction medium 736. The first reactor 720a contains the first part of the reaction medium 736a, while the second reactor 720b contains the second part of the reaction medium 736b. The first separate 20% continuous volume 737 of the reaction medium 736 is shown to be defined inside the first reactor 720a, while the second separate 20% continuous volume 739 of the reaction medium 736 is shown to be defined inside the second reactor 720b.

Turning now to FIG. 31, a plurality of reactors is illustrated as comprising a first reactor 820a, a second reactor 820b, and a third reactor 820c. Reactors 820a, b, c together comprise a total volume of reaction medium 836. The first reactor 820a contains a first portion of reaction medium 836a; the second reactor 820b comprises a second portion of the reaction medium 836b; and the third reactor 820c contains a third of the reaction medium 836c. A first separate 20% continuous volume 837 of reaction medium 836 is shown as being defined inside the first reactor 820a; a second separate 20% continuous volume 839 of reaction medium 836 is shown as being defined inside the second reactor 820b; and a third separate 20% continuous volume 841 of reaction medium 836 is shown as being defined inside the third reactor 820c.

A step change in the availability of oxygen in the reaction medium can be quantified by mentioning a 20% continuous volume of the reaction medium having the most excess molar fraction of oxygen in the gas phase, and by referring to a 20% continuous volume of the reaction medium having the most depleted molar fraction of oxygen in the gas phase. In the gas phase of a separate 20% continuous volume of the reaction medium containing the highest concentration of oxygen in the gas phase, the concentration of oxygen averaged over time and volume, in the wet state, is preferably in the range of about 3 to about 18 mol%, more preferably in the range of about 3.5 to about 14 mol%, and most preferably in the range of 4 to 10 mol%. In the gas phase of a single 20% continuous volume of the reaction medium containing the lowest concentration of oxygen in the gas phase, time averaged and volume averaged oxygen concentration, in the wet state, preferably is in the range of about 0.3 to about 5 mol%, more preferably, in the range of about 0.6 to about 4 mol%, and most preferably, in the range of 0.9 to 3 mol%. Moreover, the ratio of the time-averaged and volume-averaged oxygen concentration, in the wet state, in the most excess 20% continuous volume of the reaction medium, compared to the most depleted 20% continuous volume of the reaction medium, is preferably in the range of about 1.5: 1 to about 20: 1, more preferably in the range of about 2: 1 to about 12: 1, and most preferably in the range of 3: 1 to 9: 1.

The stepwise change in the rate of oxygen consumption in the reaction medium can be quantified in terms of oxygen oxygen STR, originally described above. The oxygen STR is previously described in the general sense (that is, from the point of view of the average oxygen STR for the reaction medium as a whole); however, an oxygen STR can also be considered in the local sense (i.e., for a part of the reaction medium) to quantify the stepwise change in the rate of oxygen consumption in the reaction medium.

The authors found that it is most preferable to use to cause a change in STR in oxygen in the reaction medium in combination with the desired gradients described here with respect to the pressure in the reaction medium and to the molar fraction of molecular oxygen in the gas phase of the reaction medium. Thus, it is preferable that the ratio of STR for oxygen of the first separate 20% of the continuous volume of the reaction medium compared to STR for oxygen of the second separate 20% of the continuous volume of the reaction medium is in the range from about 1.5: 1 to about 20: 1, more preferably in the range of about 2: 1 to about 12: 1, and most preferably in the range of 3: 1 to 9: 1. In one embodiment, the “first separate 20% continuous volume” is closer than the “second separate 20% continuous volume” to the position where molecular oxygen is initially introduced into the reaction medium. These large oxygen oxygen gradients are desirable, whether the partial oxidation reaction medium is contained in a bubble column reactor-type oxidation or in any other type of reaction vessel in which pressure gradients and / or mole fractions of molecular oxygen are generated in the gas phase of the reaction medium (for example, in a mechanically agitated container having a plurality of vertically arranged agitated zones obtained by using a plurality of impellers having a strong The flow, possibly expandable by means of generally horizontal deflector assemblies, with an upward flow of oxidizing agent, as a rule, is from bottom to top from the inlet near the lower part of the reaction vessel, despite the fact that significant back mixing of the oxidizer flow can occur within each vertically located zone mixing and that some back-mixing of the oxidizing stream may occur between adjacent upright mixing zones). That is, when there is a gradient of pressure and / or molar fraction of molecular oxygen in the gas phase of the reaction medium, as the authors have discovered, it is desirable to create such a gradient of chemical consumption of dissolved oxygen using the means described here.

Preferred means for forcibly changing the local oxygen STR is to monitor the introduction positions of the oxidizable compound and control the mixing of the liquid phase reaction medium to control the concentration gradients of the oxidizable compound in accordance with other descriptions of the present invention. Other usable means of forcibly changing the local oxygen STR by force include forcing the reaction activity by forcing the temperature locally and by changing the mixture of catalyst and solvent components locally (for example, by introducing additional gas to cause evaporative cooling in a specific part of the reaction medium, and by adding a stream of solvent containing a higher amount of water to reduce activity in a specific part of the reaction medium).

As discussed above with reference to FIG. 30 and 31, the partial oxidation reaction can be carried out in a suitable manner in a plurality of reaction vessels, where at least a portion, preferably at least 25%, more preferably at least 50%, and most preferably at least 75% of the molecular oxygen leaving the first reaction vessel is transferred to one or more subsequent reaction vessels to consume an additional small additive, preferably more than 10%, more preferably more than 20%, and most preferably titelno more than 40% of the molecular oxygen exiting the first / the first reaction vessel during the process. When such a sequential flow of molecular oxygen is used from one reactor to another, it is desirable that the first reaction vessel operate at a higher reaction rate than at least one of the subsequent reaction vessels, preferably with a ratio of oxygen averaged over the vessel STR over the inside the first reaction vessel to the oxygen averaged STR vessel in the subsequent reaction vessel in the range of about 1.5: 1 to about 20: 1, more preferably in the range of about 2: 1 approximately up to 12: 1, and most preferably in the range from 3: 1 to 9: 1.

As discussed above, all types of the first reaction vessel (for example, a bubble column, mechanically stirred, with reverse mixing, with internal step change, with piston flow, and so on) and all types of sequential reaction vessels, which may or may not be of types other than the first reaction vessel, are suitable for use for the sequential flow of molecular oxygen into subsequent reaction vessels in accordance with the present invention. Means of forcibly reducing the oxygen averaged over the oxygen capacity STR for decreasing within subsequent reaction vessels, suitable for use, include decreasing the temperature, decreasing the concentration of the oxidizable compound, and decreasing the reactivity of a particular mixture of catalytic components and solvent (e.g., decreasing the concentration of cobalt, increasing water concentrations and the addition of a catalytic inhibitor, such as small amounts of ionic copper).

When flowing from the first reaction vessel to the subsequent reaction vessel, the oxidizing stream can be processed by any means known in the art, such as compressing or depressurizing, cooling or heating and removing the mass or adding mass of any quantity or any type. However, the use of reducing the oxygen averaged over the oxygen capacity STR in the subsequent reaction vessel is particularly suitable for use when the absolute pressure at the top of the first reaction vessel is less than about 2.0 MPa, more preferably less than about 1.6 MPa and most preferably less than 1.2 MPa. Moreover, the use of decreasing the oxygen averaged over the oxygen capacity STR in the subsequent reaction vessel is particularly suitable for use when the ratio of the absolute pressure in the upper part of the first reaction vessel compared to the absolute pressure in the upper part of the at least one subsequent reaction vessel in the range of about 0.5: 1 to 6: 1, more preferably in the range of about 0.6: 1 to about 4: 1, and most preferably in the range of 0.7: 1 to 2: 1. Lowering the pressure in the subsequent tank below these lower limits unnecessarily reduces the availability of molecular oxygen, and increasing the pressure above these upper limits is all the more expensive compared to using a fresh oxidizing agent.

When a sequential stream of molecular oxygen is used in a subsequent reaction vessel having a reduced oxygen averaged STR over the oxygen vessel, streams of fresh starting materials and an oxidizable compound, solvent, and oxidizing agent can flow into the subsequent reaction vessel and / or into the first reaction vessel. Flows of the liquid phase and solid phase, if present, of the reaction medium can flow in any direction between the reaction vessels. The entire gas phase or part thereof, leaving the first reaction vessel and entering the subsequent reaction vessel, can proceed separately or mix with a part of the liquid or solid phase, if present, of the reaction medium from the first reaction vessel. The flowing stream of a product containing a liquid phase and a solid phase, if present, can be removed from the reaction medium in any reaction vessel in the system.

Referring again to FIG. 1-29, in one embodiment of the present invention, an oxidation reactor type bubble column 20 has a significantly higher production rate than conventional oxidation reactor type bubble columns, specifically ordinary reactor type bubblers used to produce terephthalic acid. To ensure an increase in the production rate, the size of the bubble column reactor type 20 should be increased. However, the dynamics of the multiphase fluid flow with natural convection for the reaction medium in such a scaled bubble column reactor-type column is significantly different for the worse from the flow dynamics in conventional smaller reactors. It was found that certain design and operating parameters are important for the proper functioning of a scaled bubble column reactor-type oxidation with a high production rate.

When the bubbler column 20 is a scalable reactor oxidation type bubbler with a high production rate in accordance with one embodiment of the present invention, the height “H” of the reaction medium 36 is preferably at least about 30 m, more preferably in the range of about 35 to about 70 m, and most preferably in the range of 40 to 60 m. The density and height of the reaction medium 36 in the scaled bubble column reactor type 20 causes a significant difference phenomena between the upper part of the reaction medium 36 and the lower part of the reaction medium 36. Preferably, the pressure differential between the upper and lower parts of the reaction medium 36 is at least about 1 bar, more preferably at least about 1.4 bar, and most preferably in the range of 1.6 to 3 bar. The maximum width “W” of the reaction medium 36 is preferably at least about 2.5 m, more preferably in the range of about 3 to about 20 m, even more preferably in the range of about 3.25 to about 12 m, and most preferably in the range of 4 to 10 m. The H: W ratio of the reaction medium 36 is preferably at least about 6: 1, more preferably in the range from about 8: 1 to about 20: 1, and most preferably in the range from 9: 1 to 15: 1. The total volume of the reaction medium 36 and / or reaction zone 28 is preferably at least about 250 m ³ , more preferably at least about 500 m ³ , and most preferably at least 1000 m ³ .

During operation of the scaled oxidation reactor bubbler column 20, it is preferable that the time-averaged surface velocity of the gas phase of the reaction medium 36 at one fourth height, half height and / or three fourth height is in the range of about 0.8 to about 5 m / s, more preferably in the range of about 0.9 to about 3 m / s, and most preferably in the range of 1 to 2 m / s. When a scaled oxidation type reactor bubble column 20 is used to produce terephthalic acid by partial oxidation of paraxylene, it is preferred that paraxylene is introduced into reaction zone 28 at a rate of at least about 11,000 kg / h, more preferably at a speed ranging from about 20,000 up to about 100,000 kg / h, and most preferably in the range from 30,000 to 80,000 kg / h. The terephthalic acid production rate for the scaled bubble column 20 is preferably at least about 400 tons per day, more preferably at least about 700 tons per day, and most preferably in the range of 750 to 3000 tons per day. Other structural and operational parameters of the scaled oxidation reactor bubble column 20 may be substantially the same as originally described above with reference to FIG. 1-29.

It has been found that varying the horizontal cross-sectional area of the reaction zone of a reactor-type bubble column along the height of the reactor can help improve fluid dynamics for the reaction medium, especially in the scaled oxidation-type bubble column design discussed above. Turning now to FIG. 32, in one embodiment of the present invention, the oxidation reactor type bubbler column housing 22 includes a wide lower section 23, a narrow upper section 25, and a transition section 27. Preferably, the lower and upper sections 23, 25 are essentially cylindrical in shape and are located along a common central vertically axis. The transition section 27 may have a variety of usable shapes (e.g., horizontal flat shape, 2: 1 elliptical shape, hemispherical shape, and so on). Preferably, the transition section 27 is typically a diverging conical element that extends into the housing 22 from the wide lower section 23 to the narrow upper section 25. The lower section 23 of the housing 22 defines a wide lower reaction zone 29. The upper section 25 of the housing 22 defines a narrow the upper reaction zone 31. The transition section 27 defines a transition zone located between the lower and upper reaction zones 29, 31. The lower reaction zone 29, the upper reaction zone 31 and the transition zone together form a common reaction the zone of the bubble column of the reactor type of oxidation 20, which receives a multiphase reaction medium 36.

In a preferred embodiment of the present invention, the maximum horizontal cross-sectional area of the reaction medium 36 in the lower reaction zone 29 is at least about 10% larger than the minimum cross-sectional area of the reaction medium 36 in the upper reaction zone 31. More preferably, the maximum horizontal area is the cross section of the reaction medium 36 in the lower reaction zone 29 is higher, in the range of about 25 to about 210%, than the minimum horizontal cross-sectional area I of the reaction medium 36 in the upper reaction zone 31. Most preferably, the maximum horizontal cross-sectional area of the reaction medium 36 in the lower reaction zone 29 is higher, ranging from 35 to 160%, than the minimum horizontal cross-sectional area of the reaction medium 36 in the upper reaction zone 31 .

As illustrated in FIG. 32, the lower reaction zone 29 has a maximum diameter “D ₁ ” that is larger than the minimum diameter “D _u ” of the upper reaction zone 31. Preferably, D ₁ is at least about 5% greater than D _u . More preferably, D _{1 is} greater, in the range of about 10 to about 100%, than D _u . Most preferably, D _{1 is} greater, in the range of 15 to 50%, than D _u . 32 also illustrates that the lower reaction zone 29 has a maximum height of "L ₁ ", the upper reaction zone 31 has a maximum height of "L _u " and the transition zone has a maximum height of "L _t ". It should be noted that the height of the portion of the reaction medium 36 contained in the lower reaction zone 29 is L ₁ , and the height of the portion of the reaction medium 36 contained in the transition zone is L _t . In one embodiment, the height of the portion of the reaction medium 36 located in the upper reaction zone 31 is L _u . However, in certain cases, the height of the reaction medium 36 in the upper reaction zone 31 may be less than L _u . In other cases, the total height of the reaction medium 36 may extend above the upper edge 50 of the upper reaction zone 31 (i.e., the total height of the reaction medium 36 is greater than the sum of L _l plus L _t plus L _u ). Preferably, the total height of the reaction medium 36 is within 50, 25, or 10% of L _u , as measured up or down from the upper level of the surface 50 of the upper reaction zone 31. Preferably, the oxidation reactor type bubble column 20 has an L _l : L _u ratio in the range of about 0.05: 1 to about 5: 1, more preferably in the range of about 0.1: 1 to about 2.5: 1, and most preferably in the range of 0.15: 1 to 1.5: one. Preferably, the oxidation reactor type bubble column 20 has an L ₁ : D ₁ ratio _of greater than about 0.5: 1, more preferably in the range of about 1: 1 to about 10: 1, and most preferably in the range of 1, 5 to 8: 1. The oxidation reactor type bubble column 20 preferably has a L _u : D _u ratio of greater than about 2: 1, more preferably in the range of about 2.5: 1 to about 20: 1, and most preferably in the range of 3: 1 up to 15: 1. In a preferred embodiment of the present invention, L ₁ is at least about 2 m, more preferably L ₁ is in the range of about 4 to about 50 m, and most preferably in the range of 6 to 40 m. Preferably, L _u is, at least about 4 m, more preferably in the range of about 5 to about 80 m, and most preferably in the range of 10 to 60 m. Preferably, D ₁ is in the range of about 2 to about 12 m, more preferably is within approximately from 3.1 to about 10 m, and most preferably in the range from 4 to 9 m.

32 also illustrates that the oxidation reactor type bubble column 20 has a separation section 26 located above the upper reaction section 31. The separation section 26 defines the separation zone 30. As shown in FIG. 32, the separation zone 30 has a maximum height of “Y” and maximum width is "X". It is preferred that the oxidation reactor type bubble column 20 have an X: D ₁ ratio in the range of about 0.8: 1 to about 4: 1, most preferably in the range of 1.1: 1 to 2: 1. The oxidation reactor type bubble column 20 preferably has a L _u : Y ratio in the range of about 1: 1 to about 10: 1, most preferably in the range of 2: 1 to 5: 1. The bubble column reactor oxidation type 20 preferably has a L _u : Y ratio in the range of about 0.4: 1 to about 8: 1, most preferably in the range of 0.6: 1 to 4: 1. The transition section 27 of the housing 22 has a maximum diameter "D ₁ " next to the bottom section 23 and a minimum diameter "D _u " next to the upper section 25. The bubble column reactor oxidation type 20 preferably has a ratio of L _t : D ₁ ranging from about 0, 05: 1 to about 5: 1, most preferably in the range of 0.1: 1 to 2: 1.

The vertically varying horizontal cross-sectional area of the oxidation reactor type bubble column 20 illustrated in FIG. 32 provides a portion of the reaction medium 36 contained in the upper reaction zone 31 with a higher surface gas velocity and a higher gas content than part of the reaction medium 36 contained in the lower reaction zone 29. Preferably, the time-averaged surface gas velocity for a portion of the reaction medium 36 contained in the upper reaction zone 31 per floor the pore height of the reaction medium in the upper reaction zone 31 is at least about 15% greater than the time-averaged surface gas velocity of a portion of the reaction medium 36 contained in the lower reaction zone 29 at half the height of the reaction medium in the lower reaction zone 29. More preferably, a portion of the reaction medium 36 contained in the upper reaction zone 31 has a time-averaged surface gas velocity at half the height of the reaction medium in the upper reaction zone 31, which is greater, within approximately 25% to about 210% than the time-averaged surface gas velocity of a portion of the reaction medium 36 contained in the lower reaction zone 29 at half the height of the reaction medium in the lower reaction zone 29. Most preferably, the time-averaged gas surface velocity of a portion of the reaction medium 36 contained in the upper reaction zone 31 at half the height of the reaction medium in the upper reaction zone 31 is greater, ranging from 35 to 160%, than the time-averaged surface velocity of the gas of part of the reaction medium 3 6 contained in the lower reaction zone 29 at half the height of the reaction medium in the lower reaction zone 29. Preferably, the time-averaged and volume-averaged gas content of a portion of the reaction medium 36 contained in the upper reaction zone 31 is at least about 4% more than the time-averaged and volume-averaged gas content of a portion of the reaction medium 36 contained in the lower reaction zone 29. More preferably, the time-averaged and volume-averaged gas content of a portion of the reaction medium the water 36 contained in the upper reaction zone 31 is greater, ranging from about 6 to about 80%, than the time-averaged and volume-averaged gas content of a portion of the reaction medium 36 contained in the lower reaction zone 29. Most preferably, averaged over time and the volume-averaged gas content of the portion of the reaction medium 36 contained in the upper reaction zone 31 is greater, ranging from 8 to 70%, than the time-averaged and volume-averaged gas content of the portion of the reaction medium 36 contained in the lower reaction Zone 29.

Although FIG. 32 illustrates a very specific two-stage bubble column design with cylindrical side walls, it should be understood that many other structures may fall within the scope of this embodiment of the present invention. For example, the narrow upper and wide lower sections of the reactor can be formed from one or more inclined side walls and / or from a plurality of segments of the side walls with a stepwise change in diameter. In any case, the essence of the present embodiment comprises the claims, and not the drawing.

As discussed above, in certain cases, it may be desirable to use larger reactor bubbler columns to allow higher production rates for a single reactor. However, when the size of the bubbler columns of the reactor type increases, the behavior of the fluid flow of the multiphase reaction medium contained therein is significantly different from the behavior of the flow in smaller reactors. It has been found that a change in the behavior of the fluid flow in large bubbler columns of a reactor type can be compensated by contacting the multiphase reaction medium contained in the bubbler column of the reactor type with additional vertically arranged surfaces. Accordingly, FIG. 33-44 illustrate various methods for creating additional vertically arranged surface area in the reaction zone of a scaled bubble type reactor column. Each of the bubbler columns of the reactor type illustrated in FIG. 33-44, contains one or more vertically located internal elements contained in the reaction zone of a bubble column reactor type. These vertically arranged internal elements are created in addition to the vertically supported pressure walls of the reactor. As previously discussed, the reaction medium contained in the bubble column reactor type has a maximum height “H” and a maximum width “W”. The minimum width of the reaction medium created above the height H / 5 is referred to herein as “Wmin”. In a preferred embodiment of the present invention, the total vertical surface area of the reactor-type bubbler column that comes into contact with the reaction medium is preferably greater than about 3.25 WminH, more preferably in the range of about 3.5 WminH to about 20 WminH, even more preferably in the range of about 4WminH to about 15WminH, and most preferably in the range of 5WminH to 10WminH. This total vertically located surface area that comes into contact with the reaction medium includes the entire area of vertically located surfaces, including vertically located surfaces of the pressure-holding side walls of the reactor and vertically located surfaces of any internal elements present within the reaction zone. As used here, the term "located vertically" means less than 45 ° from the vertical. Preferably, vertically contacting surfaces in a bubbler column reactor extend at an angle within 30 ° from the vertical, more preferably within 15 ° from the vertical, even more preferably within 5 ° from the vertical, and most preferably are essentially vertical.

In addition, it is preferable that the total value of the vertically located surface area coming into contact with the reaction medium, which is attributed to the internal elements that do not hold pressure, be at least about 10% of the total value of the vertically located surface area that comes into contact with contact with the reaction medium, which is attributed to the pressure-holding side walls of the vessel. More preferably, the total exposed vertical surface area represented by the internal elements and coming into contact with the reaction medium is in the range of about 15 to about 600% of the total exposed vertical surface area represented by the pressure-holding side walls and coming into contact with the reaction medium, even more preferably in the range of from about 25 to about 400%, and most preferably in the range of from 35 to 200%. The presence of an additional vertically arranged surface area during oxidation in a bubble column can make lower H: W ratios possible than would be possible with conventional reactor-type bubble columns having a small additional vertical surface area or without it. Thus, it is preferable that the H: W ratio of the reactor type bubble column using an additional vertically arranged surface area is in the range of about 3: 1 to about 20: 1, more preferably in the range of about 3.5: 1 to about 15: 1, and most preferably in the range from 4: 1 to 12: 1.

Turning now to the embodiment illustrated in FIG. 33 and 34, the oxidation reactor type bubble column 20 may comprise one separation wall 33 extending diametrically from one side of the side wall 46 to the opposite side of the side wall 46. The separation wall 33 is located above the distribution device 34, so that substantially the entire flow an oxidizing agent is introduced below the lower part of the separation wall 33. Thus, approximately half of the gas phase of the reaction medium 36, which contains the undissolved part of the oxidizing stream, flows from bottom to top for each side of the separation wall 33. Preferably, approximately half of the feed stream is introduced through the feed input 32a on one side of the separation wall 33, and the other half of the feed stream is introduced through the feed input 32b on the other side of the separation wall 33. Height of the separation wall 33 essentially equal to the height of the cylindrical side wall 46.

The separation wall 33 divides the reaction zone 28 into approximately two halves, wherein the reaction medium 36 is located on each side of the separation wall 33. Essentially, the entire vertical surface area of the reactor 20, which comes into contact with the reaction medium 36, is attributed to the internal exposed surfaces of the side wall 46 and the exposed exposed surfaces of the separation wall 33. If the separation wall 33 is not present in the reaction zone 28, then substantially all of the vertically located In fact, the surface area in contact with the reaction medium 36 would be attributed to the side walls 46 of the pressure-holding body 22. The separation wall 33 provides an additional surface area that affects the dynamics of the fluid flow of the reaction medium 36 and makes it possible to scale the bubble column of the oxidation type reactor 20 without significant negative impact on the performance of the reactor.

Turning now to the embodiment illustrated in FIG. 35, an oxidation reactor type bubbler column 20 is illustrated as having a shortened separation wall 35. The separation wall 35 in FIG. 35 is similar to the separation wall 33 of FIG. 33 and 34; however, the dividing wall 35 in FIG. 35 has a height that is significantly less than the total height of the reaction medium 36. In the configuration illustrated in FIG. 35, it is preferable that substantially the entire feed stream and oxidizer stream are introduced into the reaction zone 28 below the lower part of the partition wall 35. The upper part of the partition wall 35 is preferably at a sufficient distance from the upper surface 44 of the reaction medium 36. In this configuration, both halves of the reaction medium 36 are arranged as desired. side of the divider wall 35 can mix with one another above and below the partition wall 35.

Turning now to the embodiment illustrated in FIG. 36 and 37, the oxidation reactor type bubble column 20 may comprise a non-planar separation wall 37, which allows a significant amount of vertical surface area to be added to the reactor 20, without the need for many additional internal elements in the reaction zone 28. Like the separation walls 33, 35 in FIG. 33-35, the separation wall 37 in FIG. 36 and 37 are connected to and diametrically opposed inner surfaces of the side wall, side wall 46, and extend between them.

In the embodiment illustrated in FIG. 38 and 39, an oxidation reactor type bubble column 20 comprises a vertically disposed inner element 41 having a generally X-shaped configuration. The outer vertical edges of the inner element 41 are some distance inward from the inner surfaces of the side wall 46, so that the reaction medium 36 can flow between the individual quadrants defined by the X-shaped inner element 41. The various exposed vertical surfaces of the inner element 71 add a significant amount of area a surface that comes into contact with the reaction medium 36.

FIG. 40-42 illustrate an embodiment where a part of the reaction zone 28 is divided into 4 vertical quadrants by the inner element 53, while another part of the reaction zone 28 is divided into 8 vertical wedge-shaped sections by the inner element 55. As illustrated in FIG. 40, the reaction zone 28 varies vertically from dividing into 4 vertical quadrants with the help of inner element 53 and up to 8 vertical wedge-shaped sections by means of inner element 55.

Turning now to FIG. 43 and 44, the oxidation reactor type bubble column 20 is illustrated as having a plurality of internal elements, generally helical, 61a, b, c, d and a vertical X-shaped spacer 63 located above the helical elements 61. The helical elements 61 provide oblique outer surfaces that create a vortex flow pattern in the upwardly flowing portion of the reaction medium 36. It is preferable that the tilt direction of the screw elements 61 is such that adjacent screw-like elements 61 caused a swirl of the reaction medium 36, in general, in opposite directions. Thus, if the screw element 61a causes the reaction medium to rotate clockwise when the reaction medium 36 rises in the reaction zone 28, the screw element 61b (located directly above the screw element 61a) causes the bottom-up reaction medium 36 to spin counterclockwise. The vertical internal separation element 93 adds an additional vertically arranged surface area to the oxidation reactor type bubble column 20, and can also function to minimize swirl / turbulence of the reaction medium 36 as the gas phase rises toward the upper surface 44 of the reaction medium 36.

No matter what configuration is illustrated in FIG. 33-44, used in an oxidation type reactor bubble column 20, it is preferred that the oxidizing stream and the feed stream are introduced into the reaction zone 28 in such a way that a substantial portion of molecular oxygen and the oxidizable compound are introduced into the reaction zone 28 below a significant portion of the vertically arranged inner element or elements. Preferably, at least about 50% wt. the total molecular oxygen and oxidizable compounds introduced into the reaction zone 28 enter the reaction zone 28 below at least 50% of the vertically exposed external surface area of the inner element or elements, more preferably at least about 75% wt. molecular oxygen and the oxidizable compound enter the reaction zone below at least 50% of the vertically exposed external surface area of the inner element (s), even more preferably at least about 90% wt. molecular oxygen and the oxidizable compound enter the reaction zone below at least 50% of the vertically exposed outer surface area of the inner element (s), and most preferably, substantially all molecular oxygen and the oxidizable compound enter the reaction zone below at least at least 50% of the vertically exposed outer surface area of the inner element (s). In addition, it is preferable that the oxidizing stream and the feed stream are introduced so that a substantial portion of the gas phase of the reaction medium 36 flows from bottom to top on all sides of the additional exposed external surface area created by the internal element or elements. In addition, it is preferred that oxidant and feed streams are introduced into reaction zone 28 in accordance with the radial and azimuthal distribution schemes described above.

Although some oxidation reactors known from the literature may use heat transfer surfaces that come into contact with the reaction medium in a manner similar to the internal element (s) described herein, it should be noted that it is undesirable for the internal element (s) of the present invention to create or a significant level of heating or cooling of the reaction medium. Thus, it is preferable that the heat flux of the exposed (i.e., coming into contact with the reaction medium) vertically arranged surfaces of the inner element (s) described herein is less than about 30,000 W / m ² .

FIG. 45-53 illustrate embodiments of the present invention where an oxidation reactor type bubble column 20 is provided with one or more deflectors that come into contact with a multiphase reaction medium 36 to facilitate improved oxidation with minimal impurity formation. The baffles are particularly suitable for use in the scaled reactor bubble column designs described above. In each of the deflector bubbler columns of the reactor type 20 illustrated in FIG. 45-53, it is preferred that the baffle or baffles have an open area in the range of about 10 to about 90%. As used here, the percentage of open area of the deflector means the minimum percentage of the horizontal cross-sectional area of the reaction zone, which is open (that is, not filled with the structure of the deflector), in the vertical position of the deflector. More preferably, the open area of the deflector or deflectors illustrated in FIG. 45-53, ranges from about 25 to about 75%, most preferably from 35 to 65%.

An important feature of the deflectors illustrated in FIGS. 45-53 is the ability of the deflectors to resist the formation of sediment. As previously discussed, an oxidation reactor type bubbler column 20 is preferably used in the course of precipitation oxidation, when solid products are formed in the reaction medium 36 during oxidation. Deflectors having a significant amount of approximately horizontal, upward looking flat surface area are prone to the formation of sediment in a reactor operating under deposition conditions. When deflectors are coated with sediment, solid products are deposited on upward-facing surfaces of the deflectors, and when the amount of solid products deposited on the deflectors increases, pieces of precipitated solid products can leave the deflectors and fall towards the bottom of the reactor. These pieces of fallen solid products can be deposited at the bottom of the reactor and cause a number of problems, including, for example, slowing the release of the suspension from the bottom of the reactor.

In view of the above, it is preferable, in one embodiment, that the baffle or baffles used in the bubble column reactor type oxidation 20, do not provide upward looking flat outer surfaces (for example, the baffle may be constructed of tubular materials having a circular cross section ) Unless otherwise specified here, an upward facing surface is a surface having a normal vector projecting above the horizontal. In another embodiment, a small amount of substantially flat surfaces can be used insofar as less than about 50% of the total upwardly exposed (i.e., coming into contact with reaction medium 36) outer surface area of the deflector or deflectors is attributed essentially flat surfaces deviated less than 30 ° or 20 °, or even 10 ° from the horizontal. More preferably, less than about 35% of the total exposed upwardly exposed outer surface area of the deflector or deflectors is attributed to flat surfaces deviated less than 30 °, or 20 °, or even 10 ° from the horizontal. Most preferably, less than 25% of the total exposed upwardly exposed external surface area of the deflector or deflectors is attributed to substantially flat surfaces deflected less than 30 °, or 20 °, or even 10 ° from the horizontal. In addition, it is preferable that the upwardly exposed exposed surfaces of the deflector or deflectors have a substantially smooth finish so as to further resist the formation of sediment. Preferably, at least a substantial portion of the exposed exposed surfaces of the baffle or baffles have a surface roughness of less than about 125 μm RMS (rms), more preferably less than about 64 μm RMS, and most preferably less than 32 μm RMS. The electropolishing finish and the “2B” smooth finish in the roll mill are particularly suitable for use.

In addition to the design of the deflector or deflectors where no precipitate is formed, illustrated in FIG. 45-53, it is preferred that the baffle (s) are located at an appropriate distance from the upper and lower surface levels of the reaction medium 36, so as to provide optimized efficiency. In a preferred embodiment of the present invention, the baffle or baffles are spaced at least 0.5 W and / or 0.05 H from both the upper and lower surface levels of the reaction medium 36, where “W” is the maximum width of the reaction medium 36 and “H” represents the maximum height of the reaction medium 36. More preferably, the deflector or deflectors are spaced at least 1W and / or 0.1H from both the upper and lower surface levels of the reaction medium 36. Most preferably preferably, the baffle or baffles are spaced at least 1.5 W and / or 0.15 H from both the upper and lower surface levels of the reaction medium 36. The presence of a baffle or baffles in a bubble column of an oxidation type reactor 20 may make it possible lower H: W ratios than would be possible in similar reactors not equipped with baffles. Thus, it is preferred that the H: W ratio of reactor-type bubbler columns equipped with baffles is in the range of about 3: 1 to about 20: 1, more preferably in the range of about 3.5: 1 to about 15: 1, and most preferably in the range of 4: 1 to 12: 1.

Turning now to FIG. 45-47 in more detail, an oxidation reactor-type bubbler column 20 is illustrated as containing a plurality of vertical deflectors 71a, b, c, d separated from each other. Preferably, the oxidation reactor type bubble column 20 comprises within the range of 2 to 20 deflectors separated by a certain vertical distance, most preferably from 3 to 8 deflectors separated by a certain vertical distance. Preferably, each deflector 71 comprises a plurality of elongated individual elements of the deflector 73. In this embodiment, each individual element of the deflector 73 provides a substantially cylindrical exposed outer surface that comes into contact with the reaction medium 36. In the embodiment illustrated in FIG. 45-47, the deflectors 71a, b, c, d rotate with respect to each other, so that the individual elements of the deflectors 73 for adjacent deflectors 71 extend substantially perpendicular to each other.

Turning now to FIG. 48-50, in more detail, alternative deflectors 81a, b, c, d are illustrated as having a plurality of elongated individual deflector elements 83. In this embodiment, each deflector element 83 is formed from an L-shaped element and represents a generally inverted V-shaped inverse upward exposed surface. The configuration of the exposed outer surfaces of the individual elements of the deflectors 83 helps to prevent the formation of sediment during the deposition of solid products in the reaction medium 36. The number of iron elements of the deflector 83, the distance between them and the orientation angle can be essentially the same as described above for the cylindrical elements of the deflectors 73 in FIG. 45-47.

Turning now to FIG. 51-53 in more detail, an oxidation reactor-type bubbler column 20 is illustrated as containing a separate monolithic deflector 91 having a generally shaped two vertical cones 93a, b extending in opposite directions connected at their base. The inclination of the exposed upwardly facing outer surface of the monolithic baffle 91 helps to prevent the formation of sediment on the baffle 91 due to solid products settling in the reaction medium 36.

The various deflector configurations illustrated in FIG. 45-53 are only examples, and many other deflector configurations may fall within the scope of the present invention. It should also be noted that the deflector configurations illustrated in FIG. 45-53, can be used in combination.

Although some oxidation reactors known from the literature may use heat transfer tubes that come into contact with the reaction medium in a manner similar to the deflector (s) described herein, it should be noted that it is undesirable for the deflectors of the present invention to produce any significant level of heat or cooling the reaction medium. Thus, it is preferable that the heat flux of the exposed (i.e., coming into contact with the reaction medium) external surfaces of the baffles described herein be less than about 30,000 W / m ² .

Referring again to FIG. 1-53, the oxidation is preferably carried out in a bubble column reactor type 20 under conditions that are noticeably different than for conventional oxidation reactors, in accordance with the preferred embodiments described herein. When a reactor type bubbler column 20 is used to effect liquid phase partial oxidation of paraxylene to crude terephthalic acid (CTA) in accordance with the preferred embodiments described herein, spatial profiles of local reaction intensity, local evaporation rate and local temperature, in combination with liquid flow patterns in reaction medium and preferred relatively low oxidation temperatures contribute to the formation of CTA particles having unique nye and advantageous properties.

FIG. 54A and 54B illustrate CTA core particles obtained in accordance with one embodiment of the present invention. Fig. 54A shows the main particles of the CTA at a 500-fold increase, while Fig. 54B enlarges one of the main particles of the CTA and shows this particle at a 2000-fold magnification. As probably best illustrated in FIG. 54B, each CTA core particle is typically formed from a large number of small agglomerated CTA subparticles, thereby producing a CTA core particle with a relatively large surface area, high porosity, low density and good solubility. CTA core particles typically have an average particle size in the range of about 20 to about 150 microns, more preferably in the range of about 30 to about 120 microns, and most preferably in the range of 40 to 90 microns. CTA subparticles typically have an average particle size in the range of about 0.5 to about 30 microns, more preferably about 1 to about 15 microns, and most preferably in the range of 2 to 5 microns. The relatively high surface area of the main CTA particles illustrated in FIG. 54A and 54B can be quantified using the Braunauer-Emmett-Teller (BET) surface area measurement method. Preferably, the basic CTA particles have an average BET surface of at least about 0.6 m ² / g. More preferably, the basic CTA particles have an average BET surface area in the range of about 0.8 to about 4 m ² / g. Most preferably, the basic CTA particles have an average BET surface area in the range of 0.9 to 2 m ² / g. Physical properties (e.g., particle size, BET surface area, porosity and solubility) of the main CTA particles formed by the optimized oxidation method of the preferred embodiment of the present invention make it possible to clean the CTA particles by more efficient and / or economical methods, as described in more detail below, in connection with Fig. 57.

The average particle sizes given above are determined using polarized light microscopy and image analysis. Particle size analysis equipment includes a Nikon E800 optical microscope with a 4 × Plan Flour NA 0.13 lens, a Spot RT ™ digital camera and a personal computer using Image Pro Plus ™ V4.5.0.19 image analysis software . The method for analyzing particle sizes includes the following main steps: (1) dispersing CTA powders in mineral oil; (2) preparing a dispersion preparation between a microscope slide and a coverslip; (3) object research using polarized light microscopy (conditions of crossed polarizers - particles appear as bright objects on a black background); (4) recording various images for the preparation of each sample (field size = 3 × 2.25 mm; pixel size = 1.84 μm / pixel); (5) image analysis using Image Pro Plus ™ software; (6) transferring particle measurements to the table; and (7) performing statistical processing of the table. Stage (5) of "performing image analysis using Image Pro Plus ™ software" includes the steps of: (a) setting an image threshold for detecting white particles against a dark background; (b) creating a binary image; (c) triggering a one-pass open filter to filter out pixel noise; (d) measuring all particles in the image; and (e) recording the average diameter measured for each particle. Image Pro Plus ™ software defines the average diameter of individual particles as the number-average length of particle diameters measured at 2 ° intervals and when passing through a centroid of a particle. Stage 7 "implementation of the statistical characterization in the table" includes calculating the volume-weighted average particle size as follows. The volume of each of n particles in the sample is calculated as if they were spherical using the expression pi / 6 · d _i ^ 3; multiplying the volume of each particle by its diameter, with the expression pi / 6 · d _i ^ 4; summing for all particles in the sample the values pi / 6 · d _i ^ 4; summing the volumes of all particles in the sample and calculating the volume-weighted particle diameter as the sum of the expressions (pi / 6 · d _i ^ 4) for all n particles in the sample divided by the sum of the expressions (pi / 6 · d _i ^ 3) for all n particles in the sample. As used herein, “average particle size” refers to a volume-weighted average particle size determined in accordance with the above-described research method; and it is also referred to as D (4.3).

In addition to this, step 7 includes finding particle sizes for which various fractions of the total volume of the sample are smaller. For example, D (v, 0.1) is the particle size for which 10% of the total sample volume is smaller and 90% are large; D (v, 0.5) is the particle size for which half of the sample volume is large and half smaller; D (v, 0.9) is the particle size for which 90% of the total sample volume is smaller; and so on. In addition to this, step 7 includes calculating the value of D (v, 0.9) minus D (v, 0.1), which is defined here as a “particle size spread”; and step 7 includes calculating a particle size dispersion value divided by D (4.3), which is defined herein as “relative particle size dispersion”.

In addition, it is preferable that D (v, 0.1) CTA particles, as measured above, be in the range of about 5 to about 65 microns, more preferably in the range of about 15 to about 55 microns, and most preferably in the range from 25 to 45 microns. It is preferred that the D (v, 0.5) CTA particles, as measured above, are in the range of about 10 to about 90 microns, more preferably in the range of about 20 to about 80 microns, and most preferably in the range of 30 to 70 microns. It is preferred that the D (v, 0.9) CTA particles, as measured above, are in the range of about 30 to about 150 microns, more preferably in the range of about 40 to about 130 microns, and most preferably in the range of 50 to 110 microns. It is preferred that the relative variation in particle size is in the range of about 0.5 to about 2.0, more preferably in the range of about 0.6 to about 1.5, and most preferably in the range of 0.7 to 1.3 .

The BET surface area values given above are measured on a Micromeritics ASAP2000 (available from Micromeritics Instrument Corporation of Norcross, GA). In the first stage of the measurement process, 2 to 4 grams of a sample of particles are weighed and dried in vacuum at 50 ° C. The sample is then placed in a gas manifold for analysis and cooled to 77 K. The nitrogen adsorption isotherm is measured at a minimum of five equilibrium pressures by exposing the sample to known volumes of nitrogen gas and measuring the pressure decrease. The equilibrium pressures are approximately in the range of P / P ₀ = 0.01-0.20, where P is the equilibrium pressure and P ₀ is the vapor pressure of liquid nitrogen at 77 K. The resulting isotherm can then be plotted as with the following BET equation:

where V _a represents the volume of gas adsorbed by the sample at P, V _m represents the volume of gas required to cover the entire surface of the sample with a gas monolayer, and C is a constant. V _m and C are determined from this graph. Then V _{m is} converted to the surface area using the nitrogen cross-sectional area at 77 K, by:

where σ is the nitrogen cross-sectional area at 77 K, T is 77 K and R is the gas constant.

As is clear from the above, the CTA obtained in accordance with one of the embodiments of the present invention, demonstrates excellent dissolution properties compared with conventional CTA obtained by other methods. This increased dissolution rate makes it possible to purify the CTA of the present invention by more efficient and / or more productive purification methods. The following description describes a method by which the dissolution rate of CTA can be quantified.

The dissolution rate of a known amount of solid products in a known amount of solvent in a stirred mixture can be measured by various protocols. As used here, a measurement method called "dissolution time dependence study" is defined as follows. During the study of the time dependence of dissolution, an ambient pressure of about 0.1 MPa is used. The ambient temperature used during the dissolution time study is approximately 22 ° C. In addition, solids, solvent, and the entire dissolution apparatus are thermally balanced completely at this temperature prior to the start of the study, and there is no noticeable heating or cooling of the beaker or its contents during the dissolution time period. A portion of solvent from fresh tetrahydrofuran, analytical grade for HPLC (purity> 99.9%), then 250 grams THF, is placed in a clean, 400 ml KIMAX glass beaker (Kimble® part number 14020, Kimble / Kontes, Vineland, NJ ), which is non-insulated, has smooth walls and, in general, a cylindrical shape. A teflon-coated magnetic stir bar (VWR part number 58948-230, about 1 inch long, 3/8-inch diameter, octagonal cross section, VWR International, West Chester, PA 19380) is placed in a glass where it naturally settles to the bottom. The sample was mixed using a Variomag® multipoint 15 magnetic stirrer (H&P Labortechnik AG, Oberschleissheim, Germany) while adjusting the magnetic stirrer 800 rpm. This mixing starts no more than 5 minutes before the addition of solid products and continues continuously for at least 30 minutes after adding solid products. A solid sample of crude or purified TPA particles in an amount of 250 milligrams is weighed on a non-stick weighing pallet. At the initial time, denoted as t = 0, the suspended solids are poured, all together, into the stirred THF, and at the same time, a timer is started. Properly prepared THF wets solids very quickly and forms a diluted, well-mixed suspension within 5 seconds. Subsequently, samples of this mixture are obtained at the following time points, measured in minutes from t = 0: 0.08, 0.25, 0.50, 0.75, 1.00, 1.50, 2.00, 2.50, 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00 and 30.00. Each small sample was recovered from the diluted, well-mixed mixture using a new disposable syringe (Becton, Dickinson and Co, 5 milliliter, REF 30163, Franklin Lakes, NJ 07417). As soon as approximately 2 ml of the clear sample is removed from the beaker, liquids are rapidly released through a new unused syringe filter (25 mm diameter, 0.45 μm, Gelnan GHP Acrodisc GF®, Pall Corporation, East Hills, NY 11548) into a new labeled glass sample vial . The duration of each filling of the syringe, replacing the filter and releasing into the vial for the sample is noticeably shorter than about 5 seconds, and this interval, respectively, begins and ends within about 3 seconds on either side of each target time of sampling. Within about five minutes of each filling, the sample vials are closed and maintained at approximately constant temperature until the next chemical analysis. After the final sample was taken at time 30 minutes after time t = 0, all sixteen samples were analyzed for the amount of dissolved TPA using the HPLC-DAD (matrix diode detector) method, in general, as described elsewhere in the present description. However, in this study, both calibration standards and recorded results are reported as milligrams of dissolved TPA per gram of THF solvent (hereinafter “ppm in THF”). For example, if all 250 milligrams of solid products are very pure TPA and if all this amount is completely dissolved in 250 grams of THF solvent before a specific sample is taken, a correctly measured concentration would be about 1000 ppm. in thf.

When a CTA according to the present invention is subjected to the dissolution time study described above, it is preferable that a sample taken one minute after t = 0 dissolves to a concentration of at least about 500 ppm. in THF, more preferably at least 600 ppm. in thf. For a sample taken two minutes after t = 0, it is preferred that the CTA in accordance with the present invention dissolves to a concentration of at least about 700 ppm. in THF, more preferably at least 750 ppm. in thf. For a sample taken four minutes after t = 0, it is preferred that the CTA in accordance with the present invention dissolves to a concentration of at least about 840 ppm. in THF, more preferably at least 880 ppm. in thf.

The authors found that a relatively simple model of negative exponential growth is suitable for describing the time dependence of a complete data set on a complete study of the time dependence of dissolution, despite the complexity of the particle samples and the dissolution process. The form of the equation, hereinafter referred to as the “model of the time dependence of dissolution”, is as follows:

S = A + B · (1- exp (-C · t)), where

t = time in units of minutes;

S = solubility, in units of ppm in THF, at time t;

exp = exponential function based on the natural logarithm of 2;

A, B = regression constants in ppm units in THF, where A refers mainly to the rapid dissolution of smaller particles at very short times, and where the sum of A + B relates mainly to the total amount of dissolution near the end of the indicated study period; and

C = temporary regression constant in units of inverse minutes.

Regression constants are selected by minimizing the sum of the mean square error between the points of the real data and the corresponding model values, this method is usually called the “least squares” fit. A preferred software package for performing this data regression is JMP Release 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary, NC 27513).

When the CTA according to the present invention is investigated by dissolution time study and adjusted to the dissolution time dependence model described above, it is preferable that the CTA has a time constant “C” greater than about 0.5 min ⁻¹ , more preferably greater than about 0.6 min ^-1 , and most preferably, greater than 0.7 min ^-1 .

FIG. 55A and 55B illustrate a conventional CTA particle obtained using a conventional high temperature oxidation method in a continuous mixing reactor (CSTR). Fig. 55A shows a conventional CTA particle at 500 times magnification, while Fig. 55B enlarges and shows a CTA particle at 2000 times magnification. A visual comparison of the CTA particles of the present invention illustrated in FIG. 54A and 54B, and the conventional CTA particle illustrated in FIG. 55A and 55B show that a conventional CTA particle has a higher density, lower surface area, lower porosity and larger particle size than the CTA particle of the present invention. In fact, the conventional CTA shown in FIG. 55A and 55B, has an average particle size of about 205 μm and a BET surface area of about 0.57 m ² / g.

Fig. 56 illustrates a conventional method for producing purified terephthalic acid (PTA). In a conventional PTA preparation process, paraxylene is partially oxidized in a mechanically stirred high temperature oxidation reactor 700. The suspension containing CTA is recovered from reactor 700 and then purified in purification system 702. The PTA product from purification system 702 is introduced into separation system 706 to separate and dry particles PTA. Purification system 702 corresponds to most of the costs associated with obtaining PTA particles by conventional methods. The purification system 702 typically comprises a water addition / replacement system 708, a dissolution system 710, a hydrogenation system 712, and three separate crystallization vessels 704a, b, c. In the 708 water addition / replacement system, a substantial portion of the mother liquor is replaced by water. After adding water, the water / CTA slurry is introduced into the dissolution system 710, where the water / CTA mixture is heated until the CTA particles are completely dissolved in water. After the CTA is dissolved, the CTA solution in water is hydrogenated in the hydrogenation system 712. The hydrogenated effluent from the hydrogenation system 712 is then subjected to three stages of crystallization in crystallization vessels 704a, b, c, followed by isolation of PTA in the separation system 706.

Fig. 57 illustrates an improved method for producing PTA using an oxidation reactor type bubble column 800 configured in accordance with an embodiment of the present invention. The initial suspension containing solid particles of CTA and liquid mother liquor is recovered from reactor 800. Typically, the initial suspension may contain from about 10 to about 50 wt%. particulate matter CTA, wherein the liquid residue is a mother liquor. Solid CTA particles present in the initial suspension typically contain at least about 400 ppm. wt. 4-carboxybenzaldehyde (4-CBA), more often at least about 800 ppm. wt. 4-CBA, and most often in the range from 1000 to 15000 ppm. wt. 4-CBA. The initial slurry recovered from reactor 800 is introduced into purification system 802 to reduce the concentration of 4-CBA and other impurities present in the CTA. A cleaner / purified suspension is obtained from the purification system 802 and is subjected to separation and drying in the separation system 804 to obtain solid particles of purer terephthalic acid containing less than about 400 ppm. wt. 4-CBA, more preferably less than about 250 ppm. wt. 4-CBA, and most preferably in the range of 10 to 200 ppm. wt. 4-CBA.

The cleaning system 802 of the PTA receiving system illustrated in FIG. 57 provides a number of advantages over the cleaning system 802 of the prior art system illustrated in FIG. 56. Preferably, the purification system 802 generally comprises a mother fluid replacement system 806, a digester 808, and a separate crystallizer 810. In the mother fluid replacement system 806, at least about 50 wt. the mother liquor present in the initial suspension is replaced with a fresh replacement solvent, thereby obtaining a suspension with a replaced solvent containing CTA particles and a replacement solvent. The solvent-replaced suspension leaving the mother liquor replacement system 806 is introduced into the digester (or secondary oxidation reactor) 808. In the digester 808, the secondary oxidation reaction is carried out at slightly higher temperatures than is used in the initial / primary oxidation reaction carried out in a bubble column reactor type 800. As discussed above, the high surface area, small particle size, and low density of CTA particles produced in reactor 800 make certain impurities trapped in CTA particles available for oxidation in the digester 808 without the need for complete dissolution of the CTA particles in the digester 808. Thus, the temperature in the digester 808 may be lower than in many similar methods known from the literature. Secondary oxidation carried out in digester 808 preferably reduces the concentration of 4-CBA in the CTA by at least 200 ppm. wt., more preferably at least about 400 ppm. wt., and most preferably in the range from 600 to 6000 ppm. wt. Preferably, the secondary oxidation temperature in the digester 808 is at least about 10 ° C higher than the primary oxidation temperature in the bubble column reactor type 800, more preferably higher, in the range of about 20 to about 80 ° C, than the primary oxidation temperature in the reactor 800, and most preferably higher, in the range of 30 to 50 ° C than the primary oxidation temperature in the reactor 800. The secondary oxidation temperature is preferably in the range of about 160 to about 240 ° C, more preferably in the pre crystals of about 180 to about 220 ° C, and most preferably in the range from 190 to 210 ° C. The purified product from digester 808 requires only one crystallization step in the crystallizer 810 before separation in the separation system 804. Suitable secondary oxidation / digestion technologies are discussed in more detail in US Patent Application Publication No. 2005/0065373, the full description of which is explicitly incorporated herein by as a link.

Terephthalic acid (e.g., PTA) obtained by the system illustrated in FIG. 57 preferably consists of PTA particles having an average particle size of at least about 40 microns, more preferably in the range of about 50 to about 2000 microns, and most preferably in the range of 60 to 200 microns. The PTA particles preferably have an average BET surface area of less than about 0.25 m ² / g, more preferably in the range of about 0.005 to about 0.2 m ² / g, and most preferably in the range of 0, 01 to 0.18 m ² / g. PTA obtained by the system illustrated in FIG. 57 is suitable for use as starting materials in the preparation of PET. Typically, PET is prepared by esterification of terephthalic acid with ethylene glycol, followed by polycondensation. Preferably, terephthalic acid obtained through one of the embodiments of the present invention is used as starting materials for producing PET in a tubular reactor described in US patent application serial No. 10/013318, registered December 7, 2001, the full description of which is included here as a reference.

Particles of CTA with the preferred morphology described herein are particularly suitable for use in the oxidative digestion method described above to lower the content of 4-CBA. In addition, these preferred CTA particles provide advantages in a wide range of other, following processes, including dissolution and / or chemical reaction of the particles. These additional following processes include, but are not limited to, reacting with at least one hydroxyl-containing compound to form ester compounds, in particular the reaction of CTA with methanol to form dimethyl terephthalate and impurity esters; reaction with at least one diol to form ester monomeric and / or polymeric compounds, in particular a reaction of CTA with ethylene glycol to form polyethylene terephthalate (PET); and complete or partial dissolution in solvents, including, but not limited to, water, acetic acid and N-methyl-2-pyrrolidone, which may include further processing, including, but not limited to, re-precipitation of purer terephthalic acid and / or selective chemical reduction of carbonyl groups other than carboxylic acid groups. Explicitly included is essentially the dissolution of CTA in a solvent containing water, combined with partial hydrogenation, which reduces the amount of aldehydes, in particular 4-CBA, fluorenones, phenons and / or anthraquinones.

The authors also suggest that CTA particles having the preferred properties described herein can be obtained from CTA particles that do not coincide in the preferred properties with those described herein (mismatched CTA particles) by means of, but not limited to, thereby, mechanically mixing mismatched CTA particles and completely or partially dissolving the mismatching CTA particles, followed by complete or partial reprecipitation.

In accordance with one embodiment of the present invention, there is provided a method for partially oxidizing an oxidizable aromatic compound to one or more types of aromatic carboxylic acid, wherein the purity of a portion of the solvent of the starting materials (i.e., “the starting materials of the solvent”) and the purity of a portion of the oxidizable compound of the starting materials (i.e. there are "starting materials of oxidizable compounds") are controlled within certain limits described below. Together with other embodiments of the present invention, this makes it possible to control the purity of the liquid phase and the solid phase, if present, and the combined phase of the suspension (i.e. solids plus liquid) of the reaction medium within certain preferred ranges indicated below.

Regarding the starting materials of the solvent, it is known to oxidize an oxidizable aromatic compound (s) to produce an aromatic carboxylic acid, where the starting materials of the solvent introduced into the reaction medium are a mixture of acetic acid of analytical purity and water, as is often used when working on a laboratory scale and pilot scale. Similarly, it is known to carry out the oxidation of an oxidizable aromatic compound to an aromatic carboxylic acid, where the solvent leaving the reaction medium is separated from the obtained aromatic carboxylic acid and then recycled back to the reaction medium as solvent starting materials, primarily for reasons of production costs . This recycling of the solvent causes the accumulation over time of certain impurities of the starting materials and by-products of the process in the recycled solvent. Various means are known in the art to help purify the recycled solvent before reintroduction into the reaction medium. Typically, a higher level of purification of the recycled solvent leads to significantly higher production costs than a lower level of purification using similar means. One of the embodiments of the present invention relates to the understanding and determination of the preferred limits of a large number of impurities in the starting materials of the solvent, many of which are still considered mostly harmful, to find the optimal balance between the total cost of production and the overall purity of the product.

"Recyclable solvent starting materials" are defined herein as solvent starting materials containing at least about 5% wt. from the mass that previously passed through a reaction medium containing one or more oxidizable aromatic compounds that undergo partial oxidation. For reasons of solvent stock and flow time in a production unit, it is preferred that portions of the recycled solvent pass through the reaction medium at least once a working day, more preferably at least once a day, for at least seven consecutive working days, and most preferably at least once a day, for at least 30 consecutive working days. For economic reasons, it is preferable that at least about 20% wt. the solvent starting materials in the reaction medium of the present invention were a recycled solvent, more preferably at least about 40% by weight, even more preferably at least about 80% by weight, and most preferably at least 90 % wt.

The authors found that, for reasons of the reactivity of metal impurities remaining in the oxidation product, the concentrations of the selected metals of variable valency in the recycled solvent starting materials are preferably within the ranges indicated immediately below. The concentration of iron in the recycled solvent is preferably below about 150 ppm. wt., more preferably below about 40 ppm. wt., and most preferably is in the range between 0 and 8 ppm. wt. The nickel concentration in the recycled solvent is preferably below about 150 ppm. wt., more preferably, below about 40 ppm. wt., and most preferably is in the range between 0 and 8 ppm. wt. The chromium concentration in the recycled solvent is preferably below about 150 ppm. wt., more preferably, below about 40 ppm. wt., and most preferably, is in the range between 0 and 8 ppm. wt. The concentration of molybdenum in the recycled solvent is preferably below about 75 ppm. wt., more preferably below about 20 ppm. wt., and most preferably is in the range between 0 and 4 ppm. wt. The titanium concentration in the recycled solvent is preferably below about 75 ppm. wt., more preferably below about 20 ppm. wt., and most preferably is in the range between 0 and 4 ppm. wt. The concentration of copper in the recycled solvent is preferably below about 20 ppm. wt., more preferably below about 4 ppm. wt., and most preferably is in the range between 0 and 1 ppm. wt. Other metallic impurities are also typically present in the recycled solvent, typically varying at lower levels with respect to one or more of the metals listed above. Monitoring the above metals within the preferred range will maintain other metallic impurities at appropriate levels.

These metals can occur as impurities in any of the starting materials entering the process (for example, in the incoming oxidizable compound, solvent, oxidizing agent, and catalyst compounds). Alternatively, metals can occur as corrosion products from any of the process plants coming into contact with the reaction medium and / or coming into contact with the recycled solvent. Means for controlling metals within the concentration ranges described include appropriate quality and purity monitoring of various starting materials, and the proper use of structural materials, including, but not limited to, many commercial types of titanium and stainless steels, including those known as duplex stainless steels and high molybdenum stainless steels.

The authors also found preferred limits for selected aromatic compounds in a recycled solvent. They include both precipitated and dissolved aromatic compounds in a recycled solvent.

Unexpectedly, even the precipitated product (e.g. TPA) from the partial oxidation of paraxylene is an impurity that must be controlled in a recycled solvent. Since preferred limits unexpectedly arise for levels of solid products in the reaction medium, any precipitated product in the solvent starting materials is directly subtracted from the amount of oxidizable compound that can be introduced with it. In addition, the introduction of precipitating TPA solids in a recyclable solvent at elevated levels was found to adversely affect the nature of the particles formed in the precipitation oxidation medium, leading to the undesirable nature of the following operations (e.g., filtering the product, washing with a solvent, oxidative digestion of the crude product dissolving the crude product for further processing and so on). Another undesirable characteristic of precipitated solids in recycled solvent feeds is that they often contain very high levels of precipitating impurities compared to the concentration of impurities in the volume of solids in TPA suspensions, from which most of the recycled solvent is obtained. Possibly, the increased levels of impurities observed in solid products suspended in the recycled filtrate may be associated with nucleation times for the deposition of certain impurities from the recycled solvent and / or with cooling of the recycled solvent, either intentionally or caused by environmental losses. For example, concentrations of highly colored and undesired 2,6-dicarboxyfluorenone are observed at much higher levels in the solids present in the recycled solvent at 80 ° C than is observed in TPA solids isolated from the recycled solvent at 160 ° C. Similarly, isophthalic acid concentrations are observed at much higher levels in the solids present in the recycled solvent, compared to levels observed in TPA solids from the reaction medium. How exactly the particular precipitated impurities trapped in the recycled solvent behave when re-introduced into the reaction medium seems to vary. This probably depends on the relative solubility of the impurity in the liquid phase of the reaction medium, probably on how the precipitated impurity is distributed over the layers inside the precipitated solid products, and probably on the local TPA deposition rate, when the solid product is first reintroduced into the reaction medium . Thus, the authors found that it is useful to control the level of certain impurities in the recycled solvent, as described below, regardless of whether these impurities are present in the recycled solvent in dissolved form or are particles trapped therein.

The amount of precipitated solid products present in the recycled filtrate is determined using the gravimetric method as follows. A representative sample is recovered from the solvent supplied to the reaction medium, while the solvent flows in a passage toward the reaction medium. A usable sample size is about 100 g, enclosed in a glass container having about 250 ml of internal volume. Before being released to atmospheric pressure, but continuously flowing towards the sample container, the recycled filtrate is cooled to less than 100 ° C; this cooling is carried out to limit the evaporation of the solvent for a short period before sealing in a closed glass container. After capturing the sample at atmospheric pressure, the glass container is immediately sealed. The sample is then allowed to cool to about 20 ° C, while it is surrounded by air at about 20 ° C and without forced convection. After reaching about 20 ° C, the sample is kept in this state for at least about 2 hours. The sealed container is then shaken vigorously until a visually uniform distribution of solid products is obtained. Immediately thereafter, a magnetic stirrer is added to the sample container and rotated at a sufficient speed to maintain an efficiently uniform distribution of solid products. A 10 ml aliquot of the mixed liquid with suspended solids is removed by pipette and weighed. Then the volume of the liquid phase from this aliquot is separated by vacuum filtration, still at about 20 ° C, and effectively, without loss of solid products. Wet solids filtered from this aliquot are then dried efficiently without sublimation of the solids, and these dried solids are weighed. The ratio of the mass of dried solids to the mass of the original aliquot of the suspension is the fraction of solids, usually expressed as a percentage and referred to here as the content of precipitated solids in the recycled filtrate at 20 ° C.

The authors found that aromatic compounds dissolved in the liquid phase of the reaction medium and containing aromatic carboxylic acids in which there are no non-aromatic hydrocarbon groups (e.g., isophthalic acid, benzoic acid, phthalic acid, 2,5,4'-tricarboxybiphenyl) are unexpectedly harmful components . Although these compounds have much lower chemical activity in the considered reaction medium, compared with oxidizable compounds having non-aromatic hydrocarbon groups, the authors found that these compounds, however, undergo numerous harmful interactions. Thus, it is preferable to control the content of these compounds in the preferred range, in the liquid phase of the reaction medium. This leads to the appearance of the preferred limits of the selected compounds in the starting materials of the recycled solvent, as well as the preferred limits of the selected precursors in the starting materials of the oxidizable aromatic compound.

For example, during the liquid-phase partial oxidation of paraxylene to terephthalic acid (TPA), the authors found that highly colored and undesirable impurities of 2,7-dicarboxyfluorenone (2,7-DCF) are essentially not detected in the reaction medium and in the product sample when metasubstituted aromatic the compounds are in the reaction medium at very low levels. The authors found that when an admixture of isophthalic acid is present in the starting materials of the solvent at increasing levels, the formation of 2,7-DCF increases almost in direct proportion. The authors also found that when methaxylene is present in the paraxylene starting materials, the formation of 2,7-DCF again increases in almost direct proportion. In addition, even if the starting materials of the solvent and the starting materials of the oxidizable compound are free of metasubstituted aromatic compounds, the authors found that some isophthalic acid is formed during the typical partial oxidation of very pure paraxylene, in particular when benzoic acid is present in the liquid phase of the reaction medium. This spontaneously formed isophthalic acid can, due to its higher solubility than TPA, in a solvent containing acetic acid and water, accumulate over time in industrial plants using a recycled solvent. Thus, the amount of isophthalic acid in the starting materials of the solvent, the amount of methaxylol in the starting materials of the oxidizable aromatic compound and the rate of spontaneous formation of isophthalic acid in the reaction medium are all considered in balance with each other and in balance with any reactions that consume isophthalic acid. Isophthalic acid was found to undergo additional reactions with its consumption, in addition to the formation of 2,7-DCF, as described below. In addition, the authors found that there are other subjects to consider when appropriate limits are set for metasubstituted aromatic particles in the partial oxidation of paraxylene to TPA. Other highly colored and undesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF), are apparently strongly associated with dissolved para-substituted aromatic particles, which are always present together with the paraxylene starting materials for liquid phase oxidation. Thus, the suppression of 2.7-DCF is best viewed in perspective with the level of other colored impurities that are produced.

For example, during the liquid phase partial oxidation of paraxylene to TPA, the authors found that the formation of trimellitic acid increases when the levels of isophthalic acid and phthalic acid increase in the reaction medium. Trimellitic acid is a trifunctional carboxylic acid that leads to branching of the polymer chains during the production of PET from TPA. In many PET applications, branching levels should be controlled at low levels, and therefore trimellitic acid should be controlled at low levels in purified TPA. In addition to producing trimellitic acid, the presence of metasubstituted and orthosubstituted particles in the reaction medium also leads to the appearance of other tricarboxylic acids (e.g., 1,3,5-tricarboxybenzene). In addition, the increased presence of tricarboxylic acids in the reaction medium increases the formation of tetracarboxylic acid (e.g., 1,2,4,5-tetracarboxybenzene). Controlling the total production of all aromatic carboxylic acids having more than two groups of carboxylic acids is one of the factors in establishing the preferred levels of metasubstituted and orthosubstituted particles in the starting materials of the recycled solvent, in the starting materials of the oxidizable compound and in the reaction medium in accordance with the present invention.

For example, during the liquid-phase partial oxidation of paraxylene to TPA, the authors found that increased levels in the liquid phase of the reaction medium of several dissolved aromatic carboxylic acids that do not have non-aromatic hydrocarbon groups directly lead to an increase in the formation of carbon monoxide and carbon dioxide. This increased formation of carbon oxides leads to a loss of yield, both for the oxidizing agent and for the oxidizable compound, the latter being due to the fact that many aromatic carboxylic acids obtained together, which, on the one hand, can be considered as impurities, on the other hand, also have industrial value. Thus, the appropriate removal of relatively soluble carboxylic acids that do not have non-aromatic hydrocarbon groups from the recycled solvent is of economic value in preventing loss of yield of oxidizable aromatic compounds and oxidizing agents, in addition to inhibiting the formation of very undesirable impurities such as various fluorenones and trimellitic acid.

For example, during the liquid-phase partial oxidation of paraxylene to TPA, the authors found that the formation of 2,5,4'-tricarboxybiphenyl is apparently inevitable. 2,5,4'-Tricarboxybiphenyl is an aromatic tricarboxylic acid formed by joining two aromatic rings, probably by linking the dissolved para-substituted aromatic particles to an aryl radical, probably to an aryl radical formed by decarboxylation or decarbonylation of para-substituted aromatic particles. Fortunately, 2,5,4'-tricarboxybiphenyl usually forms at lower levels than trimellitic acid, and does not usually lead to a significant increase in the branching problems of polymer molecules during PET production. However, the authors found that elevated levels of 2,5,4'-tricarboxybiphenyl in a reaction medium containing oxidation of alkyl aromatic compounds in accordance with preferred embodiments of the present invention leads to increased levels of highly colored and undesired 2,6-DCF. An elevated level of 2,6-DCF may arise from 2,5,4'-tricarboxybiphenyl by ring closure with the loss of a water molecule, although the exact reaction mechanism is not known with certainty. If 2,5,4'-tricarboxybiphenyl, which is more soluble in a solvent containing acetic acid and water than TPA, is able to accumulate to a level too high in a recycled solvent, conversion rates to 2,6-DCF may become unacceptably high .

For example, during the liquid phase partial oxidation of paraxylene to TPA, the authors found that aromatic carboxylic acids that do not have non-aromatic hydrocarbon groups (for example, isophthalic acid) tend to mildly suppress the chemical activity of the reaction medium when they are present in the liquid phase at sufficient concentration.

For example, during the liquid phase partial oxidation of paraxylene to TPA, the authors found that precipitation is very often imperfect (i.e., nonequilibrium) with respect to the relative concentrations of various chemical particles in the solid phase and in the liquid phase. This is probably due to the fact that the deposition rate is very fast at the spatio-temporal reaction rates preferred here, leading to imperfect co-precipitation of impurities or even occlusion. Thus, if it is desirable to limit the concentration of certain impurities (for example, trimellitic acid and 2,6-DCF) in crude TPA, due to the configuration of the following plants, it is preferable to control their concentration in the starting materials of the solvent, as well as the rate of their formation in the reaction medium.

For example, the authors found that benzophenone compounds (e.g., 4,4'-dicarboxybenzophenone and 2,5,4'-tricarboxybenzophenone) obtained during the partial oxidation of paraxylene have undesirable effects in the reaction medium to produce PET, even though that benzophenone compounds are not strongly colored in TPA per se, like fluorenones and anthraquinones. Accordingly, it is desirable to limit the presence of benzophenones and to select precursors in the recycled solvent and in the starting materials of the oxidizable compound. In addition, the authors found that the presence of elevated levels of benzoic acid, either introduced in a recycled solvent or formed in the reaction medium, leads to an increase in the rates of formation of 4,4'-dicarboxybenzophenone.

As a review, the authors found and quantified a rather unexpected series of reactions for aromatic compounds that do not have non-aromatic hydrocarbon groups that are present in the liquid phase partial oxidation of paraxylene to TPA. Considering just a separate case of benzoic acid, the authors found that increased levels of benzoic acid in the reaction medium in some embodiments of the present invention lead to a strong increase in the formation of highly colored and undesirable 9-fluorenone-2-carboxylic acid, to a strong increase in levels of 4.4 '-dicarboxybiphenyl, to an increase in the levels of 4,4'-dicarboxybenzophenone, to a mild suppression of the chemical activity of the alleged oxidation of paraxylene and to an increase in the levels of carbon oxides, and accordingly Leica Geosystems yield loss. The authors found that increasing levels of benzoic acid in the reaction medium also leads to an increase in the formation of isophthalic acid and phthalic acid, the levels of which it is desirable to control in a low range, in accordance with similar aspects of the present invention. The number and importance of reactions involving benzoic acid are probably even more surprising since some recent inventors have suggested using benzoic acid instead of acetic acid as the primary component of the solvent (see, for example, US Pat. No. 6,562,997). In addition, the authors of the present invention observe that benzoic acid spontaneously forms during the oxidation of paraxylene at rates that are obviously relevant to its formation from impurities such as toluene and ethylbenzene, which are usually found in the starting materials of the oxidized compound containing paraxylene industrial purity.

On the other hand, the authors found little benefit from further adjusting the composition of the recycled solvent with respect to the presence of an oxidizable aromatic compound and with respect to the presence of aromatic intermediate reaction products, which both contain non-aromatic hydrocarbon groups, are also relatively soluble in the recycled solvent. As a rule, these compounds are either introduced or formed in the reaction medium with fractions substantially larger than their presence in the recycled solvent; and the rate of consumption of these compounds in the reaction medium is sufficiently large, while maintaining one or more non-aromatic hydrocarbon groups, to appropriately limit their accumulation in a recycled solvent. For example, during the partial oxidation of paraxylene in a multiphase reaction medium, paraxylene evaporates to a limited extent along with large amounts of solvent. When this evaporated solvent leaves the reactor as part of the off-gas and condenses to be recovered as a recycled solvent, a substantial part of the vaporized para-xylene also condenses in it. It is not necessary to limit the concentration of this paraxylene in a recycled solvent. For example, if the solvent is separated from the solid products when the suspension leaves the paraxylene oxidation reaction medium, this recovered solvent will contain a concentration of dissolved paratoluylic acid similar to that which is present at the point of removal from the reaction medium. Although it may be important to limit the constant concentration of paratoluylic acid in the liquid phase of the reaction medium, see below, it is not necessary to separately regulate paratoluylic acid in this part of the recycled solvent, due to its relatively good solubility and its low mass flow rate, with respect to the formation of paratoluylic acids in the reaction medium. Similarly, the authors found few reasons for limiting in a recycled solvent the concentrations of aromatic compounds with methyl substituents (e.g. toluyl acids), aromatic aldehydes (e.g. terephthaldehyde), aromatic compounds with hydroxymethyl substituents (e.g. 4-hydroxymethylbenzoic acid) and brominated aromatic compounds retaining at least one non-aromatic hydrocarbon group (e.g., alpha-bromo-paratoluylic acid) lower than those originally but they are in liquid phase outgoing from the reaction medium occurring in the partial oxidation of xylene according to preferred embodiments of the present invention. Unexpectedly, the authors also found that there is also no need to control the concentration of selected phenols in the recycled solvent, which are formed during the partial oxidation of xylene, since these compounds form and decompose in the reaction medium at much larger proportions than their presence in the recycled solvent. For example, the authors found that 4-hydroxybenzoic acid has a relatively small effect on chemical activity in preferred embodiments of the present invention when coadministered at proportions of more than 2 g of 4-hydroxybenzoic acid per 1 kg of paraxylene, much larger than the natural presence in the recycled solvent despite this, others report it as an important poison in such a reaction medium (see, for example, W. Partenheimer, Catalysis Today 23 (1995) p. 81).

Thus, there are numerous reactions and numerous considerations in setting the preferred limits of various aromatic impurities in the starting materials of the solvent, as is now described. These data are formulated in terms of the total weighted average composition of all solvent streams that are introduced into the reaction medium over a given period of time, preferably one day, more preferably one hour, and most preferably one minute. For example, if the starting materials of a single solvent flow substantially continuously with a composition of 40 ppm. wt. of isophthalic acid at a flow rate of 7 kilograms per minute, the starting materials of the second solvent flow essentially continuously with a composition of 2000 ppm. wt. isophthalic acid at a flow rate of 10 kilograms per minute, and there are no other solvent feed streams entering the reaction medium, then the total weighted average composition of the solvent starting materials is calculated as (40 × 7 + 2000 × 10) / (7 + 10) = 1.193 m .d. wt. isophthalic acid. Note that the mass of any starting materials of the oxidizable compound or any starting materials of the oxidizing agent, which are likely to be mixed with the starting materials of the solvent before being accumulated in the reaction medium, is not considered when calculating the total weighted average composition of the starting materials of the solvent.

Table 1 below lists preferred values for specific components in the starting materials of the solvent introduced into the reaction medium. The components of the solvent starting materials listed in Table 1 are the following compounds: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxystilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6- DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), fluorenones in general, including other fluorenones not individually listed (fluorenones in general), 4,4'-dicarboxybiphenyl (4,4'-DCB ), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA), zophthalic acid (IPA), benzoic acid (BA), trimellitic acid (TMA), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4'-dicarboxybenzyl (4,4'-DCBZ), 4,4 ' -dicarboxybenzophenone (4,4'-DCBP), 2,5,4'-tricarboxybenzophenone (2,5,4'-TCBP), terephthalic acid (TPA), precipitated solids at 20 ° C and aromatic carboxylic acids in general, which do not have non-aromatic hydrocarbon groups. Table 1, below, gives the preferred amounts of these impurities in the CTA obtained in accordance with an embodiment of the present invention.

Table 1
Components of solvent starting materials introduced into the reaction medium Component Identification Preferred Amount
(ppm wt.) A more preferred amount (ppm wt.) Most preferred
actual amount
(ppm wt.) 4-CBA <1200 30-600 60-300 4.4'-DCS <3 <2 <1 2,6-DCA <6 0.1-3 0.2-1 2,6-DCF <20 0,1-10 0.5-5 2,7-DCF <10 0.1-5 0.5-2 3,5-DCF <10 <5 <2 9F-2CA <10 0.1-5 0.5-2 9F-4CA <5 <3 <1 Fluorenones in general <40 <20 1-8 4,4'-DCB <45 <15 0.5-5 2,5,4'-TCB <45 0.1-15 0.5-5 PA <1000 15-400 40-150 IPA 2500 40-1200 120-400 BA <4500 50-1500 150-500 TMA <1000 15-400 40-150 2,6-DCBC <40 <20 <5 4,4'-DCBZ <40 <20 <5 4,4'-DCBP <40 <20 <5 2,5,4'-TCBP <40 <20 0.5-5 TPA <9000 200-6000 400-2000 Precipitated solids at 20 ° C <9000 200-6000 600-2000 Aromatic carboxylic acids in general, which have no non-aromatic hydrocarbon groups <18000 300-9000 450-3000

Many other aromatic impurities, as a rule, are also present in the recycled solvent, usually changing at even lower levels and / or in proportion to one or more of the described aromatic compounds. Methods for controlling the described aromatic compounds within the preferred range will typically maintain other aromatic impurities at appropriate levels.

When bromine is used in the reaction medium, a large number of ionic and organic forms of bromine are known to exist in dynamic equilibrium. These various forms of bromine have different stability characteristics when they leave the reaction medium and go through various plant operations related to the recycled solvent. For example, alpha-bromo-paratoluylic acid may exist as such under certain conditions, or may rapidly hydrolyze under other conditions to form 4-hydroxymethylbenzoic acid and hydrogen bromide. In the present invention, it is preferred that at least about 40% by weight, more preferably, at least about 60% by weight, and most preferred, at least about 80% by weight. of the total mass of bromine present in the aggregated solvent starting materials entering the reaction medium, were in one or more of the following chemical forms: ionic bromine, alpha-bromo-paratoluic acid, and bromoacetic acid.

Although the importance and value of controlling the overall weighted average purity of the solvent starting materials within the described desirable ranges of the present invention have not been previously discovered and / or described, appropriate means for controlling the purity of the solvent starting materials may be collected from various methods already known in the art. Firstly, any solvent evaporated from the reaction medium is generally of appropriate purity, provided that the liquid or solid products from the reaction medium are not captured together with the evaporated solvent. The introduction of droplets of solvent reflux into the space for separating the flue gas above the reaction medium, as described herein, appropriately limits such capture; and a recycled solvent of appropriate purity with respect to the aromatic compound may condense from such flue gas. Secondly, the more complex and expensive purification of the starting materials of the recycled solvent, as a rule, relates to a solvent selected from the reaction medium in liquid form, and to a solvent which subsequently comes into contact with the liquid and / or solid phases of the reaction medium recovered from a reaction vessel (for example, to a recycled solvent obtained from a filter in which solid products are concentrated and / or washed, to a recycled solvent obtained from a centrifuge in which solid products are concentrated and / or washed, to a recycled solvent taken from the crystallization operation, and so on). However, agents are also known in the art for performing the necessary purification of these recycled solvent streams using one or more of the previous descriptions. Regarding the control of precipitated solid products in a recycled solvent so that they are within the specified limits, appropriate controls include, but are not limited to, gravimetric sedimentation, mechanical filtration using filter cloth on rotary belt filters and rotary drum filters, mechanical filtration using stationary filter medium in high pressure tanks, hydrocyclones and centrifuges. Regarding the control of dissolved aromatic particles in a recycled solvent so that they are within the indicated limits, controls include, but are not limited to, those described in US Pat. No. 4,939,297 and in the publication of US Patent Application No. 2005-0038288, incorporated herein. as links. However, none of these previous inventions detects or describes preferred purity levels in the total solvent starting materials, as described herein. Instead, these previous inventions merely provided means for purifying selected and partial recycled solvent streams without considering the optimal composition values of the aggregated weighted average solvent starting materials introduced into the reaction medium of the present invention.

Turning now to the purity of the starting materials of the oxidizable compound, it is known that certain levels of isophthalic acid, phthalic acid and benzoic acid are present and are acceptable at low levels in the purified TPA used to make the polymer. In addition, it is known that these particles are relatively more soluble in many solvents and can advantageously be removed from purified TPA by crystallization methods. However, from one of the embodiments of the present invention described herein, it is now known that controlling the level of several relatively soluble aromatic particles, including explicitly isophthalic acid, phthalic acid and benzoic acid, in the liquid phase of the reaction medium is unexpectedly important for controlling the level of polycyclic and colored aromatic compounds formed in the reaction medium to control compounds with more than 2 functional groups of carboxylic acids per molecule, to control p promotional activity in the reaction medium for partial oxidation and oxidant yield loss control and an aromatic compound.

It is known in the art that isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium as follows. An admixture of metaxylene starting materials is oxidized with good conversion and yield to IPA. An admixture of orthoxylene starting materials is oxidized with good conversion and yield to phthalic acid. Impurities of the starting materials of ethylbenzene and toluene are oxidized with good conversion and yield to benzoic acid. However, the authors observe that significant amounts of isophthalic acid, phthalic acid and benzoic acid are also formed in a paraxylene-containing reaction medium by means other than the oxidation of methaxylene, orthoxylene, ethylbenzene and toluene. These other valid chemical routes may include decarbonylation, decarboxylation, rearrangement of transition states, and addition of methyl and carbonyl radicals to aromatic rings.

In determining the preferred limits of impurities in the starting materials of the oxidizable compound, many factors are important. Any impurity in the starting materials is likely to cause immediate loss of yield and purification costs if the requirements for the purity of the oxidized product are high enough (for example, in a reaction medium for the partial oxidation of paraxylene, toluene and ethylbenzene, usually found in industrial grade paraxylene , lead to the formation of benzoic acid, and this benzoic acid is mostly removed from most industrial varieties of TPA). When the product of the partial oxidation of the impurity of the starting materials is involved in additional reactions, factors other than simply loss of yield and removal become important when considering what the costs of purifying the starting materials should be (for example, in the reaction medium for the partial oxidation of paraxylene, ethylbenzene leads to the formation of benzoic acid, and benzoic acid subsequently leads to the formation of highly colored 9-fluorenone-2-carboxylic acid, isophthalic acid, phthalic acid and an increase in the amount of Twa carbon oxides, among others). When the reaction medium spontaneously forms additional amounts of impurities by chemical mechanisms that are not directly related to the impurities of the starting materials, the analysis becomes even more complicated (for example, in the reaction medium for the partial oxidation of paraxylene, benzoic acid also spontaneously forms from paraxylene itself). In addition, subsequent processing of the crude oxidation product may affect considerations regarding the preferred purity of the starting materials. For example, the cost of removing direct impurities (benzoic acid) and subsequent impurities (isophthalic acid, phthalic acid, 9-fluorenone-2-carboxylic acid and the like) to the appropriate levels may be the same, may differ from others, and may differ from removal requirements for the most part unrelated impurities (for example, the product of incomplete oxidation of 4-CBA during the oxidation of paraxylene to TPA).

The purity limits for paraxylene starting materials described below are preferred when paraxylene is introduced, together with a solvent and an oxidizing agent, into a partial oxidation reaction medium to give TPA. These limits are more preferred in a TPA production process having subsequent oxidation steps to remove impurities other than an oxidizing agent and a solvent (e.g., metal catalysts) from the reaction medium. These limits are even more preferable in methods for producing TPA that remove additional 4-CBA from CTA (e.g., by converting CTA to dimethyl terephthalate, plus impurity esters, and then separating 4-CBA methyl ester by distillation, by means of oxidative digestion, to convert 4-CBA to TPA, by means of hydrogenation methods to convert 4-CBA to paratoluylic acid, which is then separated by fractional crystallization methods). These limits are most preferred in methods for producing TPA, which remove additional 4-CBA from CTA via oxidative digestion methods to convert 4-CBA to TPA.

Using new knowledge of the preferred limits for the recycling of aromatic compounds and the relative amounts of aromatic compounds generated directly from the oxidation of the impurities of the starting materials compared to other actual chemical routes, improved limits are found for the impurities for contaminating paraxylene, which is the starting materials introduced into the partial oxidation processes to get TPA. Table 2 below provides preferred values for the amount of metaxylene, orthoxylene and ethylbenzene + toluene in paraxylene starting materials.

table 2
Contamination components of paraxylene starting materials Component Identification Preferred amount (ppm wt.) A more preferred amount (ppm wt.) Most preferred amount
(ppm wt.) methaxylol 20-800 50-600 100-400 orthoxylene 10-300 20-200 30-100 ethylbenzene + toluene * 20-700 50-500 100-300 Generally 50-900 100-800 200-700 * Specification for ethylbenzene + toluene is given for each separately and in total

Specialists in this field will now notice that the above impurities in contaminated paraxylene can have the greatest effect on the reaction medium after the products of their partial oxidation are accumulated in a recycled solvent. For example, the introduction of an upper amount from the most preferred range of metaxylene, 400 ppm. wt., will directly lead to about 200 ppm wt. isophthalic acid in the liquid phase of the reaction medium, if you work with about 33% wt. solid products in the reaction medium. This is comparable to 400 ppm from the upper amount from the most preferred range for isophthalic acid in a recycled solvent. wt., which, after being given the opportunity to cool the reaction medium by evaporation of a typical solvent, results in about 1200 ppm. wt. isophthalic acid in the liquid phase of the reaction medium. Thus, over time, partial oxidation products accumulate in the recycled solvent, which is the most likely impact of impurities of methaxylene, orthoxylene, ethylbenzene and toluene in the starting materials of contaminated paraxylene. Accordingly, it is preferable that the above ranges of impurities in the contaminated paraxylene starting materials be maintained for at least half a day, on any day of operation of any partial oxidation reaction medium in a particular industrial plant, more preferably for at least three quarters of each day for at least seven consecutive working days, and most preferably, when the weighted average parameters of the composition of the source materials of contaminated steam xylene is in the preferred range for at least 30 consecutive working days.

Means for producing contaminated paraxylene of preferred purity are already known in the art and include, but are not limited to, distillation, fractional crystallization methods at temperatures below ambient temperature, and molecular sieve methods using pore-selective adsorption. However, the preferred purity limits indicated here are, in their highest degree, more demanding and expensive than what is characteristic of the industrial supply of paraxylene; and to the lowest degree, the preferred purity limits exclude the unnecessarily expensive purification of paraxylene intended to introduce partial oxidation into the reaction medium by detecting and describing where the combined effects of spontaneous formation of impurities from paraxylene itself and reactions involving the consumption of impurities in the reaction medium become more important, than the rate of introduction of impurities in contaminated paraxylene.

When the xylene containing feed stream includes selected impurities, such as ethylbenzene and / or toluene, the oxidation of these impurities can form benzoic acid. As used here, the term “impurity-formed benzoic acid” should mean benzoic acid obtained from any source other than xylene during the oxidation of xylene.

As described here, a portion of the benzoic acid formed during the oxidation of xylene is obtained from xylene itself. This production of benzoic acid from xylene is clearly complementary to any part of the production of benzoic acid, which may be a benzoic acid formed by impurities. Not wanting to be limited by theory, it is assumed that benzoic acid is obtained from xylene in a reaction medium when various intermediate xylene oxidation products are spontaneously decarbonylated (lose carbon monoxide) or decarboxylated (lose carbon dioxide), producing aryl radicals. These aryl radicals can then remove a hydrogen atom from any of the many available sources in the reaction medium and form spontaneously formed benzoic acid. As a chemical mechanism, the term "spontaneously formed benzoic acid", as used here, shall mean benzoic acid obtained from xylene during the oxidation of xylene.

As also described here, when paraxylene is oxidized to produce terephthalic acid (TPA), the production of spontaneously formed benzoic acid causes a loss in paraxylene yield and a loss in oxidizer yield. In addition, the presence of spontaneously formed benzoic acid in the liquid phase of the reaction medium correlates with an increase in many undesirable side reactions, which explicitly include the formation of highly colored compounds called monocarboxyfluorenones. The spontaneously formed benzoic acid also contributes to the undesired accumulation of benzoic acid in the recycled filtrate, which further increases the concentration of benzoic acid in the liquid phase of the reaction medium. Thus, the formation of spontaneous formation of benzoic acid is desirable to be minimized, but it is also considered accordingly simultaneously with the formation of impurities of benzoic acid, with factors affecting the consumption of benzoic acid, with factors related to other issues of reaction selectivity, and with the general economy.

The authors found that spontaneous formation of benzoic acid can be controlled at low levels by appropriate selection, for example, temperature, xylene distribution and oxygen availability in the reaction medium during oxidation. Without wishing to be limited by theory, lower temperatures and improved oxygen availability seem to suppress decarbonylation and / or decarboxylation rates, thereby eliminating the aspect of loss of yield of spontaneously formed benzoic acid. The sufficient availability of oxygen apparently directs aryl radicals towards other more harmful products, in particular hydroxybenzoic acids. The distribution of xylene in the reaction medium may also affect the balance between the conversion of the aryl radical to benzoic acid or to hydroxybenzoic acids. Whatever the chemical mechanisms, the authors found reaction conditions that, although gentle enough to reduce the production of benzoic acid, are stringent enough to oxidize a high proportion of the resulting hydroxybenzoic acid to carbon monoxide and / or carbon dioxide, which are easily removed from the oxidation product.

In a preferred embodiment of the present invention, the oxidation reactor is configured and operated so that the formation of spontaneously formed benzoic acid is minimized and the oxidation of hydroxybenzoic acids to carbon monoxide and / or carbon dioxide is maximized. When an oxidation reactor is used to oxidize paraxylene to terephthalic acid, it is preferred that paraxylene is replenished by at least about 50% by weight. of total xylene in the feed stream introduced into the reactor. More preferably, paraxylene is made up of at least about 75% wt. from total xylene in the feed stream. Even more preferably, paraxylene is replenished by at least 95% wt. from total xylene in the feed stream. Most preferably, paraxylene is made up for substantially all of the xylene in the feed stream.

When the reactor is used to oxidize paraxylene to terephthalic acid, it is preferable that the rate of terephthalic acid production is maximized, while the rate of formation of spontaneously formed benzoic acid is minimized. Preferably, the ratio of the rate of formation (by weight) of terephthalic acid to the rate of formation (by weight) of spontaneously formed benzoic acid is at least about 500: 1, more preferably at least about 1000: 1, and most preferably, at least 1500: 1. As will be seen below, the rate of formation of spontaneously formed benzoic acid is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is below 2000 ppm. wt., more preferably below 1000 ppm. wt., and most preferably below 500 ppm. wt., because these low concentrations are suppressed to the corresponding low reaction rates, which convert benzoic acid to other compounds.

When combining spontaneously formed benzoic acid and formed by impurities of benzoic acid, the ratio of the rate of formation (by weight) of terephthalic acid to the rate of formation (by weight) of benzoic acid as a whole is preferably at least about 400: 1, more preferably at least , about 700: 1, and most preferably at least 1100: 1. As will be seen below, the total rate of formation of spontaneously formed benzoic acid, plus benzoic acid formed by impurities, is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is below 2000 ppm. wt., more preferably below 1000 ppm. wt., and most preferably below 500 ppm. wt., because these low concentrations are suppressed to the corresponding low reaction rates, which convert benzoic acid to other compounds.

As described here, an increased concentration of benzoic acid in the liquid phase of the reaction medium leads to an increase in the formation of many other aromatic compounds, several of which are harmful impurities in TPA; and, as described here, increased concentrations of benzoic acid in the liquid phase of the reaction medium lead to an increase in the formation of gaseous carbon oxides, the formation of which represents a loss of yield in the oxidizing agent and in the aromatic compounds and / or solvent. In addition, it is now described that the authors found that a significant portion of this increase in the formation of other aromatic compounds and carbon oxides is due to reactions that convert some of the benzoic acid molecules on their own, as opposed to benzoic acid, which catalyzes other reactions, without spending her own. Accordingly, “the formation of benzoic acid as a whole” is defined here as the time-averaged mass of all benzoic acid leaving the reaction medium, minus the time-averaged mass of all benzoic acid entering the reaction medium over the same period of time. This formation of benzoic acid as a whole is often positive, due to the rate of formation of benzoic acid formed by impurities and spontaneous formation of benzoic acid. However, the authors found that the rate of conversion of benzoic acid to carbon oxides and several other compounds appears to increase approximately linearly when the concentration of benzoic acid increases in the liquid phase of the reaction medium, when measured when other reaction conditions, including temperature, oxygen availability, STR and reactivity are maintained, respectively, constant. Thus, when the concentration of benzoic acid in the liquid phase of the reaction medium is sufficiently large, probably due to the increased concentration of benzoic acid in the recycled solvent, then the conversion of benzoic acid molecules to other compounds, including carbon oxides, may become equal to or greater than the chemical the formation of new benzoic acid molecules. In this case, the formation of benzoic acid as a whole can balance near zero or even be negative. The authors found that when the formation of benzoic acid as a whole is positive, then the ratio of the rate of formation (by weight) of terephthalic acid in the reaction medium compared to the rate of formation of benzoic acid as a whole in the reaction medium is preferably above about 700: 1, more preferably above about 1100 : 1, and most preferably above 4000: 1. The authors found that when the formation of benzoic acid as a whole is negative, the ratio of the rate of formation (by weight) of terephthalic acid in the reaction medium, compared with the rate of formation of benzoic acid as a whole in the reaction medium, is preferably above about 200: (- 1), more preferably above about 1000: (- 1), and most preferably above 5000: (- 1).

The authors also found preferred limits for the composition of the suspension (liquid + solids) recovered from the reaction medium and for the solid portion of the CTA suspension. The preferred suspension and preferred CTA compositions are unexpectedly excellent and suitable for use. For example, purified TPA obtained from this preferred CTA by oxidative digestion has a fairly low level of total impurities and colored impurities, so that purified TPA is suitable, without hydrogenation of additional 4-CBA and / or colored impurities, for a wide range of applications in fibers of PET and in PET packaging applications. For example, a preferred suspension formulation provides the liquid phase of the reaction medium in which the concentration of important impurities is relatively low, and this significantly reduces the formation of other, even more undesirable impurities, as described herein. In addition to this, the preferred suspension formulation greatly helps the further processing of the liquid from the suspension to obtain an appropriate pure recycled solvent, in accordance with other embodiments of the present invention.

CTA obtained in accordance with one of the embodiments of the present invention contains fewer impurities of the selected types than CTA obtained using conventional methods and devices, in particular those that use a recycled solvent. Impurities that may be present in the CTA include the following: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxylstilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylic acid (9F- 2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB) , phthalic acid (PA), isophthalic acid (IPA), benzoic acid (BA), trimellitic acid (TMA), paratoluylic acid (PTAC), 2,6-dicarboxybenzo-coumarin (2,6-DCBC), 4,4'-dicar hydroxybenzyl (4,4'-DCBZ), 4,4'-dikarboksibenzofenon (4,4'-DCBP), 2,5,4'-trikarboksibenzofenon (2,5,4'-TCBP). Table 3 below lists preferred amounts of these impurities in a CTA prepared in accordance with one embodiment of the present invention.

Table 3
Impurities CTA Impurity Identification Preferred amount (ppm wt.) A more preferred amount (ppm wt.) Most preferred amount
(ppm wt.) 4-CBA <15000 100-8000 400-2000 4.4'-DCS <12 <6 <3 2,6-DCA <9 <6 <2 2,6-DCF <100 2-50 5-25 2,7-DCF <30 <15 <5 3,5-DCF <16 <8 <2 9F-2CA <16 <8 <4 9F-4CA <8 <4 <2 Fluorenones in general <100 2-60 4-35 4,4'-DCB <64 1-32 2-8 2,5,4'-TCB <24 <12 <8 PA <200 3-100 5-50 IPA <800 10-400 20-200 BA <600 5-300 15-100 TMA <800 10-400 20-200 PTAC <2000 10-1000 50-500 2,6-DCBC <64 <32 <8 4,4'-DCBZ <12 <8 <4 4,4'-DCBP <40 <30 <20 2,5,4'-TCBP <32 <16 <4

In addition, it is preferable that the CTA obtained in accordance with one embodiment of the present invention has a reduced color content with respect to the CTA obtained by conventional methods and devices, namely those that use a recycled solvent. Thus, it is preferable that the CTA obtained in accordance with one embodiment of the present invention has a transmittance at 340 nm, at least about 25%, more preferably at least about 50%, and most preferably at least 60%. In addition, it is preferable that the CTA obtained in accordance with one of the embodiments of the present invention, had a transmittance at 400 nm, at least about 88%, more preferably at least about 90%, and most preferably at least 92%.

A study for percent transmittance involves measuring the concentration of colored, light-absorbing impurities present in a TPA or CTA. As used here, the study refers to measurements taken per serving of a solution obtained by dissolving 2.00 g of dry solid TPA or CTA in 20.0 ml of dimethyl sulfoxide (DMSO), analytical or even better. A portion of this solution is then placed in a Hellma flow-through cell, PN 176.700, which is made of quartz and has a light absorption thickness of 1.0 cm and a volume of 0.39 ml (Hellma USA, 80 Skyline Drive, Plainview, NY 11803). An Agilent 8453 diode array spectrophotometer is used to measure the transmission of various wavelengths of light through this filled flow cell. (Agilent Technologies, 395 Page Mill Road, Palo Alto, CA 94303). After appropriate adjustment for background absorption, including, but not limited to, the cell used and the solvent, the percent transmission results characterizing the fraction of incident light that passes through the solution are recorded directly by the machine. The transmission percentages at light wavelengths of 340 nanometers and 400 nanometers are especially suitable for use in order to detect the difference between pure TPA and many impurities, as a rule, contained in it.

The preferred limits of various aromatic impurities in the suspension phase (solids + liquid) of the reaction medium are given below in table 4.

Table 4
Impurities in suspension Impurity Identification Preferred Amount
(ppm wt.) A more preferred amount (ppm wt.) The most preferred amount (ppm wt.) 4-CBA <8000 <5000 <2500 4.4'-DCS  <4 <2 <1 2,6-DCA <6 <3 <1 2,6-DCF <70 2-40 4-20 2,7-DCF <12 <8 <4 3,5-DCF <12 <8 <4 9F-2CA <12 <8 <4 9F-4CA <8 <4 <2 Fluorenones in general <90 2-60 5-30 4,4'-DCB <64 1-16 2-4 2,5,4'-TCB <60 2-40 4-20 PA <3000 25-1500 75-500 IPA 9000 75-4500 225-1500 BA <15000 100-6000 300-2000 TMA <3000 25-1500 75-500 PTAC <8000 100-4000 200-2000 4,4'-DCBZ <5 <4 <3 4,4'-DCBP <240 <160 <80 2,5,4'-TCBP <120 <80 <40

These preferred suspension formulations embody the preferred composition of the liquid phase of the reaction medium, while at the same time beneficially eliminating the experimental difficulties related to the deposition of additional components of the liquid phase from the reaction medium into the components of the solid phase during sampling from the reaction medium, separation of liquid and solid products and a change in analytical conditions.

Many other aromatic impurities, as a rule, are also present in the suspension phase of the reaction medium and in CTA from the reaction medium, usually changing at lower levels and / or in proportion to one or more of the described aromatic compounds. Control of the described aromatic compounds within the preferred range will maintain other aromatic impurities at appropriate levels. These advantageous formulations for the suspension phase in the reaction medium and for solid CTA taken directly from the suspension make it possible to work with the embodiments of the present invention described here for the partial oxidation of paraxylene to TPA.

Measurement of the concentration of components with low levels in a solvent, a recycled solvent, CTA, suspension from the reaction medium and PTA are carried out using liquid chromatography methods. Two interchangeable embodiments are now described.

The method referred to as HPLC-DAD herein includes high pressure liquid chromatography (HPLC) combined with a matrix diode detector (DAD) to allow separation and quantification of various molecular particles within a given sample. The instrument used in this measurement is the DAD-supplied Model 1100 HPLC supplied by Agilent Technologies (Palo Alto, CA), although other usable instruments are also commercially available from other suppliers. As is known in the art, both the elution time and the response of the detector are calibrated using known compounds present in known amounts, these compounds and their amounts are in accordance with those found in real unknown samples.

The method, referred to herein as HPLC-MS, includes high pressure liquid chromatography (HPLC) coupled to mass spectrometry (MS) to allow separation, identification and quantification of various molecular particles within a given sample. The instrument used in this measurement is the Alliance HPLC and ZQ MS supplied by Waters Corp. (Milford, MA), although other usable tools are also commercially available from other suppliers. As is known in the art, both the elution time and the response of the detector are calibrated using known compounds present in known amounts, these compounds and their amounts are in accordance with those found in real unknown samples.

Another embodiment of the present invention relates to the partial oxidation of an aromatic oxidizable compound with an appropriate balance of suppressing harmful aromatic impurities, on the one hand, compared to the formation of carbon dioxide and carbon monoxide, collectively, carbon oxides (COx), on the other. These carbon oxides typically leave the reaction vessel in the flue gas, and they correspond to the destructive losses of the solvent and oxidizable compound, including, ultimately, the preferred oxidized derivatives (e.g., acetic acid, paraxylene and TPA). The authors found lower limits for the formation of carbon oxides, below which, apparently, a high formation of harmful aromatic impurities, as described below, and a low overall level of conversion are inevitably too bad to be economically used. The authors also found upper boundaries of carbon oxides, above which the formation of carbon oxides continues to increase with a small additional value provided by a decrease in the formation of harmful aromatic impurities.

The authors found that lowering the liquid-phase concentrations of the starting materials of the aromatic oxidizable compound and the aromatic intermediate particles in the reaction medium leads to lower formation rates for harmful impurities during the partial oxidation of the aromatic oxidizable compound. These harmful impurities include bound aromatic rings and / or aromatic molecules containing more than the desired number of carboxylic acid groups (for example, in the oxidation of paraxylene, harmful impurities include 2,6-dicarboxyanthraquinone, 2,6-dicarboxyfluorenone, trimellitic acid , 2,5,4'-tricarboxybiphenyl and 2,5,4'-benzophenone). Aromatic intermediates include aromatic compounds falling out of the starting materials of the oxidizable aromatic compound and still retaining non-aromatic hydrocarbon groups (for example, in the oxidation of paraxylene, aromatic intermediates include paratolualdehyde, terephthaldehyde, paratoluylic acid, 4-CBA, 4-hydroxymethylbenzoic acid and alpha bromo-paratoluylic acid). The starting materials of the oxidizable aromatic compound and aromatic intermediate particles holding non-aromatic hydrocarbon groups when they are present in the liquid phase of the reaction medium apparently lead to the appearance of harmful impurities in a manner similar to that already described here for dissolved aromatic particles that do not have non-aromatic hydrocarbon groups (e.g. isophthalic acid).

Confronting this need for higher reactivity to suppress the formation of harmful aromatic impurities during the partial oxidation of an oxidizable aromatic compound, the authors found that an undesirable concomitant result was an increase in the formation of carbon oxides. It is important to note that these carbon oxides represent losses in the yield of the oxidizable compound and oxidizing agent, and not just the solvent. In fact, a significant, and sometimes the main, proportion of carbon oxides comes from the oxidizable compound and its derivatives, and not from the solvent; and often an oxidizable compound costs more per unit carbon than a solvent. In addition, it is important to note that the desired carboxylic acid product (e.g., TPA) also undergoes superoxidation to carbon oxides when present in the liquid phase of the reaction medium.

It is also important to note that the present invention relates to reactions in the liquid phase of a reaction medium and to concentrations of reactants therein. This contrasts with some previous inventions, which relate directly to the formation of an aromatic compound in precipitated solid form, retaining non-aromatic hydrocarbon groups. Specifically, for the partial oxidation of paraxylene to TPA, certain previous inventions relate to the amount of 4-CBA precipitated in the solid phase of CTA. However, the authors of the present invention found a spread greater than two to one for the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase, using the same conditions of temperature, pressure, catalysis, solvent composition and the spatio-temporal reaction rate of paraxylene, depending on whether partial oxidation is carried out in a well-stirred autoclave or in a reaction medium with a stepwise change in the concentration of oxygen and paraxylene in accordance with the present invention. In addition, the inventors have found that the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase can also change by more than two to one, either in a well-mixed reaction medium or in a reaction medium with a stepwise change in concentration, depending from the spatio-temporal reaction rate of paraxylene, under otherwise similar conditions of temperature, pressure, catalysis and solvent composition. In addition, 4-CBA in the solid phase of CTA does not seem to contribute to the formation of harmful impurities, and 4-CBA in the solid phase can be extracted and oxidized to TPA simply and in high yield (for example, by oxidative digestion of a suspension of CTA, as described here); while the removal of harmful impurities is much more complex and expensive than the removal of the solid phase of 4-CBA, and the formation of carbon oxides represents a permanent loss of yield. Thus, it is important to distinguish that this aspect of the present invention relates to liquid phase compositions in a reaction medium.

Whether they come from a solvent or an oxidizable compound, but the authors found that during the conversion of industrial value, the formation of carbon oxides is strongly related to the level of general reaction activity, despite the wide scatter for each specific combination of temperature, metals, halogens, temperature, acidity of the reaction medium, measured by pH, the concentration of water used to obtain the level of overall reactivity. The authors find it useful for the partial oxidation of xylene to assess the level of total reaction activity using the liquid-phase concentration of toluic acids at the mid-height of the reaction medium, in the lower part of the reaction medium and in the upper part of the reaction medium.

Thus, there is an important simultaneous balancing to minimize the formation of harmful impurities by increasing the reactivity, and to minimize the formation of carbon oxides by reducing the reaction activity. That is, if the overall formation of carbon oxides is suppressed to a too low value, then excessive levels of harmful impurities are formed, and vice versa.

In addition, the inventors of the present invention have found that the solubility and relative reactivity of the desired carboxylic acid (e.g., TPA), and the presence of other dissolved aromatic particles that do not have non-aromatic hydrocarbon groups, introduces a very important factor in this balancing of carbon oxides compared to harmful impurities. The desired carboxylic acid product is usually dissolved in the liquid phase of the reaction medium, even if it is also present in solid form. For example, at temperatures in the preferred range, TPA is soluble in a reaction medium containing acetic acid and water, at levels ranging from about one thousand ppm. wt. to more than 1% wt., in the present, the solubility increases when the temperature increases. Despite the fact that there are differences in reaction rates with respect to the formation of various harmful impurities from the starting materials of an oxidizable aromatic compound (e.g., paraxylene), from aromatic intermediate reaction products (e.g., paratoluylic acid), from the desired aromatic carboxylic acid product (e.g. TPA ) and of aromatic particles that do not have non-aromatic hydrocarbon groups (for example, isophthalic acid), the presence and reactivity of the last two groups sets the region It is justified to reduce, in relation to further suppression of the first two groups of starting materials of oxidizable aromatic compounds and aromatic intermediate reaction products. For example, in the case of partial oxidation of paraxylene to TPA, if in the liquid phase of the reaction medium under these conditions the dissolved TPA reaches 7000 ppm. wt., dissolved benzoic acid reaches 8000 ppm. wt., dissolved isophthalic acid reaches 6000 ppm. wt. and dissolved phthalic acid reaches 2000 ppm wt., then the value for a further decrease in the concentration of harmful compounds as a whole begins to decrease when the reaction activity increases, with the suppression of the liquid-phase concentration of paratoluylic acid and 4-CBA below similar levels. That is, the presence and concentration in the liquid phase of the reaction medium of aromatic particles that do not have non-aromatic hydrocarbon groups changes very little by increasing the reaction activity, and their presence serves to expand upward the range of beneficial reduction, to lower the concentration of intermediate reaction products in order to suppress the formation of harmful impurities.

Thus, one of the embodiments of the present invention provides preferred ranges of carbon oxides, limited from below by low reactivity and excessive formation of harmful impurities, and from above by excessive losses of carbon, but at levels lower than those previously found and described as industrially suitable . Accordingly, the formation of carbon oxides is preferably controlled as follows. The ratio of the resulting moles of carbon oxides in general to the moles of the introduced oxidizable aromatic compound is preferably greater than about 0.02: 1, more preferably greater than about 0.04: 1, even more preferably greater than 0.05: 1, and most preferably more than 0.06: 1. At the same time, the ratio of the resulting moles of carbon oxides in general to the moles of the introduced oxidizable aromatic compound is preferably less than about 0.24: 1, more preferably less than about 0.22: 1, even more preferably less, about than 0.19: 1, and most preferably less than 0.15: 1. The ratio of moles of carbon dioxide formed to moles of introduced oxidizable aromatic compound is preferably greater than about 0.01: 1, more preferably greater than about 0.03: 1, even more preferably greater than about 0.04: 1, and most preferably more than 0.05: 1. At the same time, the ratio of moles of carbon dioxide formed to moles of introduced oxidizable aromatic compound is preferably less than about 0.21: 1, more preferably less than about 0.19: 1, even more preferably less than 0.16: 1, and most preferably less than 0.11. The ratio of moles of carbon monoxide formed to moles of the oxidizable aromatic compound introduced is preferably greater than about 0.005: 1, more preferably greater than about 0.010: 1, even more preferably greater than about 0.015: 1, and most preferably more than 0.020: 1. At the same time, the ratio of the moles of carbon monoxide formed to the moles of the oxidizable aromatic compound introduced is preferably less than about 0.09: 1, more preferably less than about 0.07: 1, even more preferably less than than 0.05: 1, and most preferably less than 0.04: 1.

The carbon dioxide content of the dry flue gas from the oxidation reactor is preferably greater than about 0.10 mol%, more preferably greater than about 0.20 mol%, even more preferably greater than about 0.25 mol% ., and most preferably more than 0.30 mol%. At the same time, the carbon dioxide content of the dry flue gas from the oxidation reactor is preferably less than about 1.5 mol%, more preferably less than about 1.2 mol%, even more preferably less than than 0.9 mol%, and most preferably less than 0.8 mol%. The carbon monoxide content in the dry flue gas from the oxidation reactor is preferably greater than about 0.05 mol%, more preferably greater than about 0.10 mol%, even more preferably greater than 0.15, and most preferably more than 0.18 mol%. At the same time, the carbon monoxide content in the dry flue gas from the oxidation reactor is preferably less than about 0.60 mol%, more preferably less than about 0.50 mol%, even more preferably less than than 0.35 mol%, and most preferably less than 0.28 mol%.

The inventors of the present invention have found that an important factor for reducing the formation of carbon oxides to these preferred ranges is to improve the purity of the recycled filtrate and the starting materials of the oxidizable compound to lower the concentration of aromatic compounds that do not have non-aromatic hydrocarbon groups, in accordance with the descriptions of the present invention reduces the formation of carbon oxides and harmful impurities. Another factor is the improvement in the distribution of paraxylene and oxidizing agent in the reaction vessel in accordance with the descriptions of the present invention. Other factors that make possible the above preferred levels of carbon oxides are the work with gradients in the reaction medium, as described here, of pressure, temperature, concentration of the oxidizable compound in the liquid phase and the oxidizing agent in the gas phase. Other factors that make possible the above preferred levels of carbon oxides are the work within the scope of the descriptions given here, preferred for the spatio-temporal reaction rate, pressure, temperature, solvent composition, catalyst composition, and mechanical geometry of the reaction vessel.

An important advantage of working within the preferred range of formation of carbon oxides is that the use of molecular oxygen can be reduced, although not to stoichiometric values. Despite a good stepwise change in the concentration of the oxidizing agent and the oxidizable compound in accordance with the present invention, the excess oxygen must be maintained above the stoichiometric value, as calculated only for the starting materials of the oxidizable compound, in order to allow some loss of carbon oxides and provide an excess of molecular oxygen to control the formation of harmful impurities. Specifically, for the case where xylene is the starting materials of the oxidizable compound, the ratio of molecular oxygen to xylene in the starting materials is preferably greater than about 0.91: 1.00, more preferably greater than about 0.95: 1, 00, and most preferably more than 0.99: 1.00. At the same time, the ratio of molecular oxygen to xylene in the starting materials is preferably less than about 1.20: 1.00, more preferably less than about 1.12: 1.00, and most preferably less than 1.06: 1.00. Specifically for xylene starting materials, the time-averaged molecular oxygen content of the dry flue gas from the oxidation reactor is preferably greater than about 0.1 mol%, more preferably greater than about 1 mol%, and most preferably more than 1.5 mol% At the same time, the time-averaged molecular oxygen content of the dry flue gas from the oxidation reactor is preferably less than about 6 mol%, more preferably less than about 4 mol%, and most preferably less than 3% pier

Another important advantage of working within the preferred range of carbon monoxide production is that less aromatic compound is converted to carbon oxides and other, less valuable forms. This advantage is measured using the sum of the moles of all aromatic compounds leaving the reaction medium divided by the sum of the moles of all aromatic compounds entering the reaction medium over a continuous period of time, preferably one hour, more preferably one day, and most preferably 30 consecutive days. This ratio is hereinafter referred to as the “molar survival ratio” for aromatic compounds when passing through the reaction medium and is expressed as a numerical percentage. If all incoming aromatic compounds leave the reaction medium as aromatic compounds, although for the most part in the oxidized forms of incoming aromatic compounds, then the molar survival ratio has its maximum value of 100%. To be precise, 1 out of every 100 incoming aromatic molecules is converted to carbon oxides and / or other non-aromatic molecules (e.g., acetic acid) while it is passing through the reaction medium, then the molar survival ratio is 99 % . Specifically, for the case where xylene is the main starting material of an oxidizable aromatic compound, the molar survival ratio for aromatic compounds through the reaction medium is preferably greater than about 98%, more preferably greater than about 98.5%, and most preferably less. than 99.0%. At the same time, and in order for a sufficient overall reaction activity to be present, the molar survival ratio for aromatic compounds when passing through the reaction medium is preferably less than about 99.9%, more preferably less than about 99, 8%, and most preferably less than 99.7%, when xylene is the main starting material of an oxidizable aromatic compound.

Another aspect of the present invention includes the production of methyl acetate in a reaction medium containing acetic acid and one or more oxidizable aromatic compounds. This methyl acetate is relatively volatile compared to water and acetic acid, and thus seeks to follow the flue gas, unless additional cooling or other plant operation is used to extract and / or destroy it before releasing the flue gas back into the surrounding Wednesday Thus, the formation of methyl acetate represents the cost of the work, as well as capital costs. Methyl acetate is probably formed by, first, combining the methyl radical, probably from the decomposition of acetic acid with oxygen, to obtain methyl hydroperoxide, through subsequent decomposition to form methanol and through the final reaction of the obtained methanol with the remaining acetic acid to form methyl acetate. Whatever the chemical route, the authors found that when the rate of formation of methyl acetate is too low, then the formation of carbon oxides is also too low, and the formation of harmful aromatic impurities is too high. If the methyl acetate formation rate is too high, then the formation of carbon oxides is also excessively high, resulting in loss of yield of solvent, oxidizable compound and oxidizing agent. When the preferred embodiments described herein are used, the ratio of the formation of moles of the resulting methyl acetate to the moles of the introduced oxidizable aromatic compound is preferably greater than about 0.005: 1, more preferably more than about 0.010: 1, and most preferably more than 0.020: 1. At the same time, the ratio of the formation of moles of the formed methyl acetate to the moles of the introduced oxidizable aromatic compound is preferably less than about 0.09: 1, more preferably less than about 0.07: 1, even more preferably less than 0.05: 1, and most preferably less than 0.04: 1.

Certain embodiments of the present invention may be further illustrated by the following examples, although it should be understood that these examples are included for illustrative purposes only and are not intended to limit the scope of the present invention, unless specifically indicated otherwise.

Examples 1-10

Example 1 is a numerical example of an oxidation reactor type bubble column for the oxidation of paraxylene in the liquid phase of a three-phase reaction medium. The reactor bubbler column of Example 1 is a tested industrial design with an accumulation rate of paraxylene starting materials of 7000 kg / h. Examples 2-10 are numerical examples for bubble columns of an oxidation type reactor having working vessels 7 times larger than the reactor of Example 1. FIG. 58 provides a table showing various parameters of bubble columns of an oxidation type reactor that vary in Examples 1- 10.

Example 1

This example uses a bubble column reaction vessel having a vertical cylindrical section with an inner diameter of 2.44 m. The height of the cylindrical section is 32 m from the lower bend start line (TL) to the upper TL of the cylindrical section. The tank is connected to the elliptical cover and the bottom 2: 1 in the upper and lower parts of the cylindrical section, respectively. The height from the bottom of the reaction medium to the top of the cylindrical section is approximately 32.6 m, and the total height of the reaction vessel is approximately 33.2 m. The operating level is approximately 25.6 m above the bottom of the reaction medium.

Paraxylene is introduced into the reactor at a constant speed of 7000 kg / h. The filtered solvent, containing mainly acetic acid, is introduced, carefully mixed with paraxylene, at a constant speed of 70,000 kg / h. The starting materials are distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using a horizontal distribution unit designed to provide substantially uniform release of the starting materials over a cross-sectional area lying within a radius of 0.45 · (internal diameter) . The concentration of catalyst components in the filtered solvent is such that the composition in the liquid phase of the reaction medium is 1800 ppm. wt. cobalt, 1800 ppm wt. bromine and 100 ppm wt. Manganese A separate solvent reflux stream is introduced in the form of droplets into the gas separation zone above the working level of the reaction medium at a constant speed of 49000 kg / h, and it is distributed over essentially the entire cross-sectional area of the separation zone. This solvent reflux does not contain significant levels of catalyst components. The combined water content of the starting materials of the filtered solvent and the starting materials of the reflux of the solvent is such that the concentration of water in the liquid phase of the reaction medium is 6% by weight. The feed speed of the air is constant at a speed of 35,000 kg / h through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The working gas pressure in the upper part of the reaction vessel is constantly equal to 0.52 MPa in the sensor. The reaction vessel works essentially adiabatically, so that the heat of reaction raises the temperature of the incoming feed and evaporates most of the incoming solvent. Measured near the mid-height of the reaction medium, the operating temperature is approximately 160 ° C. The reaction medium containing crude terephthalic acid (CTA) is removed from the side of the reaction vessel at a height of 15 m at a constant speed, using an external deaeration vessel.

The L: D ratio is 13.4, and the H: W ratio is 10.5. The volume occupied by the reaction medium is approximately 118 m ³ and the reaction vessel contains approximately 58,000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.3 hours. The spatio-temporal intensity of the reaction is approximately 59 kg of introduced paraxylene per cubic meter of reaction medium per hour.

The pressure difference from the lower part of the reaction medium to the exhaust gas in the upper part of the column leaving the reaction vessel is approximately 0.12 MPa. The surface velocity of the gas phase at a mid-height of the reaction medium is approximately 0.8 m / s. The vertically located surface area in contact with the reaction medium is approximately 197 m ² , which is approximately 3.16 Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is approximately 0.60 m. Under the conditions of Example 1, it was found that the decomposition of acetic acid in the reaction medium is approximately 0.03 kg per kilogram of CTA obtained. This represents a significant cost in the overall economy of production.

Useful indicators of flow rates from one level to another inside the reaction vessel of the bubble column are the maximum time-averaged upward speeds of the suspension and phase of the oxidizing agent. In the cylindrical reaction vessel of the bubble column, these maximum upward velocities arise near the vertical axis of symmetry of the cylindrical section. Calculated by a method taken from the Gas-Liquid Recirculation Model of Gupta (reference Churn-turbulent bubble columns: experiments and modeling; Gupta, Puneet; Washington Univ., St. Louis, MO, USA. Avail. UMI Order No. DA3065044; ( 2002), 357 pp. From: Diss. Abstr. Int., In 2003, 63 (9), 4269. Dissertation written in English. CAN 140: 130325 AN 2003: 424187 CAPLUS), and using the appropriate database the maximum averaged in time, the upward velocity of the gas phase near the mid-height of the reaction medium is approximately 3.1 m / s. The similarly calculated maximum upward velocity averaged over time for a suspension near the median height of the reaction medium is approximately 1.4 m / s.

Another useful indicator of flow rates from one end to the other inside the reaction vessel of the bubbler column is the maximum time-averaged downward velocity of the suspension in parts of the reaction vessel located away from the central core. In a cylindrical reaction vessel, this maximum downward velocity usually occurs in a region lying outside a radius of 0.35 · (internal diameter) from the vertical axis of symmetry of the cylindrical section. Calculated using a method taken from the Gas-Liquid Recirculation Model of Gupta and using the appropriate database, the maximum time-averaged downward velocity of the suspension in the outer annular gap near the median height of the reaction medium is approximately 1.4 m / s.

Example 2

In the present example, paraxylene is introduced into the bubbler column of the reactor type at an increased speed of 49,000 kg / h — 7 times more than in Example 1. The surface gas velocity, often considered as an important scaling variable for bubbler columns, is maintained approximately equal to Example 1 by increasing the cross-sectional area of the reaction vessel is about 7 times, compared with Example 1. The H: W and L: D ratios, often considered important scaling variables for bubble columns, are also n dderzhivayutsya approximately equal to Example 1.

Other feed streams increase at the same ratio of 7: 1, as compared to Example 1. The feed compositions are the same as in Example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium as in Example 1. The working gas pressure in the upper part of the reaction vessel is again 0.52 MPa in the sensor, and the working temperature is again approximately 160 ° C, measured near the mid-height of the reaction medium. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 40 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 6.46 m. The L: D ratio is maintained the same as in example 1, and the height from the lower part of the reaction medium to the upper part of the cylindrical section is thus 86, 3 m. The upper and lower parts of the cylindrical section are connected to the elliptical bottom and the cover 2: 1, and the total height of the reaction vessel is very large, approximately 88.0 m. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom s part of the reaction medium using a horizontal distributor assembly designed for a uniform, substantially release starting materials by cross-sectional area lying inside radius 0.45 · (internal diameter). Air starting materials are again introduced through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The phlegm of the solvent is distributed in the form of droplets essentially over the entire cross-sectional area of the separation zone.

The H: W ratio of the reaction medium is maintained approximately equal to Example 1, and thus the working level is approximately 67.8 m of the reaction medium. This leaves a compartment height of approximately 18.5 m in the cylindrical section plus approximately 1.6 m in the upper elliptical cover. This separation height is approximately 10 m redundant. Thus, scaling a container with a constant L: D gives a mechanical design that is unnecessarily expensive in terms of capital costs (for example, excessive costs for a high-pressure tank, for foundations due to mass and wind loads, structural steel, technological and service pipelines and / or tools and electric cables).

The volume occupied by the reaction medium is approximately 2,200 m ³ , and the reaction vessel contains approximately 1,100,000 kg of suspension. The ratio of the mass of the suspension to the speed of the starting materials of paraxylene is approximately 22 hours, which is much larger compared to example 1. The spatio-temporal intensity of the reaction is only about 22 kg of introduced paraxylene per cubic meter of reaction medium per hour, which is much lower compared to example 1 The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel increases to about 0.33 MPa, which is much larger compared to example 1. Location ennaya upright surface area in contact with the reaction medium is about 1,393 m ^2, which is about 3,18 · Wmin · H. The ratio of the volume of the reaction medium to the vertically arranged surface area in contact with the reaction medium is approximately 1.58 m, which is much larger compared to Example 1.

Thus, scaling the size of the reaction medium to maintain both surface velocity and the H: W ratio, which are approximately constant between this example and example 1, produces very large changes in the reaction conditions. In general, these changes are very unfavorable. In the present example, there are positive effects (for example, more dilute concentrations of the oxidizable compound, lower requirements for mass transfer rate per unit volume for molecular oxygen from the gas phase to the liquid phase, and so on), resulting in less undesirable colored by-products per CTA unit . However, there are severe economic burdens related to the decomposition of acetic acid and the pressure and energy necessary to supply air to the lower part of the bubble column reaction vessel. The decomposition of acetic acid is approximately proportional to the mass of acetic acid in the reaction medium when the operating temperature and composition of the liquid phase are kept approximately constant and when paraxylene is introduced with excess molecular oxygen. Due to the much greater mass of acetic acid in the reaction vessel, compared with the amount of CTA obtained in this example, it was found that the decomposition of acetic acid increases to about 0.09 kg per kg of CTA obtained. In addition, in the present example, the air compressor should deliver air to the reaction medium at a pressure that is 0.85 MPa in the sensor, while in example 1, the delivery pressure is 0.64 MPa in the sensor. For an air delivery rate of 245,000 kg / h and with a typical tolerance for various compression and delivery efficiencies, the additional energy requirement for a higher delivery pressure in the present example is approximately 3000 kW, continuously. Thus, scaling the reaction medium in the present example with approximately constant surface gas velocities and H: W ratios provides an unacceptable economy, despite the good expected CTA quality.

Example 3

This example scales the method of Example 1 using surface velocity and the spatio-temporal intensity of the reaction. This leads to poor product quality because, in simple words, the natural convection flow patterns initially give a poor vertical reaction profile.

In the present example, the paraxylene feed speed is again equal to 49000 kg / h — 7 times more than in Example 1. The surface gas velocity is again maintained approximately equal to Example 1, but the L: D and H: W ratios are not maintained equal. Instead, STR is maintained at approximately equal to Example 1. This provides a basic column pressure and decomposition ratio of acetic acid, which are approximately equal to Example 1. Other feed streams increase at the same ratio of 7: 1, for example 1. The feed compositions are same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working gas pressure in the upper part of the reaction vessel is again 0.52 MPa in the sensor, and the working temperature is again approximately 160 ° C, measured near the median height of the reaction medium. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 15 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 6.46 m, maintaining the surface velocity of the gas phase approximately constant, compared with examples 1 and 2. To maintain the same STR as in example 1, the working level varies slightly, up to about 26.1 m of reaction medium. The height from the lower part of TL to the upper part of TL of the cylindrical section is 32 m, as in Example 1, providing approximately the same height of separation with free walls between the upper part of the reaction medium and the gas outlet from the upper part of the column. The upper and lower parts of the cylindrical section are connected to the elliptical bottom and cover, 2: 1. The height from the bottom of the reaction medium to the top of the cylindrical section is approximately 33.6 m, and the total height of the reaction vessel is approximately 35.2 m. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using horizontal distribution unit, designed for a uniform, essentially, release of the starting materials over a cross-sectional area lying inside a radius of 0.45 · (internal diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The phlegm of the solvent is distributed in the form of droplets essentially over the entire cross-sectional area of the separation zone.

The H: W ratio of the reaction medium decreases markedly to 4.0. The L: D ratio of the reaction vessel decreases markedly to 5.2. The volume occupied by the reaction medium is approximately 828 m ³ and the reaction vessel contains approximately 410000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.3 hours. STR is approximately 59 kg of paraxylene injected per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.13 MPa. The vertically located surface area in contact with the reaction medium is approximately 546 m ² , which is approximately 3.24 Wmin · H. The ratio of the volume of the reaction medium to the vertical surface area in contact with the reaction medium is approximately 1.52 m. Under the conditions of this example, it was found that the decomposition of acetic acid in the reaction medium desirably returns to the lower level of Example 1, approximately 0.03 kg / kg of obtained CTA.

However, the larger diameter of the reaction vessel used in the present example, combined with its lower H: W ratio, leads to very undesirable changes in flow rates, mixing, and stepwise changes in concentration in the reaction medium. This leads to a significant increase in the loss of paraxylene in the flue gas from the top of the column and to the formation of undesirable colored by-products. Simply put, axial flow velocities caused by forces of natural convection increase more strongly in bubbler columns of larger diameter, even if the surface velocity is kept constant. Calculated by the method taken from the Gas-Liquid Recirculation Model of Gupta and using the appropriate database, the maximum upward velocity averaged over time of the gas phase near the median height of the reaction medium is approximately 3.9 m / s. Similarly calculated, the maximum time-averaged velocity of upward movement of the suspension near the median height of the reaction medium is approximately 2.2 m / s. Similarly calculated, the maximum time-averaged downward velocity of the suspension downward in the outer annular gap near the median height of the reaction medium is approximately 2.3 m / s.

Since the height of the reaction medium varies slightly between this example and example 1, this increase in time-averaged vertical speeds causes a significant reduction in mixing times from one edge to the other in the present example, compared with example 1. This causes undesirable increases in the amount of paraxylene migrating in the direction top of the tank before oxidation. This leads to undesirable losses in the yield of paraxylene leaving the upper part of the reaction vessel together with the flue gas, and more strongly changes the need for dissolved molecular oxygen near the upper part of the reactor, where the molar fraction of molecular oxygen in the gas phase is relatively depleted. In addition, an increase in the time-averaged downward velocity of the suspension in the regions closer to the walls of the container in the present example causes a larger number of gas phase bubbles and larger ones to move downward, against their natural rise in the gravitational field. This leads to an undesirable increase in gas phase recirculation, partially depleted of molecular oxygen, which, in turn, leads to a decrease in the availability of dissolved oxygen in these areas. Among other effects, this reduced availability of dissolved oxygen in various parts of the reaction medium leads to a significant increase in the formation ratio of undesirable colored by-products in the present example, compared to example 1, and this increased level of undesired, colored by-products makes the product unsuitable for many applications PET

Thus, examples 2 and 3 demonstrate the inadequacy of the current state of the art in the construction of large-scale oxidation bubble columns using mainly the surface gas velocity (Ug), the L: D ratio, and the average spatio-temporal reaction rate (STR).

Example 4

In the present example, the pressure-holding elements of the bubble column reaction vessel are the same as in Example 3, but surfaces vertically arranged in the reaction medium are added to impart vertical thrust in order to establish stepwise reaction profiles more similar to Example 1, thereby restoring the quality of the products and the yield of paraxylene, but without increasing the degree of decomposition of acetic acid, as in example 2.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction with food. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 15 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 6.46 m. The height from the lower part of the TL to the upper part of the TL of the cylindrical section is 32 m, and the working level is approximately 26.3 m of the reaction medium. The upper and lower parts of the cylindrical section are connected to the elliptical bottom and cover, 2: 1. The height from the bottom of the reaction medium to the top of the cylindrical section is approximately 33.6 m, and the total height of the reaction vessel is approximately 35.2 m. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using horizontal distribution unit, designed for a uniform, essentially, release of the starting materials over a cross-sectional area lying inside a radius of 0.45 · (internal diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The phlegm of the solvent is distributed in the form of droplets essentially over the entire cross-sectional area of the separation zone.

The reaction capacity of the bubbler column further comprises two orthogonal flat surfaces arranged so that their intersection line coincides with the vertical axis of symmetry of the cylindrical section. Conveniently, these flat surfaces contain metal plates and surface finishes of the same type as used in the cylindrical section of the reaction vessel. Each flat surface starts at a lower height of 3 m above the lower TL and extends upwards by 20 m. Two flat surfaces, each, extend horizontally substantially along the entire length of the wall of the cylindrical section (i.e., the width of each is equal to the inner diameter) and is supported from the side cylindrical section. The thickness and support of flat surfaces are designed to withstand the various forces that can occur in normal and extreme working conditions. Due to the volume occupied by the metal of the plates, the working level is set slightly higher than 26.3 m above the bottom of the reaction medium to maintain the same STR as in example 1.

Thus, in the present example, the reaction medium is divided into 4 subvolumes of the same shape and size, 20 m each, from the total height of the reaction medium. These 4 subvolumes communicate with each other both below and above flat surfaces. Due to the relatively uniform distribution of the starting materials of the oxidizing agent and paraxylene below the lower point of the flat surfaces, each of the 4 subvolumes has a similar surface gas phase velocity and a similar reaction intensity profile. 4 subvolumes can be considered as an approximation of 4 reaction vessels of bubbler columns of smaller sizes in the shell of one vessel that holds pressure.

The H: W ratio of the reaction medium is 4.1. The L: D ratio of the reaction vessel is 5.2. The volume occupied by the reaction medium is approximately 828 m ³ and the reaction vessel contains approximately 410000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.3 hours. STR is equal to about 59 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.13 MPa. The vertically located surface area in contact with the reaction medium is about 1066 m ² , which is about 6.29 · Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is 0.78 m. This value is intermediate between examples 1 and 3, and is preferably close to the value of example 1. Under the conditions of this example, it was found that the decomposition of acetic acid in the reaction medium desirably returns to the lower level of Example 1, approximately 0.03 kg / kg of the obtained CTA.

The maximum time-averaged up and down velocities of the gas phase and suspension are reduced in the present example compared to Example 3. This provides useful improvements in the vertical profile of paraxylene, and this leads to improved availability of dissolved oxygen in the liquid phase near the surfaces of vertically arranged walls. Together, these changes improve the yield of paraxylene and reduce the formation of undesirable colored by-products in the present example, compared with example 3.

Example 5

In the present example, the pressure-holding elements of the reaction vessel are the same as in Example 3, but the sediment-free deflector elements are added to the reaction medium to restore stepwise reaction profiles more similar to Example 1, thereby restoring product quality and paraxylene yield, but without increasing the decomposition ratio of acetic acid, as in example 2.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and the working temperature is again approximately 160 ° C, measured near the middle height of the reactionary Reda. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 15 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 6.46 m. The height from the lower part of the TL to the upper part of the TL of the cylindrical section is 32 m, and the working level is approximately 26.1 m of the reaction medium. The upper and lower parts of the cylindrical section are connected to the elliptical bottom and cover, 2: 1. The height from the bottom of the reaction medium to the top of the cylindrical section is approximately 33.6 m, and the total height of the reaction vessel is approximately 35.2 m. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium, using horizontal distribution unit, designed for a uniform, essentially, release of the starting materials on a cross-sectional area lying inside a radius of 0.45 · (internal diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The phlegm of the solvent is distributed in the form of droplets essentially over the entire cross-sectional area of the separation zone.

The reaction capacity of the bubble column further comprises a horizontal deflector assembly located in the bubble column at a height of 12 m above the lower TL of the reaction vessel. This places the deflector assembly at approximately 13.6 m, or 2.1 · D, above the bottom of the reaction medium. This deflector assembly consists of 15 individual deflector elements. Each deflector element consists of an extruded or prefabricated L-shaped corner, where both sides of the L-shaped are 0.15 m wide and the internal angle between the two sides is 90 °. All L-shaped corners are located horizontally, parallel to each other, with corners looking up, and are located at the same height. The two end edges of each L-shaped corner are at the same height, below the angles looking up. Seen from the end, each element looks like an inverted letter V. Thus, the percentage of the deflector assembly containing upward looking flat surfaces inclined less than 5 ° from the horizontal is effectively zero. The gap between the lower surface levels of each element and its closest neighbor is always 0.21 m. The longest element has a length effectively equal to the inner diameter of the cylindrical container, extending diametrically through the container from wall to wall. The remaining 14 separate deflector elements, all of them, are necessarily shorter in length. All deflector elements are supported on each edge, extending, all the time, to the cylindrical wall and attaching to it. Thus, the open area at the height of the deflector assembly is approximately 16 m ² , which is approximately 50% of the cross-sectional area of the reaction vessel at this height. Deflector elements are designed to withstand the various forces that can occur under normal and extreme operating conditions. The elements are constructed from the same metal as the components of the pipelines used in the air distribution device, in which the metal is suitably selected so that it resists corrosion and erosion. However, the surfaces of the deflector elements are polished to a surface finish of 125 RMS or thinner. Despite the deposition of approximately 76,000 kg / h of CTA in the reaction vessel, the deflector assembly does not accept excess sediment or pieces of solid products slide off it.

The H: W ratio of the reaction medium is 4.0. The L: D ratio of the reaction vessel is 5.2. The volume occupied by the reaction medium is approximately 828 m ³ and the reaction vessel contains approximately 410000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.3 hours. STR is equal to about 59 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.13 MPa. The vertically located surface area in contact with the reaction medium is about 546 m ² , which is about 3.24 · Wmin · H. The ratio of the volume of the reaction medium to the vertical surface area in contact with the reaction medium is approximately 1.52 m. Under the conditions of this example, it was found that the decomposition of acetic acid in the reaction medium desirably returns to the lower level of Example 1, approximately 0.03 kg / kg of obtained CTA.

The effect of the horizontal node of the deflectors is to perturb the vertical velocity of the gas phase and the suspension in the reaction vessel. This slows down the progression of paraxylene towards the upper surface of the reaction medium, leading to a useful reduction in the loss of paraxylene output in the off-gas from the top of the column. In addition to this, a stepwise change in the concentration of molecular oxygen and an oxidizable compound is improved, thereby reducing the formation of undesired colored by-products in the present example, as compared with Example 3.

Example 6

In the present example, the reaction vessel is designed for very high surface gas velocities and gas contents in accordance with the present invention. The use of lower D values makes it possible to have a higher L: D ratio without resorting to an excessively high reaction capacity and without carrying out excessive decomposition of the solvent, acetic acid.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction with food. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 28 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 5.00 m. The height from the lower part of the TL to the upper part of the TL of the cylindrical section is 70 m. The upper and lower parts of the cylindrical section are connected to the elliptical bottom and cover, 2: 1. The height from the lower part of the reaction medium to the upper part of the cylindrical section is approximately 71.3 m, and the total height of the reaction vessel is approximately 72.5 m. The operating level is approximately 61.3 m above the lower part of the reaction medium. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using a horizontal distribution unit designed to substantially uniformly release the starting materials over a cross-sectional area lying within a radius of 0.45 * (inner diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidant are below the lower TL of the cylindrical section. The phlegm of the solvent is distributed in the form of droplets over the entire substantially cross-sectional area of the separation zone.

The H: W ratio of the reaction medium is 12.3. The L: D ratio of the reaction vessel is 14.3. The volume occupied by the reaction medium is approximately 1,190 m ³ and the reaction vessel contains approximately 420,000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.7 hours. STR is about 41 kg of paraxylene injected per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.21 MPa. The vertically located surface area in contact with the reaction medium is about 975 m ² , which is about 3.18 · Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is approximately 1.22 m. The relatively small value of D creates the surface velocity of the gas phase at a median height of the reaction medium, which exceeds approximately 1.7 times the surface velocity used in the examples 1-5. The gas content at the mid-height of the reaction medium is more than 0.6. Under the conditions of example 6, it was found that the decomposition of acetic acid in the reaction medium desirably decreases below 0.03 kg / kg of the resulting CTA. This is due to a decrease in the amount of suspension, more specifically acetic acid, in the reaction vessel, compared with example 3.

In this example, the H: W ratio is favorable to reduce mixing from one edge to the other and to advantageous stepwise change in the concentrations of molecular oxygen and the oxidizable compound. However, the axial speeds are higher than in example 1, which accelerates mixing from one edge to the other, for a given H: W. Fortunately, a lower STR reduces the volumetric need for the transfer of molecular oxygen from gas to liquid; and an increase in gas content serves to increase the ability to transfer molecular oxygen from gas to liquid. As a result, the level of production of undesirable colored by-products was found to be comparable to Example 1.

Example 7

In this example, the reaction vessel is designed for even higher values of surface gas velocity and gas content in accordance with the present invention. The enlarged, upstream separation zone is used to limit the capture of the suspension in the flue gas from the top of the column.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction with food. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 28 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains a vertical cylindrical section with an inner diameter of 4.60 m. The height from the lower part of TL to the upper edge of the cylindrical section is 60 m. At the upper edge of this cylindrical section, the conical section diverges to an internal diameter of 7 m, at the same time, rising in height by 2 m. The slope of the conical wall is thus approximately 31 ° from the vertical. Above the conical section is a cylindrical section for separating gas with an inner diameter of 7 m. The height of the upper cylindrical section is 7 m. The capacity is connected to the elliptical bottom and cover, 2: 1, in the upper and lower parts. Thus, the combined height of the reaction vessel is approximately 71.9 m. The operating level is approximately 61.2 m above the bottom of the reaction medium, located near the interface between the main cylindrical body and the diverging conical section. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using a horizontal distribution unit designed to substantially uniformly release the starting materials over a cross-sectional area lying within a radius of 0.45 * (internal diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidizing agent are below the lower TL of the lower cylindrical section. The phlegm of the solvent is distributed in the form of droplets over essentially the entire cross-sectional area of the expanded section of the compartment.

The ratio H: W of the reaction medium and the ratio L: D of the reaction vessel are 13.3. The X: D ratio of the reaction vessel is 1.5. The L: Y ratio of the reaction vessel is 5.7. The volume occupied by the reaction medium is approximately 1000 m ³ and the reaction vessel contains approximately 320,000 kg of suspension. The ratio of the mass of the suspension to the speed of the starting materials of paraxylene is approximately 6.5 hours. STR is equal to about 49 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.19 MPa. The vertically located surface area in contact with the reaction medium is about 896 m ² , which is about 3.19 · Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is approximately 1.12 m. The relatively small value of D creates the surface velocity of the gas phase at a median height of the reaction medium, which is approximately 2 times higher than the surface velocity used in examples 1- 5. However, the surface velocity of the gas phase in the expanded separation section decreases to approximately 0.85 from the surface velocity used in Examples 1-5. The gas content at the mid-height of the reaction medium is about 0.7. Under the conditions of this example, it was found that the decomposition of acetic acid in the reaction medium desirably decreases below 0.03 kg / kg of the obtained CTA. This is due to a decrease in the amount of suspension, more specifically acetic acid, in the reaction vessel, as compared with Example 3. It was found that the level of undesirable colored by-products is lower than in Example 6 due to the improvement of stepwise concentration changes and a higher gas content.

Example 8

In this example, the reaction vessel is the same as in example 7, but the operating level rises to about 63.2 m above the bottom of the reaction medium, located near the interface of the diverging conical section and the expanded cylindrical section for gas separation. This provides various advantages over level control in the main cylindrical section of the housing, including a decrease in the tendency, for the upper part of the reaction medium, both to foaming and to too poor circulation.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction with food. The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 28 m at a constant speed using an external deaeration vessel.

The H: W ratio of the reaction medium is approximately 13.7, and the L: D ratio of the reaction vessel is 13.3. The volume occupied by the reaction medium is approximately 1060 m ³ , and the reaction vessel contains approximately 330,000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 6.8 hours. STR is equal to about 46 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.20 MPa. The vertically located surface area in contact with the reaction medium is about 953 m ² , which is about 3.39 · Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is approximately 1.11 m. The surface velocity of the gas phase at the mid-height of the reaction medium is approximately 2 times higher than the surface velocity used in Examples 1-5. The gas content at the mid-height of the reaction medium is approximately 0.7. Under the conditions of this example, it was found that the decomposition of acetic acid in the reaction medium desirably decreases below 0.03 kg / kg of the obtained CTA. This is due to a decrease in the amount of suspension, more specifically acetic acid, in the reaction vessel, as compared with Example 3. It was found that the level of undesirable colored by-products is lower than in Example 6, due to the improvement of the stepwise change in the concentration and gas content.

Example 9

In this example, the pressure-holding elements of the reaction vessel are the same as in Example 7, but the internal elements used to introduce the oxidizing agent and paraxylene are significantly modified, providing a plurality of inlets separated by a certain vertical distance for each of them.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction medium . The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 28 m at a constant speed using an external deaeration vessel.

The reaction capacity of the bubble column contains the same oxidizer distributor at the bottom of the reaction vessel as in Example 8. However, only 70% of the total gas-phase oxidizer stream is introduced through this lower distributor. Another 30% of the gas phase oxidizer is introduced through a raised oxidizer inlet. This flow ratio is superimposed by flow control loops using control valves and flow sensors conveniently located on the supply passages for compressed air outside the reaction vessel. The raised oxidizer dispenser contains a horizontal, beveled, square-shaped flow passage, instead of an octagonal passage, as used in the bottom. Conveniently, the square-shaped passageway contains nominally 14-inch Schedule 10S pipe materials. The distance from the centroid on one side to the centroid on the opposite side is 1 m. The raised oxidizer distributor contains about sixty outlets for a gas phase oxidizer, all 0.03 m in diameter, and near the bottom of the aisle, approximately 14 m above the bottom part of the reaction medium. This elevated oxidizing agent distributor serves at least two useful functions. Firstly, the oxidizer flow injected down into the reaction medium destroys the axial velocity profile, rising along the vertical axis of symmetry of the cylindrical section. It performs the useful function of a hydraulic deflector to slow down the distribution of paraxylene in the upper regions of the reaction medium, associated with output losses in the upper part of the column, and to reduce the flow rate in dissolved oxygen in the upper regions. A few meters above the elevated inlet for the oxidizer, the natural convection flow structure reorganizes spontaneously, rising up along the central axis of symmetry, but, nevertheless, the function of the hydraulic deflector is carried out efficiently. Secondly, most of the heat of the reaction is removed from the reaction medium by evaporation of the solvent, and most of this evaporation is carried out in positions near the input of the oxidizing materials. Separating vertically the positions of the introduction of the parts of the gas-phase flow of the oxidizing agent, the vertical temperature profile in the reaction medium is regulated.

The bubble column reaction vessel contains two inlet dispensers for paraxylene, similar to that of Example 8. A lower inlet dispenser for paraxylene is arranged to provide substantially uniform release of 50% of the liquid phase feed materials over a cross-sectional area lying within a radius of 0.45 (inner diameter) at a lower height of 2 m above the bottom of the reaction medium. An upper inlet distributor for paraxylene is arranged to provide substantially uniform release of 50% of the liquid phase feed materials over a cross-sectional area lying within a radius of 0.45 * (internal diameter) at a higher height of 10 m above the lower part of the reaction medium. This flow ratio is superimposed by flow control loops using control valves and flow sensors conveniently located on the feed passages for liquid phase feed materials outside the reaction vessel.

In this example, the operating level is raised to about 63.7 m above the bottom of the reaction medium, located just above the diverging conical section and in the expanded cylindrical section for separating gas. The H: W ratio of the reaction medium is approximately 13.8, and the L: D ratio of the reaction vessel is 13.3. The volume occupied by the reaction medium is approximately 1070 m ³ and the reaction vessel contains approximately 340,000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 6.9 hours. STR is equal to about 46 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.20 MPa. The vertically located surface area in contact with the reaction medium is about 975 m ² , which is about 3.47 · Wmin · H. The ratio of the volume of the reaction medium to the vertically located surface area in contact with the reaction medium is approximately 1.10 m. The surface velocity of the gas phase at the mid-height of the reaction medium is approximately 2 times greater than the surface velocity used in Examples 1-5.

Example 10

In this example, the reaction vessel is constructed with three different cylindrical diameters at different heights, with the diameter at the average height being the smallest. This configuration wins thanks to the lower cylindrical section, where the liquid feed stream and the gas phase oxidizer first enter, with a relatively large mass of the liquid phase for the initial dilution and interaction of paraxylene, where oxygen is still more redundant; thanks to the middle cylindrical section, where molecular oxygen is becoming increasingly depleted, with a relatively higher gas content and mass transfer rate from gas to liquid; and thanks to the upper cylindrical section, which is a gas separation zone, with a relatively reduced velocity of the gas phase, to limit the capture of the suspension in the exhaust gas from the top of the column.

The speed of paraxylene starting materials is again equal to 49000 kg / h - 7 times more than in example 1. Other streams of starting materials increase with the same ratio of 7: 1, for example 1. The compositions of the starting materials are the same as in example 1, providing the same concentrations of water, cobalt, bromine and manganese in the liquid phase of the reaction medium, as in example 1. The working pressure of the gas reaction vessel in the upper part is again 0.52 MPa in the sensor, and again the working temperature is approximately 160 ° C, measured near the median height of the reaction medium . The reaction medium containing CTA is removed from the side of the reaction vessel at a height of 28 m at a constant speed using an external deaeration vessel.

The bubble column contains three vertical cylindrical sections with different diameters. The lowest cylindrical section has an inner diameter of 6.46 m, giving the surface velocity of the gas phase in this section approximately equal to the surface velocity of Example 1. The height of this lower cylindrical section from the lower TL to the upper edge is 8 m. On the upper edge of this lower cylindrical section the conical section diverges to an internal diameter of 4.5 m, while rising to a height of 1 m. The slope of this conical wall, thus, is approximately 44 ° from the vertical. The middle cylindrical section located above the lower conical section has an inner diameter of 4.5 m, giving the surface velocity of the gas phase in this section, approximately two times the surface velocity in the lowest cylindrical section. The height of the middle cylindrical section is 45 m. On the upper edge of the middle cylindrical section, the conical section diverges to an inner diameter of 7 m, while rising 2 m in height. The inclination of the conical wall is thus approximately 32 ° from the vertical. Above the upper conical diverging section is a cylindrical section for separating gas with an inner diameter of 7 m. The height of the upper cylindrical section is 7 m. The capacity is connected to the elliptical bottom and cover, 2: 1, of the upper and lower parts. Thus, the combined height of the reaction vessel is approximately 66.4 m. The operating level is at an altitude of approximately 57.6 m above the lower part of the reaction medium, located near the interface of the diverging conical section and the upper cylindrical section. The starting materials are again distributed in the reaction vessel approximately 2 m above the bottom of the reaction medium using a horizontal distribution unit designed to substantially uniformly release the starting materials over a cross-sectional area lying within a radius of 0.45 * (internal diameter). Air starting materials again pass through an oxidizer inlet distributor, similar to that shown in FIG. 12-15, and all openings for the accumulation of oxidizing agent are below the lower TL of the lowest cylindrical section. The phlegm of the solvent is distributed in the form of droplets essentially over the entire cross-sectional area of the upper cylindrical section.

The volume occupied by the reaction medium is approximately 1080 m ³ and the reaction vessel contains approximately 400,000 kg of suspension. The ratio of the mass of the suspension to the speed of the paraxylene starting materials is approximately 8.1 hours. STR is equal to about 45 kg of paraxylene starting materials introduced per cubic meter of reaction medium per hour. The pressure difference from the lower part of the reaction medium to the exhaust gas from the upper part of the column leaving the reaction vessel is approximately 0.20 MPa. The vertically located surface area in contact with the reaction medium is about 944 m ² , which is about 3.64 · Wmin · H and about 2.34 · Wmax · H. The ratio of the volume of the reaction medium to the vertical surface area in contact with the reaction medium is approximately 1.14 m. The ratio L _l : D ₁ is approximately 1.5: 1. The ratio of L _u : D _u is approximately 10: 1. The ratio of L _l : L _u is approximately 0.21: 1. The ratio of X: D ₁ is approximately 1.1: 1. The ratio of L _u : Y is approximately 4.2: 1. The ratio of L _t : D ₁ is approximately 0.15: 1.

The larger diameter at the base of the reactor provides a large mass of slurry near the injection zone for the paraxylene starting materials, where liquid flows and mixing is absolutely essential to ensure dilution of the starting materials to eliminate associated colored aromatic impurities. Also, this larger diameter places a larger fraction of the reaction medium under a higher pressure from the suspension from above, increasing the partial pressure of oxygen and the mass transfer of molecular oxygen from gas to liquid. Having a smaller diameter, the elongated middle cylindrical section provides a stepwise change in the concentration of reagents and a high gas content; this improves the coordination of the flow rate of the reaction in dissolved oxygen with the supply by mass transfer from the rising gas phase, where oxygen is consumed more and more and its partial pressure decreases. The present invention is described in detail with specific references to its preferred embodiments, but it will be understood that variations and modifications can be made within the spirit and scope of the present invention.

Claims (30)

1. The method of liquid-phase catalytic oxidation of an aromatic compound, comprising introducing a source of oxidizable material into the reaction zone of the initial reactor, introducing a gas-phase stream of an oxidizing agent containing molecular oxygen, oxidizing the oxidizing compound in the liquid phase of the reaction medium contained in a stirred reactor, which is a bubble column reactor type where the specified reaction medium has a gas content of at least about 0.6, averaged over time and averaged over to lift. 2. The method according to claim 1, in which the specified reaction medium contains at least about 3 wt.% Solid products, with averaging over time and averaging over volume. 3. The method according to claim 1, in which the specified initial reactor is a bubble column reactor type, where the specified reaction medium has a gas content in the range from 0.65 to 0.85, averaged over time and averaged over the volume, where the specified reaction medium contains from about 5 to about 40 wt.% solid products, with averaging over time and averaging over volume. 4. The method according to claim 1, in which the time-averaged flow rate of the gas phase of the reaction medium at half the height of the specified reaction medium is in the range from about 0.8 to about 5 m / s. 5. The method according to claim 1, wherein said oxidizable compound is an aromatic compound. 6. The method according to claim 1, in which the specified oxidizable compound is selected from the group consisting of para-xylene, meta-xylene, para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluic acid, acetaldehyde and combinations of two or more of them. 7. The method according to claim 1, wherein said oxidizable compound is para-xylene. 8. The method according to claim 1, wherein said reaction medium is a three-phase reaction medium. 9. The method according to claim 1, wherein said oxidation causes the formation of solid products in said reaction medium. 10. The method according to claim 1, wherein said oxidation causes at least about 10% by weight of said oxidizable compound to form solid products in said reaction medium. 11. The method according to claim 1, in which the specified initial reactor is a bubble column reactor type. 12. The method according to claim 11, in which the specified oxidizable compound contains para-xylene. 13. The method according to claim 1, in which the specified oxidation is carried out in the presence of a catalyst system containing cobalt. 14. The method of claim 13, wherein said catalyst system further comprises bromine and manganese. 15. The method according to claim 1, wherein said reaction medium has a maximum height (H), a maximum width (W), and an H: W ratio of at least about 3: 1. 16. The method according to clause 15, in which the specified ratio H: W is in the range from about 8: 1 to about 20: 1. 17. The method according to clause 15, in which the specified method further includes introducing a mainly gas-phase flow of an oxidizing agent containing molecular oxygen, in the specified reaction zone. 18. The method according to 17, in which most of the specified molecular oxygen enters the specified reaction zone within about 0.25 W from the bottom of the specified reaction zone. 19. The method according to 17, in which most of the specified molecular oxygen enters the specified reaction zone within about 0.025H from the bottom of the specified reaction zone. 20. The method of claim 17, wherein said method further comprises introducing a substantially liquid phase feed stream containing said oxidizable compound into said reaction zone. 21. The method according to claim 20, in which the specified initial reactor is a bubble column reactor type. 22. The method according to claim 20, in which at least about 50 wt.% The specified oxidizable compounds enters the specified reaction zone within about 2.5 W from the lowest position, where the specified molecular oxygen enters the specified reaction zone. 23. The method according to claim 20, wherein said feed stream is introduced into said reaction zone through a plurality of inlets, where at least two of said inlets are vertically separated from each other by a distance of at least about 0, 5W. 24. The method according to clause 34, in which at least two of these inlets are separated from each other vertically by a distance of at least about 1.5W. 25. The method according to claim 20, in which 80 wt.% Of the oxidizable compound of the feed stream is introduced into one vertical quadrant of the reaction zone, determined theoretically by a pair of intersecting vertical planes dividing the reaction zone into 4 vertical quadrants with equal volume. 26. The method according to claim 20, in which at least a portion of the specified reaction zone is determined by one or more vertically arranged side walls of the specified reactor, where at least about 25 wt.% Of the specified oxidizable compounds enters the specified reaction zone in one or more positions distant inward by at least 0.05D from said vertically positioned side walls, where said reaction zone has a maximum diameter (D). 27. The method according to claim 1, wherein said oxidation in said initial reactor causes terephthalic acid to form in said reaction medium, wherein said method further comprises exposing at least a portion of said terephthalic acid to oxidation in a secondary oxidation reactor. 28. The method according to item 27, in which the specified oxidation in the specified secondary oxidation reactor is carried out at an average temperature of at least about 10 ° C greater than the specified oxidation in the specified initial reactor. 29. The method according to item 27, in which the specified oxidation in the specified secondary oxidation reactor is carried out at an average temperature exceeding, in the range of from about 20 to about 80 ° C., The average temperature of the specified initial reactor, where the specified oxidation in the specified initial reactor is carried out at an average temperature in the range of about 140 to about 180 ° C, where said oxidation in said secondary oxidation reactor is carried out at an average temperature of about 180 to about 220 ° C. 30. The method according to claim 1, wherein said oxidation causes the formation of crude terephthalic acid particles in said reaction medium, wherein a representative sample of said crude terephthalic acid particles has one or more of the following characteristics: (i) contains less than about 12 m .d. wt. 4,4-dicarboxylstilbene (4,4-DCS), (ii) contains less than about 800 ppm. wt. isophthalic acid (IPA), (iii) contains less than about 100 ppm. wt. 2,6-dicarboxyfluorenone (2,6-DCF), (iv) has a transmittance percentage at 340 nm of about 25% more.

Images