RU2384563C2 - Optimised liquid-phase oxidation - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: method involves oxidation of para-xylene in the liquid phase of a multiphase reaction medium contained in one or more primary oxidation reactors, in which the reaction medium contains a first separate 20% continuous volume having a first oxygen space-time rate (STR), and a second a second separate 20% continuous volume having a second oxygen STR, where the ratio of the first oxygen STR to the second oxygen STR is at least 1.5:1. The second version of the method involves feeding para-xylene into a reaction zone of a reactor-type bubble column, feeding a stream of oxidising agent, and oxidation of at least a portion of the para-xylene in the liquid phase of the multiphase reaction medium contained in the reaction zone of the primary oxidation reactor in which oxidation takes place such that, if the said reaction zone was theoretically divided into 30 horizontal layers of the same volume, one layer O₂-max has maximum concentration of oxygen, and one layer O₂-min has minimum concentration of oxygen, with their ratio approximately equal to 2:1.

EFFECT: efficient reaction at relatively low temperature, increased cost-effectiveness and efficiency of the method.

37 cl, 7 tbl, 35 dwg, 11 ex

Description

FIELD OF THE INVENTION

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

BACKGROUND OF THE INVENTION

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), oxidize cyclohexane to adipic acid and oxidize aromatic alkyls to alcohols, acids or diacids. A particularly significant industrial method for the oxidation of the latter category (oxidation of aromatic alkyls) is the liquid-phase catalytic partial oxidation of para-xylene to terephthalic acid. Terephthalic acid is an important compound having numerous uses. The main application of terephthalic acid is its use as a raw material for the production of polyethylene terephthalate (PET). PET is a well-known plastic used in large quantities worldwide for the manufacture of products such as bottles, fibers and packaging.

In a typical liquid phase oxidation process, including partial oxidation of para-xylene to terephthalic acid, a liquid phase feed stream and a gaseous oxidizing 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., para-xylene), and the gaseous 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, which makes oxygen available for the liquid phase reaction. If the liquid phase of the multiphase reaction medium has an insufficient concentration of molecular oxygen (i.e. if certain parts of the reaction medium are oxygen depleted), impurities can form as a result of unwanted side reactions and / or target reactions can slow down. If the liquid phase of the reaction medium contains too little oxidizable compound, the reaction rate may be undesirably low. Further, if the liquid phase of the reaction medium contains an excess concentration of the oxidizable compound, additional unwanted side reactions can form impurities.

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

Stirring of the reaction medium subjected to liquid-phase oxidation is often provided by mechanical means of mixing in devices such as, for example, a continuous mixing reactor (RNND). Although RNND can provide thorough mixing of the reaction medium, they have several disadvantages. For example, RNND entail relatively high capital costs due to the fact that they require more expensive motors, impervious to liquids and gases supports and drive shafts and / or complex mixing mechanisms. In addition, the rotating and / or oscillating mechanical components of conventional RDCs require regular maintenance. The time and downtime associated with such maintenance also increases the operating costs of the RNND. However, even with regular maintenance, the mechanical agitation systems used in the RDC are prone to mechanical breakdowns and may require regular replacement at relatively short intervals.

Reactor bubblers are an attractive alternative to RNND and other mechanically agitated oxidation reactors. Reactor bubble columns provide mixing of the reaction medium without requiring expensive and unreliable mechanical equipment. Reactor bubbler columns typically include an elongated vertical reaction zone in which the reaction medium is contained. Stirring of the reaction medium in the reaction zone is ensured mainly by the natural emergence of gas bubbles rising in the liquid phase of the reaction medium. This mixing due to natural floating, available in bubbler columns of the reactor type, reduces capital and maintenance costs compared to mechanically stirred reactors. In addition, the essentially absence of moving mechanical parts in bubbler columns of a reactor type gives an oxidizing system that is less prone to mechanical breakdowns than mechanically stirred reactors.

When the liquid-phase partial oxidation of para-xylene is carried out in a conventional oxidation reactor (RND or bubble column), the product taken from the reactor is usually a suspension containing technical terephthalic acid (TTFA) and mother liquor. Technical TPA contains relatively high levels of impurities (for example, 4-carboxybenzaldehyde, para-toluic acid, fluorenones and other colored substances), which make it unsuitable as a raw material for the production of PET. Thus, the TTFA obtained in conventional oxidation reactors is usually subjected to a purification process that converts TTFA to pure terephthalic acid (CTFA) suitable for the production of PET.

One typical purification method for converting TTFA to CTFA involves the following steps: (1) replacing the mother liquor of the suspension containing TTFA with water, (2) heating the suspension of TTFA / water to dissolve the TTFA in water, (3) catalytic hydrogenation of the TTFA / water solution converting impurities to more desirable and / or easily separable compounds, (4) precipitating the resulting PTFC from the hydrogenated solution through several crystallization steps, and (5) separating the crystallized PTFC from the remaining liquids. Being effective, this conventional cleaning method can, however, be very expensive. Some of the factors leading to the high cost of conventional TTFA purification methods include, for example, the thermal energy required to accelerate the dissolution of TTFA in water, the catalyst required for hydrogenation, the hydrogen stream required for hydrogenation, loss of yield due to hydrogenation of a portion of terephthalic acid, and the need in several devices for multi-stage crystallization. Thus, it would be desirable to provide such a commodity TTFC that could be cleaned without requiring heat-dissolving in water, hydrogenation and / or multi-stage crystallization.

OBJECTS OF THE INVENTION

Thus, one aim of the present invention is to provide a more efficient and economical liquid phase reactor and oxidation process.

Another objective of the invention is to provide a more economical and efficient reactor and method for the liquid-phase catalytic partial oxidation of para-xylene to terephthalic acid.

Another objective of the invention is to provide a bubble column reactor, which contributes to the improvement of liquid phase oxidation reactions with reduced formation of impurities.

Another objective of the invention is to provide a more efficient and economical system for the production of pure terephthalic acid (CTFA) by liquid-phase oxidation of para-xylene to obtain technical terephthalic acid (TTFA), followed by purification of TTFA in CTFA.

A further object of the invention is to provide a bubble column reactor for the oxidation of para-xylene and the production of commercial TTFC, which can be purified without the need for thermal activation of the dissolution of TTFC in water, hydrogenation of the dissolved TTFC and / or multi-stage crystallization of hydrogenated CTFC.

It should be noted that the scope of the present invention, as defined in the attached claims, is not limited to methods or apparatuses suitable for implementing all of the above objectives. On the contrary, the scope of the claimed invention may cover many systems that do not satisfy all or do not satisfy any of the above objectives. Additional objectives and advantages of the present invention will become readily apparent to those skilled in the art upon consideration of the following detailed description and the attached drawings.

The essence of the invention

One embodiment of the present invention relates to a method comprising oxidizing an oxidizable compound in the liquid phase of a multiphase reaction medium contained in one or more mechanical stirring reactors, the reaction medium comprising a first separate 20% continuous volume, which corresponds to a first space-time oxygen velocity ( STR oxygen), and the second separate 20% continuous volume, which corresponds to the second STR oxygen, and the ratio of the first STR oxygen to the second STR oxygen is at least about 1.5: 1.

Another embodiment of the present invention relates to a method comprising the following steps: (a) introducing a feed stream containing an oxidizable compound into the first reaction zone of the first reactor, wherein the first reaction zone has a maximum diameter (D), at least a portion of the first reaction zone limited to one or more vertical side walls of the first reactor, with at least about 25% wt. the oxidizable compound enters the first reaction zone at a distance at least 0.05 D inward from the vertical side walls; and (b) oxidizing at least a portion of the oxidizable compound in the liquid phase of the multiphase reaction medium contained in one or more reactors, wherein at least a portion of the reaction medium is contained in the first reactor, the reaction medium comprising a first separate 20% continuous volume to which the first space-time oxygen velocity (STR oxygen), and the second separate 20% continuous volume, which corresponds to the second STR oxygen, and the ratio of the first STR oxygen to the second oxygen STR is at least about 1.5: 1.

Another embodiment of the present invention relates to a method comprising oxidizing an oxidizable compound in the liquid phase of a multiphase reaction medium contained in separate upper and lower oxidation reactors, wherein the first part of the reaction medium is contained in the upper reactor and the second part of the reaction medium is contained in lower reactor, and the first part of the reaction medium has a first oxygen averaged over the apparatus STR, and the second part of the reaction medium has a second averaged over a arata of oxygen STR, wherein the ratio of the first oxygen averaged over the STR apparatus to the second oxygen averaged over the STR apparatus is in the range from about 2: 1 to about 25: 1.

Another embodiment of the present invention relates to a method comprising the following steps: (a) introducing a feed stream containing an oxidizable compound into the reaction zone of a bubble column reactor type; (b) introducing a stream of oxidizing agent containing molecular oxygen into the reaction zone; and (c) oxidizing at least a portion of the oxidizable compound in the liquid phase of the multiphase reaction medium contained in the reaction zone, the reaction medium having a maximum width (W), a maximum height (H), and an H: W ratio of at least about 3: 1 moreover, most of the molecular oxygen enters the reaction zone within about 0.25W from the bottom of the reaction zone, and the oxidation is carried out so that if the reaction zone is theoretically divided into 30 horizontal layers of equal volume, one horizontal layer (O ₂ -max) has a maximum ksimalnuyu oxygen concentration of all the 30 horizontal slices and one horizontal layer (O ₂ -min) has the minimum oxygen concentration of all the horizontal slices located above the horizontal bed O ₂ -max, wherein the oxygen concentration measured in the gas phase of reaction medium as the time-averaged and by volume the molar concentration, based on the wet state, wherein the ratio of the oxygen concentration in the horizontal O ₂ -max layer to the oxygen concentration in the horizontal O ₂ -min layer is at least about 2: 1.

A further embodiment of the present invention relates to a method comprising the following steps: (a) introducing a para-xylene-containing feed stream into the reaction zone of a first reactor-type bubble column, wherein the reaction zone has a maximum diameter (D), at least a portion of the reaction zone is limited to one or more vertical side walls of the bubble column reactor type, with at least about 25% wt. the oxidizable compound enters the first reaction zone at a distance at least 0.05 D inward from the side walls; (b) oxidizing at least a portion of para-xylene in the liquid phase of a three-phase reaction medium contained in a bubble column reactor to thereby form crude terephthalic acid, the reaction medium comprising a first separate 20% continuous volume, which corresponds to the first space-time the oxygen rate (STR oxygen), and the second separate 20% continuous volume with the second oxygen STR, the ratio of the first STR oxygen to the second STR oxygen is at least about 1.5: 1; and (c) oxidizing at least a portion of the crude terephthalic acid in a secondary oxidation reactor to thereby obtain a purer terephthalic acid.

Brief Description of the Figures

Below, preferred embodiments of the invention are described in detail with reference to the attached drawings, moreover:

figure 1 is a side view of an oxidation reactor constructed in accordance with one embodiment of the present invention, showing, in particular, the introduction of flows of raw materials, oxidizing agent and reflux into the reactor, the presence of a multiphase reaction medium in the reactor and the withdrawal of gas and slurry from above and from the bottom of the reactor, respectively ;

figure 2 is a side view with an increase in part of the bubble column reactor type in section along line 2-2 in figure 3, showing, in particular, the position and configuration of the oxidizer distributor used to introduce the oxidant stream into the reactor;

FIG. 3 is a plan view of the oxidizer distributor of FIG. 2, showing, in particular, openings for the oxidizer at the top of the oxidizer distributor;

FIG. 4 is a bottom view of the oxidizer distributor of FIG. 2, showing, in particular, openings for the oxidizer at the bottom of the oxidizer distributor;

5 is a side view of the oxidizer distributor in section along line 5-5 of FIG. 3, showing, in particular, the orientation of the oxidizer holes at the top and bottom of the oxidizer distributor;

6 is a side view in enlargement of the lower part of the bubble column reactor type, showing, in particular, a system for introducing the flow of raw materials into the reactor in several places spaced apart in height;

Fig.7 is a side view in section according to 7-7 in Fig.6, showing, in particular, how the raw material injection system shown in Fig.6, distributes the flow of raw materials in the preferred radial loading zone (FZ) and more than one azimuth - quadrant (Q ₁ , Q ₂ , Q ₃ , Q ₄ );

Fig.8 is a top view in section, similar to Fig.7, but showing an alternative means of discharging the feed stream into the reactor using pipes with bayonet closures, each of which has several small supply openings;

Fig.9 is an isometric representation of an alternative system for introducing a feed stream into the reaction zone in several places spaced apart in height, which does not require inserting several inputs into the apparatus, showing, in particular, that the feed distribution system can at least partially rely on the oxidizer distributor;

figure 10 is a side view of the feed distribution system with one insert and a view of the oxidizer distributor shown in figure 9;

11 is a top sectional view taken along line 11-11 of FIG. 10, further illustrating a single-sided feed distribution system based on an oxidizing agent distributor;

12 is an isometric view of an alternate oxidizing agent distributor in which all of the oxidizing holes are at the bottom of the ring member;

Fig.13 shows an alternative oxidizer dispenser of Fig.12 in a top view;

FIG. 14 is a bottom view of the alternate oxidizing agent distributor of FIG. 12, showing in particular the position of the lower openings for introducing the oxidizing agent stream into the reaction zone;

Fig is a side view of the distributor of the oxidizing agent in section along the line 15-15 in Fig.13, showing, in particular, the orientation of the lower holes for the oxidizing agent;

Fig is a side view of a bubble column reactor type equipped with an internal deaerator near the lower discharge opening of the reactor;

FIG. 17 is a side elevational view of the lower portion of the reactor type bubble column of FIG. 16, in section along line 17-17 of FIG. 18, showing, in particular, the configuration of an internal deaerator located at the lower discharge opening of the reactor type bubble column;

Fig is a top view in section along the line 18-18 in Fig.16, showing, in particular, a flow stabilizer installed in the deaerator;

Fig is a side view of a bubble column reactor equipped with an external deaerator, showing the way in which part of the deaerated suspension emerging from the bottom of the deaerator can be used to flush the discharge line connected to the bottom of the reactor;

Fig. 20 is a side view of a bubble column reactor equipped with a hybrid internal / external deaerator for discharging a gas phase from a reaction medium discharged from an elevated side of the reactor;

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

FIG. 22 is a side elevational view of the bottom of the reactor type bubble column of FIG. 21, showing, in particular, the use of an alternative oxidizer distributor using inlet lines that receive an oxidizer stream through the bottom of the reactor;

Fig is a side section in magnification, similar to Fig.22, showing, in particular, an alternative means for introducing the flow of oxidant into the reactor through several holes in the lower base of the reactor and, alternatively, the use of chippers to more evenly distribute the flow of oxidant in the reactor;

Fig. 24 shows a side view of a bubble column reactor using an internal streamline to facilitate the distribution of the oxidizable compound by returning part 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, which uses an internal line to facilitate the distribution of oxidizable compounds by returning part of the reaction medium from the upper part of the reactor to the lower part of the reactor;

Fig is a side view in section of a horizontal ejection device, which can be used to improve the distribution of the oxidizable compound in the oxidation reactor, showing, in particular, an ejection device that uses incoming liquid feed to draw the reaction medium into it and releases the feed mixture at high speed and the reaction medium into the reaction zone;

Fig. 27 is a sectional side view of a vertical ejection device that can be used to improve the distribution of the oxidizable compound in the oxidation reactor, showing, in particular, an ejection device that combines the liquid feed and the inlet gas and uses the combined two-phase fluid to draw the reaction medium and at high speed releases a mixture of liquid feed, incoming gas and the reaction medium into the reaction zone;

Fig is a side view of a bubble column reactor containing a multiphase reaction medium, illustrating, in particular, that the reaction medium is theoretically divided into 30 horizontal layers of equal volume to quantify some gradients in the reaction medium;

Fig is a side view of a bubble column reactor containing a multiphase reaction medium, showing, in particular, the first and second discrete 20% continuous volumes of the reaction medium, which have substantially different oxygen concentrations and / or oxygen flow rates;

30 is a side view of two tier-mounted reaction apparatuses, with or without alternative mechanical stirring, containing a multiphase reaction medium, showing, in particular, that the apparatuses contain discrete 20% continuous volumes of the reaction medium having substantially different oxygen concentrations and / or flow rates oxygen;

Fig. 31 is a side view of three adjacent reaction apparatuses, with or without alternative mechanical stirring, containing a multiphase reaction medium, showing, in particular, that the apparatuses contain discrete 20% continuous volumes of the reaction medium having substantially different oxygen concentrations and / or flow rates oxygen;

FIG. 32A and 32B are particle enlargement images of technical terephthalic acid (TTFA) particles obtained according to one embodiment of the present invention, showing, in particular, that each TTFA particle is a low density particle with a high specific surface area consisting of several weakly bound TTFA subparticles;

FIG. 33A and 33B are enlarged images of TTFC produced by conventional methods, showing, in particular, that traditional TTFC particles have a larger size, lower density and lower specific surface than TTFC particles according to the invention (FIGS. 32A and 32B);

Fig is a simplified diagram of a technological method for producing purified terephthalic acid (CTFA) according to the prior art; and

Fig. 35 is a simplified flow diagram of a process method for producing CTFC in accordance with one embodiment of the present invention.

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 mixing reactors. Suitable mixing reactors include, for example, gas bubble mixed reactors (e.g. reactor bubblers), mechanically mixed reactors (e.g. continuous mixing reactors) and flow-through mixed reactors (e.g. jet reactors). In one embodiment of the invention, liquid phase oxidation is carried out in a single reactor type bubbler column.

As used herein, the term “reactor type bubbler column” means a reactor to facilitate chemical reactions in a multiphase reaction medium, the stirring of the reaction medium being mainly provided by the movement of gas bubbles upstream of the reaction medium. As used herein, the term “mixing” means work dispersed in a reaction medium, causing flow and / or mixing of a liquid or gas. As used herein, the terms “primarily,” “primarily,” and “primarily” mean more than 50%. The term “mechanical agitation” as used herein means agitation of the reaction medium caused by the physical movement of the rigid or flexible element (s) against or in the direction of the reaction medium. For example, mechanical agitation may be provided by rotation, vibrations and / or vibrations of internal mixers, vanes, vibrators or acoustic diaphragms located in the reaction medium. As used herein, the term “flow mixing” means mixing the reaction medium caused by high-speed injection and / or circulation of one or more fluids in 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 mixing of the reaction medium in a bubble column reactor during oxidation is provided by mechanical mixing and / or mixing of the stream, more preferably less than about 20% of the mixing is provided by mechanical mixing and / or mixing of the stream, and most preferably less than 5% mixing is provided by mechanical stirring and / or mixing of the stream. Preferably, the level of mechanical stirring and / or mixing of the stream introduced into the multiphase reaction medium during oxidation corresponds to 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 bubbler column of a reactor type 20 is shown, including a vessel 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 separation zone 30. Advantageously, a liquid phase feed stream is introduced into the reaction zone 28 through feed inlets 32a, b, c, d. Advantageously, a gaseous oxidant stream is introduced into the reaction zone 28 through an oxidizer distributor 34 located at the bottom of the reaction zone 28. The liquid-phase feed stream and the gaseous oxidizer stream together form a multiphase reaction medium 36 in the reaction zone 28. The multiphase reaction medium 36 contains liquid and gas phases. 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 in the reaction zone 28 as a result of the oxidation reaction in the liquid phase of the reaction medium 36. The bubble column reactor 20 includes an opening 38 for discharging a suspension located near the bottom of the reaction zone 28 and an opening 40 for venting gas, located near the top of separation zone 30. An effluent stream of a suspension containing liquid-phase and solid-phase components of the reaction medium 36 is discharged from the reaction zone 28 through an opening 38 for discharging a suspension, and a predominantly gaseous effluent is discharged from the separation zone 30 through an outlet 40 for discharging gas.

The liquid phase feed stream introduced into the bubble column reactor 20 through feed inlets 32a, b, c, d preferably comprises 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 with at least one attached substituted hydrocarbon group, or with at least one attached heteroatom, or with 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, each attached group containing from 1 to 5 carbon atoms. Even more preferably, the oxidizable compound is an aromatic compound containing exactly two attached groups, each attached group containing exactly one carbon atom and formed by methyl groups and / or substituted methyl groups and / or not more than one carboxylic acid group. Even more preferably, the oxidizable compound is para-xylene, meta-xylene, para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluic acid and / or acetaldehyde. Most preferably, the oxidizable compound is para-xylene.

As defined herein, a “hydrocarbon group” is at least one carbon atom that is bonded only to hydrogen atoms or 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 except carbon and hydrogen atoms. Aromatic compounds, as defined here, include an aromatic ring containing preferably at least 6 carbon atoms, even more preferably containing only 6 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 fused carbon-based 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 a combination of the above aliphatic components plus other heteroatoms (e.g., 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 (e.g., toluene, ethylbenzene, isopropylbenzene, n-propylbenzene, neopentylbenzene, para-xylene, meta-xylene, ortho-xylene, all isomers trimethylbenzene, all isomers of tetramethylbenzene, pentamethylbenzene, hexamethylbenzene, all isomers of ethyl methylbenzene, all isomers of diethylbenzene, all isomers of ethyl dimethylbenzene, all isomers of dimethylnaphthalene, all isomers of ethylmethylnaphthalene, all isomers diethylnaphthalene, all isomers of dimethylbiphenyl, all isomers of ethylmethylbiphenyl and all isomers of diethylbiphenyl, stilbene and stilbene with one or more attached hydrocarbon groups, fluorene and fluorene with one or more attached hydrocarbon groups, anthracene and anthracene with one or more attached hydrocarbon groups and diphenyl ethane and di 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, benzyl methyl ether, all isomers of bromophenols, bromobenzene, all isomers of bromatoluenes, including alpha-bromtoluene, dibromobenzene, cobalt naphthenate, and all bromine isomers ombiphenyls); 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, and / or with one or more attached substituted hydrocarbon groups (for example, benzaldehyde, all isomers bromobenzaldehyde, all isomers of brominated tolualdehydes, including all isomers of alpha bromo tolualdehydes, all isomers of hydroxybenzaldehydes, all isomers of bromohydroxybenzaldehydes, all isomers of benzold carboxaldehyde, all isomers benzoltrikarboksaldegidov, para-tolualdehyde, meta-tolualdehyde, ortho-tolualdehyde, all isomers toluoldikarboksaldegidov, all isomers toluoltrikarboksaldegidov, all isomers toluoltetrakarboksaldegidov, all isomers dimetilbenzoldikarboksaldegidov, all isomers dimetilbenzoltrikarboksaldegidov, all isomers dimetilbenzoltetrakarboksaldegidov, all isomers trimetilbenzoltrikarboksaldegidov, all isomers etiltolualdegidov, all isomers of trimethylbenzene dicarboxaldehyde, tetramethylbenzoldic arboxaldehyde, hydroxymethylbenzene, all isomers of hydroxymethyltoluene, all isomers of hydroxymethylbromtoluenes, all isomers of hydroxymethyltoluene aldehydes, all isomers of hydroxymethylbromtolualdehydes, benzylhydroperoxide, benzoylhydroperoxide, all isomers of tomethylmethides, all methyl isomers various benzene rings, naphthalene rings, biphenyls, terphenyls and other aromatic groups with one or more attached selected groups, the selected groups being hydrocarbon groups and / or attached heteroatoms and / or substituted hydrocarbon groups and / or carboxylic acid groups, and / or peracid groups (e.g. benzoic acid, para-toluic acid, meta-toluic acid, ortho-toluic acid, all isomers of ethylbenzoic acid, all isomers of propylbenzoic acid, all isomers of butyl benzoic acid, all isomers of pentylbenzoic acid, all isomers of dimethylbenzoic acid, all isomers of ethylmethylbenzoic acid, all isomers of trimethylbenzoic acid, all isomers of tetramethylbenzoic acid, pentamethylbenzoic acid benzoic acid, all isomers of diethylbenzoic acid, isomeric, all , all isomers of dimethylbenzene dicarboxylic acids, all isomers of methylbenzene tricarboxylic acids, all isomers b ombenzoic acid, all isomers of dibromobenzoic acid, all isomers of bromotoluylic acid, including alpha bromotoluylic acid, tolyl acetic acid, all isomers of hydroxybenzoic acid, all isomers of hydroxymethylbenzoic acid, all isomers of hydroxymethylbenzoic acid, all of the isomers of hydroxytoluenyl hydroxymethyloxymethoxymethoxybenzoic acid, of all isomers of hydroxybromobenzoic acid, all isomers of hydroxybromobtoluic acid, all isomers of hydroxymethyl bromobenzoic acid you all isomers carboxybenzaldehyde, all isomers dikarboksibenzaldegidov, perbenzoic acid, all isomers gidroperoksimetilbenzoynoy acids, all isomers gidroperoksimetilgidroksibenzoynoy acids, all isomers gidroperoksikarbonilbenzoynoy acids, all isomers gidroperoksikarboniltoluola, all isomers metilbifenilkarbonovyh acids, all isomers dimetilbifenilkarbonovyh acids, all isomers metilbifenildikarbonovyh acids, all isomers bifeniltrikarbonovyh acids, all isomers of stilbene with one or more its attached 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 or more attached wound 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 which is solid under normal conditions (i.e., solid at standard temperature and pressure), it is preferred that the oxidizable compound be substantially dissolved in the solvent when introduced into reaction zone 28. Preferably, 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 from about 80 to about 400 ° C, and most preferably in the range of from 125 to 155 ° C. The amount of oxidizable compound present in the liquid phase feed is preferably from about 2 to about 40% by weight, more preferably from about 4 to about 20% by weight, and most preferably from 6 to 15% by weight.

We now note that the oxidizable compound present in the liquid phase feed may contain a combination of two or more different oxidizable reagents. These two or more different chemical materials can be fed mixed in a liquid phase feed stream or can be fed separately in multiple feed streams. For example, an oxidizable compound containing para-xylene, meta-xylene, para-tolualdehyde, para-toluic acid and acetaldehyde may be fed to the reactor through one inlet or through several separate inlets.

The solvent present in the liquid phase feed stream preferably contains an acid component and an aqueous component. The solvent is present in the liquid phase feed stream in a concentration of preferably in the range of from about 60 to about 98% by weight, more preferably in the range of from about 80 to about 96% by weight, and most preferably in the range of 85 to 94% by weight. The acid component of the solvent is preferably essentially an organic low molecular weight monocarboxylic acid containing 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the solvent 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 rest being mainly water. The solvent introduced into the bubble column reactor 20 may contain small amounts of impurities, such as, for example, para-tolualdehyde, terephthaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid, para-toluene aldehyde, alpha bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatics and / or suspended particles. Preferably, 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 capable of promoting the oxidation (including partial oxidation) of an oxidizable compound. More preferably, the catalyst system comprises at least one transition metal of variable valency. Even more preferably, the transition metal of variable valency contains cobalt. Even more preferably, the catalyst system contains cobalt and bromine. Most preferably, the catalyst system comprises 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 in the range of about 300 to about 6000 ppm by weight (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 kept in the range of from about 300 to about 5000 ppm. wt., more preferably in the range 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 preferable 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 kept in the range of about 20 to about 1000 ppm. wt., more preferably in the range from about 40 to about 500 ppm. wt., most preferably in the range from 50 to 200 ppm wt.

The above concentrations of cobalt, bromine and / or manganese in the liquid phase of the reaction medium 36 are expressed as averaged over time and volume. As used herein, the term "time averaged" means an average of at least 10 measurements taken at regular intervals over a continuous period of at least 100 seconds. As used herein, the term “volume averaged” means an average of at least 10 measurements taken at equal three-dimensional distances over a given volume.

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

The liquid phase feed stream introduced into the bubble column reactor 20 may include small amounts of impurities, such as, for example, toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid, steam α-toluyl aldehyde, alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended particulates. If a bubble column reactor type 20 is used to produce terephthalic acid, meta-xylene and ortho-xylene are also considered impurities. Preferably, the total amount of impurities in the liquid-phase feed stream introduced into the bubble column reactor 20 is less than about 3% by weight.

Although FIG. 1 shows an embodiment where the oxidizable compound, solvent and catalyst system are mixed together and introduced into the bubble column reactor 20 as a single feed stream, in an alternative embodiment of the present invention, the oxidizable compound, solvent and catalyst may be introduced into the bubble column reactor 20 apart. For example, a stream of pure para-xylene can be supplied to the bubble column reactor 20 through an inlet separated from the inlet (s) of the solvent and catalyst.

Advantageously, the gaseous oxidant stream introduced into the bubble column reactor 20 through the oxidizer distributor 34 contains molecular oxygen (O ₂ ). The oxidizing stream preferably contains from about 5 to about 40% molar molecular oxygen, more preferably from about 15 to about 30% molar molecular oxygen, and most preferably from 18 to 24% molar molecular oxygen. Preferably, the remainder of the oxidizing stream contains mainly gas or gases, such as nitrogen, which is inert to oxidation. More preferably, the oxidizing stream consists mainly of molecular oxygen and nitrogen. Most preferably, the oxidizing agent stream is dry air, which contains about 21% molar molecular oxygen and about 78 to about 81% molar nitrogen. In an alternative embodiment of the present invention, the oxidizing stream may contain substantially pure oxygen.

Referring again to FIG. 1: the reactor type bubbler column 20 is preferably equipped with a reflux dispenser 42 located above the top surface 44 of the reaction medium 36. The reflux sparger 42 is operable to introduce droplets of a predominantly liquid phase reflux stream into the separation zone 30 by any droplet formation method known in the prior art. More preferably, irrigation distributor 42 produces atomized droplets directed downward toward the upper surface 44 of reaction medium 36. Preferably, these atomized droplets falling downward affect (i.e. affect and affect) at least about 50% of the maximum horizontal area of zone 30 branches. More preferably, the sprayed droplets act on at least about 75% of the maximum horizontal area of the separation zone 30. Most preferably, the atomized droplets affect at least 90% of the maximum horizontal area of the separation zone 30. These dropping droplets of liquid phlegm can help prevent foaming at or above the top surface 44 of reaction medium 36 and can also help to divert any liquid or suspension droplets entrained in the upward flowing gas that flows to the gas outlet 40. In addition, liquid phlegm can be used to reduce the amount of particulate and potentially precipitating compounds (for example, dissolved benzoic acid, para-toluic acid, 4-CBA, terephthalic acid and catalytic metal salts) leaving the gaseous effluent discharged from separation zone 30 through the gas outlet 40. In addition, the introduction of reflux droplets into the separation zone 30 can, through rectification, be used to control the composition of the gaseous effluent discharged through the gas outlet 40.

The liquid reflux stream introduced into the bubble column reactor 20 through the reflux dispenser 42 preferably has almost 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 with 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. reflux stream, more preferably at least about 80% wt. reflux stream, and most preferably from 85 to 98% wt. reflux stream, the rest being water. Since the reflux stream usually has essentially the same composition as the solvent in the liquid phase feed stream, when in this description reference is made to “all solvent” introduced into the reactor, such “all solvent” will include both the reflux stream and the fraction solvent in the feed stream.

In liquid phase oxidation in a bubble column reactor 20, it is preferred that the feed streams, oxidizing agent and reflux are introduced into the reaction zone 28 substantially continuously and the effluent gas and slurries are substantially continuously withdrawn from the reaction zone 28. The term “substantially continuous” as used herein means a period of at least 10 hours, with a break of not more than 10 minutes. During the oxidation, it is preferable that the oxidizable compound (e.g. para-xylene) is introduced into the reaction zone 28 substantially continuously at a rate of at least about 8000 kg / h, more preferably at a speed in the range of from about 13000 to about 80,000 kg / h, even more preferably in the range of from about 18,000 to about 50,000 kg / h, and most preferably in the range of from 22,000 to 30,000 kg / h. Although it is generally preferred that the feed rates of the incoming feed streams, oxidizing agent and reflux are substantially constant, it should be noted here that one embodiment of the present invention involves pulsed feed streams of feed, oxidizing agent and / or reflux in order to improve mixing and mass transfer. If the incoming feedstock, oxidizing agent and / or reflux stream are pulsed, it is preferred that their feed rate ranges from 0 to about 500% of the stationary feed rates indicated here, more preferably from about 30 to about 200% of these here stationary feed intensities, and most preferably in the range from 80 to 120% of the stationary feed intensities indicated here.

The average spatiotemporal reaction rate (STR) in a bubble column reactor of oxidation type 20 is defined as the mass of the oxidizable compound supplied per unit volume of the reaction medium 36 per unit time (e.g., kilograms of para-xylene fed per cubic meter per hour). Typically, before calculating STR, the amount of oxidizable compound not converted to product should be subtracted from the amount of oxidizable compound in the feed stream. However, for many oxidizable compounds (e.g., para-xylene) preferred here, conversions and yields are typically high, and it is preferable to define the term as described above. For reasons of capital expenditure and production resources, among others, it is usually preferable that the reaction be carried out with a high STR. However, carrying out a reaction with an excessively high STR can affect the quality or yield of partial oxidation. A reactor type bubbler column 20 is particularly advantageous when the STR of an oxidizable compound (e.g., para-xylene) is in the range of from 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 to about 250 kg / m ³ · h, even more preferably in the range 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.

STR oxygen in a bubble column of an oxidation reactor type 20 is defined as the weight 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, among other things, capital costs and solvent consumption for oxidation, it is usually preferred that the reaction is carried out with a high STR of oxygen. However, carrying out a reaction with excessively high STR oxygen can reduce the quality or yield of partial oxidation. Without being attached to the theory, it is believed that this is possibly related to the rate of transfer of molecular oxygen from the gas phase to the liquid at the interface and from there to the volume of the liquid. Too high STR oxygen may possibly lead to too low dissolved oxygen in the volume of the liquid phase of the reaction medium.

The global average oxygen STR is defined here as the weight 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 bubble column 20 is particularly advantageous when the global average oxygen STR is in the range of from about 25 kg / m ³ · h to about 400 kg / m ³ · h, more preferably in the range from about 30 kg / m ³ · h to about 250 kg / m ³ · h, even more preferably in the range 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

When oxidizing in a bubble column reactor 20, it is preferable that the ratio of the mass flow rate of the entire solvent (from the feed and reflux streams) to the mass flow rate of the oxidizable compound entering reaction zone 28 is kept in the range from about 2: 1 to about 50: 1 more preferably in the range of from about 5: 1 to about 40: 1, and most preferably in the range of from 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 kept in the range from about 0.5: 1 to the complete absence of the reflux stream, more preferably in the range from about 0 5: 1 to about 4: 1, even more preferably in the range of from about 1: 1 to about 2: 1, and most preferably from 1.25: 1 to 1.5: 1.

In liquid phase oxidation in a bubble column reactor 20, it is preferred that the oxidizing agent stream is introduced into the bubble column reactor 20 in such a quantity as to provide a slight excess of the amount of molecular oxygen over that required by stoichiometry. The amount of excess molecular oxygen needed for best results with a particular oxidizable compound affects the entire economy of liquid phase oxidation. In liquid phase oxidation in a bubble column reactor 20, it is preferable that the ratio of the mass flow rate of the oxidizing agent to the mass flow rate of the oxidizable organic compound (e.g., para-xylene) entering the reactor 20 is kept in the range from about 0.5: 1 to about 20 : 1, more preferably in the range of from about 1: 1 to about 10: 1, and most preferably from 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 a gas phase. As noted above, the oxidation of an oxidizable compound (eg, para-xylene) occurs predominantly 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 taking 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 openings 32a, b, c, d to feed the feed. Thus, the gas phase of the reaction medium 36 in the reactor 20 is formed mainly by an evaporated solvent and an undissolved, unreacted part of the oxidizer stream. In some prior art oxidation reactors, heat exchange tubes / fins are used to heat or cool the reaction medium. However, such heat exchange structures may not be desirable in the reactor and method of the invention described herein. Thus, it is preferable that the bubble column reactor 20 essentially had no surfaces that come into contact with the reaction medium 36 and exhibit a time-averaged heat flux of more than 30,000 W / m ² .

The concentration of dissolved oxygen in the liquid phase of the reaction medium 36 is in dynamic equilibrium between the rate of mass transfer from the gas phase and the rate of consumption in the reaction in the liquid phase (i.e., it is not set simply by the partial pressure of molecular oxygen in the feed gas phase, although this is one of factors in the rate of supply of dissolved oxygen and really affects the maximum upper concentration of dissolved oxygen). The amount of dissolved oxygen varies locally, and it is higher near the interface of the bubbles. In general, the amount of dissolved oxygen depends on the ratio of supply and consumption in different areas of the reaction medium 36. Over time, the amount of dissolved oxygen depends on the uniformity of gas and liquid mixing with respect to flow rates in chemical reactions. When designed to properly match the supply and demand for dissolved oxygen in the liquid phase of the reaction medium 36, it is preferable that the time and volume averaged oxygen concentration in the liquid phase of the reaction medium 36 is maintained at a level above about 1 molar ppm, more preferably in the range of from about 4 to about 1000 molar ppm, even more preferably in the range from about 8 to about 500 molar ppm, and most preferably in the range from 12 to 120 molar ppm.

The liquid phase oxidation reaction proceeding in a bubble column reactor type 20 is preferably a precipitation reaction in which solids are formed. More preferably, the liquid phase oxidation carried out in a bubble column reactor type 20 causes at least about 10% wt. an oxidizable compound (eg, para-xylene) introduced into the reaction zone 28 to form a solid compound (eg, particles of industrial terephthalic acid) in the reaction medium 36. Even more preferably, the liquid-phase oxidation causes at least about 50% wt. the oxidizable compound to form a solid compound in the reaction medium 36. Most preferably, liquid phase oxidation causes at least 90% wt. the oxidizable compound to form a solid compound in the reaction medium 36. It is preferable that the total amount of solids in the reaction medium 36, calculated as averaged over time and volume, be higher than about 3% wt. More preferably, the total amount of solids in the reaction medium 36 is maintained at a level of from about 5 to about 40% by weight, even more preferably from about 10 to about 35% by weight, and most preferably from 15 to 30% by weight. It is preferable that a substantial part of the oxidation product (eg, terephthalic acid) obtained in the bubble column reactor type 20, is in the reaction medium 36 as a solid phase, in contrast to the remaining part 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 entire oxidation product (solid and liquid phase) in the reaction medium 36, more preferably at least about 75% wt. from the entire oxidation product in the reaction medium 36, and most preferably at least 95% wt. from the entire oxidation product in the reaction medium 36. The above numerical ranges of the amount of solids in the reaction medium 36 are applicable for the essentially stationary operation of the bubble column 20 for a substantially continuous period of time, but not for starting, stopping or suboptimal operation of the bubble column reactor 20. The amount of solids in the reaction medium 36 is determined gravimetrically. In this gravimetric method, a representative fraction of the suspension is withdrawn from the reaction medium and weighed. Under conditions in which the complete solid phase-liquid distribution actually exists in the reaction medium, the free liquid is removed from the solid fraction by precipitation or filtration, virtually without loss of precipitated solid phase and with less than about 10% of the initial liquid mass remaining with the solid fraction. The liquid remaining with the solid phase is evaporated to dryness, practically without sublimation of solids. The remaining solids are weighed. The ratio of the weight of the proportion of solids to the weight of the initial proportion of the suspension is the proportion of solids, usually expressed as a percentage.

The deposition reaction proceeding in a bubble column reactor type 20 may cause fouling (ie, solid phase formation) of the surface of certain rigid structures that come 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 include internal heat transfer, mixing, or reflective structures in reaction zone 28, since such structures would be prone to fouling. If internal structures are present in the reaction zone 28, it is desirable to avoid internal structures having external surfaces that contain a significant level of upwardly flat surfaces, since such upwardly flat surfaces would be very prone to fouling. Thus, if there are any internal structures in reaction zone 28, it is preferred that less than about 20% of all upwardly facing outer surfaces of such internal structures is formed by substantially flat surfaces that deviate from the horizontal by less than about 15 °.

Turning again to FIG. 1: the physical configuration of the bubble column reactor 20 allows optimized oxidation of the oxidizable compound (eg, para-xylene) with minimal formation of impurities. Preferably, the elongated reaction section 24 of the apparatus body 22 includes a substantially cylindrical main body 46 and a bottom 48. The upper level of the reaction zone 28 is defined by a horizontal plane 50 intersecting the top of the cylindrical body 46. The lower level 52 of the reaction zone 28 is limited to the lowest inner surface of the bottom 48 Typically, the lower level 52 of reaction zone 28 is in close proximity to the slurry outlet 38. Thus, the elongated reaction zone 28, limited by the internal volume of the bubble column reactor 20, has a maximum length "L", measured from the upper 50 to the lower level 52 of the reaction zone 28 along the longitudinal axis of the cylindrical body 46. The length "L" of the reaction zone 28 is preferably is from about 10 to about 100 m, more preferably from about 20 to about 75 m, and most preferably from 25 to 50 m. The reaction zone 28 has a maximum diameter (width) of "D", which is typically equal to the maximum internal diameter a meter of the cylindrical main body 46. The maximum diameter “D” of the reaction zone 28 is preferably from about 1 to about 12 m, more preferably from about 2 to about 10 m, even more preferably from about 3.1 to about 9 m, and most preferably from 4 to 8 m. In a preferred embodiment of the present invention, reaction zone 28 has a length to diameter ratio of "L: D" in the range of from about 6: 1 to about 30: 1. Even more preferably, reaction zone 28 has an L: D ratio in the range of from about 8: 1 to about 20: 1. Most preferably, reaction zone 28 has an L: D ratio in the range of 9: 1 to 15: 1.

As discussed above, the reaction zone 28 of the bubble column reactor 20 has a multiphase reaction medium 36. The lower level of the reaction medium 36 coincides with the lower level 52 of the reaction zone 28, and the upper level is at the upper surface 44. The upper surface 44 of the reaction medium 36 is defined by a horizontal plane , which cuts the reaction zone 28 at that vertical position, where the contents of the reaction zone 28 goes from a state with a continuous gas phase to a state with a continuous liquid phase. The upper surface 44 is preferably located at a height position where the local, time-averaged volumetric gas content in the thin horizontal layer 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 levels. The maximum width “W” of the reaction medium 36 is usually equal to the maximum diameter “D” of the cylindrical main body 46. For liquid phase oxidation in a bubble column reactor type 20, it is preferred that H is from about 60 to about 120% L, more preferably from about 80 to about 110% L, and most preferably from 85 to 100% L. In a preferred embodiment of the present invention, reaction medium 36 has a height to width "H: W" ratio of greater than about 3: 1. More preferably, reaction medium 36 has an H: W ratio in the range of from about 7: 1 to about 25: 1. Even more preferably, reaction medium 36 has an H: W ratio in the range of from 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, L = H and D = W are such that the various sizes or ratios provided herein for L and D apply also to H and W, and vice versa.

The relatively high L: D and H: W ratios indicated in accordance with one embodiment of the invention may contribute to some important advantages of the system of the invention. As will be discussed in more detail below, it has been found that increased L: D and H: W ratios, as well as some of the other properties discussed below, may contribute to favorable gradients in elevation of molecular oxygen and / or oxidizable concentrations (e.g., para-xylene) in reaction medium 36. In contrast to the conventional generally accepted approach, which would support a well-mixed reaction medium with relatively uniform concentrations throughout, it was found that the stepwise change in height of entratsy oxygen and / or the oxidizable compound facilitates a more effective and economical oxidation reaction. Minimizing the concentrations of oxygen and an oxidizable compound near the top of the reaction medium 36 can help to avoid losses 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 selectivity of the oxidation is reduced. Thus, it is preferable that the concentrations of molecular oxygen and / or the oxidizable compound are significantly higher near the bottom of the reaction medium 36 than near the top of the reaction medium 36.

In addition, high L: D and H: W ratios cause the pressure at the bottom of the reaction medium 36 to be significantly higher than the pressure at the top of the reaction medium 36. This pressure difference in height is a consequence of the difference in height and density of the reaction medium 36. One advantage of this vertical differential pressure is that the increased pressure at the bottom of the apparatus leads to greater oxygen solubility and mass transfer than would otherwise be achieved at comparable temperatures and pressure head in low reactors. Thus, the oxidation reaction can be carried out at lower temperatures than would be required in lower apparatuses. When a bubble column reactor type 20 is used to partially oxidize para-xylene to technical terephthalic acid (TTFC), the ability to operate at lower reaction temperatures with the same or better oxygen mass transfer rates has several advantages. For example, the low oxidation temperature of para-xylene reduces the amount of solvent burned during the reaction. As discussed in more detail below, the low oxidation temperature also contributes to the formation of small, high specific surface, weakly bound, readily soluble TTFC particles, which can be subjected to more economical cleaning methods than large, low specific surface, dense TTFC particles obtained in conventional high-temperature oxidation methods.

When oxidized in the reactor 20, it is preferable that the temperature and time average temperature of the reaction medium 36 be kept in the range of from about 125 to about 200 ° C, more preferably in the range of from about 140 to about 180 ° C, and most preferably in the range of 150 up to 170 ° C. The pressure above the reaction medium 36 is preferably kept in the range of from about 1 to about 20 bar gauge (bar), more preferably in the range of from about 2 to about 12 bar bar, and most preferably in the range of 4 to 8 bar . Preferably, the pressure difference between the top of the reaction medium 36 and the bottom of the reaction medium 36 is from about 0.4 to about 5 bar, more preferably if the pressure difference is from about 0.7 to about 3 bar, and most preferably the pressure difference is from 1 to 2 bar. Although it is generally preferred that the pressure above the reaction medium 36 is maintained at a relatively constant value, one embodiment of the present invention allows a pulsed pressure regime to improve mixing and / or mass transfer in the reaction medium 36. When the pressure fluctuates, it is preferred that the pressure fluctuates in the range of about 60 to about 140% of the stationary pressure indicated here, more preferably from about 85 to about 115% of the stationary pressure indicated here, and most e preferably from 95 to 105% of the stationary pressure indicated herein.

A further advantage of the high L: D ratio of reaction zone 28 is that it can contribute to an increase in the average reduced velocity of the reaction medium 36. The terms “reduced velocity” and “reduced velocity of gas” as used herein with respect to reaction medium 36 mean volume velocity the gas phase of the reaction medium 36 at a certain height of the reactor divided by the horizontal sectional area of the reactor at this height. The increased reduced speed provided by the high L: D ratio of the reaction zone 28 can facilitate local mixing and increase the volumetric gas content in the reaction medium 36. The time-averaged reduced velocities of the reaction medium 36 are one quarter of the height, half and / or three quarters of the height of the reaction medium 36 is preferably greater than about 0.3 m / s, more preferably in the range of from about 0.8 to about 5 m / s, even more preferably in the range of from about 0.9 to about 4 m / s, and most preferably in the range of 1 to 3 m / s.

Referring again to FIG. 1: the separation section 26 in the bubble column reactor 20 is simply an expanded part of the apparatus body 22 located directly above the reaction section 24. The separation section 26 reduces the velocity of the upward flowing gas phase in the bubble column of the reactor type 20 as the gas phase rises above the upper surface 44 of the reaction medium 36 and approaches the gas outlet 40. This decrease in the velocity of the upwardly moving gas phase helps to facilitate the removal of liquids and / or solids entrained in the upwardly flowing gas phase, and thereby reduces the undesirable loss of certain components present in the liquid phase of the reaction medium 36.

The separation section 26 preferably includes a mating wall 54, typically in the form of a truncated cone, a generally cylindrical wide side wall 56 and a cover 58. A narrow lower level of the mating wall 54 is connected to the top of the cylindrical body 46 of the reaction section 24. A wide upper level of the mating wall 54 is connected to the bottom of the wide side wall 56. It is preferable that the interface wall 54 extends up and out from its narrow lower edge at an angle to the vertical in the range from about 10 to about 70 °, more preferably from about 15 to about 50 ° to the vertical, and most preferably from 15 to 45 ° to 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 smaller diameter than the total maximum diameter of the reaction section 24, then X may actually be less than D. In a preferred embodiment of the present invention, the ratio “X: D” of the diameter of the wide side wall 56 to the maximum diameter of the reaction section 24 is from about 0.8: 1 to about 4: 1, most preferably from 1.1: 1 to 2: 1. The cover 58 is connected to the top of the wide side wall 56. The cover 58 is preferably a generally elliptical top element defining a central opening that allows gas to exit the separation zone 30 through the gas outlet 40. Alternatively, cover 58 may be of any shape, including conical. The separation zone 30 has a maximum height "Y", measured from the top 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 from about 2: 1 to about 24: 1, more preferably from about 3: 1 to about 20: 1, and most preferably from 4: 1 up to 16: 1.

Now referring again to FIG. 1-5, consider in more detail the location and configuration of the oxidizer distributor 34. In FIG. 2 and 3, it is shown that the oxidizer distributor 34 may include an annular element 60, a cross member 62, and a pair of inlet channels 64a, b into the oxidizer. Conveniently, these oxidant inlets 64a, b can enter the apparatus higher in height of the annular element 60 and then turn down, as shown in FIG. 2 and 3. Alternatively, the oxidizer inlet channel 64a, b may enter the apparatus below the annular element 60 or at approximately the same horizontal plane as the annular element 60. Each oxidizer inlet channel 64a, b includes a first end connected to a corresponding oxidant inlet 66a, b formed in the apparatus body 22 and a second end hydraulically connected to the annular element 60. The annular element 60 is preferably formed of nozzles, more preferably of several sections of straight nozzles, and Lee preferably from several 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 ring member 60 is formed of 8 pipe sections, it usually has an octagon configuration. The crossbeam 62 is preferably formed from a substantially straight pipe section that is hydraulically connected to and extends diagonally between opposing pipe sections of the annular member 60. The pipe section used for the crossbar 62 preferably has substantially the same diameter as the pipe sections, used to obtain the ring element 60. It is preferable that the pipe segments forming the inlet channels 64a, b for the oxidizing agent, the ring element 60 and the transverse crossbar 62 have a nominal the meter is 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 illustrated in FIG. 3, and the ring element 60 and transverse bar 62 have a plurality of upper openings 68 for discharging the oxidant stream upward into the reaction zone 28. As best illustrated in FIG. 4, the annular element 60 and / or the transverse bar 62 may have one or more lower openings 70 to release the oxidant stream down into the reaction zone 28. The bottom openings 70 for the oxidizing agent can also be used to supply liquids and / or solids that could enter the annular element 60 and / or the crossbar 62. In order to prevent the formation of a solid phase inside the oxidizer distributor 34, the distributor 34 can be continuously or periodically bypass fluid flow to flush out any accumulated solids.

According to FIG. 1-4, during oxidation in a bubbler column of a reactor type 20, the oxidizer flows through the inlet 66a, b for the oxidizer into the inlet channels 64a, b for the oxidizer, respectively. Then, the oxidizer flows are transported through the oxidizer inlet channels 64a, b to the ring element 60. After the oxidizer stream enters the ring element 60, it is distributed along the internal volumes of the ring element 60 and the cross bar 62. Then, the oxidizer stream is displaced from the oxidizer distributor 34 into the reaction zone 28 through the upper and lower holes 68, 70 for the oxidizing agent in the annular element 60 and the transverse crossbar 62.

The outlets of the upper oxidant openings 68 are spaced apart from each other and are substantially at the same height in the reaction zone 28. Thus, the outlets of the upper oxidizer openings 68 are usually located along a substantially horizontal plane bounded by the top of the oxidizer distributor 34. The outlets of the bottom oxidizer holes 70 are spaced apart from each other and are substantially at the same height in the reaction zone 28. Thus, the outlets of the bottom oxidizer holes 70 are typically located along a substantially horizontal plane defined by the bottom of the oxidizer distributor 34.

In one embodiment of the present invention, at least about 20 upper oxidant openings 68 are made in the oxidizer distributor 34. More preferably, about 40 to about 800 upper oxidant openings are made in the oxidizer distributor 34. Most preferably, 60 to 400 top oxidizer openings 68 are made in the oxidizer distributor 34. In the oxidizer distributor 34, at least about 1 lower oxidant hole 70 is preferably made. More preferably, from the oxidizing agent distributor 34, about 2 to about 40 lower oxidizing holes 70 are made. Most preferably, in the oxidizer distributor 34, 8 to 20 lower oxidant holes 70 are made. The ratio of the number of upper oxidant holes 68 to the number of lower oxidant holes 70 in the oxidizer distributor 34 is preferably in the range of from about 2: 1 to about 100: 1, more preferably from about 5: 1 to about 25: 1, and most preferably from 8: 1 to 15: 1. The diameters of essentially all of the upper and lower oxidizer openings 68, 70 are preferably substantially the same, so that the ratio of the volumetric flow rate of the oxidant flow from the upper and lower openings 68, 70 is essentially the same as the above ratios between the number of upper and lower openings 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. As for the upper oxidizing holes 68, it is preferred that at least a portion of the upper oxidizing holes 68 release an oxidizing stream at an angle “A” that deviates from the vertical. Preferably, the percentage of the upper oxidizing holes 68 that are angled “A” from the vertical is in the range of from about 30 to about 90%, more preferably from about 50 to about 80%, even more preferably from 60 to 75% , and most preferably was about 67%.

The angle "A" preferably lies in the range from about 5 to about 60 °, more preferably in the range from about 10 to about 45 °, and most preferably from 15 to 30 °. As for the lower oxidant holes 70, it is preferred that substantially all of the lower oxidant holes 70 are located near the lowermost part of the annular element 60 and / or the cross-bar 62. Thus, any liquids and / or solids that can spontaneously enter the oxidizer distributor 34, can be easily removed from the oxidizer distributor 34 through the bottom oxidizer openings 70. Preferably, the bottom oxidant openings 70 discharge the oxidant stream downward at a substantially vertical angle. For the purposes of the present description, the top hole for the oxidizing agent can be any hole that releases the oxidizer stream as a whole in an upward direction (i.e., at an angle above the horizontal), and the bottom hole for the oxidizing agent can be any hole that releases the oxidizer stream as a whole in downward direction (i.e. 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 distributor (or other mechanism for introducing the oxidizer stream into the reaction zone) has a very low gas volumetric value. As is known in the art, a "volumetric gas content" is simply the volume fraction of a multiphase medium in a gaseous state. Zones with a low volumetric gas content in the medium can also be called "non-aerated" zones. In many conventional reactor-type suspension bubblers, a substantial portion of the total volume of the reaction medium is below the oxidizer distributor (or other mechanism for introducing the oxidizer stream into the reaction zone). Thus, a substantial portion of the reaction medium below conventional bubbler reactors is unaerated.

It was found that by minimizing the fraction of unaerated zones in the reaction medium that undergoes oxidation in a bubble column reactor type, the formation of certain types of undesirable impurities can be minimized. Unaerated zones of the reaction medium contain relatively few oxidizing bubbles. This low 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 promote undesirable side reactions rather than the desired oxidation reaction. For example, when para-xylene is partially oxidized to form terephthalic acid, insufficient oxygen availability in the liquid phase of the reaction medium can cause the formation of undesirably high amounts of benzoic acid and associated aromatic rings, including, in particular, very undesirable colored molecules known as fluorenones and anthraquinones.

According to one embodiment of the present invention, the liquid phase oxidation is carried out in a bubble column reactor formed and operating in such a way that the volume fraction of the reaction medium with low gas volumes is reduced to a minimum. This minimization of unaerated zones can be quantified by theoretically dividing the entire volume of the reaction medium into 2000 separate horizontal layers of the same volume. With the exception of the uppermost and lowermost horizontal layers, each horizontal layer is a separate volume, bounded laterally by the side wall of the reactor and bounded by imaginary horizontal planes above and below. The highest horizontal layer is bounded below by an imaginary horizontal plane, and above by the upper surface of the reaction medium. The lowest horizontal layer is bounded above by an imaginary horizontal plane, and below - by the lower level of the apparatus. After the reaction medium was theoretically divided into 2,000 separate horizontal layers of equal volume, it is possible to determine the volumetric gas content averaged over time and volume in each horizontal layer. When this method of quantifying the level of unaerated zones is applied, it is preferable that the number of horizontal layers having a time and volume average volumetric gas content of less than 0.1 is less than 30, more preferably less than 15, even more preferably less than 6, even more preferably less 4 and most preferably less than 2. Preferably, the number of horizontal layers with a volumetric gas content of less than 0.2 is less than 80, more preferably less than 40, even more preferably less than 20, even more than edpochtitelno less than 12, and most preferably less than 5. Preferably, the number of horizontal layers with a volumetric 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 lower placement of the oxidizing agent distributor 34 in the reaction zone 28 provides several advantages, including reducing the level of unaerated zones in the reaction medium 36. For a given height “H” of the reaction medium 36, the length “L” of the reaction zone 28 and a maximum diameter "D" of reaction zone 28, it is preferred that most (i.e.> 50% by weight) of the oxidizing agent stream is introduced into reaction zone 28 within about 0.025H, 0.022L and / or 0.25D from the lower level 52 of reaction zone 28. More preferably, most of the oxidizing agent stream is introduced into the reaction zone 28 within about 0.02H, 0.018L and / or 0.2D from the lower level 52 of the reaction zone 28. Most preferably, most of the oxidant stream is introduced into the reaction zone 28 within 0.015H, 0.013L and / or 0.15D from the lower level 52 of reaction zone 28.

In the embodiment shown in FIG. 2, the height distance “Y ₁ ” between the lower level 52 of the reaction zone 28 and the outlet of the upper opening 68 of the oxidizing agent distributor 34 is less than about 0.25H, 0.022L and / or 0.25D, so that essentially the entire oxidant stream enters reaction zone 28 within about 0.25H, 0.022L and / or 0.25D from the lower level 52 of reaction zone 28. More preferably, Y _{1 is} less than about 0.02H, 0.018L and / or 0 , 2D. Most preferably, Y _{1 is} less than 0.015H, 0.013L and / or 0.15D, but greater than 0.005H, 0.004L and / or 0.06D. Figure 2 shows the tangent 72 at the place where the lower level of the cylindrical part 46 of the apparatus body 22 is connected to the upper level of the elliptical bottom 48 of the apparatus body 22. Alternatively, bottom 48 can be of any shape, including conical, but still tangent is defined as the lowest level of the cylindrical body 46. A distance adjustment "Y _2" between the tangent 72 and the top of oxidant distributor 34 is preferably at least about 0,0012H, 0,001 L and / or 0.01D; 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 height distance “Y ₃ ” between the lower level 52 of reaction zone 28 and the outlet of the lower opening 70 of the oxidizing agent distributor 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 greater than 0.003H, 0.002L and / or 0.025D.

In a preferred embodiment of the present invention, the openings through which the oxidizing agent stream and the feed stream are discharged into the reaction zone are formed so that the amount (by weight) of the oxidizing agent or feed stream discharged from the hole is directly proportional to the open surface area of the hole. Thus, for example, if 50% of the total open surface defined by all openings for the oxidizing agent is concentrated at a level within 0.15D from the bottom of the reaction zone, then 50% wt. the oxidizing agent stream enters the reaction zone within 0.15D from the bottom of the reaction zone, and vice versa.

In addition to the benefits provided by minimizing unaerated zones (ie, zones with a low volumetric gas content) in the reaction medium 36, it has been found that oxidation can be enhanced by maximizing the volumetric gas content in the entire reaction medium 36. The reaction medium 36 preferably has an averaged in time and volume, the volumetric gas content is at least about 0.4, more preferably from about 0.6 to about 0.9, and most preferably from 0.65 to 0.85. Some of the physical and performance characteristics of the bubble column reactor type 20 contribute to the high volumetric gas content discussed above. For example, for a given reactor size and oxidizer flow stream, a high L: D ratio of reaction zone 28 gives a smaller diameter, which increases the reduced speed in reaction medium 36, which in turn increases the volumetric gas content. In addition, it is known that the actual diameter of the bubble column and the L: D ratio affect the average volumetric gas content even for a given constant reduced velocity. In addition, minimizing unaerated zones, especially at the bottom of reaction zone 28, increases the volumetric gas content. Further, the pressure and mechanical configuration of the reactor type bubble column may affect the stability at the high speed and volumetric gas contents described herein.

In addition, the authors of the present invention have found the importance of working under optimized pressure to obtain an increased volumetric gas content and enhanced mass transfer. It may seem that working at a lower pressure, which, according to Henry's law, reduces the solubility of molecular oxygen, will lead to a decrease in the rate of mass transfer of molecular oxygen from gas to liquid. In a machine with mechanical stirring, this is indeed the case, since the levels of aeration and mass transfer rates are determined by the design of the mixer and pressure. However, for a reactor bubbler column, in accordance with a preferred embodiment of the present invention, it was found how to use a lower pressure to force a given mass of gaseous oxidant stream to occupy a larger volume, increasing the reduced speed in the reaction medium 36 and, in turn, increasing the volume content gas and molecular oxygen transfer rate.

The balance between coalescence and bubble division is an extremely complex phenomenon, which leads, on the one hand, to the formation of foam, which reduces the rate of internal circulation of the liquid phase and which may require extremely large separation zones, and, on the other hand, to the formation of a small number very large bubbles, which will give a low volumetric gas content and a reduced mass transfer rate from the oxidizing stream to the liquid phase. As for the liquid phase, it is known that its composition, density, viscosity and surface tension, among other factors, interact in a very complex way with very complex results, even in the absence of a solid phase. For example, researchers in laboratories, when reporting and evaluating observations, even for simple water-to-air bubble columns, find it useful to clarify whether “water” is tap water, distilled water, or deionized water. For complex mixtures in the liquid phase and with the addition of a solid phase, the degree of complexity increases even more. The surface irregularities of individual particles of the solid phase, the average size of the solids, the particle size distribution, the amount of solids with respect to the liquid phase and the ability of the liquid to wet the surface of the solid particles - all this, along with other things, is important in their interaction with the liquid phase and an oxidizer flow to establish what the behavior of the bubbles will be and what the resulting pattern of natural convection will be.

Thus, the ability of a bubbler column reactor to function efficiently with the high reduced velocities and high volumetric gas contents described herein depends, for example, on the appropriate choice: (1) the composition of the liquid phase of the reaction medium; (2) the amount and type of precipitated solids, both of which may depend on the reaction conditions; (3) the amount of oxidant stream fed to the reactor; (4) pressure, which affects the volumetric flow of the oxidizing stream, the stability of the bubbles and, through the energy balance, the reaction temperature; (5) the reaction temperature itself, which affects the properties of the liquid and gas, the properties of the precipitated solids, and the specific flow rate of the oxidizing agent; and (6) the geometry and mechanical details of the reaction apparatus, including the L: D ratio.

Turning again to FIG. 1: it was found that an improved distribution of an oxidizable compound (e.g., para-xylene) in the reaction medium 36 can be achieved by introducing a liquid-phase feed stream into the reaction zone 28 in several locations spaced apart in height. Preferably, a liquid phase feed stream is introduced into the reaction zone 28 through at least 3 feed openings, more preferably at least 4 feed openings. As used herein, the term “feed openings” means openings through which a liquid phase feed stream is discharged into the reaction zone 28 for mixing with the reaction medium 36. It is preferred that at least 2 of the feed openings are at least about 0 apart from each other. , 5D, more preferably at least about 1.5D, and most preferably at least 3D. However, it is preferable that the highest feed opening is separated in height from the lowest oxidizer hole by no more than about 0.75H, 0.65L and / or 8D, more preferably no more than about 0.5H, 0 , 4L and / or 5D, and most preferably not more than 0.4H, 0.35L and / or 4D.

Although it is desirable to introduce a liquid phase feed stream in several places along the height, 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 feed stream 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 in the lower half of the reaction medium 36 and / or reaction zone 28. Most preferably, at least 90% wt. a liquid phase feed stream is introduced in the lower half of the reaction medium 36 and / or reaction zone 28. In addition, it is preferable that at least about 30% wt. the liquid phase feed stream was introduced into the reaction zone 28 within about 1.5 D from the lowest point in height where the oxidant stream is introduced into the reaction zone 28. This is the lowest place in height where the oxidizer stream is introduced into the reaction zone 28, usually located at the bottom of the distributor an oxidizing agent, however, many alternative configurations are contemplated for introducing an oxidizing agent stream into the reaction zone 28 by the preferred embodiment of the present invention. Preferably, at least about 50% wt. a liquid-phase feed was introduced within about 2.5 D from the lowest elevation where the oxidant stream is introduced into reaction zone 28. It is preferred that at least about 75% wt. a liquid phase feed stream was introduced within about 5D from the lowest elevation where the oxidant stream is introduced into reaction zone 28.

Each feed opening has an open surface through which the feed is discharged. Preferably, at least about 30% of the entire open surface of all the power inlets is within about 1.5 D from the lowest elevation where the oxidant stream is introduced into the reaction zone 28. It is preferred that at least about 50% of the entire open surface all power inputs were located within about 2.5 D from the lowest elevation where the oxidant stream is introduced into reaction zone 28. It is preferable that at least about 75% of the entire open surface of all power inputs Laga within about 5D of the lowest space height, where the oxidant stream is introduced into reaction zone 28.

Referring again to FIG. 1: in one embodiment of the present invention, the power inputs 32a, b, c, d are located just adjacent to the height-aligned holes along one side of the apparatus body 22. These feed openings preferably have substantially similar diameters of less than about 7 cm, more preferably in the range of from about 0.25 to about 5 cm, and most preferably in the range of from 0.4 to 2 cm. The bubble column reactor 20 is preferably equipped with a control system the flow rate of the liquid-phase feed stream from each feed opening. Such a flow control system preferably includes an individual valve — a flow regulator 74a, b, c, d for each respective loading input 32a, b, c, d. In addition, it is preferable that the bubble column reactor 20 is equipped with a flow control system that allows at least a portion of the liquid phase feed stream to be introduced into the reaction zone 28 with an increased inlet 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 from 8 to 20 m / s. As used herein, the term “reduced input velocity” means a time-averaged volumetric rate of flow of a feed stream from a feed opening divided by the area of the feed opening. Preferably, at least about 50% wt. the feed stream is introduced into the reaction zone 28 with an increased reduced input speed. Most preferably, substantially the entire feed stream is introduced into the reaction zone 28 at an increased reduced inlet velocity.

Turning now to FIG. 6-7, an alternative system for introducing a liquid phase feed stream into the reaction zone 28 is shown. In this embodiment, the feed stream is introduced into the reaction zone 28 at four different heights. Each height is equipped with a corresponding feed distribution system 76a, b, c, d. Each feed distribution system 76 includes a main supply line 78 and a distributor 80. Each distributor 80 is provided with at least two outlets 82, 84 connected to respective plug-in pipes 86, 88 that enter the reaction zone 28 of the apparatus body 22. Each insert pipe 86, 88 has a corresponding feed opening 87, 89 for discharging a feed stream into the reaction zone 28. The feed openings 87, 89 preferably have substantially similar diameters of less than about 7 cm, more preferably from about 0.25 to about 5 cm, and most preferably from 0.4 to 2 cm. It is preferable that the feed openings 87, 89 of each feed distribution system 76a, b, c, d are diametrically opposed to introduce the feed stream into the reaction zone 28 in opposite directions. Further, it is preferable that the diametrically opposite feed openings 86, 88 of the adjacent feed distribution systems 76 are oriented at a rotation angle of 90 ° with respect to each other. During operation, the liquid-phase feed stream is fed into the main feed line 78 and then enters the distributor 80. The distributor 80 evenly distributes the feed stream for simultaneous introduction on opposite sides of the reactor 20 through the feed openings 87, 89.

FIG. 8 shows an alternative configuration, with each feed distribution system 76 equipped with tubes 90, 92 with bayonet locks, rather than plug-in nozzles 86, 88 (shown in FIG. 7). Pipes 90, 92 with bayonet closures protrude into the reaction zone 28 and have many small feed openings 94, 96 for discharging liquid phase feed into the reaction zone 28. Preferably, the small feed openings 94, 96 of the pipes 90, 92 with bayonet closures have substantially the same diameters less 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 feed distribution system 100. The feed distribution system 100 introduces a liquid phase feed stream through a plurality of vertically spaced and spaced apart locations without requiring multiple tapping into the side wall of a reactor type bubble column 20. The feed distribution system 100 typically includes one inlet pipe 102, a manifold 104, and several vertical distribution pipes 106, lateral support mechanism 108 and vertical support mechanism 110. The inlet line 102 passes through the side wall of the main part 46 of the apparatus body 22. The inlet line 102 is hydraulically connected to the manifold 104. The manifold 104 evenly distributes the feed stream received from the inlet line 102 to the vertical distribution pipes 106. Each distribution pipe 106 has a series of height-spaced feed openings 112a, b, c, d for discharging the raw material stream into the reaction zone 28. A lateral support mechanism 108 is connected to each distribution pipe 106 and prevents the relative lateral movement of the distribution pipes 106. The vertical support mechanism 110 is preferably connected to a larger a supporting support mechanism 108 and with the top of the oxidizer distributor 34. The vertical support mechanism 110 substantially impedes the vertical movement of the distribution pipes 106 in the reaction zone 28. It is preferred that the feed 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 height distance between the feed openings 112 of the feed distribution system 100 shown in FIG. 9-11 may be substantially the same as described above with respect to the feed distribution system of FIG. 1.

It has been found that the flow pattern of the reaction medium in many bubble column reactor-type columns can allow a heterogeneous azimuthal distribution of the oxidizable compound in the reaction medium, especially when the oxidizable compound is introduced mainly along one side of the reaction medium. As used herein, the term “azimuthal” means an angle or distance around the vertical longitudinal axis of the reaction zone. As used here, the term "vertical" means within 45 ° from the vertical. In one embodiment of the present invention, a feed stream containing an oxidizable compound (e.g., para-xylene) is introduced into the reaction zone through several azimuthally spaced feed openings. These azimuthally spaced loading openings can help get rid of areas with too high and too low concentrations of oxidizable compounds in the reaction medium. The various feed systems shown in FIG. 6-11 are examples of systems that provide proper azimuthal spacing of feed openings.

Referring again to Fig. 7: in order to quantitatively characterize the azimuthally spaced introduction of a liquid-phase feed stream into the reaction medium, the reaction medium can theoretically be divided into four vertical azimuth quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ " of approximately equal volume. These azimuth quadrants "Q ₁ , Q ₂ , Q ₃ , Q ₄ " are defined by a pair of imaginary intersecting perpendicular vertical planes "P ₁ , P ₂ " that extend beyond the maximum vertical size and maximum radial size of the reaction medium. When the reaction medium is contained in a cylindrical unit, the intersection of the imaginary intersecting vertical planes P _1, P ₂ is almost the same as the vertical centerline of the cylinder line, and each azimuth-quadrant Q _1, Q _2, Q _3, Q ₄ is generally tapered vertical volume with a height equal to the height of the reaction medium. Preferably, a substantial portion of the oxidizable compound is discharged into the reaction medium through loading openings located in at least two different azimuth quadrants.

In one preferred embodiment, the implementation of the present invention is not more than about 80% wt. the oxidizable compound is discharged into the reaction medium through feed openings, which may be in the same azimuth quadrant. More preferably, not more than about 60% wt. the oxidizable compound is discharged into the reaction medium through feed openings, which may be in the same azimuth quadrant. Most preferably, not more than 40% wt. the oxidizable compound is discharged into the reaction medium through feed openings, which may be in the same azimuth quadrant. These parameters for the angular distribution of the oxidizable compound are measured when the azimuth quadrants are oriented in azimuth so that the maximum possible amount of oxidizable compound is released into one of the azimuth quadrants. For example, if the entire feed stream is discharged into the reaction medium through two feed openings, which are 89 ° apart, then from the point of view of determining the angular distribution in four azimuth quadrants, it turns out that all 100% wt. the feed stream is released into the reaction medium in a single azimuth quadrant, since the azimuth quadrants can be so azimuthally oriented that both feed openings are in the same azimuth quadrant.

In addition to the advantages associated with proper azimuthal spacing of the feed openings, it has also been found that correct radial spacing of the feed openings in the reactor type bubbler may also be important. Preferably, a substantial portion of the oxidizable compound introduced into the reaction medium is discharged through loading openings which are spaced radially inward from the side wall of the apparatus. Thus, in one embodiment of the present invention, a substantial portion of the oxidizable compound enters the reaction zone through feed openings located in the “preferred radial feed zone” which is spaced inward from the vertical side walls defining the reaction zone.

7, again, the preferred radial loading zone “FZ” may take the form of a theoretically vertical cylinder located in the center of the reaction zone 28 and having an outer diameter “D _o ” of 0.9 D, where “D” is the diameter of the reaction zone 28. Thus, an OA outer ring having a thickness of 0.05D lies between the preferred radial loading zone FZ and the inner surface of the side wall defining the reaction zone 28. It is preferable that the reaction zone 28 through the feed openings located in this bunker the bottom ring of OA, little oxidizable compound was introduced, or not at all.

In another embodiment, it is preferable that little oxidizable compound is introduced into the center of reaction zone 28 or not at all. Thus, as shown in FIG. 8, the preferred radial loading zone FZ may take the form of a theoretically vertical ring located in the center of the reaction zone 28 having an outer diameter D _{o of} 0.9D and an inner diameter of D _{i of} 0.2D . Thus, in this embodiment, the inner cylinder IC having a diameter of 0.2D is “cut out” from the center of the preferred radial loading zone FZ. Preferably, little oxidizable compound is introduced into the reaction zone 28 through the feed openings located in this inner cylinder IC, or not 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 feed openings located in the preferred radial feed zone, regardless of whether the preferred radial feed zone has the cylindrical or ring shapes described above. More preferably, at least about 25% wt. the oxidizable compound is discharged into the reaction medium 36 through feed openings located in a preferred radial feed zone. Even more preferably, at least about 50% wt. the oxidizable compound is discharged into the reaction medium 36 through feed openings located in a preferred radial feed zone. Most preferably, at least 75% wt. the oxidizable compound is discharged into the reaction medium 36 through feed openings located in a preferred radial feed zone.

Although the theoretical azimuth quadrants and the theoretical preferred radial loading zone shown in FIGS. 7 and 8 are described in connection with the distribution of the liquid phase feed stream, it has been found that proper azimuth and radial distribution of the gaseous oxidizer stream can also provide certain advantages. Thus, in one embodiment of the present invention, the above azimuthal and radial distribution of the liquid phase feed stream is also applicable to the method by which a gaseous oxidant stream is introduced into the reaction medium 36.

Turning now to FIG. 12-15, where an alternative oxidizing agent distributor 200 is shown comprising a generally annular element 202 and a pair of oxidizer inlet channels 204, 206. The oxidizer dispenser 200 of FIG. 12-15 is similar to the oxidizing agent distributor 34 of FIG. 1-11 with the following three main differences: (1) the oxidizing agent distributor 200 does not contain a diagonal transverse beam; (2) the upper part of the annular element 202 has no openings for the release of the oxidizing agent in the upward direction; and (3) the oxidizing agent distributor 200 has many more openings in the lower portion of the annular member 202.

As perhaps best illustrated in FIG. 14 and 15, the lower part of the oxidizer distributor ring 202 has a plurality of oxidizer holes 208. The openings 208 for the oxidizing agent are preferably shaped so that at least about 1% of the total open area formed by the openings 208 for the oxidizing agent is below the midline 210 (FIG. 15) of the ring element 202, with the middle line 210 at the height of the volume center of gravity the annular element 202. More preferably, at least about 5% of the total open area formed by all of the oxidizing holes 208 is below the midline 210, with at least about 2% of the total open area being formed by the holes 208 through which the oxidant stream is discharged in running in a generally downward direction within about 30 ° of vertical. Even more preferably, at least about 20% of the total open area formed by all of the oxidant openings 208 is below the midline 210, with at least about 10% of the total open area formed by openings 208 through which the oxidant stream is discharged in the substantially down direction within approximately 30 ° of the vertical. Most preferably, at least about 75% of the total open area formed by all of the oxidant openings 208 is below the midline 210, with at least about 40% of the total open area formed by openings 208 through which the oxidant stream is discharged in a substantially downward flow direction within approximately 30 ° from the vertical. The fraction of the total open area formed by all oxidizing holes 208 that are above the midline 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 shown in FIG. 14 and 15, oxidizer holes 208 include downward-facing holes 208a and inclined holes 208b. The downwardly facing openings 208a are configured to discharge the oxidant stream substantially 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 discharge the oxidant stream substantially outward and downward at an angle “A” that is from about 15 to about 75 ° from the vertical, more preferably angle A is from about 30 to about 60 ° from the vertical, and most preferably, the angle A is from 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 hole 208 is preferably from about 2 to about 300 mm, more preferably from about 4 to about 120 mm, and most preferably from 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 oxidizing holes 208 formed in the ring member 202 is at least about 10, more preferably the total number of oxidizing holes 208 is in the range of from about 20 to about 200, and most preferably the total number of oxidizing holes 208 is in the range of 40 to 100.

Although FIG. 12-15 depict a very specific configuration of an oxidizing agent distributor 200, it should be noted here that many distributor configurations may be used to achieve the advantages described herein. For example, the oxidizing agent distributor does not need to have the octagonal configuration of the ring member shown in FIG. 12-13. On the contrary, it is possible for the oxidizing agent distributor to be formed by any configuration of streamlines in which there are several openings for discharging the oxidizing agent stream. The size, number and direction of discharge of the openings for the oxidizing agent in the current lines are preferably in the above ranges. In addition, the oxidizing agent distributor is preferably shaped to provide the azimuthal and radial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidizing agent distributor, it is preferred that the oxidizing agent distributor is physically configured and operating so as to minimize the pressure drop associated with the release of the oxidant stream from the streamlines through the oxidant openings into the reaction zone. This pressure drop is calculated as the time-averaged static pressure of the oxidizer flow inside the streamline at the inlets 66a, b of the oxidizer distributor minus the time-averaged static pressure in the reaction zone at a height where one half of the oxidizer flow is introduced higher in height of this place, and the other half of this oxidizer flow is introduced below this place in height. 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 distributor 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 oxidizer flow pressure within the streamlines of the oxidizer distributor is preferably in the range of about 0.35 to about 1 MPa, more preferably in the range of about 0.45 to about 0.85 MPa, and most preferably from 0.5 to 0.7 MPa.

As indicated above with respect to the oxidizer distributor configuration shown in FIG. 2-5, it may be desirable to continuously or periodically flush the oxidizing agent distributor with a liquid (e.g., acetic acid, water and / or para-xylene) to prevent the oxidizing agent distributor from becoming clogged with solids. When such a liquid flushing is used, it is preferable that an effective amount of the liquid (i.e., not just a small amount of liquid droplets that may naturally be present in the oxidizing stream) pass through the oxidizing agent outward from the oxidizing holes for at least one period of more than one minute every day. When the liquid is continuously or periodically discharged from the oxidizing agent distributor, it is preferred that the time-averaged ratio of the mass flow rate of the liquid through the oxidizing agent to the mass flow rate of molecular oxygen through the oxidizing agent be in the range of from about 0.05: 1 to about 30: 1, or from about 0.1: 1 to about 2: 1, or even from 0.2: 1 to 1: 1.

In one embodiment of the present invention, a significant proportion of the oxidizable compound (e.g., para-xylene) can be introduced into the reaction zone through an oxidizing agent distributor. In such a configuration, it is preferred that the oxidizable compound and molecular oxygen are discharged from the oxidizer distributor through the same openings. As noted above, the oxidizable compound is usually a liquid at standard temperature and pressure. Therefore, in this embodiment, a two-phase stream can be discharged from the oxidizer distributor, the liquid phase containing the oxidizable compound and the gas phase containing molecular oxygen. However, it should be understood that at least a portion of the oxidizable compound may be in a gaseous state upon exiting the oxidizing agent distributor. In one embodiment, the liquid phase exiting the oxidizing agent distributor is formed primarily from an oxidizable compound. In another embodiment, the liquid phase exiting the oxidizing agent distributor has substantially the same composition as the feed stream described above. When the liquid phase leaving the oxidizing agent distributor 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 in relation to the composition of the feed stream.

In one embodiment of the present invention, it is preferable that at least about 10% wt. of the entire oxidizable compound introduced into the reaction zone was introduced through an oxidizing agent distributor, more preferably at least about 40% wt. the oxidizable compound is introduced into the reaction zone through an oxidizing agent distributor, and most preferably at least 80% by weight. the oxidizable compound is introduced into the reaction zone through an oxidizing agent distributor. When all or part of the oxidizable compound is introduced into the reaction zone through an oxidizing agent distributor, it is preferred that at least about 10% wt. all molecular oxygen introduced into the reaction zone was introduced through the same oxidizing agent distributor, more preferably at least about 40 wt%. the oxidizable compound is introduced into the reaction zone through the same oxidizing agent distributor, and most preferably at least 80% by weight. the oxidizable compound is introduced into the reaction zone through the same oxidizing agent distributor. When a substantial portion of the oxidizable compound is introduced into the reaction zone through an oxidizing agent distributor, it is preferred that one or more temperature sensors (for example, thermocouples) are located in the oxidizing agent distributor. These temperature sensors can be used to help ensure that the temperature in the oxidizer distributor has not become dangerously high.

Turning now to FIG. 16-18, which shows a bubble column reactor 20, including an internal deaerator 300 located at the bottom of the reaction zone 28 near the opening 38 for outputting the suspension. It has been found that by deaerating the reaction medium 36 at a relatively high speed, side reactions occur that form impurities. As used herein, the term “deaeration” means the evolution of a gas phase from a multiphase reaction medium. When the reaction medium 36 is highly aerated (gas volume content> 0.3), the formation of impurities is minimal. When the reaction medium 36 is not strongly aerated (volumetric gas content <0.01), the formation of impurities is also minimal. However, when the reaction medium is partially aerated (gas volume 0.01–0.3), undesirable side reactions are activated and an increased amount of impurities is formed. The deaerator 300 solves these and other problems by minimizing the volume of the reaction medium 36 in a partially aerated state and the time it takes to deaerate the reaction medium 36. At the bottom of the deaerator 300, a substantially deaerated suspension is obtained that exits the reactor 20 through the opening 38 for withdrawing the suspension. The substantially 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.

On Fig shows a bubble column reactor type 20, including a level controller 302 and a flow controller 304. The level controller 302 and the flow controller 304 work together to maintain the reaction medium 36 in the reaction zone 28 at a substantially constant height. The level controller 302 is designed to detect (for example, by measuring the pressure drop or by radioisotope level measurement) the height of the upper surface 44 of the reaction medium 36 and generate a control signal 306 in response to the height of the reaction medium 36. The flow controller 304 receives the control signal 306 and sets the flow rate of the suspension through line 308 to withdraw the suspension. Thus, the flow rate of the suspension from the opening 38 for withdrawing the suspension can vary between the maximum volumetric flow rate of the suspension (F _max ) when the rise of the reaction medium 36 is too high and the minimum volumetric flow rate (F _min ) when the rise of the reaction medium 36 is too low .

To remove the solid-state oxidation product from reaction zone 28, the fraction must first pass through a deaerator 300. The deaerator 300 gives a slightly turbulent internal volume 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 phase flow down to the hole 38 to output the suspension. The rise of the gas phase from the liquid and solid phases is caused by the natural lifting force of the gas phase in the liquid and solid phases. When a deaerator 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, where the deaerator 300 includes a generally vertical side wall 308 defining within the deaeration zone 312. Preferably, the side wall 308 extends upwardly 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 separated from the reaction zone 28 and has a height "h" and a diameter of "d". The upper level 310 of the side wall 308 is open to receive the reaction medium from the reaction zone 28 into the internal volume 312. The lower level of the side wall 308 is hydraulically connected through the transition section 314 to the opening 38 for withdrawing the suspension. In some cases, for example, when the slurry outlet 38 is large or when the diameter “d” of the side wall 308 is small, a transition portion 314 can be dispensed with. As best illustrated in FIG. 18, the deaerator 300 may also include a flow stabilizer 316 located in the deaeration zone 312. The flow stabilizer 316 may be any structure that acts to prevent the formation of vortices when the solid and liquid phases flow down to the outlet 38 for withdrawing the suspension.

In order to enable the proper exit of the gas phase from the solid and liquid phases in the deaerator 300, carefully select the height "h" and the horizontal sectional area of the inner zone of deaeration 312. The height "h" and the horizontal sectional area of the inner zone of deaeration 312 must provide sufficient distance and time, so that even when the maximum amount of slurry is removed (i.e., when the slurry is discharged at F _max ), substantially all of the volume of gas bubbles may float out of the solid and liquid phases before the gas bubbles reach the bottom the discharge opening of the deaerator 300. Thus, it is preferable that the cross-sectional area of the deaeration zone 312 be such that the maximum velocity (V _dmax ) of the passage of liquid and solid phases through the deaeration zone 312 is substantially less than the natural velocity (V _u ) of the ascent of the gas bubble phase through the liquid and solid phases. The maximum rate (V _dmax ) of lowering of the liquid and solid phases through the deaeration zone 312 takes place at the maximum volumetric flow rate of the suspension (F _max ), discussed above. The natural ascent rate (V _u ) of gas bubbles through the liquid and solid phases varies depending on the size of the bubbles, however, the natural ascent rate (V _u0.5 ) of gas bubbles with a diameter of 0.5 cm through the liquid and solid phases can be used as a critical value as essentially all bubbles initially present in reaction medium 36 will have a size greater than 0.5 cm. Preferably, the cross-sectional area of deaeration zone 312 is such that V _{dmax is} less than about 75% of V _u0.5 , more preferably V _dmax less than about 40% of V _u0.5 , n most preferably, V _{dmax is} less than 20% of V _u0.5 .

The rate of lowering of the liquid and solid phases in the deaeration zone 312 of the deaerator 300 is calculated as the volumetric flow rate of the deaerated suspension through the opening 38 for withdrawing the suspension divided by the minimum cross-sectional area of the deaeration zone 312. The rate of lowering of the liquid and solid phases in the deaeration zone 312 of the deaerator 300 is preferably less than about 50 cm / sec, more preferably less than about 30 cm / sec, and most preferably less than 10 cm / sec.

It should be noted here that although the vertical side wall 308 of the deaerator 300 is shown as having a cylindrical configuration, the side wall 308 may include several side walls that form a variety of configurations (for example, triangular, square or oval), provided that the walls limit the internal space, 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 from about 0.2 to about 2 m, more preferably in the range of from about 0.3 to about 1.5 m, and most preferably in the range of from 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 sectional area of the deaeration zone 312 is substantially constant over the entire height “h” of the deaeration zone 312. Preferably, the maximum horizontal 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 of the deaeration zone 312 is in the range from 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 cross-sectional area of the reaction zone 28. Preferably, the maximum horizontal horizontal cross-sectional area of deaeration 312 is from about 0.02 to about 3 m ² , more preferably from about 0.05 to about 2 m ² , and most preferably from 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 of 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 straight approximately 0.01 to approximately 1 m ^3, and most preferably from 0.05 to 0.5 m ^3.

Turning now to FIG. 19, a reactor type bubbler column 20 is shown having an external deaerator 400. In this configuration, the aerated reaction medium 36 is discharged from the reaction zone 28 through an opening located at a height in the side of the apparatus body 22. The remote aerated medium is transferred to the external deaerator 400 through an exhaust line 402 to separate the gas phase from the solid and liquid phases. The separated gas phase leaves the deaerator 400 via line 404, and the substantially deaerated suspension leaves the deaerator 400 via line 406.

In Fig.19, the exhaust line 402 is shown approximately straight, horizontal and at right angles to the housing 22 of the apparatus. This is just one convenient configuration, and the discharge line 402 may be different in any respect, provided that it practically connects the bubble column 20 to the external deaerator 400. With regard to line 404, it is advantageous that this line connects to the deaerator 400 at the top or near top to control safety problems associated with a stagnant gas zone containing an oxidizable compound and an oxidizing agent. In addition, lines 402 and 404 may advantageously comprise flow isolating means, such as a valve.

When the reaction medium 36 is discharged from the reactor 20 through an outlet at a height as shown in FIG. 19, it is preferable that the bubble column reactor 20 is provided with a lower outlet 408 near the lower level 52 of the reaction zone 28. The lower outlet 408 and the lower one connected thereto line 410 may be used to discharge contents (i.e., emptying) of reactor 20 during shutdowns. Preferably, one or more lower exits 408 are made at the lower third of the height of the reaction medium 36, more preferably at the bottom of a quarter of the reaction medium 36, and most preferably at the lowest point of the reaction zone 28.

With the elevated suspension and deaeration system shown in FIG. 19, the bottom line 410 and outlet 408 are not used to remove the suspension from reaction zone 28 during oxidation. It is known in the art that the solid phase tends to settle under the action of gravity in non-aerated and otherwise unmixed parts of the suspension, including idle streamlines. In addition, precipitated solids (e.g. terephthalic acid) may tend to solidify as large agglomerates with continued precipitation and / or rearrangement of the crystal structure. Thus, in order to avoid clogging of the bottom streamline 410, a portion of the deaerated slurry from below the deaerator 400 can be used to continuously or periodically flush the bottom line 410 during normal operation of the reactor 20. The preferred means of providing such a flush with the suspension of line 410 is to periodically open valve 412 in line 410 and allow parts of the deaerated slurry to flow along line 410 to reaction zone 28 through the lower opening 408. Even when valve 412 is fully or partially open, only part of the deaerated slurry the slurry flows through the bottom line 410 back to the reaction zone 28. The remainder of the deaerated slurry, not used to flush the bottom line 410, is withdrawn through line 414 from the reactor 20 for further processing (eg, purification).

During normal operation of the bubbler column reactor type 20 for a long period of time (for example,> 100 hours), it is preferable that the amount of deaerated suspension used for washing line 410, was less than 50 wt.%. of the entire deaerated suspension obtained from the bottom of the deaerator 400, more preferably less than about 20 wt.%, and most preferably less than 5% wt. Further, it is preferable that for a substantial period of time the average mass flow rate of the deaerated slurry used to flush the bottom line 410 is less than about 4 times the average mass flow rate of the oxidizable compound to reaction zone 28, more preferably less than about 2 times the average mass flow rate of the oxidizable compound to the reaction zone 28, even more preferably was less than the average mass flow rate of the oxidizable compound to the reaction zone 28, and it is preferably less than half the average mass flow rate of the oxidizable compound into reaction zone 28.

Referring again to FIG. 19: the deaerator 400 includes a substantially vertical, preferably cylindrical side wall 416 defining a deaeration zone 418. The deaeration zone 418 has a diameter “d” and a height “h”. The height “h” is measured as the height distance from the place where the aerated reaction medium enters the deaerator 400 to the bottom of the side wall 416. The height “h”, the diameter “d”, the area and volume of the deaeration zone 418 are preferably substantially the same as described above with respect to the deaeration zone 312 of the deaerator 300 shown in FIG. 16-18. In addition, the deaerator 400 includes an upper section 420 formed by extending the side wall 416 above the deaeration zone 418. The upper section 420 of the deaerator 400 can be of any height, although preferably it extends up to or above the level of the reaction medium 36 in the reaction zone 28. The upper section 420 ensures that the gas phase has room to escape from the liquid and solid phases before leaving the deaerator 400 via line 404. It should be noted here that although line 404 is shown as returning the evolved gas phase to the separation zone of reactor 20, alternatively line 404 can be connected to the apparatus body 22 at any height above the outlet 402. Alternatively, line 404 can be connected to a gas exhaust line 40 so that the gas phase released from the deaerator 400 is combined with the remote overhead vapor stream in line 40 and sent further according to the scheme for further processing.

Turning now to FIG. 20, a reactor type bubbler column 20 is shown including a hybrid internal-external deaerator 500. In this configuration, part of the reaction medium 36 is discharged from the reaction zone 28 through a relatively wide opening 502 in the side wall of the apparatus body 22. The withdrawn reaction medium 36 is then transported through a relatively large diameter elbow 504 and enters the deaerator 500 from above. In FIG. 20, the elbow 504 is shown connecting at right angles to the side wall of the apparatus body 22 and including a smooth rotation at an angle of about 90 °. This is just one convenient configuration, and the elbow 504 may be different in all respects provided that it practically connects the bubbler column 20 to the external deaerator 500 as described. In addition, elbow 504 may practically comprise a means for cutting off the flow, such as a valve.

In the deaerator 500, the gas phase moves up, and the solid and liquid phases move down. The rising gas phase can again enter the elbow 504 and then through the opening 502 enter the reaction zone 28. Thus, counterflow of the incoming reaction medium 36 and the outgoing separated gas can take place near the opening 502. The deaerated slurry exits the deaerator 500 via line 506. The deaerator 500 includes a substantially vertical, preferably cylindrical side wall 508 defining a deaeration zone 510. The deaeration zone 510 has a height “h” and a diameter “d”. Preferably, the opening 502 is at a height of the elbow 504 had the same or a larger diameter 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 respect to the deaeration zone 312 of the deaerator 300 shown of FIG. 16-18.

In FIG. 19 and 20 show an embodiment of a bubble column reactor 20, where a solid product (e.g., technical terephthalic acid) obtained in reaction zone 28 is discharged from reaction zone 28 through an outlet located at a height. The removal of the aerated reaction medium 36 from a height above the bottom of the bubble column reactor 20 can help to avoid the accumulation and stagnation of poorly aerated reaction medium 36 in the lower level 52 of reaction zone 28. According to other aspects of the present invention, the concentration of oxygen and oxidizable compound (for example, para-xylene) in the reaction medium 36 near the top of the reaction medium 36 is preferably lower than near the bottom. Thus, removing the reaction medium 36 from an elevated location can increase the yield by reducing the amount of unreacted reactants discharged from the reactor 20. Also, the temperature of the reaction medium 36 varies significantly in the vertical direction when the bubble column reactor 20 operates at high STR and gradients chemical composition as described here. Under such conditions, the temperature of the reaction medium 36 typically will have local minima near the lower and upper levels of the reaction zone 28. Near the lower level, the minimum is associated with evaporation of the solvent near the point where all or part of the oxidizing agent enters. Near the upper level, the minimum is again due to the evaporation of the solvent, although here it is due to a drop in pressure in the reaction medium. In addition, other local minima may occur between the upper and lower levels where additional raw materials or an oxidizing agent are introduced into the reaction medium. Thus, between the lower and upper levels of reaction zone 28, there is one or more temperature maxima due to the heat of the exothermic oxidation reaction. Removal of the reaction medium 36 from the outlet located at a height in places with an elevated temperature can be particularly advantageous when subsequent processing is carried out at elevated temperatures, since the energy costs associated with heating the remote medium for subsequent processing are reduced.

Thus, in a preferred embodiment of the present invention, and especially when subsequent processing occurs at elevated temperatures, the reaction medium 36 is discharged from the bubble column reactor 20 through high-level outlets located above the place (s) where the reaction zone 28 is introduced at least 50% wt. liquid phase feed stream and / or gaseous oxidant stream. More preferably, the reaction medium 36 is discharged from the bubbler column of the reactor type 20 through high-level outlets located above the point (s) where substantially all the liquid-phase feed stream and / or gaseous oxidant stream are introduced into the reaction zone 28. Preferably, at least 50% wt. solid-phase and liquid-phase components removed from the bubble column reactor type 20, is removed through the exits located at a height. More preferably, substantially all of the solid phase and liquid phase components discharged from the bubble column reactor type 20 are removed through outlets at a height. Preferably, the exits located at a height are located at least about 1D above the lower level 52 of the reaction zone 28. More preferably, the exits located at a height of at least about 2D above the lower level 52 of the reaction zone 28. Most preferably, the exits are at a height located at least 3D above the lower level 52 of reaction zone 28. For a given height “H” of reaction medium 36, it is preferred that the exits located at a height are in height 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 the height of the exit from the reaction zone 28 be at least 1 ° C higher than the temperature of the reaction medium 36 at the lower level 52 of the reaction zone 28. More preferably, the temperature of the reaction medium 36 at the at the outlet height of the reaction zone 28, about 1.5-16 ° C higher than the temperature of the reaction medium 36 at the lower level 52 of the reaction zone 28. Most preferably, the temperature of the reaction medium 36 at the outlet of the reaction zone 28 and 2-12 ° C higher than the temperature of reaction medium 36 at lower end 52 of reaction zone 28.

Referring now to FIG. 21, a reactor-type bubbler column 20 is shown including an alternative hybrid deaerator 600 located at the bottom of reactor 20. In this configuration, aerated reaction medium 36 is discharged from reaction zone 28 through a relatively wide opening 602 in the lower level 52 of apparatus vessel 22 . Hole 602 defines the upper open end of the deaerator 600. In the deaerator 600, the gas phase moves up, and the solid and liquid phases move down. The upwardly moving gas phase can again enter the reaction zone 28 through the opening 602. Thus, counterflow of the incoming reaction medium 36 and the outgoing separated gas can take place at the opening 602. The deaerated slurry exits the deaerator 600 via line 604. The deaerator 600 includes a substantially vertical, preferably cylindrical side wall 606 defining a deaeration zone 608. The deaeration zone 608 has a height “h” and a diameter “d”. Preferably, the aperture 602 has the same or larger diameter than the diameter "d" of the deaeration zone 608. The height "h", diameter "d", the area and volume of the deaeration zone 608 are preferably substantially the same as described above with respect to the deaeration zone 312 of the deaerator 300 shown in FIG. 16-18.

Turning now to FIG. 22, the reactor type bubbler column 20 of FIG. 21 is shown including an alternate oxidizer distributor 620. The oxidizer distributor 620 includes an annular element 622 and a pair of inlets 624, 626. The annular element 622 preferably has substantially the same configuration as the annular element 202 described above in connection with FIG. 12-15. The inlet pipes 624, 626 pass upward through the openings in the bottom 48 of the apparatus body 22 and provide an oxidizer flow into the annular element 622.

Turning now to FIG. 23, the reactor type bubble column 20 of FIG. 21 is shown including a dispenser-free device for introducing an oxidizer stream into the reaction zone 28. In the configuration of FIG. 23, an oxidizer stream is supplied to the reactor 20 via an input line 630, 632 oxidizing agent. The oxidizer input lines 630, 632 are connected to the corresponding oxidizer holes 634, 636 in the bottom 48 of the apparatus body 22. The oxidizing agent stream is introduced directly into the reaction zone 28 through the oxidizing holes 634, 636. Optional bumpers 638, 640 may be installed to deflect the flow of the oxidant stream as soon as it enters the reaction zone 28.

As noted above, it is preferable that the oxidation reactor be designed and operate so as to avoid zones with a high concentration of oxidizable compounds in the reaction medium, since such zones can lead to the formation of impurities. One way to improve the initial distribution of an oxidizable compound (e.g., para-xylene) in the reaction medium is to dilute the oxidizable compound with a liquid. The liquid used to dilute the oxidizable compound may be part of the reaction medium located at a considerable distance from the place (s) where the oxidizable compound is fed into the reaction zone. This liquid from a remote part of the reaction medium can be returned to a place close to the entry point of the oxidizable compound along a stream line that is located inside and / or outside the main reaction apparatus.

FIG. 24 and 25 illustrate two preferred methods for circulating liquid between the distal portion of the reaction medium and the site near the entrance of the oxidizable compound, using internal (Fig. 24) or external (Fig. 25) channels. Preferably, the length of the streamline from its inlet (i.e., the hole (s) where the liquid enters the line) to its exit (i.e., the hole (s) where the liquid enters the line) is greater than about 1 m, more preferably more than about 3 m, even more preferably more than about 6 m, and most preferably more than 9 m. However, the actual line length becomes less important if the fluid is obtained from a separate apparatus, possibly immediately above or next to the apparatus into which the oxidizable compound is initially introduced. Liquid from any single apparatus containing at least some portion of the reaction medium is the preferred source for the initial dilution of the oxidizable compound.

Whatever the source, it is preferred that the fluid flowing through the line has a lower steady state concentration of the oxidizable compound than the reaction medium immediately adjacent to at least one outlet of the line. In addition, it is preferable that the fluid flowing through the line has a concentration of oxidizable compounds in the liquid phase 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., and the concentration was measured before adding to the line an additional supply of oxidizable compounds and any (optional) separate supply of solvent. When measured after adding an additional supply of oxidizable compound and an alternative supply of solvent, it is preferable that the combined liquid stream entering the reaction medium has a concentration of oxidizable compound in the liquid phase below about 300,000 ppm. wt., more preferably below about 50,000 ppm. wt. and most preferably below 10,000 ppm. wt.

It is desirable to keep the flow through the line at a sufficiently low speed so that the circulating liquid does indeed lower 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, where an additional amount of the oxidizable compound is initially admitted, to the mass flow rate of the liquid flowing along the line, be greater than about 0.3 minutes, more preferably more than about 1 minute , even more preferably would be from about 2 minutes to about 120 minutes, and most preferably from 3 minutes to 60 minutes.

There are many ways to make fluid flow through the line. Preferred means include gravity, any type of ejection device using gas or liquid or both as a motive fluid, and mechanical pumps of any type. When using an ejection device, in one embodiment of the invention, at least one medium selected from the group consisting of a feed of an oxidizable compound (liquid or gas), a feed of an oxidizing agent (gas), a feed of a solvent (liquid), and pumped fluid is used as a driving fluid pump source of the reaction medium (suspension). In another embodiment, at least two media selected from the group consisting of a feed of an oxidizable compound, a feed of an oxidizing agent, and a feed of a solvent are used as a driving fluid. In yet another embodiment, a combination of an oxidizable compound feed, an oxidizer feed and a solvent feed is used as the driving fluid.

The appropriate diameter or diameters of the circulation line can vary in accordance with the quantity and properties of the material being conducted, the energy available to carry out the flow, and for capital costs reasons. Preferably, the minimum diameter for such a line is greater than about 0.02 m, more preferably from about 0.06 m to about 2 m, and most preferably from 0.12 to 0.8 m.

As noted above, it is desirable to control the flow through the line in certain preferred ranges. Many means are known in the art for influencing this regulation by establishing a suitable specific configuration when creating current channels. Another implementation option is the use of configurations that change during operation, including, in particular, valves of all types and designs, including manual or actuated by any means, including feedback circuits from sensors or without them. Another preferred means of controlling the flow of dilute liquid is to change the energy input between the input and output of the line. Preferred means include changing the flow rate of one or more moving media into the ejection device, changing the energy supply to the pump drive, and changing the density difference or increasing the difference when gravity is used. These preferred agents may also be used in any combination.

The line used to circulate liquid from the reaction medium may be of any type known in the art. In one embodiment, a line is used, in whole or in part, made from conventional piping materials. In another embodiment, a line is used made in whole or in part using the wall of the reaction apparatus as part of the line. The line can be made entirely within the boundaries of the reaction apparatus (Fig.24), or it can be made completely outside the reaction apparatus (Fig.25), or it can include sections both inside and outside the reaction apparatus.

The authors of the present invention believe that, especially in larger reactors, it may be desirable to have several lines of different designs for movement of fluid along them. Further, it may be desirable to make several outputs in several positions in one or all lines. Design features should balance the desired overall gradient of stationary concentrations of the oxidizable compound with the desired initial dilution of the feed of the oxidizable compound in accordance with other aspects of the present invention.

Both of FIGS. 24 and 25 illustrate a construction in which a deaerator connected to the line is used. This deaerator ensures that part of the reaction medium used to dilute the incoming oxidizable compound is a substantially deaerated suspension. It should be noted here, however, that the liquid or suspension used to dilute the incoming oxidizable compound may be in either an aerated form or a deaerated form.

The use of fluid flowing through the line to dilute the feed of the oxidizable compound is particularly useful in bubble column reactors. In addition, in reactor-type bubbler columns, a large gain for the initial dilution of the feed of the oxidizable compound can be achieved even without adding the feed of the oxidizable compound directly to the line, provided that the output of the line lies close enough to the place of addition of the oxidizable compound. In such an embodiment, it is preferable that the line outlet is within about 27 diameters of the line outlet from the closest point of addition of the oxidizable compound, more preferably within about 9 diameters of the line outlet, even more preferably within about 3 diameters of the line outlet, and most preferably within 1 diameter of the outlet of the line.

It has also been found that, according to one embodiment of the present invention, an ejection flow device may be useful for initially diluting the supply of an oxidizable compound to bubble oxidation columns, even without using liquid dilution lines from a distant portion of the reaction medium. In such cases, the ejection device is placed in the reaction medium and has an open passage from the reaction medium to the inlet of the ejection device, where low pressure draws in the nearest reaction medium. Examples of two possible configurations of ejection devices are shown in FIG. 26 and 27. In a preferred embodiment of these ejection devices, the closest feed point of the oxidizable compound is within about 4 m, more preferably within about 1 m, and most preferably 0.3 m from the throat of the ejection device. In another embodiment, the oxidizable compound is supplied under pressure as a driving fluid. In yet another embodiment, together with the oxidizable compound, either a solvent or an oxidizing agent is supplied under pressure as an additional driving fluid. In yet another embodiment, both the solvent and the oxidizing agent are supplied together with the oxidizable compound as an additional driving fluid under pressure.

The inventors believe that, especially in larger reactors, it may be desirable to have several ejection devices of different designs located at different places in the reaction medium. A balance of the desired overall gradient of stationary concentrations of the oxidizable compound with the desired initial dilution of the feed of the oxidizable compound, according to other aspects of the present invention, will be ensured by design features. In addition, the authors of the present invention believe that the stream of the output stream from the ejection device can be oriented in any direction. When using multiple ejection devices, each ejection device can be oriented independently, again in any direction.

As noted above, certain physical and operational properties of the bubble column reactor 20 described above in connection with FIG. 1-27, provide gradients in the heights of pressure, temperature, and concentrations of reagents (ie, oxygen and an oxidizable compound) in the reaction medium 36. As discussed above, these height gradients can provide a more efficient and cost-effective oxidation method than conventional oxidation methods that maintain a well-mixed reaction medium with relatively uniform pressure, temperature and concentration of reagents throughout the medium. Further, gradients of oxygen along the height, oxidizable compound (for example, para-xylene) and temperatures made possible by the use of the oxidation system according to an embodiment of the present invention will be described in more detail.

According to FIG. 28, in order to quantify the concentration gradients of the reactants present in the reaction medium 36 during oxidation in a bubble column reactor type 20, the total volume of the reaction medium 36 can theoretically be divided into 30 separate horizontal layers of equal volume. The principle of dividing the reaction medium 36 into 30 separate horizontal layers of equal volume is illustrated in FIG. With the exception of the uppermost and lowermost horizontal layers, each horizontal layer in a separate volume is bounded above and below by imaginary horizontal planes, and on the sides it is bounded by the walls of the reactor 20. The uppermost horizontal layer is bounded below by an imaginary horizontal plane, and above by the upper surface of the reaction medium 36 The lowest horizontal layer is bounded above by an imaginary horizontal plane, and from below by the bottom of the apparatus body. After the reaction medium 36 has been theoretically divided into 30 separate horizontal layers of equal volume, the average concentration in time and volume in each horizontal layer can be determined. An individual horizontal layer having a maximum concentration of all 30 horizontal layers can be referred to as a “C-max horizontal layer”. An individual horizontal layer above the horizontal C-max layer and having a minimum concentration of all horizontal layers above the horizontal C-max layer can be called a "horizontal C-min layer". Then you can calculate the gradient for the height of the concentration as the ratio of the concentration in the horizontal layer C-max to the concentration in the horizontal layer C-min.

As for the quantitative determination of the oxygen concentration gradient, when the reaction medium 36 was theoretically divided into 30 separate horizontal layers of equal volume, the horizontal O ₂ -max layer is defined as the layer having the maximum oxygen concentration of all 30 horizontal layers, and the horizontal O ₂ -min layer is determined as a layer having a minimum oxygen concentration of horizontal layers above the horizontal O ₂ -max layer. The oxygen concentration in the horizontal layers is measured in the gas phase of the reaction medium 36 as a molar concentration averaged over time and volume, calculated on the wet state. Preferably, the ratio of the oxygen concentration in the horizontal O ₂ -max layer to the oxygen concentration in the horizontal O ₂ -min layer is in the range from about 2: 1 to about 25: 1, more preferably from about 3: 1 to about 15: 1, and most preferably from 4: 1 to 10: 1.

Typically, the horizontal O ₂ -max layer will be near the bottom of the reaction medium 36, and the horizontal O ₂ -min layer will be near the top of the reaction medium 36. Preferably, the horizontal O ₂ -min layer is one of the 5 topmost horizontal layers of 30 separate horizontal layers. Most preferably, horizontal O ₂ -min layer is the uppermost of separate horizontal layers 30 as shown in Figure 28. Preferably, the horizontal layer of O ₂ -max is one of the 10 lowest horizontal layers of 30 separate horizontal layers. Most preferably, the horizontal O ₂ -max layer is one of the 5 lowest horizontal layers of 30 separate horizontal layers. For example, FIG. 28 shows the horizontal O ₂ -max layer as the third horizontal layer from the bottom of the reactor 20. Preferably, the height distance between the horizontal O ₂ -min and O ₂ -max layers is at least about 2W, more preferably at least about 4W and most preferably at least 6W. Preferably, the height distance between the horizontal layers 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 oxygen concentration, on a wet state in a horizontal layer O ₂ -min is preferably in the range of from about 0.1 to about 3 mole%, more preferably in the range of from about 0.3 to about 2% molar, and most preferably from 0.5 to 1.5% molar. The time and volume averaged oxygen concentration in the horizontal O ₂ -max layer is preferably in the range of about 4 to about 20% mole, more preferably in the range of about 5 to about 15% mole, and most preferably 6 to 12% mole . The time-averaged oxygen concentration, calculated on the dry matter, in the gaseous stream leaving the reactor 20 through the gas outlet 40 is preferably from about 0.5 to about 9% mole, more preferably from about 1 to about 7% mole and most preferably from 1.5 to 5% mole.

Since the oxygen concentration drops markedly towards the top of the reaction medium 36, it is desirable that the oxygen demand at the top 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, para-xylene), with a minimum concentration of the oxidizable compound near the top of the reaction medium 36.

As for the quantitative determination of the concentration gradient of an oxidizable compound (for example, para-xylene), when the reaction medium 36 is theoretically divided into 30 separate horizontal layers of equal volume, the horizontal layer OC-max is defined as having the maximum concentration of the oxidizable compound from all 30 horizontal layers, and the horizontal the OC-min layer is defined as having a minimum concentration of an oxidizable compound of horizontal layers above the horizontal OC-max layer. The concentration of the oxidizable compound in horizontal layers is measured in the liquid phase as a mass fraction averaged over time and volume. Preferably, the ratio of the concentration of the oxidizable compound in the horizontal layer OC-max to the concentration of the oxidizable compound in the horizontal layer OC-min is greater than about 5: 1, more preferably more than about 10: 1, even more preferably more than about 20 : 1, and most preferably was in the range from 40: 1 to 1000: 1.

Typically, the horizontal OC-max layer will be near the bottom of the reaction medium 36, and the horizontal OC-min layer will be near the top of the reaction medium 36. Preferably, the horizontal OC-min layer is one of the 5 uppermost horizontal layers of 30 separate horizontal layers. Most preferably, the horizontal layer OC-min is the uppermost of 30 separate horizontal layers, as shown in FIG. Preferably, the horizontal layer OC-max is one of the 10 lowest horizontal layers of 30 separate horizontal layers. Most preferably, the horizontal layer OC-max is one of the 5 lowest horizontal layers of 30 separate horizontal layers. For example, in FIG. 28, the horizontal OC-max layer is shown as the fifth horizontal layer from the bottom of the reactor 20. Preferably, the height distance between the horizontal layers OC-min and OC-max is at least about 2W, where “W” is the maximum the width of the reaction medium is 36. More preferably, the height distance between the horizontal layers 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 height distance between the horizontal layers 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 concentration of an oxidizable compound (e.g., para-xylene) averaged over time and volume in the liquid phase of the horizontal layer 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 lies in the range from 1 ppm wt. up to 100 ppm wt. The concentration of oxidizable compound averaged over time and volume in the liquid phase of the horizontal layer OC-max is preferably from about 100 ppm. wt. up to about 10,000 ppm wt., more preferably from about 200 ppm. wt. up to about 5000 ppm wt., and most preferably from 500 ppm wt. up to 3000 ppm wt.

Although it is preferable that the bubble column reactor 20 provides gradients of the concentration of the oxidizable compound in height, it is also preferable that the volume fraction of the reaction medium 36 having a concentration of the oxidizable compound in the liquid phase of more than 1000 ppm. wt., was minimal. Preferably, the time-averaged volume fraction of the reaction medium 36 having a concentration of an oxidizable compound in the liquid phase is higher than 1000 ppm. wt. was less than about 9%, more preferably less than about 6%, and most preferably less than 3%. Preferably, the time-averaged volume fraction of the reaction medium 36 having a concentration of an oxidizable compound in the liquid phase of more than 2500 ppm. wt., is less than about 1.5%, more preferably less than about 1%, and most preferably less than 0.5%. Preferably, the time-averaged volume fraction of the reaction medium 36 having a concentration of an oxidizable compound in the liquid phase of more than 10,000 ppm. wt., is less than about 0.3%, more preferably less than about 0.1%, and most preferably less than 0.03%. Preferably, the time-averaged volume fraction of the reaction medium 36 having a concentration of an oxidizable compound in the liquid phase of more than 25,000 ppm. wt., is less than about 0.03%, more preferably less than about 0.015%, and most preferably less than 0.007%. The inventors note that the volume of reaction medium 36 having elevated levels of oxidizable compound does not necessarily lie in a single continuous volume. At many points in time, the chaotic flow regime in the bubbler reaction apparatus yields simultaneously two or more continuous but separated parts of the reaction medium 36 having elevated levels of oxidizable compound. At each time used in time averaging, all such continuous but separated volumes, comprising more than 0.0001% by volume of the entire reaction medium, 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 above gradients of the concentration of oxygen and oxidizable compounds, it is preferable that in the reaction medium 36 there was a temperature gradient. Referring again to FIG. 28, this temperature gradient can be quantified in a manner analogous to the concentration gradients, theoretically dividing the reaction medium 36 into 30 separate horizontal layers of equal volume and measuring the temperature of each layer averaged over time and volume. The horizontal layer with the lowest temperature of the 15 lower horizontal layers can in this case be defined as the horizontal T-min layer, and the horizontal layer above the horizontal T-min layer and having the maximum temperature of all layers above the horizontal T-min layer can in this case be defined as "horizontal T-max layer." Preferably, the temperature of the horizontal layer T-max is at least about 1 ° C higher than the temperature of the horizontal layer T-min. More preferably, the temperature of the horizontal T-max layer is about 1.25-12 ° C higher than the temperature of the horizontal T-min layer. Most preferably, the temperature of the horizontal T-max layer is 2-8 ° C higher than the temperature of the horizontal T-min layer. The temperature of the horizontal T-max layer is preferably from about 125 to about 200 ° C, more preferably from about 140 to about 180 ° C, and most preferably from 150 to 170 ° C.

Typically, the horizontal T-max layer will be near the center of the reaction medium 36, and the horizontal T-min layer will be near the bottom of the reaction medium 36. Preferably, the horizontal T-min layer is one of the 10 lowest horizontal layers of the 15 lower horizontal layers. Most preferably, the T-min horizontal layer is one of the 5 lowest horizontal layers of the 15 lower horizontal layers. For example, in FIG. 28, the horizontal T-min layer is shown as the second horizontal layer from the bottom of the reactor 20. Preferably, the horizontal T-max layer is one of 20 middle horizontal layers of 30 separate horizontal layers. Most preferably, the T-min horizontal layer is one of 14 middle horizontal layers of 30 separate horizontal layers. For example, in FIG. 28, the horizontal T-max layer is shown as the twentieth horizontal layer from the bottom of the reactor 20 (i.e., one of the 10 middle horizontal layers). Preferably, the height distance between the horizontal layers T-min and T-max is at least about 2W, more preferably at least about 4W, and most preferably at least 6W. Preferably, the height distance between the horizontal layers 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 temperature gradient in the reaction medium 36 over the height, it may be advantageous to divert the reaction medium 36 at a height where the temperature of the reaction medium is highest, especially when the product to be removed is further processed at elevated temperatures. Thus, when the reaction medium 36 is withdrawn from the reaction zone 28 through one or more exits located at a height, as shown in FIG. 19 and 20, it is preferable that the exits located at a height are near the horizontal T-max layer. Preferably, the outlet located at a height is within 10 horizontal layers from the horizontal T-max layer, more preferably within 5 horizontal layers from the horizontal T-max layer, and most preferably within 2 horizontal layers from the horizontal T-max layer.

It should be noted here that many of the distinguishing features of the invention described herein can be used in oxidation systems with multiple reactors, and not only in systems using a single oxidation reactor. In addition, some of the distinguishing features of the invention described herein can be used in oxidation reactors with mechanical stirring and / or with stirring with the flow, and not just in reactors with mixing with gas bubbles (i.e., bubble-type bubble columns). For example, the authors of the present invention have found certain advantages associated with the gradual change / variation of the oxygen concentration and / or its flow rate in the reaction medium. The benefits obtained by gradually changing the oxygen concentration / flow rate in the reaction medium can be realized regardless of whether the entire volume of the reaction medium is contained in one apparatus or in several apparatuses. Further, the benefits realized by gradually changing the oxygen concentration / flow rate in the reaction medium can be realized regardless of whether the reaction apparatus (s) is mechanically agitated, agitated by flow and / or agitated by gas bubbles.

One way to quantify the degree of gradual change in oxygen concentration and / or flow rate 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 limited to any particular form. However, each 20% continuous volume may be formed from adjacent volumes of the reaction medium (i.e., each volume is “continuous”), and 20% continuous volumes should not overlap with each other (i.e., the volumes are “separate”). In FIG. 29-31 show that these individual 20% continuous volumes can be in the same reactor (Fig. 29) or in several reactors (Figs. 30 and 31). It should be noted that the reactors shown in FIG. 29-31 may be mechanically agitated, jet agitated and / or bubble agitated reactors. In one embodiment, it is preferred that the reactors shown in FIG. 29-31 were bubble stirred reactors (i.e., bubble type bubble columns).

Referring again to FIG. 29, a reactor 20 is shown containing the reaction medium 36. The reaction medium 36 includes a first separate 20% continuous volume 37 and a second separate 20% continuous volume 39.

Next, FIG. 30 shows a cascade of reactors including a first reactor 720a and a second reactor 720b. The reactors 720a, b together contain the full volume of the reaction medium 736. The first reactor 720a contains the first part 736a of the reaction medium, and the second reactor 720b contains the second part 736b of the reaction medium. A first separate 20% continuous volume 737 of the reaction medium 736 is shown enclosed within the first reactor 720a, and a second separate 20% continuous volume 739 of the reaction medium 736 is shown enclosed within the second reactor 720b.

Turning now to FIG. 31, a multi-reactor system is shown including a first reactor 820a, a second reactor 820b, and a third reactor 820c. Reactors 820a, b, c together contain the full volume of the reaction medium 836. The first reactor 820a contains the first part 836a of the reaction medium, the second reactor 820b contains the second part 836b of the reaction medium, and the third reactor 820c contains the third part 836c of the reaction medium. The first separate 20% continuous volume 837 of the reaction medium 836 is shown enclosed within the first reactor 820a, the second separate 20% continuous volume 839 of the reaction medium 836 is shown enclosed inside the second reactor 820b, and the third separate 20% continuous volume 841 of the reaction medium 836 is enclosed inside the third reactor 820c.

The stepwise change in the oxygen concentration in the reaction medium can be quantified by considering a 20% continuous volume of the reaction medium having the largest molar fraction of oxygen in the gas phase, and considering a 20% continuous volume of the reaction medium having the smallest molar fraction of oxygen in the gas phase. In the gas phase of an individual 20% of the continuous volume of the reaction medium having the highest oxygen concentration in the gas phase, the oxygen concentration averaged over time and volume, calculated on the wet state, is preferably from about 3 to about 18% molar, more preferably from about 3.5 to about 14% molar, and most preferably from 4 to 10% molar. In the gas phase of an individual 20% of the continuous volume of the reaction medium having the lowest oxygen concentration in the gas phase, the concentration of oxygen averaged over time and volume, based on the wet state, is preferably from about 0.3 to about 5% mole, more preferably from about 0.6 to about 4% molar, and most preferably from 0.9 to 3% molar. Further, the ratio of the concentration of oxygen averaged over time and volume, based on the wet state, in the most enriched 20% continuous volume of the reaction medium and the most depleted 20% continuous volume of the reaction medium is preferably from about 1.5: 1 to about 20: 1 , more preferably from about 2: 1 to about 12: 1, and most preferably from 3: 1 to 9: 1.

The stepwise change in the rate of oxygen flow in the reaction medium can be quantified in terms of the oxygen STR described above at the beginning. STR oxygen has been previously described in global terms (i.e., in terms of the average STR oxygen throughout the reaction medium); however, the STR of oxygen can also be considered in a local value (i.e., part of the reaction medium) in order to quantify the degree of stepwise change in the rate of oxygen consumption in the reaction medium.

The inventors of the present invention have found that it is very advantageous for the oxygen STR to vary across the reaction medium in perfect harmony with the desired gradients described here with respect to the pressure in the reaction medium and the mole fraction of molecular oxygen in the gas phase of the reaction medium. Thus, it is preferable that the ratio of STR oxygen in the first separate 20% continuous volume of the reaction medium to STR oxygen in the second separate 20% continuous volume of the reaction medium is in the range from about 1.5: 1 to about 20: 1, more preferably in the range from about 2: 1 to about 12: 1, and most preferably in the range from 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 point where molecular oxygen is initially introduced into the reaction medium. These large oxygen gradients STR are desirable regardless of whether the partial oxidation reaction medium is in a bubble column oxidation reactor or in any other type of apparatus in which pressure and / or molar fractions of molecular oxygen in the gas phase of the reaction medium are created (e.g. in an apparatus with mechanical stirring, having several mixing zones located along the height, achieved by using several mixers having a strong radial flow, possibly which is mainly provided by horizontal guides of the reflective baffles, and the oxidizer flow rises mainly upward from the feed near the lower portion of the reaction apparatus, despite the fact that in each mixing zone located at the highest height, substantial countercurrent mixing of the oxidizing stream can occur and that some countercurrent mixing of the oxidizing stream can occur between adjacent high-altitude mixing zones). That is, the inventors of the present invention have found that when there is a gradient of pressure and / or molar fraction of molecular oxygen in the gas phase of the reaction medium, it is desirable to use the methods described here to create a similar gradient in the chemical need of dissolved oxygen.

A preferred way of causing the local oxygen STR to change is to control the feed points of the oxidizable compound and control the mixing of the liquid phase of the reaction medium to adjust the concentration gradients of the oxidizable compound according to other descriptions of the present invention. Another suitable way to make the local STR of oxygen change is to cause a change in the activity of the reaction by creating a local change in temperature, cause a change in the local mixture of the catalyst and solvent components (for example, introducing additional gas to evaporate to cool in a specific part of the reaction medium, and adding a stream of solvent, containing a higher amount of water to lower activity in a particular part of the reaction medium).

As discussed above with respect to FIG. 30 and 31, the partial oxidation reaction can advantageously be carried out in several reaction apparatuses, with at least a portion, preferably at least 25%, more preferably at least 50%, and most preferably at least 75% of the molecular oxygen leaving from the first reaction apparatus, is carried out in one or more of the following reaction apparatus to absorb an additional additive, preferably more than 10%, more preferably more than 20%, and most preferably more than 40% molecular oxygen a, leaving the first / top according to the scheme of the reaction apparatus. If such a sequential flow of molecular oxygen from one reactor to another is used, it is desirable that the first reaction apparatus operate at a higher reaction rate than at least one of the following reaction apparatuses, preferably with a ratio of the oxygen averaged over the STR apparatus in the first reaction apparatus to the averaged by oxygen apparatus STR in the following reaction apparatus, comprising from about 1.5: 1 to about 20: 1, more preferably from about 2: 1 to about 12: 1, and most preferably respectfully from 3: 1 to 9: 1.

As discussed above, in accordance with the present invention, any type of first reaction apparatus (e.g., a bubble column, with mechanical stirring, with back mixing, with internal steps, piston, etc.) is suitable for the sequential flow of molecular oxygen into the following reaction apparatuses and all types of the following reaction apparatus, which may be of the same or different type than the first reaction apparatus. Means of causing STR averaged oxygen to decrease in subsequent reaction apparatuses effectively include lowering the temperature, lowering the concentration of the oxidizable compound, and lowering the reaction activity of the particular mixture of catalyst components and solvent (e.g., lowering the concentration of cobalt, increasing the concentration of water, and adding a catalytic moderator, such as a small amount ionic copper).

When flowing from the first reaction apparatus to the next reaction apparatus, the oxidizing stream can be treated by any method known in the art, such as increasing or decreasing pressure, cooling or heating, and removing the mass or adding mass of any type or in any quantity. However, the use of a decreasing average STR apparatus oxygen in subsequent reaction apparatuses is particularly useful when the absolute pressure at the top of the first reaction apparatus is less than about 2.0 MPa, more preferably less than about 1.6 MPa, and most preferably less than 1.2 MPa. In addition, the use of decreasing oxygen averaged over the STR apparatus in subsequent reaction apparatuses is particularly useful when the ratio of the absolute pressure in the upper part of the first reaction apparatus to the absolute pressure in the upper part of at least one subsequent reaction apparatus is in the range of from about 0.5: 1 up to 6: 1, more preferably in the range of from about 0.6: 1 to about 4: 1, and most preferably in the range of from 0.7: 1 to 2: 1. Decreasing the pressure in subsequent apparatuses below these lower limits excessively reduces the availability of molecular oxygen, and increasing the pressure above these upper limits is increasingly expensive compared to using a fresh oxidizing agent.

When a sequential flow of molecular oxygen is used in subsequent reaction apparatuses with decreasing oxygen averaged over the STR apparatus, fresh streams of oxidizable compound, solvent and oxidizing agent may be introduced into subsequent reaction apparatuses and / or into the first reaction apparatus. The flows of the liquid phase and the solid phase (if any) of the reaction medium can flow in any direction between the reaction apparatuses. All or part of the gas phase leaving the first reaction apparatus and entering the next reaction apparatus can flow separately from or mixed with fractions of the liquid phase or solid phase (if any) of the reaction medium from the first reaction apparatus. A product stream containing a liquid phase and a solid phase (if any) can be removed from the reaction medium in any reaction apparatus of the system.

According to FIG. 1-29, the oxidation is preferably carried out in a bubbler column of a reactor type 20 under conditions that, according to the preferred embodiments described herein, differ markedly from conditions in conventional oxidation reactors. When a reactor type bubbler column 20 is used to conduct liquid-phase partial oxidation of para-xylene to technical terephthalic acid (TTFC) according to the preferred embodiments described here, the spatial profile of the local reaction intensity, local evaporation rate and local temperature together with the flow regime of the liquid in the reaction medium and preferably Relatively low oxidation temperatures favor the formation of TTFA particles with unique and beneficial properties.

FIG. 32A and 32B show TTFC base particles prepared in accordance with one embodiment of the present invention. On figa shows the basic particles of TTFC in a 500-fold increase, and figv looks at one of the base particles of TTFC and shows this particle in 2000-fold magnification. As perhaps best illustrated in FIG. 32B, each TTFC base particle is typically composed of a large number of small agglomerated TTFC subparticles, thereby producing TTFC base particles with a relatively high specific surface area, high porosity, low density and good solubility. Unless otherwise indicated, the various properties of TTFC of the invention described below are measured using a representative TTFC sample, the representative sample weighing at least 1 g and / or formed from at least 10,000 individual TTFC particles. TTFC base particles typically have an average size in the range of from about 20 to about 150 microns, more preferably in the range of from about 30 to about 120 microns, and most preferably from 40 to 90 microns. The TTFC subparticles typically have an average size in the range of from about 0.5 to about 30 microns, more preferably from about 1 to about 15 microns, and most preferably from 2 to 5 microns. The relatively high specific surface area of the TTFC base particles illustrated in FIG. 32A and 32B can be quantified using the Brunauer-Emmett-Teller (BET) method for measuring specific surface area. Preferably, the base particles of TTFA have a BET average surface of at least about 0.6 m ² / g. More preferably, the base particles of TTFC have an average BET specific surface area in the range of from about 0.8 to about 4 m ² / g. Most preferably, TTFK base particles have an average BET surface area in the range from 0.9 to 2 m ² / g. Physical properties (e.g., particle size, BET specific surface area, porosity and solubility) of TTFC base particles formed in an optimized oxidation method according to a preferred embodiment of the present invention make it possible to clean TTFC particles in more efficient and / or economical ways, as described in more detail below. in connection with Fig. 35.

The above average particle sizes were determined using polarized light microscopy and image analysis. Particle size analysis equipment includes a Nikon E800 optical microscope with a 4x Plan Flour NA 0.13 lens, a Spot RT ^TM digital camera, and a personal computer with Image Pro Plus ^TM image analysis software, version 4.5.0.19. The particle size analysis method includes the following main steps: (1) dispersing the TTFC powders in mineral oil; (2) preparation of a glass slide / coverslip with dispersion; (3) the study of the drug using microscopy in polarized light (the condition of crossed nicoles - the particles are visible as bright objects on a black background); (4) recording on film different images for each preparation (field size = 3 × 2.25 mm; pixel size = 1.84 μm / pixel); (5) image analysis using Image Pro Plus ^TM ; (6) entering the measurement data for particles into a spreadsheet; and (7) statistical characterization of the table. Stage (5) "image analysis by Image Pro Plus ^TM " includes the following substages: (a) setting an image threshold for detecting white particles against a dark background; (b) creating a binary image; (c) conducting single-pass open filtering to filter out pixel noise; (d) measuring all particles in the image; and (e) reporting the average diameter measured for each particle. Software Image Pro Plus ^TM defines mean diameter of individual particles as the number-average value of particle diameters measured with an interval of 2 ° and with the passage of particles through the center of gravity. Stage 7 "conducting statistical characterization of the spreadsheet" includes calculating the volumetric mass average particle size as follows. The volume of each of the n particles in the sample is calculated as if it were spherical, i.e. pi / 6 * d _i ³ , multiplying the volume of each particle by its diameter to obtain pi / 6 * d _i ⁴ , summing the values pi / 6 * d _i ⁴ for all particles in the sample, summing the volumes of all particles in the sample and calculating the volumetric mass diameter of the particles as the sum of pi / 6 * d _i ⁴ for all n particles in the sample divided by the sum of pi / 6 * d _i ³ for all n particles. As used herein, the term "average particle size" refers to the volumetric average particle size determined according to the test method described above, and is also denoted by D (4.3).

In addition, stage 7 includes finding the particle size for which different fractions of the total volume of the sample are smaller. For example, D (v, 0,1) means the particle size at which 10% of the total sample volume are smaller particles and 90% larger; D (v, 0.5) means the particle size at which half of the sample volume is made up of larger particles and half smaller; D (v, 0.9) means the particle size at which 90% of the total sample volume is made up of smaller particles, etc. In addition, step 7 includes calculating the value of D (v, 0.9) minus D (v, 0.1), which is here defined as a “particle size spread”; and step 7 includes calculating a particle size dispersion value divided by D (4.3), which will be defined here as “relative size dispersion”.

Further, it is preferable that the size D (v, 0.1) of TTFA particles measured above is in the range of from about 5 to about 65 microns, more preferably in the range of from about 15 to about 55 microns, and most preferably in the range of 25 to 45 microns . Preferably, the D (v, 0.5) TTFA particles measured above is 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. Preferably, the above-measured D (v, 0.9) TTFA particles 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. Preferably, the relative dispersion of particle sizes is in the range of from about 0.5 to about 2.0, more preferably from about 0.6 to about 1.5, and most preferably from 0.7 to 1.3.

The BET specific surface values indicated above were measured on a Micromeritics ASAP2000 instrument (manufactured by Micromeritics Instrument Corporation, Norcross, GA). In the first stage of the measuring method, 2-4 g of a sample of particles were weighed and dried in vacuum at 50 ° C. Then the sample was placed in an analytical gas collector and cooled to 77 ° K. The nitrogen adsorption isotherm was measured at a minimum of 5 equilibrium pressures, exposing the sample to the action of known volumes of nitrogen gas and measuring the pressure drop. The equilibrium pressures were suitably in the range 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. Then the obtained isotherm was rebuilt according to the following BET equation:

where V _a is the volume of gas adsorbed on the sample at pressure P, V _m is the volume of gas needed to cover the entire surface of the sample with a gas monolayer, and C is a constant. V _m and C were determined from this graph. Then V _{m was} converted to the specific surface using the nitrogen cross-sectional area at 77 ° K, according to

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

As indicated above, TTFA formed in accordance with one embodiment of the present invention exhibits an increased dissolution ability than conventional TTFC obtained by other methods. This improved dissolution rate allows the purification of TTFA according to the invention with more efficient and / or more effective purification methods. The following description relates to a method for quantifying the rate of dissolution of TTFA.

The dissolution rate of a known amount of solids in a known amount of solvent in a stirred mixture can be measured by various methods. The measurement method used here, called the “dissolution test in time”, is defined as follows. Throughout the dissolution test over time, ambient pressure of approximately 0.1 MPa is used. Throughout the dissolution test over time, an ambient temperature of approximately 22 ° C is applied. In addition, solids, solvent, and all dissolution equipment were completely brought into thermal equilibrium at this temperature prior to the start of the test, and there was no noticeable heating or cooling of the beaker or its contents during the dissolution period. The solvent fraction is fresh, tetrahydrofuran, pure for HPLC analysis (purity> 99.9%), then THF, weighing 250 g, is placed in a clean 400 ml high KIMAX glass beaker (Kimble ^® , component code 14020, Kimble / Kontes, Vineland, NJ), which is non-insulated, with smooth side surfaces and is generally cylindrical in shape. A Teflon-coated magnetic stirrer (VWR, part number 58948-230, about 1 inch long and 3/8 inch diameter, octagonal, manufactured by VWR International, West Chester, PA 19380) is placed in a laboratory beaker, where it is naturally lowered onto bottom. The sample is stirred using a magnetic stirrer Variomag ^® multipoint 15 (H & P Labortechnik AG, Oberschleissheim , Germany) at setting 800 rev / min. This mixing starts no more than 5 minutes before the addition of solids and continues continuously for at least 30 minutes after the addition of the solid phase. A solid sample of technical or purified TPA particles in an amount of 250 mg is weighed in a weighing plate that is not sticky to the sample. At the initial moment, designated t = 0, the suspended solids fall asleep all at the same time into the mixed THF and at the same time turn on the timer. If done correctly, THF very quickly wetts solids and forms a diluted, well-mixed suspension for no more than 5 seconds. Then samples of this mixture are taken 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 is taken from a diluted, well-mixed mixture using a new disposable syringe (Becton, Dickinson & Co, 5 milliliters, REF 30163, Franklin Lakes, NJ 07417). Immediately after sampling from the beaker, approximately 2 ml of the clear liquid sample is quickly passed through a new unused syringe filter (25 mm diameter, 0.45 μm, Gelman GHP Acrodisc GF ^® , Pall Corporation, East Hills, NY 11548) into a new, labeled glass sample tube . The duration of each filling of the syringe, placing the filter and releasing it into the sample tube is, if everything is done correctly, less than about 5 seconds, and this interval appropriately starts and ends within no more than about 3 seconds on each side of each target moment of sampling. Within about five minutes of each filling, the sample tubes are lidded and held at approximately constant temperature until the next chemical analysis. After the last sample was taken at time 30 minutes after t = 0, all sixteen samples were analyzed for the amount of TFA dissolved using HPLC-DAD, usually as described elsewhere in this description. However, in this test both calibration standards and the results are presented in milligrams of dissolved TFA per gram of THF solvent (hereinafter referred to as “ppm in THF”). For example, if all of 250 mg of the solid phase were very pure TFA and if the entire amount was completely dissolved in 250 g of the THF solvent before a specific sample was taken, a correctly measured concentration would be about 1000 ppm. in THF.

When the TTFA of the present invention is subjected to the time dissolution test described above, it is preferred that the 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 TTFC of the present invention is dissolved 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 TTFC of the present invention is dissolved to a concentration of at least about 840 ppm. in THF, more preferably at least up to 880 ppm. in THF.

The inventors of the present invention have found that, despite the complexity of the particle images and the dissolution method, a relatively simple model with negative exponential growth is suitable for describing the time dependence of the entire data set of a complete dissolution test. The form of the equation, hereinafter referred to as the “time dissolution model”, is as follows:

S = A + B * (1-exp (-C * t)),

Where

t = time in minutes;

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

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

A, B = coefficients calculated by regression in units of ppm in THF, where A relates mainly to the rapid dissolution of small particles at very short times, and the sum of A + B refers mainly to the total amount of solute near the end of a certain test duration; and

C = regression-defined time constant in inverse minutes.

The regression coefficients are chosen so as to minimize the sum of the squared errors between the real data and the corresponding model values, this method is usually called the least-squares fit. The preferred software suite for performing this regression data processing is JMP, version 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary, NC 27513).

When the TTFC of the present invention passes the dissolution test in time and is fitted with the time dissolution model described above, it is preferable that the TTFC has a time constant “C” of greater than about 0.5 min ⁻¹ , more preferably greater than about 0.6 min ^-1 , and most preferably more than 0.7 min ^-1 .

In FIG. 33A and 33B show a conventional TTFA particle obtained in a conventional high temperature oxidation method in a continuous mixing reactor (RNND). Fig. 33A shows a conventional TTFC particle at a 500-fold magnification, and Fig. 33B is focused and shows a TTFC particle at a 2000-fold magnification. A visual comparison of the TTFC particles of the invention shown in FIG. 32A and 32B, and the conventional TTFC particle depicted in FIG. 33A and 33B, shows that conventional TTFC particles have a higher density, lower specific surface, lower porosity and larger particle size than the TTFC particles of the invention. In fact, the conventional TTFC shown in FIG. 33A and 33B, has an average particle size of about 205 μm and a BET specific surface area of about 0.57 m ² / g.

On Fig shows a conventional method for producing purified terephthalic acid (CTFA). In a conventional process for producing CTFC, para-xylene is partially oxidized in a high temperature mechanical stirring oxidation reactor 700. The suspension containing TTFC is discharged from the reactor 700 and then cleaned in a treatment system 702. CTFC - the purification product by system 702 is introduced into a separation system 706 to separate and dry the particles of CTFC. The purification system 702 creates a large proportion of the costs associated with obtaining particles of CTFA in conventional ways. The treatment system 702 typically includes a water addition / exchange system 708, a dissolution system 710, a hydrogenation system 712, and three separate crystallization apparatuses 704a, b, c. In the water addition / exchange system 708, a substantial portion of the mother liquor is replaced with water. After adding water, the water / TTFA suspension is introduced into the dissolution system 710, where the water / TTFA mixture is heated until the TTFA particles are completely dissolved in water. After the dissolution of TTFA, the solution of TTFA in water is hydrogenated in a hydrogenation system 712. The hydrogenated stream leaving the hydrogenation system 712 is then subjected to three crystallization stages in crystallization apparatuses 704a, b, c, followed by separation of the CTFC in the separation system 706.

Fig. 35 shows an improved process for preparing CTFA using a bubble column oxidation reactor 800 constructed in accordance with an embodiment of the present invention. The initial suspension containing solid particles of TTFA and a liquid mother liquor is discharged from the reactor 800. Typically, the initial suspension may contain from about 10 to about 50% wt. solid particles of TTFC, with the rest being a liquid mother liquor. The solid particles of TTFA present in the initial suspension typically contain at least about 400 ppm. wt. 4-carboxybenzaldehyde (4-CBA), more typically at least about 800 ppm. wt. 4-CBA, and most typically from 1000 to 15000 ppm. wt. 4-CBA. The initial slurry withdrawn from reactor 800 is introduced into purification system 802 to reduce the concentration of 4-CBA and other impurities present in the TTFA. After purification system 802, a cleaner / purified suspension is obtained which is separated and dried in separation system 804 to thereby 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 from 10 to 200 ppm. wt. 4-CBA.

The cleaning system 802 in the CTFC production system shown in FIG. 35 provides several advantages over the cleaning system 702 in the prior art system shown in FIG. 34. Preferably, the purification system 802 typically includes a fluid change system 806, an autoclave 808, and a single crystallizer 810. In the fluid exchange system 806, at least about 50 wt. The mother liquor present in the initial suspension is replaced with a fresh substitute solvent to thereby obtain a suspension with a replaced solvent containing TTFA particles and a replaced solvent. The suspension with the replaced solvent leaves the fluid exchange system 806 and is introduced into the autoclave (or secondary oxidation reactor) 808. In the autoclave 808, the secondary oxidation reaction is carried out at a slightly higher temperature than that used in the initial / primary oxidation reaction carried out in a bubbler a reactor type 800 column. As discussed above, high specific surface area, low particle size, and low density TTFC particles obtained in reactor 800 cause some impurities to trap data in TTFC particles become available for oxidation in an autoclave 808, without requiring complete dissolution of the TTFC particles in an autoclave 808. Thus, the temperature in an autoclave 808 may be lower than in many similar prior art methods. Secondary oxidation carried out in an autoclave 808 preferably reduces the concentration of 4-CBA in TTFA by at least 200 ppm. wt., more preferably at least about 400 ppm. wt. and most preferably 600-6000 ppm wt. Preferably, the secondary oxidation temperature in the autoclave 808 is at least about 10 ° C higher than the primary oxidation temperature in the bubble column reactor type 800, more preferably about 20-80 ° C higher than the primary oxidation temperature in reactor 800, and most preferably 30-50 ° C higher than the primary oxidation temperature in reactor 800. The secondary oxidation temperature is preferably in the range of about 160 to about 240 ° C, more preferably in the range of about 180 to about 220 ° C and ibolee preferably in the range from 190 to 210 ° C. The purified autoclave product 808 requires only one crystallization step in the crystallizer 810 before separation in the separation system 804. Suitable secondary oxidation / oxidative cooking methods are discussed in more detail in the publication of patent application US 2005/0065373, the full description of which is expressly incorporated herein by reference.

The terephthalic acid (e.g., CTFA) obtained in the system shown in FIG. 35 is preferably formed from particles of CTFA with an average size of at least about 40 microns, more preferably in the range of from about 50 to about 2000 microns, and most preferably in the range of 60 to 200 microns. Particles of CTFA preferably have an average BET specific surface area of less than about 0.25 m ² / g, more preferably from about 0.005 to about 0.2 m ² / g, and most preferably from 0.01 to 0.18 m ² / g. The CTFC obtained in the system shown in FIG. 35 is suitable for use as a raw material for producing PET. Typically, PET is produced by esterification of terephthalic acid with ethylene glycol followed by polycondensation. Preferably, terephthalic acid obtained according to one embodiment of the present invention is used as a feedstock for the method for producing PET in tubular reactors described in patent application US 10/013318 of December 7, 2001, the full description of which is incorporated herein by reference.

TTFA particles with the preferred morphology described herein are particularly suitable for the secondary oxidation process described above to reduce 4-CBA. In addition, these preferred TTFC particles provide advantages for a wide variety of other subsequent processes, including dissolving and / or chemically reacting the particles. These additional post-processes include, without limitation, reacting with at least one hydroxyl-containing compound to form ester compounds, in particular, reacting TTFA with methanol to form dimethyl terephthalate and esters as an impurity, reacting with at least one diol to form ester monomeric and / or polymer compounds, in particular, the reaction of TTFA with ethylene glycol to form polyethylene terephthalate (PET), and complete or partial dissolution in solvents, including without ogre water, acetic acid and N-methyl-2-pyrrolidone, which may include further processing, including, without limitation, re-precipitation of purer terephthalic acid and / or selective chemical reduction of carbonyl groups other than carboxylic acid groups. In particular, this includes the substantial dissolution of TTFA in a solvent containing water, combined with partial hydrogenation, which reduces the amount of aldehydes, especially 4-CBA, fluorenones, phenons and / or anthraquinones.

The authors of the present invention also believe that particles of TTFA having the preferred properties described here can be obtained from particles of TTFC that do not meet the preferred properties described here (inappropriate particles of TTFC) in ways that include, without limitation, mechanical grinding of inappropriate particles of TTFC and full or partial dissolving inappropriate TTFA particles followed by complete or partial re-precipitation.

In accordance with one embodiment of the present invention, there is provided a method of partial oxidation of an oxidizable aromatic compound to aromatic carboxylic acids of one or more types, the purity of the solvent fraction in the feed (i.e., “solvent feed”) and the purity of the fraction of the oxidizable compound in the feed (t ie, the "feed of the oxidizable compound") are controlled within certain ranges defined below. Together with other embodiments of the present invention, this option makes it possible to adjust, in certain preferred ranges indicated below, the purity of the liquid phase, solid phase (if any) and the combined suspension (i.e., solid plus liquid) phase of the reaction medium.

As for the solvent supply, it is known that the oxidizable aromatic compound (s) can be oxidized to produce aromatic carboxylic acid, the solvent feed being introduced into the reaction medium being a mixture of pure acetic acid and water, which is often used on a laboratory scale and on a larger scale in pilot plants. Similarly, it is known to oxidize an oxidizable aromatic compound to aromatic carboxylic acid, in which the solvent leaving the reaction medium is separated from the resulting aromatic carboxylic acid and then returned back to the reaction medium as a solvent supply, mainly for cost reasons. This solvent return over time leads to the accumulation of certain impurities in the feed and process by-products in the recycle solvent. Various means are known in the art for facilitating the purification of a recycle solvent prior to its reintroduction into the reaction medium. In general, an increased degree of purification of the recycle solvent leads to a substantial increase in the cost of production compared to pain by a low degree of purification by similar methods. One embodiment of the present invention relates to understanding and identifying preferred ranges of a large number of impurities in a solvent supply, many of which were previously considered very favorable in order to find the optimal balance between the total cost of production and the purity of the entire product.

"Recycling solvent supply" is defined herein as a solvent supply containing at least about 5% wt. a mass that has previously passed through a reaction medium containing one or more oxidizable aromatic compounds that undergo partial oxidation. For reasons of solvent supply and the operating life of the production unit, it is preferable that the fractions of the recycle solvent pass through the reaction medium at least once per 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 business days. For economic reasons, it is preferable that at least about 20% wt. feeding the solvent into the reaction medium of the present invention was a recycle solvent, more preferably at least about 40% by weight, even more preferably at least about 80% by weight, and most preferably at least 90% by weight.

The inventors of the present invention have found that, due to the activity of the reaction and from the point of view of metallic impurities remaining in the oxidation product, the concentrations of selected metals of variable valency in the feed of the recycle solvent are preferably in the ranges defined immediately below. The concentration of iron in the recycle solvent is preferably below about 150 ppm. wt., more preferably below about 40 ppm. wt. and most preferably is from 0 to 8 ppm. wt. The concentration of nickel in the recycle solvent is preferably below about 150 ppm. wt., more preferably below about 40 ppm. wt. and most preferably is from 0 to 8 ppm. wt. The concentration of chromium in the recycle solvent is preferably below about 150 ppm. wt., more preferably below about 40 ppm. wt., and most preferably ranges from 0 to 8 ppm. wt. The concentration of molybdenum in the recycle solvent is preferably below about 75 ppm. wt., more preferably below about 20 ppm. wt., and most preferably ranges from 0 to 4 ppm. wt. The concentration of titanium in the recycle solvent is preferably below about 75 ppm. wt., more preferably below about 20 ppm. wt., and most preferably ranges from 0 to 4 ppm. wt. The concentration of copper in the recycle solvent is preferably below about 20 ppm. wt., more preferably below about 4 ppm. wt., and most preferably ranges from 0 to 1 ppm. wt. Other metallic impurities are also usually present in the recycle solvent, usually varying at lower levels with respect to one or more of the above metals. Regulation of the contents of the above metals in the preferred ranges will keep other metal impurities at suitable levels.

These metals can form as impurities in any of the feeds entering the feed process (for example, in the input oxidizable compound, solvent, oxidizing agent, and catalytic compounds). Alternatively, metals can form as corrosion products of any of the process units in contact with the reaction medium and / or in contact with the recycle solvent. Means for regulating metals within the described concentration ranges include appropriate specification and monitoring of the purity of various feeds and the use of appropriate materials of construction, including, without limitation, many commercial grades of titanium and stainless steel, including grades known as duplex steel , and high molybdenum stainless steels.

The present inventors have also found preferred ranges of selected aromatic compounds in a recycle solvent. This includes both precipitated and dissolved aromatic compounds in a recycle solvent.

Unexpectedly, it turned out that even the precipitated product (for example, TFA) of the partial oxidation of para-xylene is a contaminant that should be controlled in a recycle solvent. Since it turned out that there are surprisingly preferred ranges of levels of solids in the reaction medium, any precipitated product in the solvent feed is directly subtracted from the amount of oxidizable compound that can be fed together. In addition, it was found that an increased supply of precipitated solid TFA to the recycle solvent negatively affects the properties of particles formed in the precipitating oxidizing medium, which leads to the undesirable nature of subsequent operations (for example, filtering the product, washing the solvent, secondary oxidation of the crude product, dissolution raw product for further processing, etc.). Another undesirable property of precipitated solids in the feed of the recycle solvent is that they often contain very high levels of precipitated impurities compared to the concentration of impurities in the volume of solids in TFA suspensions, from which most of the recycle solvent is obtained. Possibly, these elevated levels of impurities observed in solids suspended in the recycle filtrate may be related to nucleation time for precipitation of certain impurities from the recycle solvent and / or its cooling, it does not matter intentionally or due to heat dissipation into the environment. For example, much higher levels of the concentration of highly colored and undesired 2,6-dicarboxyfluorenone in the solid phase present in the recycle solvent at 80 ° C were observed than was observed in solid TFA isolated from the recycle solvent at 160 ° C. Also, much higher levels of isophthalic acid concentration in the solids present in the recycle solvent were observed compared with levels found in solid TFA from the reaction medium. The exact behavior of certain precipitated impurities entrained in a recycle solvent when reintroduced into the reaction medium is apparently not constant. This depends, possibly, on the relative solubility of the impurity in the liquid phase of the reaction medium, possibly on how the precipitated impurity is deposited in the precipitated solid phase, and, possibly, on the local deposition rate of TFA where the solids re-enter the reaction medium for the first time . Thus, the authors of the present invention have found it useful to control the level of certain impurities in the recycle solvent, as described below, regardless of whether these impurities are in the recycle solvent in dissolved form or entrained therein in the form of particles.

The amount of precipitated solids present in the return filtrate is determined gravimetrically as follows. A representative sample is taken from the solvent supply to the reaction medium as the solvent moves through a pipeline to the reaction medium. A sample of a suitable size weighing about 100 g is placed in a glass container with an internal volume of about 250 ml. Before opening to atmospheric pressure, but with a continuous flow to the sample container, the recycle filtrate is cooled to less than 100 ° C; this cooling is carried out in order to limit the evaporation of the solvent for a short period before sealing it in a glass container. After the sample is placed at atmospheric pressure, the glass container is immediately sealed. The sample is then allowed to cool to about 20 ° C, surrounded by air with a temperature of about 20 ° C and without forced convection. After reaching about 20 ° C, the sample remains under these conditions for at least about 2 hours. Then the sealed container is shaken vigorously until a uniform distribution of solids is obtained. Immediately after this, the core of the magnetic stirrer is placed in the container with the sample, which rotates at a sufficient speed to maintain a virtually uniform distribution of solids. An aliquot of 10 ml of liquid mixed with suspended solids is pipetted and weighed. Then the volume of the liquid phase from this aliquot is separated by vacuum filtration, still at about 20 ° C and practically without loss of solid phase. The wet solid phase, filtered from this aliquot, is then dried, virtually without sublimation of the solid phase, and these solids are weighed. The ratio of the weight of dry solids to the weight of the initial aliquot of the suspension is the fraction of the solid phase, usually expressed as a percentage and referred to herein as the amount of “precipitated solid at 20 ° C.” in the solvent feed.

The inventors of the present invention unexpectedly found that aromatic compounds dissolved in the liquid phase of the reaction medium and containing aromatic carboxylic acids having no non-aromatic hydrocarbon groups (e.g., isophthalic acid, benzoic acid, phthalic acid, 2,5,4'-tricarboxybiphenyl) are harmful components. Although these compounds are much less chemically active in the considered reaction medium compared with oxidizable compounds having non-aromatic hydrocarbon groups, the inventors of the present invention have found that these compounds nevertheless undergo numerous harmful reactions. Thus, it is preferable to control the content of these compounds in the liquid phase of the reaction medium in the preferred ranges. This leads to the preferred ranges of the selected compounds in the feed of the recycle solvent, as well as to the preferred ranges of the precursors in the feed of the oxidizable aromatic compound.

For example, for the liquid phase partial oxidation of para-xylene to terephthalic acid (TFA), the inventors of the present invention found that a strongly colored and undesirable impurity of 2,7-dicarboxyfluorenone (2,7-DCF) was practically impossible to detect in the reaction medium and the selected product when meta -substituted aromatic compounds are in the reaction medium at a very low level. The inventors of the present invention have found that when isophthalic acid impurity is present in the solvent supply at an elevated level, the formation of 2,7-DCF rises almost in direct proportion. The authors also found that when impurity meta-xylene is present in the para-xylene feed, the formation of 2,7-DCF again increases almost in direct proportion. In addition, even if the solvent and oxidizable compound did not contain meta-substituted aromatic compounds, the inventors found that some isophthalic acid is formed during the typical partial oxidation of very pure para-xylene, especially when benzoic acid is present in the liquid phase of the reaction medium. This "self-forming" isophthalic acid can, due to its higher solubility than TFA, in a solvent containing acetic acid and water, form over time in industrial plants using a recycle solvent. Thus, both the amount of isophthalic acid in the feed of the solvent, and the amount of meta-xylene in the feed of the oxidizable aromatic compound, and the rate of spontaneous formation of isophthalic acid in the reaction medium, all must be appropriately considered in balance with each other and in balance with any reaction, which consumes isophthalic acid. It was found that isophthalic acid, in addition to the formation of 2,7-DCF, undergoes the additional reactions described below in which it is consumed. In addition, the authors of the present invention found that there are other problems for discussion when establishing suitable ranges of meta-substituted aromatic compounds in the partial oxidation of para-xylene in TFA. Other highly colored and undesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF), appear to be substantially associated with dissolved para-substituted aromatic compounds, which are always present with the addition of para-xylene to liquid phase oxidation. Thus, lowering the level of 2.7-DCF is best seen in terms of the level of other colored impurities formed.

For example, for the liquid phase partial oxidation of para-xylene in TFA, the inventors of the present invention found that the formation of trimellitic acid increases when isophthalic acid and phthalic acid levels increase in the reaction medium. Trimellitic acid is a trifunctional carboxylic acid that leads to branching of polymer chains in the production of PET from TFA. In many applications of PET, branching must be controlled at low levels and, therefore, trimellitic acid must be controlled so as to be low in purified TPA. In addition to leading to trimellitic acid, the presence of meta-substituted and ortho-substituted compounds in the reaction medium also leads to other tricarboxylic acids (e.g., 1,3,5-tricarboxybenzene). In addition, the increased presence of tricarboxylic acids in the reaction medium increases the amount of tetracarboxylic acid formed (e.g., 1,2,4,5-tetracarboxybenzene). According to the present invention, the regulation of the total formation of all aromatic carboxylic acids having more than two carboxylic acid groups is one of the factors in establishing the preferred levels of meta-substituted and ortho-substituted compounds in the feed of the recycle solvent, in the feed of the oxidizable compound and in the reaction medium.

For example, in the liquid phase partial oxidation of para-xylene in TFA, the inventors of the present invention found that increased levels in the liquid phase of the reaction medium of several dissolved aromatic carboxylic acids that do not contain non-aromatic hydrocarbon groups directly lead to increased formation of carbon monoxide and carbon dioxide. This increased formation of carbon oxides is expressed by the loss of the yield of both the oxidizing agent and the oxidizable compound, the latter due to the fact that many of the aromatic side carboxylic acids, which on the one hand can be considered as impurities, on the other hand, also have market value. Thus, the corresponding removal of relatively soluble carboxylic acids that do not contain non-aromatic hydrocarbon groups from solvent recycling has, in addition to suppressing the formation of highly undesirable impurities, such as various fluorenones and trimellitic acid, economic importance to prevent a decrease in the yield of oxidizable aromatic compounds and oxidizing agents.

For example, in the liquid phase partial oxidation of para-xylene in TFA, the present inventors have found that the formation of 2,5,4'-tricarboxybiphenyl is apparently inevitable. 2,5,4'-Tricarboxybiphenyl is an aromatic tricarboxylic acid formed by combining two aromatic rings, possibly combining dissolved para-substituted aromatic components with an aryl radical, possibly an aryl radical formed by decarboxylation or decarbonylation of para-substituted aromatic compounds. Fortunately, 2,5,4'-tricarboxybiphenyl is usually formed at a much lower level than trimellitic acid, and does not always lead to a significant increase in the difficulty of branching polymer molecules in the production of PET. However, the inventors of the present invention have found that elevated levels of 2,5,4'-tricarboxybiphenyl in a reaction medium including oxidation of aromatic alkyls according to preferred embodiments of the present invention lead to elevated levels of highly colored and undesired 2,6-DCF. An increased level of 2,6-DCF is formed, possibly, by cyclization of 2,5,4'-tricarboxybiphenyl with the loss of a water molecule, although the exact reaction mechanism is not known with certainty. If too strong a formation of 2,5,4'-tricarboxybiphenyl in the recirculating solvent, which is better than TFA, is soluble in the solvent containing acetic acid and water, the conversion rates to 2,6-DCF may become unacceptably high.

For example, in the liquid phase partial oxidation of para-xylene in TFA, the inventors of the present invention found that aromatic carboxylic acids that do not contain non-aromatic hydrocarbon groups (e.g., isophthalic acid) usually result in mild suppression of the chemical activity of the reaction medium when they are present in the liquid phase in a sufficient concentration.

For example, in the liquid phase partial oxidation of para-xylene in TFA, the inventors of the present invention found that precipitation very often is imperfect (i.e., nonequilibrium) with respect to the relative concentrations of various chemical components in the solid phase and in the liquid phase. Perhaps this is due to the fact that the deposition rate at the spatio-temporal reaction rates preferred here is very fast, which leads to imperfect coprecipitation of impurities or even to occlusion. Thus, if it is desirable to limit the concentration of certain impurities (for example, trimellitic acid and 2,6-DCF) in technical TPA, due to the structure of the work of the plants following the scheme, it is preferable to control their concentration in the solvent supply, as well as the rate of their formation in the reaction environment.

For example, the inventors of the present invention have found that benzophenone compounds (e.g., 4,4'-dicarboxybenzophenone and 2,5,4'-tricarboxybenzophenone) obtained by partial oxidation of para-xylene have undesirable effects on the PET reaction medium, even if the benzophenone compounds by themselves, are not so strongly colored in TFA as fluorenones and anthraquinones. Accordingly, it is desirable to limit the presence of benzophenones and to select precursors in the recycle solvent and in the supply of the oxidizable compound. In addition, the authors found that the presence of elevated levels of benzoic acid, still introduced into the recycle solvent or formed in the reaction medium, leads to increased rates of formation of 4,4'-dicarboxybenzophenone.

By analyzing, the authors of the present invention discovered and sufficiently quantified the unexpected many reactions of aromatic compounds that do not contain non-aromatic hydrocarbon groups that are present in the liquid-phase partial oxidation of para-xylene in TFA. Considering just one case of benzoic acid, the inventors of the present invention found that elevated levels of benzoic acid in the reaction medium in certain embodiments of the present invention lead to a large increase in the formation of highly colored and undesired 9-fluorenone-2-carboxylic acid, to very elevated levels of 4, 4'-dicarboxybiphenyl, to elevated levels of 4,4'-dicarboxybenzophenone, to a weak suppression of the chemical activity of the target oxidation of para-xylene, to elevated levels of carbon dioxide as well as associated output losses. The inventors of the present invention have found that elevated levels of benzoic acid in the reaction medium also lead to increased formation of isophthalic acid and phthalic acid, the levels of which, according to related aspects of the present invention, it is desirable to control in narrow ranges. The number and importance of the reactions in which benzoic acid is involved is perhaps even more surprising, as some authors of recent inventions propose using benzoic acid instead of acetic acid as the main component of the solvent (see, for example, US Pat. No. 6,562,997). In addition, the authors of the present invention found that benzoic acid spontaneously forms during the oxidation of para-xylene at rates that are quite comparable to its formation from impurities such as toluene and ethylbenzene, which are usually found in the supply of an oxidizable compound containing technically pure para-xylene.

On the other hand, the inventors of the present invention have found little benefit from further adjusting the composition of the recycle solvent with respect to the presence of an oxidizable aromatic compound and with respect to aromatic intermediates of the reaction, both of which retain non-aromatic hydrocarbon groups and are also relatively soluble in the recycle solvent. In general, in order to appropriately limit their formation in a recycle solvent, these compounds must be supplied or formed in the reaction medium to a degree substantially greater than their presence in the recycle solvent; and the rate of consumption of these compounds in the reaction medium is quite high, while maintaining one or more non-aromatic hydrocarbon groups. For example, in the partial oxidation of para-xylene in a multiphase reaction medium, para-xylene 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 is condensed to return as a recycle solvent, a substantial part of the vaporized para-xylene also condenses there. There is no need to limit the concentration of this para-xylene in the recycle solvent. For example, if the solvent is separated from the solid phase after the suspension leaves the para-xylene oxidation reaction medium, this returned solvent will contain a concentration of dissolved para-toluic acid close to that at the point of removal from the reaction medium. Although it may be important to limit the stationary concentration of para-toluic acid in the liquid phase of the reaction medium (see below), it is not necessary to separately regulate para-toluic acid in this part of the recycle solvent because of its relatively good solubility and because of its low mass velocity flow compared with the formation of para-toluic acid in the reaction medium. Similarly, the inventors of the present invention found no particular reason to limit the concentration 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) in the recirculating solvent, and brominated aromatic compounds retaining at least one non-aromatic hydrocarbon group (e.g., alpha-bromo-para-toluic acid) below conc ntratsy naturally found in the liquid phase exiting from the reaction medium occurring in the partial oxidation of xylene according to preferred embodiments of the present invention. Unexpectedly, the inventors of the present invention also found that it is also not necessary to control the concentration of selected phenols in the recycle solvent, which are inherently formed during the partial oxidation of xylene, since these compounds form and decompose in the reaction medium at a level significantly exceeding their presence in the recycle solvent. For example, the inventors of the present invention have found that in preferred embodiments of the present invention, 4-hydroxybenzoic acid has a relatively small effect on reactivity when co-fed at a level of more than 2 g of 4-hydroxybenzoic acid per 1 kg of para-xylene, i.e. much larger than the natural presence in the recycle solvent, although others report it as a substantially dangerous compound of such a reaction medium (see, for example, W. Partenheimer, Catalysis Today 23 (1995) p. 81).

Thus, there are many reactions and many considerations for establishing preferred ranges for various aromatic impurities in the solvent supply that are currently being described. These findings are formulated in terms of the total weight-average composition of all solvent streams supplied to the reaction medium over a set period of time, preferably one day, more preferably one hour and most preferably one minute. For example, if the solvent is supplied substantially continuously at a composition of 40 ppm. wt. isophthalic acid with a flow rate of 7 kg / min, the second solvent supply is made essentially continuously at a composition of 2000 ppm wt. isophthalic acid with a flow rate of 10 kg / min, and there are no other solvent feed streams entering the reaction medium, the total mass-averaged solvent feed composition is calculated as (40 * 7 + 2000 * 10) / (7 + 10) = 1193 ppm wt. isophthalic acid. It should be noted that the weight of any supply of the oxidizable compound or any supply of oxidizing agent, which may be mixed with the supply of solvent before introduction into the reaction medium, is not taken into account when calculating the total weight-average composition of the supply of solvent.

Table 1 below lists the preferred values for some of the components in the solvent feed introduced into the reaction medium. Table 1 lists the following solvent feed components: 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), all fluorenones, including other fluorenones not listed separately (all fluorenones), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'- tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA), isophthalic acid (IPA), benzoic acid (BA), trimell itic 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 (TFA), solids precipitated at 20 ° C, and all aromatic carboxylic acids that do not contain non-aromatic hydrocarbon groups. Table 1 below lists the preferred amounts of these impurities in TTFC obtained according to an embodiment of the present invention.

Table 1 The components of the feed solvent introduced into the reaction medium Component Designation Preferred range (ppm wt.) More preferred range (ppm wt.) The most preferred range (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 All fluorenones <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 TFK <9000 200-6000 400-2000 Solids precipitated at 20 ° C <9000 200-6000 600-2000 All aromatic carboxylic acids without non-aromatic hydrocarbon groups <18000 300-9000 450-3000

Many other aromatic impurities are usually also present in the recycle solvent, usually varying at even lower levels and / or in proportions to one or more of the aromatic compounds described. Methods of controlling the described aromatic compounds in preferred ranges will typically keep other aromatic impurities at suitable levels.

It is known that when bromine is used in the reaction medium, a large number of ionic and organic forms of bromine are in dynamic equilibrium. These various forms of bromine have different stability characteristics after exiting the reaction medium and passing through various standard methods related to the recycle solvent. For example, alpha-bromo-para-toluic acid may persist as is under certain conditions or may rapidly hydrolyze under other conditions to form 4-hydroxymethylbenzoic acid and hydrogen bromide. In the present invention, it is preferable that at least about 40% by weight, more preferably, at least about 60% by weight, and most preferably, at least about 80% by weight. of the total mass of bromine present in the total solvent supply to the reaction medium, was in one or more of the following chemical forms: ionic bromine, alpha-bromo-para-toluic acid, and bromoacetic acid.

Although the importance and importance of controlling the total weight-average purity of the solvent in the ranges described above, desirable according to the present invention, have not yet been discovered and / or described, suitable means for controlling the purity of the solvent can be combined from various methods already known in the art. First, any solvent that has evaporated from the reaction medium usually has a suitable purity, provided that the liquid or solids from the reaction medium are not entrained with the evaporating solvent. The supply of droplets of solvent reflux to the off-gas outlet space above the reaction medium, as described herein, appropriately limits such entrainment, and the recycle solvent of an aromatically suitable purity can condense from such off-gas. Secondly, the more difficult and expensive cleaning of the recycle solvent feed usually relates to a solvent taken from the reaction medium in liquid form and to a solvent which is then contacted with the liquid and / or solid phases of the reaction medium withdrawn from the reaction apparatus (e.g., recycle a solvent obtained from a filter in which solids are concentrated and / or washed; a recirculating solvent obtained from a centrifuge in which solids are concentrated and / or washed; recirculation ruyuschy solvent obtained from the crystallizing operation, etc.). However, means are also known in the art for performing the necessary purification of these recycle solvent streams using one or more of the previous inventions. With regard to the retention of precipitated solids in a recirculating solvent in certain predetermined ranges, suitable control means include, without limitation, gravimetric sedimentation, mechanical filtration using filter cloth on rotary belt filters and rotary drum filters, mechanical filtration using a stationary filter medium in high-pressure apparatuses pressure, hydrocyclones and centrifuges. With regard to controlling the level of dissolved aromatic components in the recycle solvent in certain predetermined ranges, control means include, without limitation, those described in US Pat. No. 4,939,297 and the publication of US Patent Application No. 2005-0038288, incorporated herein by reference. However, none of these early inventions disclosed or described preferred purity levels of the total solvent supply, as disclosed herein. On the contrary, these earlier inventions merely provided means for purifying selected partial recycle solvent streams without deriving the optimal composition according to the present invention of the total weighted average supply of solvent to the reaction medium.

Now, with regard to the purity of supply of the oxidizable compound, it is known that certain levels of isophthalic acid, phthalic acid and benzoic acid are present in the purified TFA used to prepare the polymer. In addition, it is known that these components are relatively more soluble in many solvents and can be effectively removed from purified TFA in crystallization processes. However, it is now known from one embodiment of the invention described here that controlling the level of several relatively soluble aromatic components, including, in particular, isophthalic acid, phthalic acid and benzoic acid, in the liquid phase of the reaction medium unexpectedly turned out to be important for controlling the level of polycyclic and colored aromatic compounds formed in the reaction medium to control compounds with more than two functional groups of carboxylic acid per molecule, for re control the activity of the reaction in the partial oxidation reaction medium and to control the yield loss of the oxidizing agent and aromatic compounds.

It is known in the art that isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium as follows. The impurity meta-xylene present in the feed oxidizes with good conversion and gives IPA. The impurity ortho-xylene present in the feed is oxidized with good conversion and gives phthalic acid. The impurity ethylbenzene and toluene present in the feed are oxidized with good conversion and give benzoic acid. However, the inventors of the present invention have found that significant amounts of isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium containing para-xylene also in other ways than the oxidation of meta-xylene, ortho-xylene, ethylbenzene and toluene. These other internal chemical routes may include decarbonylation, decarboxylation, rearrangement of transition states, and the addition of methyl and carbonyl radicals to aromatic rings.

In determining the preferred ranges of impurities in the supply of the oxidizable compound, many factors are significant. Obviously, any impurity in the feed will result in a direct loss of yield and an increase in the cost of cleaning the product if the purity requirements for the oxidized product are quite stringent (for example, in the reaction medium for the partial oxidation of para-xylene, toluene and ethylbenzene, usually found in technically pure para-xylene lead to benzoic acid, and this benzoic acid is overwhelmingly removed from most of the industrial TFA). When the product of partial oxidation of an impurity in a raw material is involved in additional reactions, other factors become decisive, not just a decrease in yield and removal, when it is taken into account how much the cleaning of the raw material will take (for example, in a reaction medium for partial oxidation of para-xylene, ethylbenzene leads to benzoic acid and benzoic acid subsequently leads to highly colored 9-fluorenone-2-carboxylic acid, isophthalic acid, phthalic acid and, among other things, to an increased amount of carbon oxides). When the reaction medium itself forms additional amounts of impurities by chemical mechanisms associated with impurities not supplied directly, the analysis becomes even more complicated (for example, in the reaction medium for the partial oxidation of para-xylene, benzoic acid also spontaneously forms from para-xylene itself) . In addition, subsequent processing of the crude oxidation product may affect considerations regarding the preferred purity of the feed. For example, the cost of removal to an appropriate level of direct impurity (benzoic acid) and subsequent impurities (isophthalic acid, phthalic acid, 9-fluorenone-2-carboxylic acid, etc.) may be the same, may differ from each other, and may differ from requirements for the removal of a completely independent impurity (for example, the product of incomplete oxidation of 4-CBA during the oxidation of para-xylene in TFA).

The purity ranges described below for para-xylene are preferred where para-xylene is fed into the reaction medium together with a solvent and an oxidizing agent for partial oxidation to produce TFA. These ranges are more preferable for a process for producing TFA, in which, after oxidation, there are steps for removing impurities other than an oxidizing agent and a solvent (e.g., catalyst metals) from the reaction medium. These ranges are even more preferable for methods for producing TFA in which additional 4-CBA is removed from TTFA (for example, by converting TTFA to dimethyl terephthalate plus impurity esters and subsequent separation of 4-CBA methyl ester by distillation, by secondary oxidation methods to convert 4-CBA in TFA, by hydrogenation methods to convert 4-CBA to para-toluic acid, which is then separated by partial crystallization methods). These ranges are most preferred for TFA production processes in which additional 4-CBA is removed from TTPA by a secondary oxidation process to convert 4-CBA to TFA.

Using new knowledge about the preferred ranges of recycle aromatic compounds and the relative amounts of aromatic compounds formed directly by the oxidation of the impurities in the feed compared to other internal chemical routes, improved ranges for the impurities in the contaminated para-xylene supplied in the partial oxidation process to obtain TFK. Table 2 below shows the preferred amounts of meta-xylene, ortho-xylene and a mixture of ethylbenzene + toluene in the feed of para-xylene, expressed in wt. parts per million parts of para-xylene.

table 2 Components in the supply of impurity para-xylene Component Designation Preferred range (ppm wt.) More preferred range (ppm wt.) The most preferred range (ppm wt.) Meta-Xylene 20-800 50-600 100-400 Ortho-xylene 10-300 20-200 30-100 Ethylbenzene + toluene * 20-700 50-500 100-300 Total 50-900 100-800 200-700 * The specification for the mixture of ethylbenzene + toluene applies to each individually and to the sum

Specialists in this field should understand that the above impurities in contaminated para-xylene can have the greatest effect on the reaction medium after the products of their partial oxidation accumulate in the recycle solvent. For example, the introduction of an upper amount from the most preferred range of meta-xylene, 400 ppm. wt., immediately forms in the liquid phase of the reaction medium approximately 200 ppm wt. isophthalic acid when working with approximately 33% by weight of solid phase in the reaction medium. This should be compared to entering the upper amount from the most preferred range, 400 ppm. wt., isophthalic acid into a recycle solvent, which after a typical evaporation of the solvent to cool the reaction medium will be approximately 1200 ppm. wt. isophthalic acid in the liquid phase of the reaction medium. Thus, it is the accumulation of partial oxidation products in the recycle solvent over time that represents the most likely effect of impurity meta-xylene, ortho-xylene, ethylbenzene and toluene in the supply of contaminated para-xylene. Accordingly, the above ranges of impurities in the contaminated para-xylene feed should preferably be kept at least half of each day of the partial oxidation reaction medium in a particular plant, more preferably at least three quarters of each day for at least seven consecutive working days, and most preferably, the weighted average feed composition of the crude para-xylene is within the preferred ranges for at least 30 consecutive working days.

Means for preparing para-xylene impurities of preferred purity are already known in the art and include, but are not limited to, rectification, partial crystallization methods at low temperatures and molecular sieve methods that use selective adsorption, depending on pore size. However, the preferred purity ranges indicated here are, in their upper values, more stringent and more expensive than those commonly observed by industry suppliers of para-xylene, and in the lower values of the preferred ranges, overly expensive purification of para-xylene to feed the partial oxidation reaction medium becomes unnecessary due to the discovery and description of when the combined effects of impurity formation from para-xylene itself and reactions in which the impurity is consumed in the reaction medium become more important than levels of impurities entering para-xylene polluted.

When the xylene-containing feed stream contains selective impurities such as ethylbenzene and / or toluene, benzoic acid may form during the oxidation of these impurities. As used herein, the term “impurity benzoic acid” means benzoic acid obtained from any source other than xylene during the oxidation of xylene.

As disclosed herein, a portion of the benzoic acid formed during the oxidation of xylene is obtained from xylene itself. This formation of benzoic acid makes up a significant portion of any fraction of benzoic acid formed, which may be impurity benzoic acid. Not wishing to become attached to theory, it is believed that benzoic acid is obtained from xylene in a reaction medium when various xylene oxidation intermediates are spontaneously decarbonylated (loss of carbon monoxide) or decarboxylated (loss of carbon dioxide), thereby giving aryl radicals. These aryl radicals can then cleave the hydrogen atom from one of the many available sources in the reaction medium and produce self-generated benzoic acid. Whatever the chemical mechanism, the term “self-forming benzoic acid” as used herein means benzoic acid derived from xylene from the oxidation of xylene.

Also, as disclosed herein, when para-xylene is oxidized to produce terephthalic acid (TFA), the formation of self-generated benzoic acid causes a decrease in the yield of para-xylene and a decrease in the yield of oxidizing agent. In addition, the presence of self-forming benzoic acid in the liquid phase of the reaction medium correlates with the enhancement of many undesirable side reactions, including, in particular, the formation of highly colored compounds called mono-carboxyfluorenones. The self-forming benzoic acid also contributes to the undesirable accumulation of benzoic acid in the recycle filtrate, which further increases the concentration of benzoic acid in the liquid phase of the reaction medium. Thus, the formation of self-forming benzoic acid is desirable to minimize, but this should also be considered simultaneously with impurities of benzoic acid, with factors affecting the consumption of benzoic acid, with factors related to other problems of reaction selectivity, and overall economics.

The inventors have found that the self-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. Not wishing to become attached to theory, it is believed that lower temperatures and improved oxygen availability suppress decarbonylation and / or decarboxylation rates, thereby avoiding the problems of lowering yield due to self-generated benzoic acid. The sufficient availability of oxygen, apparently, directs aryl radicals to other, more favorable products, in particular, hydroxybenzoic acids. The distribution of xylene in the reaction medium may also affect the balance between the conversion of aryl radicals to benzoic acid or hydroxybenzoic acids. Whatever the chemical mechanisms, the present inventors have found reaction conditions that, being mild enough to reduce the formation of benzoic acid, are strict enough to oxidize a large part of the obtained 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 designed and operated so as to minimize the formation of self-generated benzoic acid and maximize the oxidation of hydroxybenzoic acids to carbon monoxide and / or carbon dioxide. When the oxidation reactor is used to oxidize para-xylene to terephthalic acid, it is preferred that the para-xylene comprises at least about 50% by weight. all xylene in the feed stream introduced into the reactor. More preferably, para-xylene is at least about 75% wt. from all xylene in the feed stream. Even more preferably, para-xylene is at least 95% wt. from all xylene in the feed stream. Most preferably, para-xylene comprises substantially all of the xylene in the feed stream.

When the reactor is used to oxidize para-xylene to terephthalic acid, it is preferable that the rate of formation of terephthalic acid is maximum and the rate of formation of self-forming benzoic acid is minimal. Preferably, the ratio of the rate of formation (by weight) of terephthalic acid to the rate of formation (by weight) of self-generated 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 self-forming 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., since these low concentrations inhibit reactions that convert benzoic acid to other compounds to suitable low levels.

If we combine the self-forming benzoic acid and the benzoic acid formed by the impurities, the ratio of the rate of formation (by weight) of terephthalic acid to the rate of formation (by weight) of the entire (self-forming and formed by the impurities) benzoic acid 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 self-forming benzoic acid plus impurities of benzoic acid is measured preferably when the concentration of benzoic acid in the liquid phase of the reaction medium is below 500 ppm. wt., since these low concentrations inhibit reactions that convert benzoic acid to other compounds to suitable low levels.

As disclosed herein, increased concentrations of benzoic acid in the liquid phase of the reaction medium lead to increased formation of many other aromatic compounds, some of which are harmful impurities in TFA, and, as described here, increased concentrations of benzoic acid in the liquid phase of the reaction medium lead to increased formation gaseous carbon oxides, the formation of which is expressed in a decrease in the yield of oxidizing agent and aromatic compounds and / or solvent. In addition, it is further described that the authors of the present invention found that a significant proportion of this increased formation of other aromatic compounds and carbon oxides is obtained in reactions in which part of the molecules of benzoic acid itself are consumed, in contrast to other reactions that benzoic acid catalyzes, they are not consumed. Accordingly, the “total benzoic acid formation" is defined here as the time-averaged weight of all benzoic acid leaving the reaction medium, minus the time-averaged weight of all benzoic acid entering the reaction medium over the same period of time. This resulting formation of benzoic acid is often positive, due to the rates of formation of self-forming benzoic acid and benzoic acid formed by impurities. However, the inventors of the present invention found that the rate of conversion of benzoic acid to carbon oxides and some other compounds appears to increase approximately linearly with an increase in the concentration of benzoic acid in the liquid phase of the reaction medium, measured when other reaction conditions, including temperature, oxygen availability, STR and reaction activity are accordingly kept constant. Thus, when the concentration of benzoic acid in the liquid-phase reaction medium is high enough, possibly due to the increased concentration of benzoic acid in the recycle solvent, the degree of conversion of benzoic acid molecules to other compounds, including carbon oxides, can become equal to or greater than the formation of new benzoic acid molecules by chemical reactions. In this case, the total formation of benzoic acid can be about zero or even be negative. The inventors of the present invention have found that when the resulting formation of benzoic acid is positive, the ratio of the rate of formation (by weight) of terephthalic acid in the reaction medium to the rate of the final formation of benzoic acid in the reaction medium is preferably higher than about 700: 1, more preferably higher than about 1100: 1, and most preferably above 4000: 1. The authors found that when the resulting formation of benzoic acid is negative, the ratio of the rate of formation (by weight) of terephthalic acid in the reaction medium to the rate of the resulting formation of benzoic acid in the reaction medium is preferably more than about 200: (- 1), more preferably more than about 1000 :( -1), and most preferably more than 5000: (- 1).

The present inventors have also found preferred ranges for the composition of the suspension (liquid + solid phase) discharged from the reaction medium and for the solid TTFA fraction in the suspension. Preferred suspension and TTFA formulations have surprisingly been found to be better and more advantageous. For example, purified TPA obtained from this preferred TTFC by secondary oxidation has a sufficiently low level of all 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 of PET fibers and for PET applications in packages. For example, a preferred suspension formulation gives the liquid phase of the reaction medium, which has a relatively low concentration of basic impurities, and this substantially reduces the formation of other, even more undesirable impurities than those described herein. In addition, the preferred suspension composition greatly facilitates the subsequent processing of the liquid from the suspension so that it becomes a sufficiently pure recycle solvent, according to other embodiments of the present invention.

TTFC obtained in one embodiment of the present invention contains fewer impurities of the selected type than TTFC obtained by conventional methods and conventional apparatuses, in particular those using a recirculating solvent. Impurities that may be present in TTFA 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), all fluorenones, including other fluorenones not listed separately (all fluorenones), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4 '-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA), isophthalic acid (IPA), benzoic acid (BA), trimellit hydrochloric acid (TMA), para-toluic acid (PTAC), 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). Table 3 below lists the preferred amounts of these impurities in TTFA obtained in one embodiment of the present invention.

Table 3 Impurities in TTFC Component Designation Preferred range (ppm wt.) More preferred range (ppm wt.) The most preferred range (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 All fluorenones <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 TTFC obtained according to one embodiment of the present invention has a less intense color than TTFC obtained by conventional methods and conventional equipment, in particular using a recycle solvent. Thus, it is preferable that the TTFC obtained in accordance with one embodiment of the present invention has a relative transmittance at 340 nm of at least about 25%, more preferably at least about 50%, and most preferably at least 60%. In addition, it is preferable that the TTFC obtained in one embodiment of the present invention has a transmittance at 400 nm of at least about 88%, more preferably at least about 90%, and most preferably at least 92%.

The transmittance test gives a measure of the colored, readily absorbing impurities present in TFA or TTFA. This test, as used here, refers to measurements taken on a portion of a solution prepared by dissolving 2.00 g of dry solid TFA or TTFA in 20.0 ml of dimethyl sulfoxide (DMSO), pure for analysis or better. Then part of this solution is placed in a Hellma PN 176.700 flow microcuvette, which is made of quartz and has an optical path 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 (Agilent Technologies, 395 Page Mill Road, Palo Alto, CA 94303) is used to measure the transmission of light of different wavelengths through this filled flow cell. After appropriate correction for background absorption, including, but not limited to, the cell and solvent used, the transmittance results characterizing the fraction of incident light that passed through the solution are immediately reported by the machine. The transmittance values at wavelengths of 340 nm and 400 nm are especially useful for distinguishing pure TFA from many impurities typically found in it.

Preferred ranges of various aromatic impurities in the suspension (solid + liquid) phase of the reaction medium are given below in table 4.

Table 4 Impurities in suspension Component Designation Preferred range
(ppm wt.) More preferred range
(ppm wt.) Most preferred range
(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 All fluorenones <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 represent the preferred composition of the liquid phase of the reaction medium, in fact allowing to circumvent experimental difficulties associated with the precipitation of additional components of the liquid phase from the reaction medium into solid phase components during sampling from the reaction medium, separation of liquids and solids, and the transition to analysis conditions .

In the suspension phase of the reaction medium and in the TTFA of the reaction medium, many other aromatic impurities are usually also present, the contents of which usually fluctuate at even lower levels and / or are in relation to one or more of the described aromatic compounds. Control of the described aromatic compounds in preferred ranges will typically keep other aromatic impurities at suitable levels. These favorable compositions of the suspension phase in the reaction medium and in solid TTFA taken directly from the suspension are resolved in accordance with embodiments of the invention described here for the partial oxidation of para-xylene in TFA.

Measurement of the concentrations of components present in small amounts in a solvent, a recycle solvent, TTFC, suspension from the reaction medium and CTFC is carried out using liquid chromatography methods. The following describes two interchangeable implementation options.

The method referred to here as HPLC-DAD includes high performance liquid chromatography (HPLC) in combination with a diode array detector (DAD) to allow separation and quantitation of various molecular components in a given sample. The instrument used in this measurement is a Model 1100 HPLC instrument supplied with DAD supplied by Agilent Technologies (Palo Alto, CA), although other suitable equipment is also 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, the compounds and amounts corresponding to those in real unknown samples.

A method, referred to herein as HPLC-MS, involves high performance liquid chromatography (HPLC) in combination with mass spectrometry (MS) to allow separation, identification and quantification of various molecular components in a given sample. The instrument used in this measurement is an Alliance HPLC instrument and ZQ MS from Waters Corp. (Milford, MA), although other suitable equipment is also available from other suppliers. As is known in the art, both the elution time and the mass spectrometric response are calibrated using known compounds present in known amounts, and these compounds and amounts are suitable for those found in real unknown samples.

Another embodiment of the present invention relates to the partial oxidation of an aromatic oxidizable compound with a proper balance, on the one hand, of suppressing harmful aromatic impurities, and the formation of carbon dioxide and carbon monoxide, together carbon oxides (COx), on the other hand. These carbon oxides typically leave the reaction apparatus in the off-gas, and they are considered to be a destructive loss of solvent and oxidizable compound, including very preferred oxidized derivatives (e.g., acetic acid, para-xylene and TFA). The authors of the present invention have identified lower limits for the production of carbon oxides, below which a high formation of harmful aromatic impurities, as described below, and a low degree of conversion, apparently, are economically unjustified. The authors of the present invention also determined the upper limits of carbon oxides, above which the formation of carbon oxides continues to increase slightly due to a decrease in the formation of harmful aromatic impurities.

The inventors of the present invention have found that a decrease in the liquid-phase concentrations of the aromatic oxidizable compound and intermediate aromatic compounds in the reaction medium leads to lower formation rates of harmful impurities during partial oxidation of the aromatic oxidizable compound. These detrimental impurities include bound aromatic rings and / or aromatic molecules containing more than desirable carboxylic acid groups (e.g., when para-xylene is oxidized, detrimental impurities include 2,6-dicarboxyanthraquinone, 2,6-dicarboxyfluorenone, trimellitic acid, 2 , 5,4'-tricarboxybiphenyl and 2,5,4'-benzophenone). Intermediate aromatic compounds include aromatic compounds formed from the feed of an oxidizable aromatic compound and still retaining non-aromatic hydrocarbon groups (for example, in the oxidation of para-xylene, intermediate aromatic compounds include para-tolualdehyde, terephthaldehyde, para-toluic acid, 4-CBA, 4-hydroxymethylbenzoic acid and alpha-bromo-para-toluic acid). The supply of aromatic oxidizable compounds and intermediate aromatic compounds with non-aromatic hydrocarbon groups, if present in the liquid phase of the reaction medium, leads to harmful impurities for reasons apparently similar to those already described here for dissolved aromatic compounds that do not contain non-aromatic hydrocarbon groups (for example, for isophthalic acid).

Faced with the need for higher reaction activity to suppress the formation of harmful aromatic impurities during partial oxidation of an oxidizable aromatic compound, the present inventors have found that an undesirable concomitant result is increased formation of carbon oxides. It is important to understand that these carbon oxides mean a decrease in the yield of the oxidizable compound and oxidizing agent, and not just the solvent. It is clear that a substantial and sometimes the main proportion of carbon oxides appears from the oxidizable compound and its derivatives, and not from the solvent, and often the contribution of the oxidizable compound is one more carbon than the solvent. In addition, it is important to understand that the desired marketable carboxylic acid (e.g., TFA) also undergoes oxidation to carbon oxides when present in the liquid phase of the reaction medium.

It is also important to understand that the present invention relates to reactions in the liquid phase of the reaction medium and to the concentration of reactants in it. In this, it differs from some earlier inventions, which relate directly to the formation in the precipitated solid form of an aromatic compound containing non-aromatic hydrocarbon groups. With particular regard to the partial oxidation of para-xylene in TFA, certain prior inventions relate to the amount of 4-CBA precipitated in the solid phase of TFA. However, the authors of the present invention, using the same conditions in temperature, pressure, catalysis, solvent composition and the spatiotemporal rate of the para-xylene reaction, found a difference of more than two to one for the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase depending on whether partial oxidation is carried out in an autoclave with good mixing or in a reaction medium with a gradual change in the concentrations of oxygen and para-xylene, according to the present invention. Further, the authors of the present invention found that the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase can also change more than two to one in a well-mixed reaction medium and in a medium with gradually changing concentrations depending on the spatio-temporal reaction rate para-xylene under similar remaining conditions in temperature, pressure, catalysis and solvent composition. In addition, 4-CBA in solid phase TTFA does not appear to contribute to the formation of harmful impurities, and 4-CBA in the solid phase can be isolated and oxidized to TFA simply and in high yield (for example, by secondary oxidation of a suspension of TTFA, as described here); whereas the removal of harmful impurities is much more difficult and costly than the removal of solid phase 4-CBA, and the formation of carbon oxides means a continuous decrease in yield. Thus, it is important to highlight that this aspect of the present invention relates to liquid phase compositions in a reaction medium.

The inventors of the present invention found that the formation of carbon oxides with industrially important conversions, regardless of whether their source is a solvent or an oxidizable compound, is strictly related to the level of activity of the entire reaction, despite the variety of particular combinations of temperature, metals, halogens, temperature, acidity the reaction medium, measured by pH, the concentration of water used to obtain this level of activity of the total reaction. The authors of the present invention have found it useful for the partial oxidation of xylene to assess the level of activity of the overall reaction using the concentration of toluic acids in the liquid phase at the middle of the height of the reaction medium, at the bottom of the reaction medium and at the top of the reaction medium.

Thus, it is important to simultaneously balance the reduction to a minimum of the formation of harmful impurities by increasing the activity of the reaction and, in addition, minimizing the formation of carbon oxides by reducing the activity of the reaction. That is, if the complete formation of carbon oxides is suppressed too weakly, 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., TFA) and the presence of other dissolved aromatic compounds that do not contain non-aromatic hydrocarbon groups are a very important core of this balancing of carbon oxides with respect to harmful impurities. The desired marketable carboxylic acid usually dissolves in the liquid phase of the reaction medium, even when in solid form. For example, at temperatures in the preferred ranges, TFA is soluble in a reaction medium containing acetic acid and water, at levels ranging from about one hundred ppm. wt. up to more than 1% by weight, and solubility increases with increasing temperature. Despite the fact that there is a difference between the reaction rates of the formation of various harmful impurities from the supply of an oxidizable aromatic compound (for example, para-xylene), from aromatic intermediate reaction products (for example, para-toluic acid), from the desired marketable aromatic carboxylic acid (for example, TFA ) and from aromatic compounds that do not contain non-aromatic hydrocarbon groups (for example, isophthalic acid), the presence and reactivity of the last two groups sets a region of diminishing effect the activity of further suppression compared with the first two groups: the supply of oxidizable aromatic compounds and aromatic intermediate reaction products. For example, with the partial oxidation of para-xylene in TFA, if the dissolved TFA is 7000 ppm. wt. in the liquid phase of the reaction medium under given conditions, the dissolved benzoic acid is 8000 ppm. wt., dissolved isophthalic acid is 6000 ppm. wt. and dissolved phthalic acid is 2000 ppm wt., the value of further reducing all harmful compounds begins to decrease when the activity of the reaction increases to lower the liquid-phase concentration of para-toluic acid and 4-CBA below similar levels. Thus, the presence in the liquid phase of the reaction medium of aromatic compounds that do not contain non-aromatic hydrocarbon groups, and their concentration changes very little with increasing activity of the reaction, and their presence serves to expand upward the region of diminishing efficiency of reducing the concentration of intermediate reaction products in order to suppress the formation of harmful impurities .

Thus, one embodiment of the present invention provides preferred ranges of carbon oxides limited at the lower boundary by low reaction activity and excessive formation of harmful impurities, and at the upper boundary by excessive carbon losses, but at levels lower than those found and described previously as having industrial significance. Accordingly, the formation of carbon oxides is preferably controlled as follows. The ratio of moles of all obtained carbon oxides to moles of the introduced oxidizable aromatic compound is preferably greater than about 0.02: 1, more preferably more than about 0.04: 1, even more preferably more than about 0.05: 1, and most preferably greater than 0.06: 1. At the same time, the ratio of moles of all formed carbon oxides to 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 than about 0.19: 1, and most preferably less 0.15: 1. The ratio of moles of carbon dioxide formed to moles of the introduced oxidizable aromatic compound is preferably greater than about 0.01: 1, more preferably more than about 0.03: 1, even more preferably more 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 the 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 about 0.16: 1, and most preferably less than 0 ,eleven. The ratio of moles of carbon monoxide formed to moles of added oxidizable aromatic compound is preferably greater than about 0.005: 1, more preferably more than about 0.010: 1, even more preferably more than about 0.015: 1, and most preferably more than 0.020: 1 . At the same time, the ratio of moles of carbon monoxide formed to 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 about 0.05: 1, and most preferably less than 0 04: 1.

The carbon dioxide content of the dry off-gas from the oxidation reactor is preferably greater than about 0.10 mol%, more preferably more than about 0.20 mol%, even more preferably more than about 0.25 mol%, and most preferably more than 0.30% molar. At the same time, the content of carbon dioxide in the dry exhaust gas from the oxidation reactor is preferably less than about 1.5% molar, more preferably less than about 1.2% molar, even more preferably less than about 0.9% molar, and most preferably less than 0, 8% molar. The carbon monoxide content in the dry off-gas from the oxidation reactor is preferably greater than about 0.05 mol%, more preferably more than about 0.10 mol%, even more preferably greater than 0.15, and most preferably greater than 0 , 18% molar. At the same time, the carbon monoxide content in the dry exhaust gas from the oxidation reactor is preferably less than about 0.60% molar, more preferably less than about 0.50% molar, even more preferably less than about 0.35% molar, and most preferably less than 0, 28% molar.

The authors of the present invention have found that, according to the descriptions of the present invention, an important factor to reduce the formation of carbon oxides to these preferred ranges is to increase the purity of the recycle filtrate and to supply an oxidizable compound to reduce the concentration of aromatic compounds containing no non-aromatic hydrocarbon groups, while reducing the formation of carbon oxides and harmful impurities. Another factor, according to the discoveries of the present invention, is to improve the distribution of para-xylene and oxidizing agent in the reaction apparatus. Another factor that makes possible the above preferred levels of carbon oxides is the work in the reaction medium with gradients, which are described here for pressure, temperature, concentration of oxidizable compounds in the liquid phase and for the oxidizing agent in the gas phase. Another factor that makes possible the aforementioned preferred levels of carbon oxides is the work within the preferred spatiotemporal reaction rates described here, pressure, temperature, solvent composition, catalyst composition and mechanical geometry of the reaction apparatus.

An important benefit of working in the preferred ranges of carbon monoxide production is that molecular oxygen can be reduced, although not to stoichiometric values. Despite a good gradual change in the concentrations of the oxidizing agent and the oxidizable compound in accordance with the present invention, an excess of oxygen may remain above the stoichiometric value calculated for supplying only the oxidizable compound to allow slightly lower levels of carbon oxides and provide excess molecular oxygen to control the formation of harmful impurities. Specifically, for the case where the feed of the oxidizable compound is xylene, the ratio in the feed of the weight of molecular oxygen to the weight of xylene is preferably greater than about 0.91: 1.00, more preferably greater than about 0.95: 1.00, and most preferably greater than 0.99: 1.00. At the same time, the ratio of molecular oxygen to weight of xylene 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 supplying xylene, the time-averaged molecular oxygen content of the dry off-gas from the oxidation reactor is preferably greater than about 0.1 molar, more preferably greater than about 1% molar, and most preferably more than 1.5% molar. At the same time, the time-averaged molecular oxygen content of the dry off-gas from the oxidation reactor is preferably less than about 6% molar, more preferably less than about 4% molar, and most preferably less than 3% molar.

Another big gain from working in the preferred ranges of carbon monoxide production is that less aromatic compounds are converted to carbon oxides and other, less valuable forms. This gain is estimated by 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 fraction of surviving" aromatic compounds in the reaction medium and is expressed as the number of percent. If all incoming aromatic compounds leave the reaction medium as aromatic compounds, even if most of them are in the form of oxidized forms of incoming aromatic compounds, then the molar fraction of survivors has a maximum value of 100%. If exactly 1 out of every 100 incoming aromatic molecules is converted to carbon oxides and / or other non-aromatic molecules (for example, acetic acid) when passing through the reaction medium, then the molar fraction of survivors is 99%. Specifically, for the case where the main supply of the oxidizable aromatic compound is xylene, the molar fraction of the surviving aromatic compounds passing 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 to have sufficient activity of the overall reaction, the molar fraction of surviving aromatic compounds that have passed 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 the main supply of oxidizable aromatic compounds is xylene.

Another aspect of the present invention includes the formation 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, therefore, tends to follow the exhaust gas, unless additional cooling or other typical methods are used to extract it and / or to neutralize it before the exhaust gas is released into the atmosphere. Thus, the formation of methyl acetate means running costs as well as capital costs. Perhaps methyl acetate is formed as follows: first, the methyl radical is combined, possibly due to the decomposition of acetic acid, with oxygen to give methyl hydroperoxide, then decomposition with the formation of methanol, and finally, the reaction of the obtained methanol with the remaining acetic acid to form methyl acetate. Regardless of the chemical route, the inventors of the present invention found that whenever methyl acetate is produced at too low a rate, the formation of carbon oxides is also too low and the formation of harmful aromatic impurities is too high. If the formation of methyl acetate proceeds at too high a rate, the formation of carbon oxides is also excessively high, which leads to a decrease in the yield of solvent, oxidizable compound and oxidizing agent. When applying the preferred embodiments described herein, the ratio of moles of methyl acetate obtained to moles of added 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 moles of methyl acetate obtained to moles of added oxidizable aromatic compound is preferably less than about 0.09: 1, more preferably less than about 0.07: 1, even more preferably less than about 0.05: 1, and most preferably less than 0 04: 1.

The authors of the present invention note that for all of the numerical ranges given here, the upper and lower bounds of the ranges may be independent of each other. For example, a numerical range from 10 to 100 means more than 10 and / or less than 100. Thus, a range from 10 to 100 provides the requirement of a limit of more than 10 (without an upper limit), provides a limit of less than 100 (without lower bound), as well as the full range from 10 to 100 (with upper and lower bounds).

Further, the invention is illustrated by the following examples of variants of its preferred implementation, although it should be understood that these examples are included only for illustration and do not mean limiting the scope of the invention, unless specifically indicated otherwise.

Example 1

This is a working example of the industrial oxidation of para-xylene in a bubble column reactor type. This example demonstrates, for example, that there are significant gradients in height of para-xylene concentrations when suitable geometric and process conditions are used in accordance with aspects of the present invention.

In this example, an industrial bubbler oxidation plant was used, having an almost vertical, substantially cylindrical body with an inner diameter of about 2.44 m. The height of the bubbler oxidation plant was about 32 m from the lowest tangent (TL) to the upper TL. Above and below the cylinder, the apparatus was equipped with an elliptical cover and a bottom of about 2: 1. The working level was approximately 25 m of reaction medium above the lower TL. The feed rate of technically pure para-xylene was practically constant and equal to approximately 81 kg / min; the feed into the reaction apparatus passed through a round hole located in the wall of the cylindrical section at an altitude of about 4.35 m above the lower TL. The internal diameter of said hole in the wall was about 0.076 m. The filter solvent was supplied at an almost constant speed of about 777 kg / min. The uncontrolled part of this filter solvent, estimated from the size of the pipeline and the pressure drop of about 20 kg / min, was supplied as a liquid stream to the oxidizer distributor. The remainder of the filter solvent, approximately 757 kg / min, was fed uniformly mixed with technically pure para-xylene. Thus, the combined liquid-phase flow of the filter solvent and technically pure para-xylene feed reached about 838 kg / min, which gives a reduced inlet flow velocity through the indicated hole in the wall of about 3 m / sec. This filter solvent was taken from a factory regeneration system and contained more than about 97% wt. acetic acid and water. The concentration of the catalytic components in the filter solvent was such that the composition in the liquid phase of the reaction medium was approximately 1777 ppm. wt. cobalt, about 1518 ppm wt. bromine and about 107 ppm wt. Manganese A separate solvent reflux stream with an almost constant speed of about 572 kg / min was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium. This solvent reflux contained more than about 99% wt. acetic acid and water, and the reflux of solvent was taken from a separate factory regeneration system that did not contain significant levels of catalytic components. The total water content in the feed of the filter solvent and the reflux of the solvent was such that the concentration of water in the liquid phase of the reaction medium was about 6.0 wt%. The oxidizing agent was compressed air, supplied at an essentially constant speed of about 384 kg / min through an oxidizer distributor similar to that shown in FIG. 2-5. This oxidizer distributor contained a current channel composed of angled segments, which was an almost equilateral octagon with a transverse crossbar connecting one side with the opposite side and passing through the vertical axis of symmetry of the reaction apparatus. The segmented current channel was made from sections of Schedule 10S pipes with a nominal diameter of 12 inches. The width of the octagon from the center of gravity of one side of the current channel to the center of gravity of the opposite side was approximately 1.83 m. The octagon lies almost horizontally, and half the elevation height of the octagonal channel was approximately 0.11 m above the lower TL of the reaction apparatus. The oxidizer distributor had 75 round holes with a diameter of about 0.025 m. The holes were distributed approximately evenly around the octagon and the transverse crossbar, located near the top of the indicated 12-inch pipes. There was only one circular hole with a diameter of about 0.012 m near the bottom of one side of the octagonal channel. The excess working gas pressure in the upper free space of the reaction apparatus was stably about 0.52 MPa. The reaction was carried out essentially adiabatically, so that the heat of reaction increased the temperature of the incoming materials and evaporated a substantial part of the incoming solvent. The operating temperature, measured near half the height of the reaction medium, was approximately 160 ° C. The effluent suspension containing technical terephthalic acid (TTFC) was removed with a practically constant speed from the reaction apparatus near the bottom of the lower elliptical base. The flow rate of the effluent suspension was approximately 408 kg / min.

Suspension samples from the reaction medium were obtained, as described below, at three points in the reaction apparatus at different heights. When determining the concentration of different components in different places of the reaction medium, it was necessary to take into account the stochastic nature of the system, taking enough samples to determine the time-averaged value with sufficient resolution.

One set of five samples was obtained from a line for withdrawing a suspension near the bottom of the elliptical bottom of the reaction apparatus. Another set of five samples was taken from a hole in the wall located about 12.4 m above the lower TL of the reaction apparatus. The third set of five samples was obtained from a hole in the wall, located at an altitude of approximately 17.2 m above the lower TL of the reaction apparatus.

All suspension samples were analyzed by calibrated gas chromatography (GC) for the composition of para-xylene and para-tolualdehyde in the liquid phase. Table 5 below shows the averages of the five results that were obtained at three different column heights. The results are given as weight parts of the analyzed sample per million wt. parts (ppm wt.) of the liquid phase.

Table 5 Sampling Place Para-xylene (ppm wt.) Para-tolualdehyde (ppm wt.) Side wall, approx. 17.2 m 21 140 Side wall, approx. 12.4 m 63 317 Bottom bottom 455 960

These results show the presence of large vertical gradients of local concentrations of para-xylene and para-tolualdehyde. For example, the concentration gradient of para-xylene, as can be seen from the data in table 5, was more than 20: 1 (455: 21). These results demonstrate that the intrinsic liquid mixing of the incoming para-xylene feed to the bubble column was significantly slower than the intrinsic reaction rates. To a lesser extent, height gradients were also observed for concentrations of other related reactive aromatic components in the reaction medium (for example, para-toluic acid and 4-carboxybenzaldehyde).

As shown in the following examples, detailed calculation models show that the real range of para-xylene concentration in the liquid phase of the reaction medium in this example was noticeably superior to 100: 1. Even without strict model calculations, specialists in this field should understand that the real maximum concentration of para-xylene takes place in the area near the injection site of para-xylene into the reaction apparatus with a bubble column through the wall of the apparatus. The height at which the concentration of para-xylene is maximum is approximately 4.35 m above the lower TL, between samples taken at an altitude of approximately 12.4 m and from the lower discharge. Likewise, the real minimum concentration of para-xylene is likely to occur at the top or very close to the top of the reaction medium, at an altitude of about 25 m, well above the highest altitude where the samples were taken.

The concentrations of para-xylene and other oxidizable compounds can be measured at other locations in the reaction medium using suitable mechanical devices for sampling at any position in height or horizontally in the reaction medium. Alternatively, concentrations for positions from which samples were not physically taken and chemically not analyzed can be calculated with reasonable accuracy using numerical models of sufficient complexity to correctly take into account the extremely complex picture of the fluid flow, kinetics of chemical reactions, energy balance, vapor-liquid equilibrium -solid and interfacial exchange rates.

Examples 2-5

Examples 2-5 are design models of bubbler reactor columns, either identical to the reactor of example 1, or generally similar, but with certain improvements. Numerical simulation of fluid dynamics (CFD), carried out to create examples 2-5, was carried out in accordance with the modeling method described in pending patent application US, reg. No. 60/594774, dated May 5, 2005, under the name "Modeling of liquid-phase oxidation", the full description of which is explicitly introduced here by reference.

In Examples 2-5, CFD modeling was performed using the CFX program, Release 5.7 (ANSYS, Inc. 275 Technology Drive, Canonsburg, PA 15317). Examples 2-5 include over about 100,000 individual three-dimensional design cells each. The time step used in examples 2-5 is less than 0.1 sec. Several bubble sizes, with diameters ranging from about 0.005 to about 0.20 m, turned out to be useful for setting up the CFD model rather close to the time-average volumetric content of bubbles estimated by measuring the differential pressure and to the vertical profile of volumetric bubble content estimated by gamma - scanning, and to horizontal profiles of the volumetric content of bubbles, estimated using computed tomography (CT). In order to select suitable bubble sizes and populations in CFD models of Examples 2-5, real factory technical data were obtained for cylindrical suspension bubble columns with inner diameters of about 2.44 m and about 3.05 m working with a reaction medium of similar composition and method conditions as described below. Basic data on the total volumetric content of bubbles was obtained by measuring the pressure drop near the base of the apparatus and up to the upper space of the exhaust gas. Basic data on the vertical profile of the volumetric content of bubbles was obtained using a gamma-emitting radioactive source and a detection method propagating to the outer part of the reaction apparatus in increments ranging from about 0.05 m to about 0.3 m. Basic data on the horizontal profile of the volumetric content of bubbles were obtained by computed tomography scan, carried out on a nine-by-nine lattice across the horizontal plane of a working bubble column using gamma - emitting radioactive source and method of detection. That is, the source was placed at a given height in nine different places, spaced approximately equal distances along the perimeter of the bubble column. For each position of the gamma radiation source, the amount of gamma radiation passed through the reaction apparatus and the reaction medium in nine different places separated by approximately equal distances around the bubble column was detected. Then, different mathematical models were used for these separate data in order to obtain an estimate of the range of changes in the volumetric content of bubbles in the reaction medium for the indicated height. Several horizontal computed tomographic scans were performed on two different days, for two different heights and at two different feed rates of para-xylene, compressed air, etc.

The model of chemical reactions for the consumption of para-xylene in this medium was modified so as to bring it into line with the profiles of the para-xylene reagents defined in Example 1, together with other data for similar temperatures, pressures, reaction intensities, catalysis, water concentration and etc. obtained from testing both on an industrial and pilot scale. As a representative approximation, the pseudo-first order time constant for the breakdown of reactive labeled para-xylene is about 0.2 sec ^-1 for about 160 ° C and about average reaction conditions used in Examples 2-4.

It is important that the CFD models of flow fields obtained in Examples 2-4 give large-scale fluctuations in the accumulation of bubbles and oscillations in the liquid level, which are generally consistent with the observed low-frequency oscillations in the operation of the reaction apparatus with a bubble column.

Example 2

In this example, calculations are carried out related to the mechanical configuration of Example 1, and a comparative basis is established for Examples 3 and 4. In this example, the mechanical configuration of the bubble column reactor is identical to Example 1, with a round inlet with a diameter of 0.076 m in the wall of the reaction apparatus for the feed stream containing para-xylene and filter solvent. The feed rate of para-xylene is about 1.84 kg / sec, higher than in Example 1. The feed rate of the filter solvent fed uniformly mixed with para-xylene is about 18.4 kg / sec. Thus, the reduced velocity of the combined stream of para-xylene plus filter solvent entering through the hole in the wall is approximately 4 m / s. The rate of reflux of solvent into the upper gas separation space is 12.8 kg / s. The compressed air supply through the oxidizer distributor is approximately 9 kg / s. The solids content in the reaction suspension is approximately 31% wt. A commodity suspension is discharged from the center of the lower base of the reaction apparatus at a practically constant speed to keep the reaction medium at an approximately stationary level of about 25 m. The average volumetric fraction of gas, expressed as averaged over the area and time, at half the height of the reaction medium is approximately 55%, moreover, the time averaging interval is at least about 100 sec. of time of the CFD model. The excess pressure in the free space above the reaction medium is approximately 0.50 MPa. The temperature, measured near half the height of the reaction medium, is approximately 160 ° C. The contents of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 1.

Example 3

In this example, calculations are carried out relating to improving the distribution of the para-xylene feed by increasing, according to one aspect of the present invention, the reduced liquid-phase feed rate containing para-xylene at its point of entry into the reaction medium. In this example, the mechanical configuration of the reactor type bubble column is identical to Example 2, except that the circular hole in the wall through which the liquid-phase feed containing para-xylene enters is reduced to a diameter of 0.025 m. The feed rate of para-xylene and other process conditions are the same as for example 2, except that the reduced velocity of the combined liquid-phase flow of para-xylene feed plus filter solvent entering through the hole in the wall is now about 36 m / s.

Calculations by CFD model of the time-averaged fraction of the reaction medium with the concentration of reactive labeled para-xylene in the liquid phase above different threshold values are presented below in table 6. According to the present invention, the volume of the reaction medium with highly concentrated reactive labeled para-xylene in the liquid phase decreases when working with a higher input velocity of the liquid-phase flow of raw materials containing para-xylene. The reduction of zones with a high concentration of para-xylene is important for limiting undesirable coupling reactions, both because they contain increased concentrations of many soluble aromatic compounds and because such high concentrations lead to a locally high consumption of dissolved molecular oxygen and, therefore, , lead to locally reduced stationary concentrations of dissolved molecular oxygen.

Table 6 Example 2 Example 3 Example 4 Diameter of the hole in the wall (meters) 0,076 0,025 changed distributor Reduced incoming pX speed + filtrate (m / s) four 36 > 15 The percentage of reaction medium with a concentration of pX above 1000 ppm wt. (%) 3.64 3.28 3.73 above 2500 ppm wt. (%) 0.640 0.378 0.130 above 10000 ppm wt. (%) 0,049 0,022 0.005 above 25000 ppm wt. (%) 0.009 0.002 0.001 The volume of the reaction medium with a concentration of pX above 1000 ppm wt. (liters) 4250 3840 4360 above 2500 ppm wt. (liters) 749 444 152 above 10000 ppm wt. (liters) 57 26 6 above 25000 ppm wt. (liters) 10 2 one

Example 4

In this example, calculations are made to improve the mechanical means of introducing an oxidizing agent and para-xylene into a bubble column reactor type. This example is performed for the same bubble column reactor type, which is used in examples 1-3. However, the reactor is modified with respect to the manner in which the oxidizing agent and para-xylene are introduced into the reaction medium. In the discussion of Example 4, attention is directed, first, to a modified device for introducing para-xylene into the reaction medium, whereby zones with high concentrations of para-xylene are reduced. Secondly, attention is directed to a modified device for introducing an oxidizing agent into the reaction medium, in accordance with which zones that are poorly aerated are reduced. It should not be considered that the results of these two modifications are completely independent, but this is just a consistent presentation.

The amount of the reaction medium with a very high concentration in the liquid phase of the reactive labeled para-xylene is reduced in Example 4 due to the use of a liquid-phase feed distribution system, mainly as shown in FIG. 9-11. It is favorable that this liquid-phase feed distribution system has four streamlines favorably located almost vertically. Each of these four streamlines is located approximately 0.75 m from the vertical axis of symmetry of the bubble column. These four streamlines are favorably made from sections of Schedule 10S pipes with a nominal diameter of 1.5 inches. The lower end of each knee in this example, respectively, has a conically converging portion with an adjacent angle, measured between opposite sides of the cone, which favorably is about 24 °; however, other shapes are also suitable to block the lower end of the streamline (for example, a conical shutter at a different adjacent angle, a flat shutter, a shutter with a pipe cap, a wedge shutter, etc.). Each of these four streamlines has a total of nine circular holes, each of which has a diameter of approximately 0.0063 m. The lowest of the nine holes in each line is at the bottom of the lower conical section. For each line, this lowest hole is approximately 0.4 m above the lower TL of the reaction apparatus. If we count each time from this lower edge of the lower truncated conical section, then the next three holes in each line are at a height of about 0.3 m, the next three holes are raised to a height of about 1.6 m, and the two most upper holes are at a height of 2 , 7 m. Thus, the height distance from the lowest hole to the highest hole in each line is about 2.7 m, or about 1.1D. The linear (not in height) distance between the farthest holes, from the bottom hole of one vertical line to the top hole of a diagonally opposite vertical line, is about 3.44 m, or about 1.4D. At each level, the holes are distributed almost evenly around the perimeter of each streamline. A feed line for supplying an oxidizable compound and a solvent to the top of four almost vertical lines is favorably almost horizontal and located at a height of about 3.60 m above the lower TL of the reaction apparatus. The feed line is favorably made from sections of Schedule 10S pipes with a nominal diameter of 3 inches. There are suitable modules of mechanical cross struts and mechanical fastening from the module to the oxidizer distributor and to the reaction apparatus to withstand both static and dynamic loads that occur during normal and disturbed operation.

Although this was not calculated in this example, many other designs of this liquid-phase feed distribution system are possible. For example, the size of the fluid flow lines may be larger or smaller, or have a different cross section, rather than approximately circular, or there may be not four. For example, each of the four essentially vertical lines can be fed independently through streamlines separately cut into the pressure wall of the reaction apparatus. For example, the connection with the supply of incoming para-xylene and the supply of solvent can take place near half the height or at a level near the bottom, or at any height, or at several heights of almost vertical lines. For example, the supply lines can be approximately vertical with distribution holes located in almost horizontal lines, or both flow directions can be inclined, or non-linear, or non-orthogonal. For example, the openings can be arranged differently radially, azimuthally, or vertically with respect to the reaction medium. For example, more or fewer holes and / or holes of different shapes and / or holes with a set of sizes and / or different shapes can be used. For example, an outlet nozzle may be used rather than outlet openings. For example, one or more flow deflecting devices may be located outside the streamline near the outlet openings and in the fluid path after being discharged into the reaction medium.

Depending on the nature and content of solids, if any, in the combined supply of para-xylene and solvent, or on the reaction medium, and depending on the start-up, shutdown, and other operating procedures used in actual technological operations, it may be necessary to clean solids inside the liquid-phase feed distribution system. Although this was not calculated in this example, the purge openings may be practical larger than the same openings shown in the current example. The hole at the bottom edge of each of the four almost vertical elbows is particularly suitable for cleaning solids, although this is not the only possible remedy. More sophisticated mechanical devices such as damper modules, check valves, overflow valves, actuated valves and the like can be used, either to prevent solids from entering or to remove accumulated solids from the liquid-phase feed distribution system.

We now turn our attention to the oxidizer distributor, which is generally the same as shown in FIG. 12-15. The annular element of this oxidizer distributor favorably contains a current channel formed from angled sections, which advantageously is an almost equiangular octagon without a transverse crossbeam. The current channel formed from sections is favorably made from sections of Schedule 10S pipes with a nominal diameter of 10 inches. The width of the octagon from the center of gravity of one side of the current channel to the center of gravity of the opposite side is approximately 1.12 m. The octagonal section favorably lies approximately horizontally, and half the height of the octagonal section is approximately 0.24 m below the lower TL of the reaction apparatus. This is the exact opposite of the ring element of the oxidizer distributor of Examples 1-3, the height of which is above the lower TL of the reaction apparatus. The octagonal part of the channel is perforated with 64 almost round holes, each with a diameter of about 0.030 m, arranged approximately evenly around the channel. About half of the holes are located around the channel in places located at an angle of approximately 45 ° below the horizon, if measured from each hole to the nearest center of gravity of the current channel section. Approximately half of the holes are located around the channel in places located approximately below the channel (i.e., at an angle of approximately 90 ° below the horizon, if measured from each hole to the nearest center of gravity of the channel section). The authors of the present invention again note, similarly to the comments made for the liquid-phase inlet distributor, that many other particular designs are possible for the oxidizer distributor, falling within the scope of certain aspects of the present invention. For example, more or less than two power lines can be cut into the pressure barrier wall. For example, the oxidizer distributor power lines should be designed without a ring element. For example, there may be more than one annular element, and any annular element may have not 8, but a different number of sides, or may have asymmetric sides. For example, the design can achieve a preferred pressure drop, or a preferred quality of aeration, or a preferred non-contaminating nature, using different numbers and sizes or sizes and placement of holes in the lines or exits from the lines. For example, lines of different diameters from preferred ranges may be used in the construction. For example, a structure can have a non-polluting nature if a liquid flush is used.

In this example, the reaction medium is discharged at a practically constant speed on the side of the reaction apparatus at a height of about 14 m through a circular hole in the wall having an internal diameter of about 0.076 m. The diverted reaction medium is divided into a commercial suspension containing technical terephthalic acid and exhaust gas using an external apparatus deaeration, which is fully described in Example 6. The separated off-gas from the external deaeration apparatus is carried out in line to connect to the main off-gas stream, you walking on top of the reaction apparatus.

The CFD modeling methods of this example are essentially the same as for examples 2 and 3, with the following exceptions. The spatial network has been changed, as is appropriate and as is known in the art, for a modernized apparatus for distributing the incoming oxidizing agent, for distributing the incoming oxidizable compound, and for removing salable suspension through the side wall of the reaction apparatus about 14 m above the lower TL.

In order to evaluate the results of the CFD model regarding the distribution of reactive labeled para-xylene, the same methods were used as in examples 2 and 3. Namely, the time-averaged fractions of the reaction medium with the concentration of reactive labeled para-xylene in liquid were determined. phase above different thresholds. To simplify the comparison, the results of this example are presented in Table 6 above. These results show that the improved distribution of reactive labeled para-xylene in this example does cause a slight increase in the amount of reaction medium above 1000 ppm. wt., but more harmful threshold levels of 2500 ppm wt., 10000 ppm wt. and 25000 ppm wt. are declining. These improvements are provided, for example, by higher feed input speeds, along with improved vertical, radial and azimuthal arrangements and a breakdown of the introduction of para-xylene into the reaction medium.

Now, returning to the quality of aeration over the reaction medium, to evaluate poorly aerated volume in the reaction medium of Examples 2-4, the method of 2000 horizontal layers of equal subvolume is used. Starting from the bottom of the reaction medium, namely the bottom of the bottom in this example, the reaction medium is divided into 2000 equal subvolumes using theoretical horizontal planes. For each time increment in the CFD model, in each of the indicated 2000 equal subvolumes, the amount of suspension and the amount of gas are determined, which are used to calculate the average volumetric gas content there. In order to take into account the stochastic nature of the method and its CFD model, the results of the CFD model are averaged over time for a model time of at least about 100 seconds to obtain time-averaged values of the volumetric gas content in each of 2000 equal subvolumes.

After the time-averaged volumetric gas content is determined for each of 2000 equal subvolumes, these values are compared with the threshold values described here. For each threshold, the total number of "abnormal" subvolumes in which a specific threshold value is not exceeded is calculated. In Table 7 below, for both example 2 and example 4, the number of 2000 horizontal layers of the reaction medium of equal volume with a time-averaged volumetric gas content of below 10% volumetric, below 20% volumetric and below 30% volumetric is given. Example 4 is significantly better compared to example 2.

Table 7 Example 2 Example 4 The height of the center of gravity of the ring element of the oxidizer distributor from the lower TL of the apparatus (meters) +0.11 -0.24 The number of 2000 horizontal layers of the reaction medium of equal volume with a time-averaged volumetric gas content below 10% vol 7 no below 20% vol 21 no below 30% vol 41 no

When comparing calculation examples 2 and 4, it should also be noted that the supply of para-xylene in example 4 to the reaction medium is lower and closer to the inlet stream of the oxidizing agent than in example 2.

Examples 5 and 6

Examples 5 and 6 are working examples, demonstrating in an industrial oxidation plant with a bubble column the importance of minimizing areas of poor aeration, improving the method of introducing technically pure para-xylene feed, to better distribute it vertically, azimuthally and radially, and lowering the feed introduction point technically pure para-xylene to make it closer to the point of greatest availability of molecular oxygen, as described in the present invention. In addition, these examples demonstrate a gain for exit from the location of the holes for outputting the suspension at a height.

In the partial oxidation of para-xylene, there are many different impurity compounds that are formed typically by combining aromatic rings. One of them is 4,4'-dicarboxystilbene. This compound has much higher light absorption than terephthalic acid, and this greatly reduces the optical permeability of the target product. In addition, 4,4'-dicarboxystilbene is a favorable impurity for use in controlling the quality of continuous oxidation, since it is selectively distributed into the solid phase of the reaction medium, therefore, in the recycle stream of the solvent in the industrial reaction apparatus with a bubble column described in examples 5 and 6 usually very little 4,4'-dicarboxystilbene is present. In Examples 5 and 6, the concentrations of 4,4'-dicarboxystilbene were measured analytically using HPLC-MS calibrated with a suitable reference mixture containing a solvent and known amounts of several analytes, including, in particular, a known amount of 4,4'-dicarboxystilbene. The HPLC-MS analysis method is described above in the Detailed Description section.

Example 5

The reactor-type bubbler column used in this example has essentially the same mechanical configuration as the reactor in Examples 1 and 2. The reactor is under process conditions comparable to Example 6 and provides a basis for comparison. The working level of the reaction medium was about 25 m. The supply of technically pure para-xylene was virtually continuous at a speed of about 81 kg / min. The filter solvent was supplied at an almost constant rate of about 793 kg / min. An unregulated proportion of it, which, according to estimates, based on line size and pressure drop, was approximately 20 kg / min, was supplied as a liquid flush to the oxidizer distributor. The rest of the filter solvent, about 773 kg / min, was supplied uniformly mixed with technically pure para-xylene. Thus, the total liquid-phase flow of filter solvent and technically pure para-xylene was thus approximately 854 kg / min. This filter solvent was taken from the factory regeneration system and contained more than about 97% wt. acetic acid and water. The concentration of catalytic components in the filter solvent was such that the composition in the liquid phase of the reaction medium was as follows: about 2158 ppm wt. cobalt, approximately 1911 ppm wt. bromine and about 118 ppm wt. Manganese A separate solvent reflux stream with an almost constant speed of about 546 kg / min was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium. This solvent reflux contained more than about 99% wt. acetic acid and water, and the phlegm of the solvent was taken from a separate factory regeneration system, which did not contain significant levels of catalytic components. The total water content in the feed of the filter solvent and the reflux of the solvent was such that the concentration of water in the liquid phase of the reaction medium was about 5.8% by weight. The oxidizing agent was compressed air supplied at an actually constant speed of about 352 kg / min. The excess working gas pressure in the upper space of the reaction apparatus was constantly at a level of about 0.42 MPa. The reaction apparatus worked essentially in an adiabatic mode, so that the heat of reaction increased the temperature of the incoming feed and evaporated a significant fraction of the incoming solvent. The operating temperature, measured near half the height of the reaction medium, was approximately 154.6 ° C. The effluent suspension containing technical terephthalic acid (TTFC) was removed near the bottom of the lower elliptical base of the reaction apparatus with an almost constant speed of about 428 kg / min.

In this example, the ratio of the rate of formation of undesired 4,4'-dicarboxystilbene to the rate of formation of the desired terephthalic acid, measured by HPLC-MS on three separate samples of the suspension product, was approximately 8.6, 9.1 and 9.2 ppm. wt., so the average value was about 9.0 ppm. wt. The concentration of para-xylene in the liquid phase of the effluent suspension, measured using calibrated GC on three different samples of the suspension product, was approximately 777, 539 and 618 ppm. wt., thus, the average value was approximately 645 ppm. wt. The concentration of para-tolualdehyde in the liquid phase of the effluent suspension, measured using calibrated GC on these individual samples of the suspension product, was about 1055, 961 and 977 ppm. wt., thus, the average value was approximately 998 ppm. wt.

Example 6

The reactor bubbler column of this example corresponds to the mechanical configuration developed in calculation example 4. The reactor of this example includes improvements in height, speed, number and breakdown of the para-xylene feed inlets, thus providing an improved distribution of para-xylene feed and an improved gradual change molecular oxygen concentration. In addition, it includes improving the quality of aeration in the reaction medium through the use of an improved oxidizing agent distributor located at a height and a method for removing a deaerating slurry exiting the reaction medium. Compared to Example 5, noticeable improvements in the yield of para-xylene are visible, and a significant reduction in the formation of impurities is seen.

The reactor of this example has an improved mechanical design compared to the CFD model of Example 4. The working level of the reaction medium was about 25 m. The supply of technically pure para-xylene was practically constant at a speed of about 81 kg / min. The filter solvent was fed uniformly mixed with technically pure para-xylene at an essentially constant speed of about 744 kg / min. Thus, the total flow of filter solvent and supply of technically pure para-xylene was approximately 825 kg / min. This filter solvent was taken from the same factory regeneration system and essentially had the same composition as in Example 5. The concentration of catalytic components in the filter solvent was such that the composition in the liquid phase of the reaction medium was as follows: about 1996 ppm. wt. cobalt, about 1693 ppm wt. bromine and about 108 ppm wt. Manganese A separate solvent reflux stream with an almost constant speed of about 573 kg / min was supplied in the form of droplets into the gas separation zone above the working level of the reaction medium. This solvent reflux contained more than about 99% wt. acetic acid and water, and the reflux of solvent was taken from a separate factory regeneration system that did not contain significant levels of catalytic components. The total water content in the feed of the filter solvent and in the reflux of the solvent was such that the concentration of water in the liquid phase of the reaction medium was about 5.7% wt. The oxidizing agent was compressed air supplied at an almost constant speed of about 329 kg / min. The excess working pressure of the gas in the upper free space of the reaction apparatus was constant at about 0.41 MPa. The reaction apparatus worked essentially in an adiabatic mode, so that the heat of reaction increased the temperature of the incoming feed and evaporated a significant fraction of the incoming solvent. The operating temperature, measured near half the height of the reaction medium, was approximately 153.3 ° C.

The reaction medium was removed from the side of the reaction apparatus at a height of about 14 m through a circular hole in the wall that had an internal diameter of about 0.076 m. The removed reaction medium was drawn through a substantially horizontal line made of sections of Schedule 10S pipes with a nominal diameter of 3 inches in a direction substantially vertical external deaeration apparatus. The external deaeration apparatus had an inner circle diameter of approximately 0.315 m and was made primarily from a Schedule 10S pipe with a nominal diameter of 12 inches. Thus, the horizontal cross-sectional area inside the external deaeration apparatus was approximately 0.0779 m ² . This compares with the horizontal cross-sectional area inside the reaction apparatus, about 4.67 m ² at the height where the reaction medium was removed. Thus, the ratio of the smaller to the larger horizontal section was approximately 0.017.

The external deaeration apparatus passed downward from the height of the inlet of the reaction medium and, approximately 1.52 m short of the lower discharge opening of the stream line, began to decrease in diameter in order to approach it of equal size. A substantially deaerated suspension containing technical terephthalic acid exited the bottom of the external deaeration apparatus at an essentially constant flow rate of about 433 kg / min. Thus, a substantially deaerated suspension at a lower height in a deaeration apparatus with a nominal diameter of 12 inches had a downwardly reduced velocity of about 0.093 m / s; and there was no harmful entrainment of the oxidizing agent by this exit suspension. The outgoing slurry was drawn along streamlines made from sections of Schedule 10S pipes with a nominal value of 3 inches in order to connect to post-treatment equipment. In this example, a means for controlling the flow rate of the discharged reaction medium was placed in a stream exiting from the bottom of the deaeration apparatus, although other places for regulation are possible and suitable.

The external deaeration apparatus went above the height at which the reaction medium entered, about 14 m before the diameter began to decrease from the nominal pipe size of 12 inches, to be consistent with the upper outlet of the current line made from sections of Schedule 10S pipes with a nominal value of 2 inches. A separate off-gas from the external deaeration apparatus was conducted through this 2-inch line to connect to the main off-gas stream exiting from the top of the reaction apparatus.

In this example, the ratio of the rate of formation of undesired 4,4'-dicarboxystilbene to the rate of formation of the desired terephthalic acid, measured by HPLC-MS on three separate samples of the suspension product, was approximately 2.3, 2.7 and 3.2 ppm. wt., with an average value of about 2.7 ppm wt. This is significantly less compared to Example 5. The concentration of para-xylene in the liquid phase of the suspension exiting the lateral outlet at a height measured by calibrated GC on three separate samples of the suspension product was approximately 86, 87 and 91 ppm. wt. with an average value of about 88 ppm wt. The concentration of para-tolualdehyde in the liquid phase of the effluent suspension, measured using calibrated GC on these individual samples of the suspension product, was approximately 467, 442 and 423 ppm. wt. with an average value of about 444 ppm wt. This means improved conversion and yield in the diverted slurry stream as compared with Example 5.

Examples 7-10

Examples 7-10 are calculated examples relating, in particular, to the initial distribution of para-xylene in the reaction medium, as well as showing other aspects of the present invention.

Example 7

This example relates to the supply of vaporized para-xylene. In this calculation example, the para-xylene feed is heated and evaporated before being introduced into the reaction medium. This contributes to the initial distribution of para-xylene. This provides an increase in input volumes and facilitates an increase in speed. In addition, it slows down the transfer of incoming para-xylene to the bulk of the liquid phase and causes the para-xylene feed to move to the reactive liquid phase in greater harmony with the supply of gaseous molecular oxygen.

In this example, the oxidation unit with a bubble column has a vertical cylindrical body with an inner diameter of 2.44 m. The height of the oxidation unit with a bubble column is 32 m from the lower tangent (TL) to the upper TL. Above and below the cylinder, the apparatus is equipped with an elliptical cover and a 2: 1 bottom. The working level of the reaction medium is approximately 25 m above the lower TL. The supply of filter solvent, which is separated from para-xylene, enters at a speed of 18.4 kg / s through a round inlet with a diameter of 0.076 m in the wall of the reaction apparatus at a height of 4.35 m above the lower TL. The feed rate of solvent reflux into the gas separation zone above the working level of the reaction medium is about 14.3 kg / s. The compressed air supply through the oxidizer distributor is substantially the same as in Examples 4 and 6, approximately 9 kg / s. A suspension containing approximately 31% wt. solids, is removed from the reaction medium through the side exhaust elbow, essentially the same as in examples 4 and 6. The excess pressure in the free space above the reaction medium is approximately 0.50 MPa. The contents of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 4.

The feed rate of para-xylene is 1.84 kg / s. Prior to isolation into the reaction medium, the pressure in the feed stream of liquid-phase para-xylene is increased and then it is evaporated at an excess pressure of about 0.69 MPa, heating from a storage temperature of about 40 ° C to a temperature of about 233 ° C. This requires heat to enter the para-xylene feed stream of about 1.3 mJ / s. For this purpose, a heat exchanger is used that uses steam at a pressure of 4 MPa, but any other energy sources with a sufficient temperature are also suitable, including waste heat from process liquids and gases. For the conversion of para-xylene to terephthalic acid, this heat input is approximately 5% of the heat of reaction. Removing this additional heat load causes a slight increase in the temperature of the reaction medium at constant pressure, compared with the supply of liquid para-xylene (see example 8). The temperature, measured near half the height of the reaction medium, is approximately 162 ° C. Alternatively, the pressure can be lowered to reduce the reaction temperature, measured near half the height of the reaction medium, to 160 ° C.

The volumetric flow of vaporized para-xylene is about 0.084 m ³ / s. This stream enters the reaction apparatus at a height of 0.1 m above the lower TL of the apparatus through 3 lines connected in parallel. In the reaction apparatus, each line is made of pipe sections with a nominal value of 1.5 inches and is connected to a round hole of equal diameter in the wall of the apparatus. 3 holes in the wall are located horizontally, with an angular offset of 120 ° from each other. The reduced velocity of each incoming para-xylene stream is approximately 21 m / s, with the incoming para-xylene being distributed in the reaction medium while simultaneously dissolving in the reactive liquid phase, where the catalyst components are mainly located.

Example 8

This example relates to the supply of partially vaporized para-xylene. In this calculation example, the para-xylene feed is partially evaporated by mixing with the feed of the oxidizing agent before being introduced into the reaction medium. This contributes to the initial distribution of para-xylene. This provides increased input volumes and facilitates an increase in speed, and it dilutes the concentration of para-xylene. In addition, it slows down the transfer of incoming para-xylene to the bulk of the liquid phase and makes the para-xylene feed move to the reactive liquid phase in better harmony with the supply of gaseous molecular oxygen.

In this example, the oxidation unit with a bubble column has a vertical cylindrical body with an inner diameter of 2.44 m. The height of the oxidation unit with a bubble column is 32 m from the lower tangent (TL) to the upper TL. Above and below the cylinder, the apparatus is equipped with an elliptical cover and a 2: 1 bottom. The working level of the reaction medium is approximately 25 m above the lower TL. The supply of filter solvent, which is separated from para-xylene, enters at a speed of 18.4 kg / s through an inlet with a circle diameter of 0.076 m in the wall of the reaction apparatus at a height of 4.35 m above the lower TL. The feed rate of solvent reflux into the gas separation zone above the working level of the reaction medium is approximately 12.8 kg / s. The feed rate of compressed air through the oxidizer distributor, similar to that in examples 4 and 6, but modified, as indicated below, is approximately 9 kg / sec. A suspension containing approximately 31% wt. solids, is removed from the reaction medium through the lateral branch elbow, essentially the same as in examples 4 and 6. The excess pressure in the free space above the reaction medium is approximately 0.50 MPa. The contents of water and cobalt, bromine and manganese in the liquid part of the reaction medium are essentially the same as in example 4.

The feed rate of para-xylene is again 1.84 kg / s. It flows like liquid along a line inside the oxidizer distributor, where the liquid is discharged into compressed air in 4 positions, using nozzles that are known in the art. Alternatively, at the point where the liquid enters the oxidizer distributor, open-ended fluid channels or nozzles for atomizing gas / liquid can be used. As a safety measure, 4 temperature sensors are located in the oxidizer distributor. These temperature sensors are connected to alarms and breakers to shut off the oxidant and para-xylene supply if high temperatures occur. When supplying compressed air at a temperature of about 80 ° C, due to the heat of compression, without a secondary cooler at the final stage of compression, and when applying para-xylene with a temperature of about 40 ° C, about 17% wt. para-xylene evaporates at a pressure inside the oxidizer distributor. The remaining liquid para-xylene is carried into the reaction medium with gas in a two-phase flow mixed with gas at speeds close to the gas flow rate. In addition, according to aspects of the invention, said remaining liquid helps to flush all solids that have passed there from the oxidizer dispenser.

The temperature, measured near half the height of the reaction medium, is approximately 160 ° C. Since no additional energy was supplied to any feed stream, this is approximately the same as in examples 4 and 6.

Alternatively, the compressed air supply or the para-xylene feed can be preheated before mixing in the oxidizer distributor to increase the proportion of para-xylene that enters the reaction medium as steam. For example, heat of 300 kJ / s supplied to para-xylene increases its temperature to about 124 ° C and increases the proportion of instantly vaporized para-xylene to about 33%. For example, heat of 600 kJ / s supplied to compressed air increases its temperature to about 146 ° C and increases the proportion of rapidly evaporated para-xylene to about 54%. In both cases, less valuable energy is required for heating than in Example 7. Indeed, waste heat from the exhaust gas from the reaction medium can be used as part or all of the heat source. However, when a certain amount of energy is added to the feed, the temperature of the reaction medium will rise slightly, settling, at constant pressure, flows and phase compositions, between 160 ° C and 162 ° C, if measured at half height. Alternatively, the pressure may be adjusted to adjust the temperature. In addition, when a certain amount of energy is supplied to the feed, the amount of solvent supplied to the reaction apparatus is adjusted if it is desired to keep the solid fraction approximately constant. For example, the solvent reflux stream in Examples 7 and 8 varies between about 12.8 and about 14.3 kg / s, depending on the amount of energy added, to keep the solid fraction approximately constant at about 31% wt.

Example 9

This example relates to the supply of para-xylene not through the wall of the reaction apparatus, but using a liquid ejection device. In this calculation example, the initial distribution of the liquid para-xylene feed is improved by using an ejection device using the liquid flow as a driving force. The reactor of this example has the same mechanical configuration and process boundary conditions as in Example 4, with the exceptions described below. The mixed liquid-phase para-xylene stream plus filter solvent enters through the wall of the reaction apparatus at the same height along the same stream line with a nominal value of 3 inches. However, instead of the internal liquid-phase feed distribution system of Example 4, the mixed liquid-phase feed is released into the reaction medium as a driving liquid stream in an ejection flow device, which is known in the art and is shown in the diagram of FIG. 26. The ejection device is designed for a pressure difference of 0.1 MPa in a moving fluid. The ejection device is arranged and oriented with a stream of flow extending vertically upward along the center line of the reaction apparatus at a height of about 4.5 m above the lower TL. The volume of the reaction medium discharged and mixed with the motive fluid varies with time depending on the random accumulations of bubbles in the bubble column at the entrance to the ejection device. However, the time-average released flow is larger than the flow of the driving fluid, which thus provides a more rapid dilution of the incoming para-xylene. Subsequent mixing and chemical reactions proceed in accordance with the usual random events in a bubble column.

Example 10

This example relates to the supply of para-xylene not through the wall of the reaction apparatus, but using an ejection device of gas and liquid. In this calculation example, the initial distribution of the para-xylene feed is improved by using an ejection device using a gas stream as a driving force. The reactor of this example has the same mechanical configuration and process boundary conditions as in Example 4, with the exceptions described below. Both the octagonal oxidizer distributor and the liquid phase feed distribution system are removed. Instead, the incoming flow of oxidizing agent and a mixed liquid-phase supply of para-xylene plus filter solvent is carried out in separate lines into the reaction apparatus. There, both streams are combined as moving fluids at the entrance to the flow ejection device, which is known in the art and shown in the diagram in FIG. 27. The ejection device is vertically aligned along the axial midline of the reaction apparatus. It is located so that the exits are facing down and are 0.2 m below the lower tangent of the reaction apparatus. The ejection device is designed for a pressure difference of 0.1 MPa on moving fluids. Two temperature sensors are located near the point where the supply of compressed air and para-xylene are combined for the first time. These temperature sensors are connected to alarms and breakers to stop the flow of oxidizing agent and para-xylene if high temperatures are detected.

The withdrawn volume of the reaction medium was increased compared to Example 9, and the initial dilution of the incoming para-xylene was further improved. In addition, the portion of the liquid phase of the reaction medium with the highest local concentrations of para-xylene has an even greater difference in concentrations compared to the gaseous portion with the highest concentration of molecular oxygen. Subsequent mixing and chemical reactions proceed in accordance with the usual random events in a bubble column.

Examples 11-13

Examples 11-13 are calculation examples relating, in particular, to the use of liquid flows from the reaction medium in lines to facilitate the initial distribution of para-xylene in the reaction medium, as well as showing other aspects of the present invention.

Example 11

This example relates to the use of streamlines within the reaction apparatus for transporting liquid to aid in the initial distribution of incoming para-xylene. The reactor in this example has the same mechanical configuration and boundary conditions of the method as in example 4, except for the exceptions described below. Consideration is given to the circuit of FIG. 24. The mixed liquid phase stream — para-xylene plus filtered solvent — enters through the wall of the reaction apparatus through a 3-inch stream line similar to the line of Example 4. However, the internal liquid-phase distribution system of Example 4 is removed, and instead, the mixed liquid stream is discharged into the line current. The streamline has an inner circle diameter of about 0.15 m over most of its length, including its lower end, which is 1 m above the lower TL of the apparatus. The current line rises vertically to a full height of 21 meters above the lower TL of the apparatus. At a height of 20 m above the lower TL of the apparatus, the streamline expands and, when raised to a height of another 1 m, has an internal section area of 0.5 m ² . This upper, with a large diameter, portion of the indicated streamline can be considered an internal deaeration apparatus, and it is indeed formed by a part of the reaction apparatus using its wall. The streamline is entirely inside the reaction apparatus. At the upper entrance to the streamline, the reaction medium is greatly depleted in para-xylene and para-tolualdehyde, although there are significant concentrations of para-toluic acid and 4-carboxybenzaldehyde. The reaction medium entering from above the specified streamline is essentially deaerated, forming a denser medium inside the specified stream line than in the rest of the reaction apparatus. The suspension inside the streamline flows downward with an estimated velocity of about 150 kg / s; at this point in the flow, the pressure drop integrated over the entire length of the specified current channel is balanced by the difference in densities inside and outside, integrated over the entire length of the specified stream line. Of this flowing down suspension, about 104 kg / s is a liquid, which is about 69% wt. A supply of uniformly mixed para-xylene and filter solvent, amounting to approximately 20.2 kg / s, is included in said streamline about 5 m above the lower TL. This mixture then passes down the streamline for another 4 m (approximately 27 line diameters) in less than 1 second and becomes substantially mixed. Thus, the concentration of para-xylene is effectively reduced to about 15000 ppm. wt. before being released into the bulk of the reaction medium in a bubble column. Subsequent mixing and chemical reactions proceed in accordance with the usual random events in a bubble column.

Example 12

This example relates to the use of streamlines external to the apparatus for transporting liquid in order to facilitate the initial distribution of incoming para-xylene. The reactor of this example has the same mechanical configuration and process boundary conditions as in Example 11, with the exceptions described below with reference to the circuit of FIG. 25. The internal current line is removed and replaced by the external current line. The line section connecting the reaction apparatus to the external deaeration section has an inner circle diameter of 0.30 m and is 20 m above the lower TL. The inner diameter of the circumference of the outer deaeration section is 1 m, and it is at a height of 2 m. The inner diameter of the streamline below the deaeration section is 0.20 m, which makes it possible to have a larger flow flowing at approximately the same hydrostatic head. A flow sensor and a flow controller are introduced into the current line to control the flow rate in the desired range. For example, flow control is set to allow 150 kg of slurry per second to be transported as much as estimated through the internal stream line of Example 11. A mixed liquid-phase stream of para-xylene and filter solvent enters the external stream line about 5 m above the bottom TL reaction apparatus. The output of the external current line is connected to the bottom of the reaction apparatus. Thus, the concentration of para-xylene is again effectively reduced to about 15,000 ppm. wt. before letting it into the bulk of the reaction medium in a bubbler column. Subsequent mixing and chemical reactions proceed in accordance with the usual random events in a bubble column. The withdrawal of the commodity suspension for further processing is carried out through the outlet in the specified streamline below the deaeration section and above the point of addition of the liquid-phase flow of para-xylene and filter solvent, thus avoiding the need for a separate system for removing and deaerating the suspension.

Example 13

This example relates to the use of streamlines consisting of both external and internal sections of the reaction apparatus to transport liquid to aid in the initial distribution of incoming para-xylene. This calculation example is identical to Example 12, except that the second branch in the outer streamline is about 3 m above the lower TL of the reaction apparatus, which is below the point of addition of the mixed liquid-phase para-xylene stream and filter solvent. The second branch of the current line also has an inner circular diameter of 0.20 m. A separate flow regulator is placed in the second branch of the current line, again to control the flow. The branch of the current channels passes through the side wall of the reaction apparatus 3 m above the lower TL, and the branch of the current line passes into the wall of the reaction apparatus by 0.4 m. Thus, the branch of the line contains sections that are both external and internal to the reaction apparatus. The stream can enter the reaction apparatus through any one or both of the line exits in the lower base or the line exits in the side wall in the internal line and in any ratio.

The invention has been described in detail with specific reference to its preferred embodiments, but it should be understood that changes and modifications can be made without departing from the spirit and scope of the invention.

Claims (37)

1. The method of liquid-phase catalytic oxidation of an aromatic compound, comprising the oxidation of para-xylene in the liquid phase of a multiphase reaction medium contained in one or more primary oxidation reactors, wherein said reaction medium contains a first separate 20% continuous volume having a first spatio-temporal oxygen velocity (STR oxygen), and a second separate 20% continuous volume having a second oxygen STR, wherein the ratio of said first oxygen STR to said second oxygen STR is at least at least about 1.5: 1. 2. The method according to claim 1, in which the ratio of the specified first STR oxygen to the specified second STR oxygen lies in the range from about 2: 1 to about 12: 1. 3. The method according to claim 1, wherein said first separate 20% continuous volume has a first oxygen concentration, and said second separate 20% continuous volume has a second oxygen concentration, said first and second oxygen concentrations being measured in the gas phase of said reaction medium as molar concentrations averaged over time and volume based on the wet state, wherein the ratio of said first oxygen concentration to said second oxygen concentration is at least about 1.5: 1. 4. The method according to claim 3, in which the ratio of the specified first oxygen concentration to the specified second oxygen concentration is in the range from about 2: 1 to about 12: 1. 5. The method according to claim 3, wherein said first oxygen concentration is from about 3 to about 18 mol%, and said second oxygen concentration is from about 0.3 to about 5 mol%. 6. The method according to claim 1, wherein said first and second separate 20% continuous volumes are in different reaction apparatuses. 7. The method according to claim 1, wherein said one or more primary oxidation reactors comprise a first reaction apparatus and a second reaction apparatus, said first and second reaction apparatuses comprising, respectively, a first and second fraction of said reaction medium. 8. The method according to claim 7, in which said first and second parts of the reaction medium have, respectively, a first and a second oxygen averaged over the STR apparatus, the ratio of said first oxygen averaged over STR apparatus to said second oxygen averaged over STR apparatus being at least about 2: 1. 9. The method of claim 8, wherein the ratio of said first oxygen averaged over apparatus STR to said second oxygen averaged over apparatus STR is in the range from about 3: 1 to about 15: 1. 10. The method according to claim 7, wherein said method further includes introducing a stream of an oxidizing agent containing molecular oxygen into said first reaction apparatus, wherein at least about 25 mol% of molecular oxygen leaving said first reaction apparatus, transferred to the specified second reaction apparatus. 11. The method of claim 10, wherein at least about 10 mol% of molecular oxygen transferred from said first reaction apparatus to said second reaction apparatus is consumed in said second reaction apparatus. 12. The method according to claim 1, wherein said first and second separate 20% continuous volumes are in the same reaction apparatus. 13. The method according to claim 1, wherein said one or more primary oxidation reactors comprise at least one reactor type bubbler column. 14. The method according to claim 1, wherein said one or more primary oxidation reactors comprise at least one mechanically stirred reactor. 15. The method according to claim 1, wherein said reaction medium is a three-phase reaction medium. 16. The method according to claim 1, wherein said oxidation causes at least about 10 wt.% Of said oxidizable compound to form solid particles in said reaction medium. 17. The method according to claim 1, in which the specified oxidation is carried out in the presence of a catalytic system containing cobalt. 18. The method according to 17, in which the specified catalytic system, in addition, includes bromine and manganese. 19. The method according to claim 1, wherein said oxidation in said one or more primary oxidation reactors leads to the formation of terephthalic acid in said reaction medium, said method further comprising oxidizing at least a portion of said terephthalic acid in a secondary oxidation reactor . 20. The method according to claim 19, in which the specified oxidation in the specified secondary oxidation reactor is carried out at an average temperature of at least about 10 ° C more than the specified oxidation in these primary oxidation reactors. 21. The method according to claim 19, in which the specified oxidation in the specified secondary oxidation reactor is carried out at an average temperature that is from about 20 to about 80 ° C more than the average temperature of these primary oxidation reactors, and the specified oxidation in these primary oxidation reactors at an average temperature in the range of from about 140 to about 180 ° C., said oxidation in said secondary oxidation reactor being carried out at an average temperature in the range of from about 180 to about 220 ° C. 22. The method according to claim 1, in which the specified oxidation leads to the formation of particles of technical terephthalic acid in the specified reaction medium, and a representative sample of these particles of technical terephthalic acid has one or more of the following characteristics:
(i) contains less than about 12 ppm. 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 at 340 nm (% T ₃₄₀₎ greater than about 25%. 23. The method according to claim 1, in which the specified oxidation leads to the formation of particles of technical terephthalic acid in the specified reaction medium, and a representative sample of these particles of technical terephthalic acid dissolves in THF to a concentration of at least about 500 ppm, when tested with a dissolution test over time. 24. The method according to claim 1, in which the specified oxidation leads to the formation of particles of technical terephthalic acid in the specified reaction medium, and a representative sample of these particles of technical terephthalic acid has a time constant "C" greater than about 0.5 min ^-1 , as determined by the described dissolution model over time. 25. The method according to claim 1, wherein said oxidation leads to the formation of particles of technical terephthalic acid in said reaction medium, wherein a representative sample of said particles of technical terephthalic acid has an average BET specific surface area of greater than about 0.6 m ² / g. 26. The method according to claim 1, in which the specified oxidation leads to the formation of particles of technical terephthalic acid in the specified reaction medium, and the specified representative sample of these particles of technical terephthalic acid has an average particle size in the range from about 20 to about 150 microns. 27. The method according to claim 1, in which the specified oxidation leads to the formation of particles of technical terephthalic acid in the specified reaction medium, and a representative sample of these particles of technical terephthalic acid has a particle size D (v, 0.9) in the range from about 30 to about 150 microns. 28. The method of liquid-phase catalytic oxidation of an aromatic compound, including
(a) introducing a stream of feed containing para-xylene into the reaction zone of a bubble column reactor-type column;
(b) introducing a stream of oxidizing agent containing molecular oxygen into said reaction zone; and
(c) oxidizing at least a portion of para-xylene in the liquid phase of the multiphase reaction medium contained in said reaction zone of the primary oxidation reactor, wherein said oxidation is carried out so that if said reaction zone is theoretically divided into 30 horizontal layers of the same volume, one horizontal volume O ₂ -max has the maximum oxygen concentration of all of said 30 horizontal slices and one horizontal layer O ₂ -min has the minimum oxygen concentration of all the horizontal slices disintegrations Proposition above a horizontal layer of said O ₂ -max, wherein said oxygen concentration measured in the gas phase of said reaction medium as a time-averaged and volume molar concentration based on the wet state, wherein the ratio of the concentration of oxygen in said horizontal layer for O ₂ -max the oxygen concentration in the specified horizontal layer of O ₂ -min is at least about 2: 1. 29. The method according to p, in which the ratio of the oxygen concentration in the specified horizontal layer of O ₂ -max to the oxygen concentration in the specified horizontal layer of O ₂ -min is in the range from about 3: 1 to about 15: 1. 30. The method according to p, in which the specified oxidizable compound contains para-xylene, and the specified oxidation is carried out so that if the specified reaction zone is theoretically divided into 30 horizontal layers of equal volume, the horizontal layer pX-max has a maximum concentration of para-xylene from all of these 30 horizontal layers, and the horizontal pX-min layer has a minimum concentration of para-xylene from all horizontal layers above the specified horizontal pX-max layer, the indicated concentration of para-xylene from heren in the liquid phase of the specified reaction medium as a mass concentration averaged over time and volume, the ratio of the concentration of para-xylene in the specified horizontal layer pX-max to the concentration of para-xylene in the specified horizontal layer pX-min is at least about 5: 1 . 31. The method of claim 28, wherein said reaction medium has a maximum width (W), a maximum height (H), and an H: W ratio of at least about 3: 1. 32. The method of claim 31, wherein said H: W ratio is at least about 6: 1. 33. The method according to p, in which the main part of the specified molecular oxygen is included in the specified reaction zone within about 0.25 W from the bottom of the specified reaction zone. 34. The method according to p, in which at least about 50 wt.% The specified oxidizable compounds are included in the specified reaction zone within about 2.5 W from the lowest point where the specified molecular oxygen is introduced into the specified reaction zone. 35. The method according to p, in which the specified flow of raw materials is introduced into the specified reaction zone through several loading holes, and at least two of these loading holes are separated from each other by a height of at least about 0.5W. 36. The method according to p, in which at least about 10 wt.% The specified oxidizable compounds in the specified reaction medium form a solid phase. 37. The method according to p, in which the specified oxidation is carried out in the presence of a catalytic system containing cobalt, bromine and manganese.

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