JP2008511653A - Optimized production of aromatic dicarboxylic acids - Google Patents

Abstract

  An optimized method and apparatus for more efficiently producing aromatic dicarboxylic acids (eg, terephthalic acid) is disclosed. In one embodiment, the method / apparatus reduces costs by recovering and purifying residual terephthalic acid present in the liquid phase of the initial oxidation slurry. In another embodiment, the process apparatus reduces the costs associated with hydrogenation by forming a final composite product containing non-hydrogenated acid particles.

Description

  The present invention relates generally to the production of aromatic dicarboxylic acids, such as terephthalic acid. One aspect of the present invention relates to a more efficient method and apparatus for producing aromatic dicarboxylic acids. Another aspect of the invention relates to a method and apparatus for controlling the purity of an aromatic dicarboxylic acid product.

  Terephthalic acid (TPA) is one of the basic building blocks in the production of linear polyester resins used in the production of polyester films, packaging materials and bottles. The TPA used in the production of such polyester resins must meet certain minimum purity requirements.

  The purified state of TPA is mainly due to the absence of significant concentrations of 4-carboxybenzaldehyde (4-CBA) and para-toluic acid (p-TAc) present in significant amounts in commercially available crude grade TPA. Refers to existence. 4-CBA and p-TAc are both partial oxidation products formed during the production of TPA by catalytic oxidation of para-xylene. The purified form of TPA also refers to the absence of color bodies that give the crude material a characteristic yellow phase. This colored body is an aromatic compound having a structure of benzyl, fluorenone and / or anthraquinone. 4-CBA and p-TAc are particularly detrimental to the polymerization process because they act as chain terminators during the condensation reaction between TPA and ethylene glycol in the production of polyethylene terephthalate (PET). .

  In a typical process for producing TPA, the slurry is removed from the primary oxidation reactor. This slurry contains a liquid mother liquor and solid crude terephthalic acid (CTA) particles. CTA particles are typically separated from a liquid mother liquor and then subjected to purification to produce purified terephthalic acid (PTA). The separated mother liquor is typically treated to remove waste material and then recycled to the primary oxidation reactor. Most of the TPA produced in the primary oxidation reactor is present as solid CTA particles, while some of the TPA produced in the primary oxidation reactor is present in the liquid mother liquor. In a typical process, TPA in a liquid mother liquor will show a potential yield drop if it is recovered and not purified.

  One common method of purifying CTA to produce PTA is by hydrotreating, in which case 4-CBA is hydrogenated to p-TAc and the color is a colorless solid. Hydrogenated to compound. In order to carry out purification by hydrogenation, solid CTA particles are typically dissolved in a solvent (eg water) and the resulting solution is subjected to liquid phase hydrogenation in the presence of a hydrogenation catalyst. . While effective in reducing yellowness, purification of CTA by hydrogenation can be expensive because it consumes energy, hydrogen, water and catalyst. Therefore, from the standpoint of operating costs, it is desirable to minimize the amount of hydrogenation required to produce a suitably pure TPA solid.

  One embodiment of the present invention provides a slurry comprising the following steps: (a) a solid phase containing a first amount of terephthalic acid and a liquid phase containing a second amount of terephthalic acid; b) producing an oxidized terephthalic acid by subjecting at least a portion of the first amount of terephthalic acid to an oxidation treatment; and (c) subjecting at least a portion of the second amount of terephthalic acid to a hydrotreatment. To a process comprising the step of producing hydrotreated terephthalic acid.

  Another aspect of the present invention is to prepare the initial terephthalic acid by the following steps: (a) oxidizing para-xylene in the initial oxidation reactor, and (b) first terephthalic acid first. The hydrotreated terephthalic acid was produced by subjecting a portion of the mixture to hydrogenation, and (c) at least a portion of the hydrotreated terephthalic acid was not subjected to the hydrogenation of step (b). The method involves the step of combining with unhydrogenated terephthalic acid, wherein the unhydrogenated terephthalic acid is derived from the second part of the initial terephthalic acid.

  Yet another aspect of the present invention provides the following steps: (a) introducing a liquid phase medium comprising mother liquor and terephthalic acid into the evaporation zone; and (b) at least part of the mother liquor from the liquid phase medium. Evaporating to form a non-evaporated portion of the mother liquor and a concentrated medium containing substantially all of the terephthalic acid, and (c) replacing at least a portion of the non-evaporated mother liquor with a washing medium, Obtaining a hydrogenation medium comprising the washing medium and terephthalic acid, and (d) forming a hydrotreated medium by subjecting at least a portion of the hydrogenated medium to a hydrotreatment. .

  Yet another aspect of the invention relates to an apparatus comprising an initial oxidation reactor, an optional secondary oxidation reactor, a solid / liquid separator, a hydrogenation system and a combined zone. The first oxidation reactor has an initial reactor outlet. An optional secondary oxidation reactor has a secondary reactor inlet and a secondary reactor outlet. The secondary reactor inlet is connected in communication with the first reactor outlet. The solid / liquid separator has a separator inlet, a separated solid outlet, and a separated liquid outlet. The separator inlet is coupled in communication with the first reactor outlet and / or the secondary reactor outlet. The hydrogenation system has a hydrogenation system inlet and a hydrogenation system outlet. The hydrogenation system inlet is coupled in communication with a separated liquid outlet. The combination zone has a hydrogenated solid inlet, a non-hydrogenated solid inlet, and a composite solid outlet. The hydrogenated solid inlet is coupled in communication with the hydrogenation system outlet, and the non-hydrogenated solid inlet is coupled in communication with the separated solid outlet.

  FIG. 1 illustrates an embodiment of the invention in which terephthalic acid (TPA) produced in a primary / first oxidation reactor 10 is purified by both oxidation and hydrotreatment. In the first step of the embodiment shown in FIG. 1, a primarily liquid phase feed stream containing an oxidizing compound (eg para-xylene), a solvent (eg acetic acid + water) and a catalyst system (eg Co + Mn + Br) is subjected to primary oxidation. Introduce into the reactor 10. Also, a primarily gas phase oxidant stream containing molecular oxygen is introduced into the primary oxidation reactor 10. The liquid phase supply stream and the gas phase supply stream form a multiphase reaction medium in the oxidation reactor 10. The oxidizing compound undergoes partial oxidation in the liquid phase of the reaction medium contained in the reactor 10.

  The primary oxidation reactor 10 is preferably a stirred reactor. Agitation of the reaction medium in the oxidation reactor 10 can be provided by any means known in the art. The term “stirring” as used herein is intended to indicate the work that is dissipated into the reaction medium to cause fluid flow and / or mixing. In one embodiment, the primary oxidation reactor 10 is a mechanically stirred reactor equipped with means for mechanically stirring the reaction medium. As used herein, the term “mechanical agitation” is intended to indicate agitation of the reaction medium caused by physical movement of the hard or soft element (s) relative to or within the reaction medium. For example, mechanical agitation can be provided by rotation, oscillation and / or vibration of an internal agitator, paddle, vibrator or acoustic diaphragm located in the reaction medium. In a preferred embodiment of the present invention, the primary oxidation reactor 10 is a bubble column reactor. As used herein, the term “bubble column reactor” is intended to indicate a reactor for easily carrying out a chemical reaction in a multiphase reaction medium, in which case the stirring of the reaction medium is mainly performed by the reaction medium. Given by the upward movement of bubbles passing through. As used herein, the terms “major”, “primarily” and “predominantly” shall mean more than 50%.

  The oxidizing compound present in the liquid phase feed stream introduced into the primary oxidation reactor 10 preferably contains at least one hydrocarbyl group. More preferably, the oxidizing compound is an aromatic compound. Even more preferably, the oxidizing compound comprises at least one bonded hydrocarbyl group or at least one bonded substituted hydrocarbyl group or at least one bonded heteroatom or at least one bonded carboxylic acid functional group. It is an aromatic compound having (—COOH). Even more preferably, the oxidizing compound is an aromatic compound having at least one bonded hydrocarbyl group or at least one bonded substituted hydrocarbyl group, each bonded group having 1 to 5 carbon atoms. is there. Even more preferably, the oxidizing compound comprises that each bonded group contains exactly one carbon atom and contains a methyl group and / or a substituted methyl group and / or at most one carboxylic acid group. It is an aromatic compound containing exactly two bonded groups. Even more preferably, the oxidizing compound is para-xylene, meta-xylene, para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluic acid and / or acetaldehyde. Most preferably, the oxidizing compound is para-xylene.

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

  The amount of oxidizing compound present in the liquid phase feed stream introduced into the primary oxidation reactor 10 is preferably in the range of about 2 to about 40% by weight, more preferably about 4 to about 20% by weight. Within the range, most preferably within the range of 6-15% by weight.

  The solvent present in the liquid phase feed stream introduced into the primary oxidation reactor 10 preferably includes an acid component and a water component. The solvent is preferably in the liquid phase feed stream in the range of about 60 to about 98% by weight, more preferably in the range of about 80 to about 96% by weight, and most preferably in the range of 85 to 94% by weight. Present at a concentration of. The acid component of the solvent is preferably an organic low molecular weight monocarboxylic acid having 1 to 6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the solvent is acetic acid. Preferably, the acid component comprises at least about 75% by weight of the solvent, more preferably at least about 80% by weight of the solvent, most preferably 85-98% by weight of the solvent, with the balance being water.

  The liquid phase feed stream introduced into the primary oxidation reactor 10 may also contain a catalyst system. The catalyst system is preferably a homogeneous liquid phase catalyst system that can promote partial oxidation of the oxidizing compound. More preferably, the catalyst system consists of at least one polyvalent transition metal. Even more preferably, the multivalent transition metal comprises cobalt. Even more preferably, the catalyst system includes cobalt and bromine. Most preferably, the catalyst system comprises cobalt, bromine and manganese.

  When cobalt is present in the catalyst system, for the amount of cobalt present in the liquid phase feed stream, the concentration of cobalt in the liquid phase of the reaction medium is about 300 to about 6,000 ppmw (parts per million by weight). ), More preferably within the range of about 700 to about 4,200 ppmw, and most preferably within the range of 1,200 to 3,000 ppmw. When bromine is present in the catalyst system, the concentration of bromine in the liquid phase of the reaction medium is within the range of about 300 to about 5,000 ppmw, more preferably about about the amount of bromine present in the liquid phase feed stream. It is preferred that it be maintained in the range of 600 to about 4,000 ppmw, most preferably in the range of 900 to 3,000 ppmw. When manganese is present in the catalyst system, the concentration of manganese in the liquid phase of the reaction medium is within the range of about 20 to about 1,000 ppmw, more preferably about about the amount of manganese present in the liquid phase feed stream. It is preferred that it be maintained in the range of 40 to about 500 ppmw, most preferably in the range of 50 to 200 ppmw.

  The weight ratio of cobalt to bromine (Co: Br) in the catalyst system introduced into the primary oxidation reactor 10 is preferably in the range of about 0.25: 1 to about 4: 1, more preferably about 0. Within the range of 5: 1 to about 3: 1 and most preferably within the range of 0.75: 1 to 2: 1. The weight ratio of cobalt to manganese (Co: Mn) in the catalyst system introduced is preferably in the range of about 0.3: 1 to about 40: 1, more preferably about 5: 1 to about 30: 1. Within the range, most preferably within the range of 10: 1 to 25: 1.

  During oxidation, an oxidizing compound (e.g., para-xylene) is fed into the primary oxidation reactor 10 at a rate of at least about 5,000 kg / hr, more preferably from about 10,000 to about 80,000 kg / hr. It is preferable to introduce continuously at a rate within the range of 20,000 to 50,000 kg / hr, most preferably within a range of 20,000 to 50,000 kg / hr. The ratio of the solvent mass flow rate to the oxidizable compound mass flow rate that enters the oxidation reactor 10 during oxidation is in the range of about 2: 1 to about 50: 1, more preferably about 5: 1. It is preferred to maintain in the range of about 40: 1, most preferably in the range of 7.5: 1 to 25: 1.

  The predominant gas phase oxidant stream introduced into the primary oxidation reactor 10 is in the range of about 5 to about 40 mole percent molecular oxygen, more preferably about 15 to about 30 mole percent. Of molecular oxygen, most preferably in the range of 18 to 24 mol%. The remainder of the oxidant stream preferably mainly comprises a gas or gas group, for example nitrogen that is inert to oxidation. More preferably, the oxidant stream consists essentially of molecular oxygen and nitrogen. Most preferably, the oxidant stream is dry air consisting of about 21 mole percent molecular oxygen and about 78 to about 81 mole percent nitrogen. In an alternative embodiment of the present invention, the oxidant stream may consist of substantially pure oxygen.

  During the liquid phase oxidation in the primary oxidation reactor 10, an oxidant stream is introduced into the reactor 10 in such an amount as to give molecular oxygen somewhat above the stoichiometric oxygen demand. Is preferred. Accordingly, the ratio of the mass flow rate of the oxidant stream (eg, air) entering reactor 10 to the mass flow rate of the oxidizing compound (eg, para-xylene) ranges from about 0.5: 1 to about 20: 1. Preferably within the range of about 1: 1 to about 10: 1, most preferably within the range of 2: 1 to 6: 1.

  The liquid phase oxidation reaction performed in the reactor 10 is preferably a precipitation reaction that produces a solid. More preferably, the liquid phase oxidation reaction performed in the reactor 10 is such that at least about 10% by weight of the oxidizing compound (eg, para-xylene) introduced into the oxidation reactor 10 is a solid ( For example, CTA particles) are formed. Even more preferably, this liquid phase oxidation causes at least about 50% by weight of the oxidizing compound to form a solid in the reaction medium. Most preferably, this liquid phase oxidation causes at least 90% by weight of the oxidizing compound to form a solid in the reaction medium. The total amount of solids in the reaction medium is maintained in the range of about 5 to about 40% by weight, still more preferably in the range of about 10 to about 35% by weight, and most preferably in the range of 15 to 30% by weight. Is preferred.

  During oxidation in oxidation reactor 10, the multiphase reaction medium is preferably in the range of about 125 to about 225 ° C, more preferably in the range of about 150 to about 180 ° C, and most preferably 155 to 165 ° C. Maintain elevated temperature within the range of. The overhead pressure in the oxidation reactor 10 is preferably in the range of about 1 to about 20 bar atmosphere, more preferably in the range of about 3.5 to about 13 bara, most preferably 5.2 to 6.9 bara. Keep within the range.

  As shown in FIG. 1, the crude product slurry is removed from the outlet of the primary oxidation reactor 10 via line 12. The solid phase of the crude product slurry in line 12 is mainly formed from solid particles of crude terephthalic acid (CTA). The liquid phase of the crude product slurry in line 12 is a liquid mother liquor containing at least a portion of the solvent, the catalyst system and a small amount of dissolved TPA. The amount of TPA present in the mother liquor is preferably less than about 10% by weight of the total TPA present in the crude product slurry exiting the primary oxidation reactor 10, more preferably about 0.1% of the total TPA. To about 5% by weight, and most preferably 0.5 to 3% by weight of total TPA. The solid content of the crude product slurry in line 12 is preferably the same as the solid content of the reaction medium in the primary oxidation reactor 10.

  In one embodiment of the invention, the crude product slurry in line 12 is transferred directly to secondary oxidation reactor 14 for purification by oxidation treatment. In an alternative embodiment, an optional liquid removal system 16 is used to remove at least a portion of the liquid mother liquor from the crude product slurry prior to introduction into the secondary oxidation reactor 14. The liquid removal system 16 can use a variety of different devices for removing / separating at least a portion of the mother liquor from the crude product slurry in line 12. For example, the liquid removal system 16 is a liquid exchange system that uses a clean replacement solvent to separate at least a portion of the mother liquor from the crude product slurry and replace at least a portion of the removed mother liquor. Good. Separation of the mother liquor from the solid in the liquid removal system 16 can be accomplished using a suitable solid / liquid separator such as a decanter centrifuge, a rotary disc centrifuge, a belt filter or a rotary vacuum filter. When using the liquid removal system 16, the removed mother liquor passes through line 18 for further processing (described in detail below) while the resulting crude product slurry passes through line 20. Then, the process proceeds to the secondary oxidation reactor 14.

  In the secondary oxidation reactor 14, the crude product slurry is subjected to purification by oxidation treatment. The secondary oxidation reactor 14 is preferably a stirred reactor, most preferably a mechanical stirred reactor. A secondary oxidant stream is fed to the secondary oxidation reactor 14 to supply the molecular oxygen necessary for secondary oxidation. If necessary, additional catalyst can be added. The crude product slurry introduced into the secondary oxidation reactor 14 contains significant amounts of impurities such as 4-carboxybenzaldehyde (4-CBA) and para-toluic acid (p-TAc). . Oxidation treatment in secondary oxidation reactor 14 preferably causes oxidation of a substantial portion of 4-CBA and p-TAc to TPA.

  The temperature at which the oxidation takes place in the secondary oxidation reactor 14 is preferably at least about 10 ° C., more preferably in the range of about 20 to about 80 ° C., higher than the temperature of the oxidation in the primary oxidation reactor 10. Most preferably, it is higher in the range of 30 to 50 ° C. The additional heat required for operation of the secondary oxidation reactor 14 can be provided by feeding the evaporated solvent to the secondary oxidation reactor and condensing the evaporated solvent therein. The oxidation temperature in the secondary oxidation reactor is preferably maintained in the range of about 175 to about 250 ° C, more preferably in the range of about 185 to about 230 ° C, and most preferably in the range of 195 to 210 ° C. The oxidation pressure in the secondary oxidation reactor 14 is preferably in the range of about 2 to about 30 bara, more preferably in the range of about 4.5 to about 18.3 bara, most preferably 13.4 to 17.2 bara. Maintained within range.

  The oxidized slurry is discharged from the secondary oxidation reactor 14 via line 22. The solid phase of the oxidized slurry is mainly formed from purified terephthalic acid (PTA) particles, while the liquid phase is formed from an oxidized mother liquor. The solid content of the oxidized slurry in line 22 is preferably in the same range as previously disclosed for the solid content of the crude product slurry in line 12.

  The oxidized slurry in line 22 is transferred to solid recovery system 24 for removal of the oxidized mother liquor and recovery of solid PTA particles. The solid recovery system preferably includes at least one solid / liquid separator and at least one dryer. The solid / liquid separator used as part of the solid recovery system 24 may be any common solid / liquid separator, such as a decanter centrifuge, a rotary disc centrifuge, a belt filter or a rotary vacuum filter. . The solid separated in the solid / liquid separator can then be dried using any suitable dryer known in the art. The recovered and dried PTA solids are discharged from the solid recovery system 24 via line 26. The separated and oxidized mother liquor is discharged from the solid recovery system 24 via line 28.

  The separated mother liquor in line 28, if present, can be combined with any separated mother liquor in line 18. The combined mother liquor stream in line 30 containing a small amount of residual TPA can then be directed to a residual TPA recovery system 32. The recovery of residual TPA preferably includes at least one evaporator 34 and solid / liquid separator 36. The evaporator 34 can be operated to remove a substantial portion of the solvent (eg, acetic acid and water) from the mother liquor. The evaporated solvent is discharged from the evaporator 34 via a line 38. Preferably, the evaporator 34 includes a first evaporator zone that operates above atmospheric pressure (eg, 1-10 atmospheres) and a second evaporator zone that operates under vacuum conditions. The second evaporator zone is preferably maintained at a temperature in the range of about 10 to about 100 ° C, more preferably in the range of about 20 to about 70 ° C, and most preferably in the range of 30 to 50 ° C. Preferably, at least about 25% by weight of the mother liquor introduced into the evaporator 34 evaporates and is discharged via line 38, more preferably at least about 50% by weight of the mother liquor introduced into the evaporator 34. Evaporate and discharge through line 38, and most preferably, 75 to 99% by weight of the mother liquor introduced into evaporator 34 is evaporated and discharged through line 38.

  The concentrated slurry is discharged from the evaporator 34 via line 40. The concentrated slurry in line 40 preferably contains greater than about 10% by weight solid TPA particles and less than 90% by weight liquid. The concentrated slurry in line 40 is less than about 10% by weight of total TPA discharged from the primary oxidation reactor via line 12, more preferably from about 0.1 to about 5% by weight of total TPA from reactor 10. % Range, most preferably in the range of 0.5 to 3% by weight of total TPA from reactor 10.

  The concentrated slurry in line 40 is introduced into solid / liquid separator 36 where substantially all of the remaining liquid mother liquor is removed from the concentrated slurry. The removed mother liquor is discharged from the solid / liquid separator 36 via line 42 and subsequently combined with the evaporated mother liquor in line 38. The combined mother liquor in line 44 can then be recycled to the primary oxidation reactor 10 in combination with the liquid phase feed stream introduced into the reactor 10, although all of the combined mother liquor Some can be removed from the process.

  The solid / liquid separator 36 preferably uses a rotary drum pressure filter device similar to the device shown in FIG. The rotary drum pressure filter of FIG. 2 includes a housing 50 and a rotary drum filter 52 that is rotatably disposed within the housing 50. An annular portion is defined between the inside of the housing 50 and the outside of the rotary drum filter 52. This annular portion is divided into various separated zones by seals 54a, 54b, 54c, 54d, 54e, 54f. Filtration zone 56 is defined in an annular portion between seals 54a and 54b. A cleaning zone 58 is defined in the annular portion between the seals 54b and 54e. A drying / dehydrating zone 60 is defined in the annular portion between the seals 54e and 54f. The housing 50 is open between the seals 54f and 54a. This open portion of the housing 50 includes a drain zone 62 and a fabric wash zone 64.

  Referring again to FIG. 2, the rotary drum filter 52 defines a plurality of filter cells 66 disposed on the periphery of the drum. The bottom of each filter cell 66 is formed from a filter media (eg, synthetic fabric, single layer metal or multilayer metal). Liquid flow through the filter media is caused by creating a pressure differential across the filter media. Each filter cell 66 has its own outlet for discharging liquid inwardly towards the axis of rotation of the rotary drum filter 52. The outlets of the filter cells 66 arranged in the axial direction are collected by a manifold. A manifold (not shown) rotates with the rotary drum filter 52 and draws liquid from the manifold in a manner that allows the liquid discharged from the zones 56, 58 and 60 to be kept separate. In communication with a collecting service / control head (not shown).

  Housing 50 defines a concentrated slurry inlet 68 in communication with filtration zone 56, a wash feed inlet 70 in communication with wash zone 58, and a dry gas inlet 72 in communication with drying / dehydration zone 60. ing. The cleaning zone 58 is divided into an initial cleaning zone 74, an intermediate cleaning zone 76, and a final cleaning zone 78 by seals 54c and 54d. The housing 50 and the rotary drum filter 52 allow filtrate discharged from the first wash zone 74 to enter the intermediate wash zone 76 and filtrate discharged from the intermediate wash zone 76 to the last wash zone. 78 is arranged to allow entry.

  In operation, the concentrated slurry in line 40 enters filtration zone 56 via slurry inlet 68 and forms a filter cake 80 in filter cell 66 on the periphery of rotary filter drum 52. Within the filtration zone 56, the liquid mother liquor is discharged radially inward from the bottom of each filter cell 66. The mother liquor collected from the filtration zone 56 can be discharged from this device via line 42. Upon obtaining a preferred height filter cake 80 within the filtration zone 56, the rotary drum filter 52 rotates and the filter cake 80 enters the wash zone 58.

  In the wash zone 58, the filter cake 80 is washed with a wash feed that enters the first wash zone 74 via the wash feed inlet 70. The cleaning feed is preferably formed primarily from water. Most preferably, the wash feed consists essentially of water. The wash filtrate from the first wash zone 74 is then transferred to the intermediate wash zone 76, and the wash filtrate from the intermediate wash zone 76 is transferred to the last wash zone 78. This washing filtrate can then be drained from the device via line 84. In one embodiment of the invention, the wash filtrate in line 84 is combined into the filtered mother liquor in line 42. After proper cleaning in the cleaning zone 58, the rotary drum filter 52 rotates so that the filter cake 80 enters the drying / dehydration zone 60.

  In the drying / dehydration zone 60, liquid is removed from the washed filter cake 80 by entering the gas inlet 72 and passing the dried gas through the washed filter cake 80. The gas stream passing through the cleaned filter cake 80 exits the apparatus as wet steam via line 85. After the filter cake 80 is dried / dehydrated in the zone 60, the rotary drum filter 52 rotates so that the dried filter cake 80 enters the discharge zone 62.

  In the discharge zone 62, the filter cake 80 leaves the rotary drum filter 52 and exits the device via line 86. The rotary filter drum 52 then rotates into the fabric wash zone 64 where any solid particles remaining in the filter cell 66 are removed.

  A suitable pressure filter that can be used as the solid / liquid separator 36 is BHS-FEST®, available from BHS-WERK, Weston Sonthofen, Sonthofen, D-8972. However, other pressure filters known in the art can perform the functions required by the solid / liquid separator 36 of FIG. Examples of other suitable devices include, for example, belt filters, filter presses, centrifuges, pressure leaf filters and cross flow filters. In one embodiment of the present invention, a solid / liquid separator is disclosed in US patent application Ser. No. 10 / 874,419 filed Jun. 23, 2004 (herein incorporated by reference in its entirety). Substantially the same shape and operating parameters as the pressure filter described in

  Referring again to FIG. 1, the solids exiting solid / liquid separator 36 via conduit 86 preferably remain from the liquid phase of the slurry exiting primary oxidation reactor 10 and / or secondary oxidation reactor 14. It is a TPA solid. These residual TPA particles in line 86 are subjected to purification by hydroprocessing in hydrogenation system 88. The hydrogenation system 88 can include one or more vessels / zones. Preferably, the hydrogenation zone 88 includes an initial dissolution zone / vessel where residual TPA solids are combined with a solvent (preferably water) at an elevated temperature, so that the residual TPA solids into the solvent. Causes dissolution. The solvent and residual TPA particles are preferably in the range of about 0.5: 1 to about 50: 1, more preferably in the range of about 1: 1 to 10: 1, most preferably 1.5: 1 to 5: 1. Combined at a solvent to TPA weight ratio within the range of 1: 1.

After the residual TPA particles are dissolved in the solvent, the resulting solution is introduced into the hydrogenation zone / vessel of the hydrogenation system 88, where the solution is hydrogenated of certain impurities present therein ( For example, contacting with hydrogen and a hydrogenation catalyst under conditions sufficient to cause 4-CBA to p-TAc and / or hydrogenation of floureneones to flourenes). In a preferred embodiment of the invention, the hydrotreatment is at a temperature in the range of about 200 to about 400 ° C, more preferably in the range of about 250 to about 350 ° C, most preferably in the range of 260 to 320 ° C. carry out. The pressure in the hydrogenation zone / vessel is preferably maintained in the range of about 5 to about 200 bara, more preferably in the range of about 10 to about 150 bara, and most preferably in the range of 46.9 to 113 bara. The average space velocity for hydrogenation is preferably in the range of about 150 to about 2,500 solution kg / hr / catalyst bed cubic meter (kg / hr / m ³ ), more preferably about 300 to about 1,500 kg / hr. within the scope of the hr / m ^3, most preferably maintained within the range of 450~850kg / hr / m ^3. The molar ratio of hydrogen fed to the hydrogenation zone / vessel to residual TPA fed to the hydrogenation zone / vessel is preferably within the range of about 5: 1 to about 500: 1, more preferably about 10: 1 to about Within the range of 300: 1, most preferably within the range of 20: 1 to 250: 1. The hydrogenation catalyst used in the hydrogenation zone / vessel is preferably a noble Group VIII metal on a common catalyst support material.

  After hydroprocessing in the hydrogenation system 88, the resulting hydrotreated solution is transferred in line 89 to the crystallization system 90, where it is in at least one crystallizer. Is subjected to crystallization. Within the crystallization system 90, the temperature of the hydrogenated solution is in the range of about 100 to about 200 ° C, more preferably in the range of about 120 to about 185 ° C, and most preferably in the range of 140 to 175 ° C. The temperature is lowered to the crystallization temperature. Due to the reduced temperature in the crystallization system 90, substantially all of the TPA dissolved in the hydrotreated solution undergoes crystallization, thereby purifying / hydrogenated terephthalic acid (ie PTA). ) Solid particles are formed.

  The biphasic (slurry) effluent from the crystallization system 90 is transferred in line 91 to a solid / liquid separator 92 for separation of the solid and liquid portions. The separated liquid portion (ie, hydrogenated mother liquor) is transferred in line 93 for further processing. The separated solid PTA from separator 92 is transferred for drying in line 94 in one or more common dryers 95. The resulting dried hydrotreated PTA particles are transferred within line 96 to a combination zone 97 where at least a portion of the hydrotreated PTA particles from line 96 are oxidized from line 26. Combine with at least a portion of the treated PTA.

  A composite PTA containing oxidized PTA and hydrotreated PTA is produced from combination zone 97 via line 98. The weight ratio of oxidized PTA to hydrotreated PTA in the composite PTA produced from combination zone 97 is preferably in the range of about 10: 1 to about 1,000: 1, more preferably about Within the range of 20: 1 to about 500: 1, most preferably within the range of 50: 1 to 250: 1. In one embodiment of the present invention, substantially all of the TPA present in the solid phase of the slurry exiting the secondary oxidation reactor 14 via line 22 is subsequently introduced into the combination zone 97, while Substantially all of the TPA present in the hydrotreated solution exiting the hydrogenation system 88 via line 89 is subsequently introduced into the combination zone 97. In such embodiments, the weight ratio of the oxidized PTA particles exiting the secondary oxidation reactor 14 to the hydrotreated PTA exiting the hydrogenation system 88 is listed above for the final composite PTA product. The weight ratio of oxidized PTA to hydrogenated PTA is the same.

  FIG. 3 illustrates an embodiment of the present invention, wherein a first portion of the initial TPA produced in the oxidation reactor 100 is subjected to hydroprocessing and the TPA produced in the oxidation reactor 100 is treated. The second part is not hydrotreated and the composite TPA product is a hydrotreated TPA (derived from the first part of the first TPA) and a non-hydrotreated TPA (second part of the TPA). In combination).

  As shown in FIG. 3, in the first step of the process, a primarily liquid phase feed stream containing an oxidizing compound (eg para-xylene), a solvent (eg acetic acid + water) and a catalyst system (eg Co + Mn + Br) And introduced into the oxidation reactor 100. A primarily gas phase oxidant stream containing molecular oxygen is also introduced into the reactor 100. The liquid phase supply stream and the gas phase supply stream form a multiphase reaction medium in the reactor 100. The oxidizing compound undergoes partial oxidation in the liquid phase of the reaction medium contained within the reactor 100.

  The oxidation reactor 100 is preferably a stirred reactor. Agitation of the reaction medium in the oxidation reactor 100 can be provided by any means known in the art. In a preferred embodiment of the present invention, the oxidation reactor 100 is a mechanically stirred reactor (eg, a continuous stirred tank reactor) fitted with means for mechanically stirring the reaction medium. In an alternative embodiment of the present invention, the oxidation reactor 100 is a bubble column reactor.

  The liquid phase feed stream and the gas phase oxidant stream introduced into the oxidation reactor 100 of FIG. 3 are preferably the liquid phase feed stream and the gas phase feed stream introduced into the primary oxidation reactor 10 of FIG. Is substantially the same. Further, the oxidation reactor 100 of FIG. 3 preferably operates in substantially the same manner as described above with reference to the primary oxidation reactor 10 of FIG. However, when the oxidation reactor 100 of FIG. 3 is a mechanically stirred reactor, the multiphase reaction medium in the oxidation reactor 100 is in the range of about 150 to about 300 ° C., more preferably about 175 to about 250. It is preferred to maintain the elevated temperature within the range of ° C, most preferably within the range of 190-225 ° C. The overhead pressure in the oxidation reactor 100 is preferably maintained in the range of about 1 to about 20 bar gauge, more preferably in the range of about 2 to about 12 barg, and most preferably in the range of 4 to 8 barg. .

  As shown in FIG. 3, a slurry containing solid particles of a crude oxidation product (for example, CTA) is taken out from the outlet of the oxidation reactor 100. The solids content of the removed slurry is preferably in the range of about 5 to about 40% by weight, even more preferably in the range of about 10 to about 35% by weight, and most preferably in the range of 15 to 30% by weight. is there. The slurry removed from reactor 100 is introduced into solid / liquid separator 102 where the slurry is subjected to solid / liquid separation. Separator 102 may be any common solid / liquid separation means including, for example, a decanter centrifuge, a rotary disc centrifuge, a belt filter, or a rotary vacuum filter.

  The liquid mother liquor discharged through the liquid outlet of the solid / liquid separator 102 is introduced into the catalyst recovery system 104. A liquid mother liquor typically comprises primarily a solvent and a catalyst system. However, this mother liquor may also contain undesirable corrosion / contamination metals such as iron, nickel and chromium and undesirable organic reaction products (formed over time). The catalyst recovery system 104 uses a general method for removing a substantial portion of undesirable components present in the liquid mother liquor. As shown in FIG. 3, the resulting clear liquid stream can be combined with a liquid phase feed stream that is introduced into the oxidation reactor 100.

The crude acid solid (e.g., CTA) discharged from the solid outlet of the solid / liquid separator 102 is typically in the form of a cake wet with solvent. One or more dryers 103 can optionally be used to evaporate the remaining solvent. The CTA typically has a 4-CBA content greater than about 600 ppmw (parts per million by weight). More typically, the 4-CBA content of the crude acid solid is in the range of about 700 to about 10,000 ppmw, most typically in the range of 800 to 7,000 ppmw. Typically, the crude acid solid has a p-TAc content greater than about 150 ppmw. More typically, the p-TAc content of the crude acid solid is in the range of about 175 to about 5,000 ppmw, most typically in the range of 200 to 1,500 ppmw. Typically, the crude acid solid has a total content of 4-CBA + p-TAc greater than about 700 ppmw. More typically, the total content of 4-CBA + p-TAc of the crude acid solid is in the range of about 850 to about 10,000 ppmw, most typically in the range of 1,000 to 5,000 ppmw. Typically, the crude acid solid has a B ^* value of at least 3, more typically in the range of about 3.5 to about 10, and most typically 4-8.

  Referring again to FIG. 3, the crude acid solid (eg, CTA) discharged from the solid / liquid separator 102 (or optional dryer 103) is introduced into the splitter 105, where the solid is Divide into a first part and a second part. The splitter 105 may be any common means for separating solids. The first portion of the crude acid solid exits from the first outlet of the splitter 105 and is subsequently subjected to purification in the hydrogenation system 106. The second portion of the crude acid solid exits the second outlet of splitter 105 and is not subjected to hydroprocessing. At least about 1% by weight of the crude acid solid (eg, CTA) produced in the oxidation reactor 100 exits the second outlet of the splitter 105 and is not subjected to hydroprocessing, more preferably about about 1% of the crude acid solid. Within the range of 3 to about 60% by weight is not subjected to hydroprocessing, most preferably within the range of 5 to about 40% by weight of the crude acid solid is not subjected to hydroprocessing. Further, the weight ratio of the second portion of the crude acid solid (not subjected to hydrogenation) to the first portion of the crude acid solid (subsequently subjected to hydrogenation) is about 0.01: 1 to about 4 Preferably within the range of about 0.05: 1 to about 2: 1 and most preferably within the range of 0.1: 1 to 1: 1.

  Hydrogenation system 106 receives a first portion of the crude acid solid from splitter 105. The hydrogenation system 106 may include one or more vessels / zones. Preferably, the hydrogenation system 106 includes an initial dissolution zone / vessel where a crude acid solid (eg, CTA) is combined with a solvent (preferably water) at an elevated temperature, thereby providing a crude acid. Dissolution of the solid in the solvent occurs. The solvent and crude acid particles are preferably in the range of about 0.5: 1 to about 50: 1, more preferably in the range of about 1: 1 to about 10: 1, and most preferably 1.5: 1. Combine at a solvent to crude acid weight ratio within the 5: 1 range.

After the crude acid particles are dissolved in the solvent, the resulting solution is introduced into the hydrogenation zone / vessel of the hydrogenation system 106 where the solution is hydrogenated of certain impurities present therein. Contact with hydrogen and a hydrogenation catalyst under conditions sufficient to cause (e.g., hydrogenation of 4-CBA to p-TAc and / or floureneone to flourene). In a preferred embodiment of the invention, the hydrotreatment is at a temperature in the range of about 200 to about 375 ° C, more preferably in the range of about 225 to about 300 ° C, and most preferably in the range of 240 to 280 ° C. carry out. The pressure in the hydrogenation zone / vessel is preferably maintained in the range of about 2 to about 50 barg. The average space velocity for hydrogenation is preferably in the range of about 150 to about 2500 solution kilograms / hour / catalyst bed cubic meter (kg / hr / m ³ ), more preferably about 300 to about 1,500 kg / hour. within the scope of the hr / m ^3, most preferably maintained within the range of 450~850kg / hr / m ^3. The molar ratio of hydrogen fed to the hydrogenation zone / vessel to the crude acid fed to the hydrogenation zone / vessel is preferably in the range of about 5: 1 to about 500: 1, more preferably about 10: 1. To about 300: 1, most preferably in the range of 20: 1 to 250: 1. The hydrogenation catalyst used in the hydrogenation zone / vessel is preferably a noble Group VIII metal on a common catalyst support material.

  After hydroprocessing in the hydrogenation system 106, the resulting hydrotreated solution is subjected to crystallization in a crystallization system 108 that includes at least one crystallizer. Within the crystallization system 108, the temperature of the hydrogenated solution is in the range of about 100 to about 200 ° C, more preferably in the range of about 120 to about 185 ° C, and most preferably in the range of 140 to 175 ° C. The temperature is lowered to the crystallization temperature. The reduced temperature in the crystallization system 108 causes substantially all of the aromatic dicarboxylic acid (eg, TPA) dissolved in the hydrotreated solution to crystallize, thereby purifying / hydrogenating. Solid particles of the formed acid (eg, PTA) are formed.

  The two-phase (slurry) effluent from the crystallization system 108 is then subjected to solid / liquid separation in a conventional separator 110. The separated, purified / hydrogenated acid solid (eg, PTA) from separator 110 is then dried in one or more conventional dryers 112.

The purified / hydrogenated acid solid (eg, PTA) discharged from the dryer 112 is preferably about 100 ppmw or less, more preferably less than about 50 ppmw, most preferably less than 25 ppmw. Has a content. The purified acid solid preferably has a p-TAc content of less than about 500 ppmw, more preferably less than about 250 ppmw, and most preferably less than 125 ppmw. The purified acid solid preferably has a combined 4-CBA + p-TAc content that is less than about 700 ppmw, more preferably less than about 500 ppmw, and most preferably less than 300 ppmw. The purified acid solid preferably has a B ^* value of less than about 3.0, more preferably less than about 2.0, and most preferably less than 1.5.

The purified / hydrogenated acid solid (eg, PTA) discharged from the dryer 112 is preferably 4 of the crude / non-hydrogenated acid solid (eg, CTA) discharged from the separator 102. Less than 80% by weight of the CBA content, more preferably 4-CBA in the range of about 5 to about 60% by weight, most preferably 4-CBA in the range of 10-40% by weight, It has a 4-CBA content. The purified acid solids are preferably less than 80% by weight of the combined 4-CBA + p-TAc content of the crude acid solids, more preferably in the range of about 5 to about 60% by weight, most preferably 10 to 10%. It has a combined 4-CBA + p-TAc content that is in the range of 40% by weight. Purified acid solids, less than 80% of B ^* value of the crude acid solids, more preferably in the range of from about 5 to about 60%, a B ^* value that is most preferably in the range of 10-40% Have.

  As shown in FIG. 3, at least a portion of the purified / hydrogenated acid solid (eg, PTA) exiting the dryer 112 is crude / hydrogen discharged from the splitter 105 in the combination zone / container 114. In combination with at least a portion of an unsolidified acid solid (eg CTA) A well-mixed composite acid (eg, composite TPA) comprising solid purified / hydrogenated acid particles and solid crude / non-hydrogenated acid particles is produced in mixing zone / vessel 114 and there Discharged from. Combination zone / container 114 is any zone having an inlet for receiving purified / hydrogenated acid, an inlet for receiving crude / non-hydrogenated acid, and an outlet for discharging composite acid. Or it may be a container. In one embodiment, the mixing zone / vessel 114 is equipped with a stirrer to ensure that the resulting composite acid is thoroughly mixed. The composite acid is sufficiently pure enough to meet product specifications, but not unnecessarily pure. Since not all of the final product acid has been subjected to hydroprocessing, the various costs associated with hydrogenation are reduced when compared to a process in which all of the final acid product is subjected to hydroprocessing in advance. To do.

  The particular amount of purified / hydrogenated acid particles and crude / non-hydrogenated acid particles combined in the mixing zone / vessel 114 depends on the level of impurities in the purified acid particles and the crude acid particles. As well as the level of impurities allowed by the final product specification. In a preferred embodiment of the present invention, the weight ratio of the crude / non-hydrogenated acid particles to the purified / hydrogenated acid particles in the composite acid is about 0.01: 1 to about 4: 1. Within the range, more preferably within the range of about 0.05: 1 to about 2: 1, most preferably within the range of about 0.1: 1 to about 1: 1.

The final composite acid (eg, composite TPA) product exiting the mixing zone / container 114 is at least about 105% by weight of the 4-CBA content of the purified / hydrogenated acid (eg, PTA) exiting the dryer 112. More preferably in the range of about 110 to about 400% by weight, most preferably in the range of about 120 to about 200% by weight, with a 4-CBA content. The composite acid product is preferably at least about 105% by weight of the p / TAc content of the purified / hydrogenated acid, more preferably in the range of about 110 to about 400% by weight, most preferably about 120 to about It has a p-TAc content that is in the range of 200% by weight. The composite acid product preferably has a purified / hydrogenated acid combination 4-CBA + p-TAc content of at least about 105% by weight, more preferably in the range of about 110 to about 400% by weight, most preferably It has a combined 4-CBA + p-TAc content that is in the range of about 120 to about 200% by weight. The composite acid product is preferably at least about 105%, more preferably in the range of about 110 to about 400%, most preferably in the range of about 120 to about 200% of the B ^* value of the purified / hydrogenated acid. Which has a B ^* value.

  We note that for all numerical ranges given herein, the upper and lower ends of the range may be independent of each other. For example, a numerical range of 10-100 means greater than 10 and / or less than 100. Thus, the range of 10-100 is support for claim limits (no upper limit) greater than 10, support for claim limits (no lower limit) less than 100, and the full range of 10-100 ( Give support for both upper and lower limits).

  We also note that “coupled in communication” as used herein refers to a direct or indirect connection that allows the flow of solids and / or liquids. For example, the outlet of the primary oxidation reactor 10 (FIG. 1) is “coupled in communication” to the solids recovery system 24 even in the presence of intermediate equipment (eg, secondary oxidation reactor 114) disposed therebetween. ing.

  The invention has been described in detail with particular reference to preferred embodiments thereof, but it will be understood that variations and modifications can be effected within the spirit and scope of the invention.

Process flow diagram showing a system for the production of terephthalic acid, wherein the first part of terephthalic acid is purified by oxidation and the remaining part of terephthalic acid is purified by hydrotreatment. Schematic representation of a pressure filter that can be used to help recover residual terephthalic acid from a liquid mother liquor produced by one or more oxidation reactors. Demonstrating a system for the production of composite terephthalic acid formed by combining an amount of purified / hydrogenated terephthalic acid with an amount of crude / unhydrogenated terephthalic acid Flow diagram.

Claims (37)

(A) providing a slurry comprising a solid phase containing a first amount of terephthalic acid and a liquid phase containing a second amount of terephthalic acid;
(B) producing an oxidized terephthalic acid by subjecting at least a portion of said first amount of terephthalic acid to an oxidation treatment; and (c) at least a portion of said second amount of terephthalic acid being hydrogenated. A method comprising producing hydrotreated terephthalic acid by subjecting to hydrotreatment.   The method of claim 1, further comprising combining at least a portion of the oxidized terephthalic acid with at least a portion of the hydrotreated terephthalic acid.   The ratio of the amount of oxidized terephthalic acid to hydrotreated terephthalic acid combined with each other is in the range of about 10: 1 to about 1,000: 1 by weight. Method.   The method of claim 1, further comprising recovering a second amount of solid particles of terephthalic acid prior to the hydrotreating of step (c).   The method of claim 4, wherein the recovery comprises evaporating at least a portion of the mother liquor from the liquid phase, thereby obtaining a concentrated slurry.   The method further comprises dissolving at least a portion of the recovered solid particles in a solvent, thereby forming a hydrogenated solution, wherein the hydrotreating of step (c) comprises converting the hydrogenated solution to liquid phase hydrogen. 5. A process according to claim 4, comprising producing a hydrogenated solution containing terephthalic acid which has been subjected to hydration and thereby hydrotreated.   7. The method of claim 6, wherein the method further comprises recovering hydrotreated terephthalic acid solid particles from the hydrogenated solution.   The method further includes recovering oxidized solid particles of terephthalic acid, and the method includes oxidizing at least a portion of the recovered particles of hydrotreated terephthalic acid with the oxidized terephthalic acid particles. 8. The method of claim 7, further comprising combining with at least a portion of the recovered particles.   The provision of step (a) comprises oxidizing the oxidizing compound in the first oxidation reactor, and step (b) is carried out in a secondary oxidation reactor separated from the first oxidation reactor. Item 2. The method according to Item 1.   The said oxidizing compound is para-xylene, the said liquid phase contains an acid component and an oxidation catalyst, the said acid component is acetic acid, and the said oxidation catalyst contains cobalt, molybdenum, and / or bromine. Method. (A) preparing a certain amount of crude acid particles;
(B) producing a hydrotreated acid by subjecting the first part of the crude acid particles to hydrotreating; and (c) a second part of the crude acid particles not subjected to hydrotreating. Producing a composite acid by combining with at least a portion of the hydrotreated acid.   The method according to claim 11, wherein the crude acid particles are crude terephthalic acid, and the hydrotreated acid is purified terephthalic acid.   The method further includes forming purified acid solid particles from at least a portion of the hydrotreated acid, wherein the combination of step (c) is purified with at least a portion of the crude acid particles. 12. The method of claim 11, comprising forming a composite acid by combining with at least a portion of the acid particles.   14. The method of claim 13, wherein the hydrotreating comprises liquid phase hydrogenation, and purified acid particles are formed by crystallization.   12. The method of claim 11, wherein the provision of step (a) comprises forming crude acid particles by oxidizing para-xylene in an oxidation reactor.   The method of claim 11, wherein the weight ratio of the crude acid particles to the hydrotreated acid in the composite acid is at least about 0.01: 1.   12. The method of claim 11, wherein the hydrotreated acid has a 4-carboxybenzaldehyde (4-CBA) content of less than about 100 ppmw and the crude acid particles have a 4-CBA content of at least about 600 ppmw. .   12. The combined 4-CBA and p-TAc content of the hydrotreated acid is less than 600 ppmw, and the crude acid has a total content of 4-CBA and p-TAc of at least about 700 ppmw. the method of. The method of claim 11, wherein the hydrotreated acid has a B ^* value less than about 3.0 and the crude acid has a B ^* value of at least about 3.0.   The method further includes dividing the crude acid particles into a first portion and a second portion, wherein the weight ratio of the second portion to the first portion is from about 0.01: 1 to about 4: The method of claim 11, which is within a range of 1. (A) producing first terephthalic acid by oxidizing para-xylene in the first oxidation reactor;
(B) producing a hydrotreated terephthalic acid by subjecting the first portion of the initial terephthalic acid to a hydrotreating; and (c) at least a portion of the hydrotreated terephthalic acid, A step comprising combining with unhydrogenated terephthalic acid that has not been subjected to hydrogenation in step (b), wherein said unhydrogenated terephthalic acid is present in the second part of the initial terephthalic acid. The method that is derived.   The method of claim 21, wherein the method further comprises subjecting at least a portion of the second portion of the initial terephthalic acid to an oxidation treatment in a secondary oxidation reactor.   The method further includes producing an initial slurry from an initial oxidation reactor, wherein the initial slurry is formed from an initial solid phase and an initial liquid phase, wherein a substantial portion of the first portion of the initial terephthalic acid is formed. 23. The method of claim 21, wherein all are contained in the first liquid phase and substantially all of the second portion of the first terephthalic acid is contained in the first solid phase.   24. The method of claim 23, wherein the method further comprises recovering residual terephthalic acid particles from the initial liquid phase.   The method further includes forming a residual solution by dissolving at least a portion of the residual terephthalic acid particles in a solvent, wherein the hydrotreating of step (b) comprises at least a portion of the residual solution as liquid phase hydrogen. 25. The method of claim 24, comprising subjecting.   The method further includes subjecting the second portion of the initial terephthalic acid to an oxidation treatment in a secondary oxidation reactor, wherein at least a portion of the initial solid phase is subjected to oxidation in the secondary oxidation reactor. 24. The method of claim 23.   The method further includes producing a secondary slurry from a secondary oxidation reactor, wherein the secondary slurry is formed from a secondary solid phase and a secondary liquid phase, and at least a first portion of the initial terephthalic acid is formed. 27. The method of claim 26, wherein a portion is derived from a secondary liquid phase.   The ratio of the amount of unhydrogenated terephthalic acid to the hydrotreated terephthalic acid combined in step (c) is within the range of about 10: 1 to about 1,000: 1 by weight. The method of claim 21. (A) introducing a liquid phase medium comprising mother liquor and terephthalic acid into the evaporation zone;
(B) evaporating at least a portion of the mother liquor from the liquid phase medium to form a concentrated medium comprising substantially all of the non-evaporated portion of the mother liquor and the terephthalic acid;
(C) obtaining a hydrogenation medium comprising at least part of the washing medium and at least part of terephthalic acid by replacing at least part of the mother liquor that has not evaporated with a washing medium; and (d) at least one of the hydrogenation medium Forming a hydrotreated medium by subjecting the part to hydrotreating.   The concentrated medium includes solid particles of terephthalic acid, the hydrogenation medium includes at least some of the solid particles of terephthalic acid, and the method dissolves at least some of the solid particles of terephthalic acid in the cleaning medium. Obtaining a hydrogenated solution, wherein the hydrotreating of step (d) comprises obtaining a hydrotreated medium by subjecting at least a portion of the hydrogenated solution to hydrogenation, 32. The method of claim 31, wherein the method comprises recovering hydrotreated terephthalic acid solid particles from the hydrotreated medium.   The method further comprises producing a slurry from an initial oxidation reactor, the slurry including an initial oxidizing liquid phase and an initial oxidizing solid phase, wherein at least a portion of the liquid phase medium of step (a) is initially 32. The method of claim 31, wherein the method is derived from a liquid phase of   The method further comprises recovering non-hydrogenated terephthalic acid particles derived from the initial oxidized solid phase, wherein the method comprises treating the hydrogenated terephthalic acid particles derived from the hydrotreated medium. And further comprising recovering the composite terephthalate by combining at least a portion of the recovered non-hydrogenated terephthalic acid particles with at least a portion of the recovered hydrotreated terephthalic acid particles. 32. The method of claim 31, further comprising producing an acid. A first oxidation reactor having a first reactor outlet;
Optionally, a secondary oxidation reactor having a secondary reactor inlet and a secondary reactor outlet coupled in communication with the first reactor outlet;
A solid / liquid separator having a separator inlet coupled in communication with the first reactor outlet and / or the secondary reactor outlet, a separated solid outlet and a separated liquid outlet;
A hydrogenation system having a hydrogenation system inlet and a hydrogenation system outlet coupled in communication with a separated liquid outlet; and a hydrogenated solid inlet coupled in communication with the hydrogenation system outlet; An apparatus comprising a combined zone having a non-hydrogenated solid inlet and a composite solid outlet coupled in communication with a solid outlet.   34. The apparatus of claim 33, wherein the secondary oxidation reactor is used in an apparatus and the separator inlet is coupled in communication with a secondary reactor outlet.   A medium further comprising an evaporation zone having an evaporation zone inlet, an evaporated medium outlet and a non-evaporated medium outlet, wherein the evaporation zone inlet is in communication with a separated liquid outlet and the hydrogenation system inlet has not evaporated 34. The apparatus of claim 33, coupled in communication with the outlet.   And further comprising a replacement zone having a replacement inlet, a replacement liquid inlet, and a hydrogenation medium outlet, wherein the replacement inlet is coupled in communication with the non-evaporated medium outlet and the hydrogenation system inlet is coupled in communication with the hydrogenation medium outlet 36. The apparatus of claim 35.   And further comprising a crystallization system having a crystallization system inlet and a crystallization system outlet, wherein the crystallization system inlet is coupled in communication with the hydrogenation system outlet, and the crystallization system outlet is coupled in communication with the hydrogenated solid inlet. 34. The apparatus of claim 33.

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