PRODUCTION PROCESS OF crude terephthalic acid

[0001] The present invention generally relates to a process for the catalytic oxidation of an aromatic compound in a liquid phase.

BACKGROUND OF THE INVENTION [0002] 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 alkylaromatic compounds to alcohols, acids, or diacids. A particularly commercially important oxidation process in the latter group (oxidation of alkylaromatic compounds) is the catalytic partial oxidation of para-xylene to terephthalic acid in the liquid phase. Terephthalic acid is an important compound with many applications. The main use of terephthalic acid is the use as a raw material in the production of polyethylene terephthalate (PET). PET is a widely known plastic that is used in huge quantities around the world to manufacture products such as bottles, fibers and packaging.

[0003] In a typical liquid phase oxidation process, including partial oxidation of para-xylene to terephthalic acid, the liquid phase feed stream and the gas phase oxidant stream are introduced into the reactor, thereby forming a multiphase reaction medium. The liquid phase feed stream entering the reactor contains at least one organic compound that can be oxidized (e.g., paraxylene), and the gas phase oxidant stream contains molecular oxygen. At least a portion of the molecular oxygen introduced into the reactor as gas dissolves in the liquid phase in the reaction medium, whereby oxygen is made available for the liquid phase reaction. If the liquid phase in the multiphase reaction medium contains an insufficient concentration of molecular oxygen (i.e. if oxygen depletion occurs in some areas of the reaction medium), impurities can form as a result of unwanted side reactions and / or the rate of desired reactions can decrease. If the liquid phase in the reaction medium contains too little compound that can be oxidized, the reaction rate may be undesirably low. In turn, if the liquid phase in the reaction medium contains an excessive concentration of a compound that can be oxidized, additional undesirable side reactions may form impurities.

[0004] Conventional liquid phase oxidation reactors are equipped with mixing devices for mixing the multiphase reaction medium contained therein. The mixing of the reaction medium is intended to assist the dissolution of molecular oxygen in the liquid phase of the reaction medium, to maintain a relatively uniform concentration of dissolved oxygen in the liquid phase of the reaction medium, and to maintain a relatively uniform concentration of the organic compound that can be oxidized in the liquid phase of the reaction medium.

[0005] EP 265 137 discloses a method for purifying terephthalic acid by oxidizing an aqueous solution of crude TPA in a fixed bed reactor with a stirrer.

[0006] Mixing of the reaction medium in which liquid phase oxidation occurs is often carried out by means of a tank mechanical stirring device, such as, for example, continuous mixing tank reactors (CSTR). Although CSTR reactors can provide thorough mixing of the reaction medium, they have a number of disadvantages. For example, CSTR reactors require relatively high capital costs because they must be equipped with expensive motors, fluid-tight bearings and shafts and / or complex mixing mechanisms. In addition, the rotating and / or oscillating mechanical components of conventional CSTR reactors require regular maintenance. The runtime and shutdown time associated with such maintenance increases the cost of operating CSTR reactors. Even if regular maintenance is performed, the mechanical mixing systems used in CSTR reactors are subject to mechanical failure and may need to be replaced at relatively short intervals.

[0007] Column bubble reactors are an attractive alternative to CSTR reactors and other mechanically stirred oxidation reactors. Column bubble reactors mix the reaction center, and at the same time do not require costly and unreliable mechanical components. Bubble column reactors usually contain an elongated vertical reaction zone in which the reaction medium is located. The reaction medium is stirred in the reaction zone mainly due to the natural buoyancy of the gas bubbles that rise through the liquid phase of the reaction medium. Mixing due to natural buoyancy occurring in bubble column reactors reduces capital and maintenance costs compared to mechanically mixed reactors. In addition, because bubble bubbling reactors generally have no moving mechanical components, an oxidation system less susceptible to mechanical failure is obtained compared to mechanically stirred reactors.

[0008] If partial oxidation of para-xylene in the liquid phase is carried out in a conventional oxidation reactor (CSTR or bubble column), the product discharged from the reactor is usually a slurry containing crude terephthalic acid (CTA) and mother liquor. CTA contains a relatively high level of impurities (e.g. 4-carboxybenzaldehyde, para-toluic acid, fluorenones and other colored compounds), as a result of which it is not suitable as a raw material for PET production. Therefore, the CTA produced in conventional oxidation reactors is usually subjected to a purification process, as a result of which the CTA is converted into purified terephthalic acid (PTA) suitable for the production of PET.

[0009] One typical purification process in which CTA is converted to PTA involves the following steps: (1) replacing the mother liquor in a suspension containing CTA with water, (2) heating the CTA suspension and water to dissolve the CTA in water, (3) catalytic hydrogenation of the CTA aqueous solution, as a result of which the impurities transform into compounds more desirable and / or easier for separation, (4) precipitation of the resulting PTA from the hydrogenated solution in a number of crystallization steps, and (5) separating the crystallized PTA from the remaining liquids. This type of conventional purification process is effective, but it can be very expensive. Separate factors causing the high costs of conventional CTA purification methods include, for example, the thermal energy necessary to dissolve the CTA in water, the catalyst necessary for hydrogenation, the hydrogen stream necessary for hydrogenation, the loss of efficiency due to the hydrogenation of some part of terephthalic acid, and a number of tanks necessary for multi-stage crystallization . Therefore, it would be advantageous to obtain a product (CTA) that could be purified without the need for heat-assisted dissolution in water, hydrogenation and / or multi-stage crystallization.

OBJECTS OF THE INVENTION [0010] It is therefore an object of the present invention to disclose a more efficient and economical liquid phase oxidation method.

SUMMARY OF THE INVENTION [0011] One embodiment relates to a crude terephthalic acid (CTA) composition comprising a plurality of CTA solid particles discharged from an oxidation reactor, wherein at least partially CTA particles have been produced, wherein a representative sample of CTA particles has one or more of the following properties: (a) contains less than about 6 ppmw of 4,4-dicarboxystilbene (4,4-DCS), (b) contains less than about 400 ppmw of isophthalic acid (IPA), (c) contains less than about 25 ppmw of 2,6-dicarboxyfluorenone ( 2,6-DCF), (d) has a percentage transmittance at 340 nanometers (% T340) greater than about 60. [0012] Another embodiment relates to a slurry composition discharged from an oxidation reactor, wherein the slurry composition comprises liquid mother liquor and crude terephthalic acid (CTA) particles, such that the CTA particles are at least partially prepared in an oxidation reactor, wherein the representative sample of the slurry has one or more of the following properties in connection with combined solid and liquid components of the suspension: (a) contains less than about 1500 ppmw of isophthalic acid (IPA), (b) contains less than about 500 ppmw of phthalic acid (PA), (c) contains less than about 500 ppmw of trimellitic acid (TMA), (d) contains less than about 2000 ppmw benzoic acid (BA).

The present invention relates to a method comprising the following steps: (a) oxidation of para-xylene in the liquid phase in a multiphase reaction medium contained in the reaction zone in at least one basic oxidation reactor constituting a bubble column reactor, and (b) draining the suspension containing liquid mother liquor and solid crude terephthalic acid (CTA) particles from the zone of the reaction, wherein the concentration of benzoic acid produced in step (a) is measured as disclosed in claim 1.

[0014] Still another embodiment relates to a method comprising the following steps: (a) introducing a recycled solvent stream into at least one oxidation reactor; (b) oxidizing the compound, which may be oxidized, in a liquid phase in a multiphase reaction medium contained in the reaction zone in an oxidation reactor; and (c) draining the slurry containing the mother liquor and solid crude terephthalic acid (CTA) particles from the reaction zone, wherein a representative sample of the CTA particles contains less than about 5 ppm in 2,7-dicarboxyfluorenone (2,7-DCF). [0015] A further embodiment relates to a method comprising the following steps: (a) liquid phase oxidation of para-xylene in a multiphase reaction medium contained in a reaction zone in at least one oxidation reactor; and (b) draining the suspension containing liquid mother liquor and solid crude terephthalic acid (CTA) particles from the reaction zone, wherein a representative sample of the suspension has one or more of the following properties in connection with the combined solid and liquid suspensions: (i) contains less than about 1500 ppmw of isophthalic acid (IPA), (ii) contains less than about 500 ppmw of phthalic acid (PA), (iii) contains less than about 500 ppmw of trimellitic acid (TMA), (iv) contains less than about 2000 ppmw benzoic acid (BA).

BRIEF DESCRIPTION OF THE DRAWINGS [0016] Preferred embodiments of the invention are described in detail below with reference to the accompanying drawings, in which:

FIG. 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 feed, oxidant and reflux streams into the reactor, the presence of a multi-phase reaction medium in the reactor and the removal of gas and suspension from the top and bottom of the reactor, respectively;

FIG. 2 is an enlarged cross-sectional side view through the bottom of a bubble column reactor along line 2-2 in FIG. 3, showing in particular the location and configuration of an oxidant misting device that introduces the oxidant stream into the reactor;

FIG. 3 is a top view of the oxidant sprinkler device of FIG. 2, showing in particular the oxidant apertures at the top of the oxidant sprinkler;

FIG. 4 is a bottom view of the oxidant sprinkler device of FIG. 2, showing in particular the oxidant holes at the bottom of the oxidant sprinkler;

FIG. 5 is a side cross-sectional view through the oxidant sprinkler along lines 5-5 in FIG. 3, showing in particular the orientation of the oxidant openings at the top and bottom of the oxidant sprinkler;

FIG. 6 is an enlarged cross-sectional side view through the bottom of a bubble column reactor, showing in particular a system for introducing the feed stream into the reactor in many locations distant in the vertical plane;

FIG. 7 is a top cross-sectional view taken along line 7-7 in FIG. 6, showing in particular how the feed introduction system shown in FIG. 6 distributes the feed stream in a favorable radial input zone (FZ) and more than one azimuth quadrant (Q₁, Q₂, Q₃, Q₄);

FIG. 8 is a top cross-sectional view similar to FIG. 7, but showing an alternative way of introducing the feed stream into the reactor by means of bayonet tubes, each having a plurality of small inlet openings;

FIG. 9 is an isometric view of an alternative system for introducing the feed stream into the reaction zone at a number of locations distant in the vertical plane, which does not require multiple holes in the tank, showing in particular that the feed distribution system may be at least partially on the oxidant sprinkler;

FIG. 10 is a side view of the single orifice feed distribution system and the oxidant sprinkler device shown in FIG. 9;

FIG. 11 is a top cross-sectional view taken along line 11-11 in FIG. 10 additionally showing a single hole feed distribution system provided on the oxidant sprinkler;

FIG. 12 is an isometric view of an alternative oxidant sprinkler in which all oxidant holes are at the bottom of the annular element;

FIG. 13 is a top view of the alternative oxidant sprinkler device of FIG. 12;

FIG. 14 is a bottom view of the alternative oxidizing sprays of FIG. 12, showing in particular the location of the lower holes for introducing the oxidant stream into the reaction zone;

FIG. 15 is a cross-sectional side view through an oxidant sprinkler along line 15-15 in FIG. 13, showing in particular the orientation of the bottom oxidant openings;

FIG. 16 is a side view showing a bubble column reactor equipped with an internal deaeration tank at the bottom outlet of the reactor;

FIG. 17 is an enlarged cross-sectional side view of the bottom of the bubble column reactor of FIG. 16 along lines 17-17 in FIG. 18, showing in particular the configuration of the internal venting tank located at the bottom outlet opening of the bubble column reactor, ·

FIG. 18 is a cross-sectional view from above along line 18-18 in FIG. 16, showing in particular a vortex device located in the vent tank;

FIG. 19 is a side view of a bubble column reactor equipped with an external vent tank, showing a method in which a portion of the vented slurry discharged from the bottom of the vent tank can be used to flush the drain line connected to the bottom of the reactor;

FIG. 20 is a side view showing a bubble column reactor equipped with a hybrid internal / external vent tank for separating the gas phase in a reaction medium discharged at a site raised from the side of the reactor;

FIG. 21 is a side view showing a bubble column reactor equipped with an alternative hybrid vent tank at the bottom of the reactor;

FIG. 22 is an enlarged cross-sectional side view of the bottom of the bubble column reactor of FIG. 21, showing in particular the use of an alternative oxidizer sprinkler, which uses inlet conduits receiving an oxidant stream through the bottom reactor head;

FIG. 23 is an enlarged cross-sectional view similar to FIG. 22, showing in particular an alternative method of introducing the oxidant stream into the reactor through a plurality of openings in the lower reactor head, in which reflective plates are optionally used to distribute the oxidant stream more evenly in the reactor;

FIG. 24 is a side view showing a bubble column reactor in which an internal channel is used to facilitate the dispersion of a compound that can be oxidized by recirculating a portion of the reaction medium from the top of the reactor to the bottom of the reactor;

FIG. 25 is a side view showing a bubble column reactor in which an external channel is used to help improve the dispersion of a compound that can be oxidized by recirculating a portion of the reaction medium from the top of the reactor to the bottom of the reactor;

FIG. 26 is a cross-sectional side view of a horizontal ejector that can be used to improve the dispersion of a compound that can be oxidized in an oxidation reactor, showing in particular an ejector that utilizes a liquid feed stream that draws the reaction medium into the ejector and discharges the feed mixture and reaction medium into the reaction zone at high speed;

FIG. 27 is a side cross-sectional view of a vertical ejector that can be used to improve the dispersion of a oxidizable compound in an oxidation reactor, showing in particular an ejector in which liquid feed and inlet gas are combined and uses a combined two-phase fluid to draw the reaction medium to an ejector and the discharge of a mixture of liquid feed, inlet gas and reaction medium to the reaction zone at a high speed;

FIG. 28 is a side view of a bubble column reactor containing a multi-phase reaction medium, showing in particular the reaction medium, which is theoretically divided into 30 horizontal sections of equal volume to quantify certain gradients in the reaction medium;

FIG. 29 is a side view of a bubble column reactor containing a multiphase reaction medium, showing in particular the first and second distinct 20% continuous volumes of the reaction medium in which oxygen concentrations and / or oxygen consumption rates are substantially different;

FIG. 30 is a side view of two stacked reaction tanks with or without a mechanical stirrer, containing a multi-phase reaction medium, showing in particular that the tanks contain separate 20-percent continuous volumes of reaction medium in which oxygen concentrations and / or oxygen consumption rates are generally different;

FIG. 31 is a side view of three reaction tanks in a row with or without a mechanical stirrer, containing a multi-phase reaction medium, showing in particular that the tanks contain separate 20% continuous volumes of reaction medium in which the oxygen concentrations and / or oxygen consumption rates are essentially different;

FIG. 32A and 32B are an enlarged view of crude terephthalic acid (CTA) particles produced according to one embodiment of the present invention, in particular showing that each CTA particle is a low density particle with a large surface area composed of many smaller, loosely bound CTA particles.

FIG. 33A and 33B are an enlarged view of CTA prepared in a conventional manner, in particular showing that a conventional CTA particle has a larger particle size, lower density and smaller surface area than the CTA particle according to the invention of FIG. 32A and 32B;

FIG. 34 is a simplified flow diagram of a prior art method for producing purified terephthalic acid (PTA);

FIG. 35 is a simplified flow diagram of a method of making PTA according to one embodiment of the present invention; and

FIG. 36 is a graph showing the results of the time dissolution test described in the Examples section, which shows in particular that the CTA crystals of the invention dissolve faster than conventional CTA crystals.

DETAILED DESCRIPTION [0017] One embodiment relates to the partial oxidation of p-xylene in the liquid phase. Such oxidation is carried out in the liquid phase in a multiphase reaction medium contained in a single bubble column reactor.

[0018] The term "bubble column reactor" herein means a reactor that allows chemical reactions to take place in a multi-phase reaction medium, whereby the reaction medium is stirred mainly due to gas bubbles moving up through the reaction medium. As used herein, "mixing" means work inserted into the reaction medium that causes fluid to flow and / or mix. The terms "majority", "mainly" and "for the most part" mean more than 50 percent in this document. As used herein, the term "mechanical mixing" means mixing of a reaction medium due to the physical movement of a rigid or flexible element or elements against or in a reaction medium. For example, mechanical mixing may occur as a result of rotations, oscillations and / or internal vibrations of agitators, blades, shakers or acoustic diaphragms in the reaction medium. As used herein, the term "flow mixing" means mixing of a reaction medium by injection and / or recirculation at high speed of one or more fluids in the reaction medium. For example, flow mixing can be caused by nozzles, ejectors and / or ejectors.

[0019] In a preferred embodiment of the present invention, mechanical and / or flow mixing provides less than about 40 percent mixing of the reaction medium in the bubble column reactor during oxidation, more preferably mechanical and / or flow mixing provides less than about 20 percent mixing, and most preferably mechanical mixing and / or flow provides less than 5 percent mixing. Preferably, the amount of mechanical and / or flow mixing to which the multiphase reaction medium is subjected during oxidation is less than about 3 kilowatts per cubic meter of reaction medium, more preferably less than about kilowatts per cubic meter and most preferably less than 1 kilowatt per cubic meter.

[0020] Further referring to FIG. 1 shows a preferred bubble column reactor 20 that includes a tank jacket 22 in which the reaction compartment 24 and the separation compartment 26 are located. The reaction interval 24 defines the internal reaction zone 28 and the separation compartment 26 defines the internal separation zone 30. The feed stream predominantly in the liquid phase it is introduced into the reaction zone 28 through the inlet openings 32a, b, c, d. The oxidant stream predominantly in the gas phase is introduced into the reaction zone 28 via an oxidant spraying device 34 located at the bottom of the reaction zone 28. The liquid stream feedstock and the gas phase oxidant stream together form the multi-phase reaction medium 36 in the reaction zone 28 The multiphase reaction medium 36 comprises a liquid phase and a gas phase. More preferably, the multi-phase reaction medium 36 comprises a three-phase medium containing solid phase, liquid phase and gas phase components. The solid phase component of reaction medium 36 preferably precipitates in reaction zone 28 as a result of the oxidation reaction carried out in the liquid phase of reaction medium 36. The bubble column reactor 20 comprises a slurry outlet 38 located at the bottom of reaction zone 28 and a gas outlet 40 near the top partition zone surface 30. The discharged stream of slurry containing the components in the liquid phase and in the solid phase of the reaction medium 36 is withdrawn from the reaction zone 28 through the suspension outlet 38, and the predominantly gaseous discharged stream is withdrawn from the separation zone 30 through the gas outlet 40.

[0021] The liquid phase feed stream introduced into the bubble column reactor 20 through the inlets 32a, b, c, d preferably contains a oxidizable compound, a solvent and a catalyst system.

[0022] The compound that can be oxidized is paraxylene.

[0023] It is preferred that the oxidizable compound is substantially dissolved in the solvent when entering the reaction zone 28. The amount of oxidizable compound present in the liquid phase feed stream is preferably in the range of from about 2 to about 40 weight percent, more preferably in the range of about 4 to about 20 weight percent, and most preferably in the range of 6 to 15 weight percent.

[0024] The solvent present in the liquid phase feed stream preferably comprises an acid component and a water component. The solvent is preferably present in the liquid phase feed stream at a concentration in the range of about 60 to about 98 weight percent, more preferably in the range of about 80 to about 96 weight percent, and most preferably in the range of 85 to 94 weight percent. The solvent acid component is preferably mainly a low molecular weight organic monocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably the solvent acid component is mainly acetic acid. The acid component is preferably at least about 75 weight percent of the solvent, more preferably at least about 80 weight percent of the solvent, and most preferably 85 to 98 weight percent of the solvent, the balance being mainly water. The solvent fed to the bubble column reactor 20 may contain small amounts of impurities such as, for example, paratolualdehyde, terephthaldehyde, 4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluylic acid, para-toluylaldehyde, alpha-bromo-para- toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended solids. It is preferred that the total content of impurities in the solvent fed to the bubble column reactor is less than about 3 weight percent.

[0025] The catalyst system present in the liquid phase feed stream is preferably a homogeneous liquid phase catalyst system capable of facilitating the oxidation (including partial oxidation) of a compound that can be oxidized. More preferably the catalyst system comprises at least one polyvalent transition metal. Even more preferably, the multivalent transition metal comprises cobalt. Even more preferably, the catalyst system comprises cobalt and bromine. Most preferably the catalyst system contains cobalt, bromine and manganese. [0026] If cobalt is present in the catalyst system, it is preferred that the cobalt content present in the liquid phase feed stream is such that the cobalt concentration in the liquid phase of the reaction medium 36 is maintained in the range from about 300 to about 6000 parts by weight ( ppmw), more preferably in the range of about 700 to about 4200 ppmw and most preferably in the range of 1200 to 3000 ppmw. If bromine is present in the catalyst system, it is preferred that the bromine content present in the liquid phase feed stream is such that the bromine concentration in the liquid phase of the reaction medium 36 is maintained in the range from about 300 to about 5000 ppmw, more preferably in the range from about 600 up to about 4000 ppmw and most preferably in the range from 900 to 3000 ppmw. If manganese is present in the catalyst system, it is preferred that the manganese content present in the liquid phase feed stream is such that the manganese concentration in the liquid phase of the reaction medium 36 is maintained in the range from about 20 to about 1000 ppmw, more preferably in the range from about 40 up to about 500 ppmw, most preferably in the range of 50 to 200 ppmw.

[0027] The concentration of cobalt, bromine and / or manganese in the liquid phase in the reaction medium 36 referred to above is expressed in time-averaged and volume-averaged form. The term "time averaged" as used herein means the average of at least 10 measurements made uniformly over a continuous period of at least 100 seconds. The term "volume averaged" as used herein means the average of at least 10 measurements made in a homogeneous 3-dimensional space in a certain volume.

[0028] The weight ratio of cobalt to bromine (Co: Br) in the catalyst system introduced into reaction zone 28 is preferably in the range of from about 0.25: 1 to about 4: 1, more preferably in the range of from about 0.5: 1 to about 3: 1 and most preferably in the range of 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 in the range of from about 0.3: 1 to about 40: 1, more preferably in the range of from about 5: 1 to about 30: 1 and most preferably in the range from 10: 1 to 25: 1.

[0029] The liquid phase feed stream fed into the bubble column reactor 20 may contain small amounts of impurities, such as, for example, toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde, 418 carboxybenzaldehyde (4-CBA), benzoic acid, paratolylic acid, para-aldehyde -toluylic acid, alpha-bromo-paratoluylic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended solids. If the bubble column reactor 20 is used to produce terephthalic acid, meta-xylene and ortho-xylene are also considered impurities. It is preferred that the total impurity content in the liquid stream feed stream entering the bubble column reactor 20 is less than about 3 weight percent.

[0030] Although FIG. 1 shows an embodiment in which the oxidizable compound, solvent and catalyst system are mixed together and fed to a bubble column reactor 20 in a single feed stream, in an alternative embodiment of the present invention the oxidizable compound, solvent and the catalyst can be fed separately to the bubble column reactor 20. For example, it is possible to introduce a stream of pure para-xylene into the bubble column reactor 20 through a hole separate from the hole or holes for the solvent and catalyst.

[0031] The oxidant stream predominantly in the gaseous phase fed to the bubble column reactor 20 through the spraying device to the oxidant 34 contains molecular oxygen (O₂). The oxidant stream preferably contains in the range of about 5 to about 40 mole percent molecular oxygen, more preferably in the range of about 15 to about 30 mole percent molecular oxygen, and most preferably in the range of 18 to 24 mole percent molecular oxygen. It is preferred that the remainder of the oxidant stream consists mainly of gas or gases, such as nitrogen, inert to the oxidation. More preferably, the oxidant stream consists essentially of molecular oxygen and nitrogen. Most preferably, the oxidant stream is dry air containing 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 contain substantially pure oxygen.

[0032] With reference again to FIG. The 1 bubble column reactor 20 is preferably equipped with a return stream divider 42 located above the top surface 44 of the reaction medium 36. The return stream divider 42 can operate in such a way that it introduces the return stream droplets predominantly in the liquid phase into the separation zone 30 by any method creating droplets known in the art. More preferably, the return stream splitter 42 produces atomized droplets directed downward toward the upper surface 44 of the reaction medium 36. Preferably, this downward atomized droplet stream comprises (i.e., occupies and affects) at least about 50 percent of the maximum horizontal cross-sectional area of ​​the separation zone 30. More preferably, the atomized droplet stream comprises at least about 75 percent of the maximum horizontal cross-sectional area of ​​the separation zone 30. Most preferably, the spray droplet stream comprises at least 90 percent of the maximum horizontal cross-sectional area of ​​the separation zone 30. This downward liquid spray return stream can help prevent foaming on the upper surface 44 of reaction medium 36 or above, and may further facilitate the separation of droplets of any liquid or suspension entrained by upward moving gas that flows towards the gas outlet 40. In addition, the liquid return stream can be used to reduce the content of solid particles and compounds that may precipitate (e.g. dissolved benzoic acid, para-toluic acid, 4-CBA, terephthalic acid and catalytic metal salts) exiting in the gaseous stream discharged from the separation zone 30 through the gas outlet 40. In addition, the introduction of return stream droplets into the separation zone 30 can be used, due to the distillation phenomenon, to correct the composition of the gaseous stream discharged through the gas outlet 40.

[0033] The liquid return stream introduced into the bubble column reactor 20 via the return stream separator 42 preferably has an approximately identical composition to the liquid phase feedstock solvent component introduced into the bubble column reactor 20 through the inlets 32a, b, c, d. It is therefore preferred that the liquid return stream contains the acid and water component. The acid component of the back stream is preferably a low molecular weight organic monocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the back stream is acetic acid. The acid component is preferably at least about 75 percent by weight of the reverse stream, more preferably at least about 80 percent by weight of the reverse stream, and most preferably 85 to 98 percent by weight of the reverse stream, the balance being water. Because the return stream usually has essentially the same composition as the solvent in the liquid phase feed stream, if this description refers to "total solvent" entering the reactor, such "total solvent" includes both the return stream and the solvent portion of the feed stream.

[0034] During the liquid phase oxidation in the bubble column reactor 20, it is preferred that the feed, oxidant and return streams are fed into reaction zone 28 substantially continuously and the discharged gas streams and suspensions are discharged from reaction zone 28 essentially continuous. The term "substantially continuous" as used herein means a period of at least 10 hours with intervals shorter than 10 minutes. During oxidation, it is preferred that the oxidizable compound (e.g. para-xylene) is introduced into reaction zone 28 substantially continuously at a rate of at least about 8000 kilograms per hour, more preferably at a rate of from about 13000 to about 80,000 kilograms per hour, even more preferably in the range of about 18,000 to about 50,000 kilograms per hour, and most preferably in the range of 22,000 to 30,000 kilograms per hour. Although it is generally preferred that the feed rates of feed, oxidizer and return streams be substantially constant, it should now be noted that one embodiment of the present invention provides for pulsed feeding of feed, oxidant and / or return feed to improve mixing and permeation mass. If the feedstock, oxidant and / or return stream is introduced in a pulsating manner, it is preferred that their flow rates vary within about 0 to about 500 percent of the steady-state flow rates given herein, more preferably, within the limits of about 30 to about 200 percent of the steady-state flow rates given in this document, and most preferably within 80 to 120 percent of the steady-state flow rates given herein.

The average reaction rate in time and space (STR) in an oxidation bubble column reactor is defined as the mass of the oxidizable compound introduced per unit volume of reaction medium 36 per unit of time (e.g. kilograms of para-xylene introduced per meter cubic per hour). In conventional use, the amount of oxidizable compound that has not been converted to the product is usually subtracted from the oxidizable compound in the feed stream before calculating the STR. The conversion and yield, however, are usually high for many oxidizable compounds that are preferred herein (e.g., para-xylene) and it is convenient to define this term herein as given above. Due to, inter alia, capital costs and operating inventory, it is generally beneficial to react with a high STR value. However, carrying out reactions at increasing STR values ​​may affect the quality or efficiency of partial oxidation. The bubble column reactor 20 is particularly useful if the STR value for a compound that can be oxidized (e.g. para-xylene) remains in the range of about 25 kilograms per cubic meter per hour to about 400 kilograms per cubic meter per hour, more preferably in the range of about 30 kilograms per cubic meter per hour to about 250 kilograms per cubic meter per hour, even more preferably from about 35 kilograms per cubic meter per hour to about 150 kilograms per cubic meter per hour and most preferably in the range of 40 kilograms per cubic meter per hour to 100 kilograms per cubic meter per hour. [0036] The oxygen STR value in the oxidation bubble column reactor is defined as the weight of molecular oxygen consumed per unit volume of reaction medium 36 per unit of time (e.g. kilograms of molecular oxygen consumed per cubic meter per hour). Due to, inter alia, capital costs and solvent consumption by oxidation, it is generally preferable to carry out the reaction at a high oxygen STR value. However, carrying out reactions at increasing oxygen STRs ultimately worsens the quality or efficiency of partial oxidation. Without theoretical restrictions, it seems that this may be related to the rate of molecular oxygen permeation from the gas phase to the liquid at the interface between it and the liquid volume. Too high an oxygen STR value may lead to too low dissolved oxygen content in the liquid phase volume in the reaction medium.

[0037] The overall average oxygen STR value is defined herein as the weight of all oxygen consumed in the entire volume of reaction medium 36 per unit of time (e.g., kilograms of molecular oxygen consumed per cubic meter per hour). The bubble column reactor 20 is particularly useful if the overall average STR STR for oxygen remains in the range from about 25 kilograms per cubic meter per hour to about 400 kilograms per cubic meter per hour, more preferably in the range from about 30 kilograms per cubic meter per hour to about 250 kilograms per cubic meter per hour, even more preferably from about 35 kilograms per cubic meter per hour to about 150 kilograms per cubic meter per hour and most preferably in the range of 40 kilograms per cubic meter per hour to 100 kilograms per cubic meter per hour.

[0038] During oxidation in the bubble column reactor, it is preferred that the ratio of the mass flow rate of all solvent (from both the feed stream and the return stream) to the mass flow rate of the oxidizable compound entering the reaction zone 28 is kept within 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 solvent introduced within the feed stream to the ratio of the mass flow rate of solvent introduced within the return stream is maintained in the range of from about 0.5: 1 to no return flow, more preferably in the range of from about 0.5: 1 to about 4: 1, even more preferably in the range of about 1: 1 to about 2: 1 and most preferably in the range of 1.25: 1 to 1.5: 1. [0039] During liquid phase oxidation in bubble column reactor 20, it is preferred that the oxidant stream is fed to bubble column reactor 20 in an amount that provides the amount of molecular oxygen slightly greater than the required stoichiometric oxygen value. The value of excess molecular oxygen necessary to obtain the best results for a given compound that can be oxidized affects the overall economics of liquid phase oxidation. During liquid phase oxidation in bubble column reactor 20, it is preferred that the ratio of the mass flow rate of the oxidant stream to the mass flow rate of the oxidizable organic compound (e.g. para-xylene) entering the reactor 20 is maintained within a range of from about 0 , 5: 1 to about 20: 1, more preferably in the range of about 1: 1 to about 10: 1 and most preferably in the range of 2: 1 to 6: 1.

[0040] With reference again to FIG. 1 feed, oxidant and return streams fed to the bubble column reactor 20 together form at least part 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 a compound that can be oxidized (e.g. para-xylene) occurs predominantly in the liquid phase of reaction medium 36. Therefore, the liquid phase in reaction medium 36 contains dissolved oxygen and a compound that can be oxidized. The exothermic nature of the oxidation reaction taking place in the bubble column reactor causes a portion of the solvent (e.g. acetic acid and water) introduced through the inlets 32a, b, c, d to boil / evaporate. Therefore, the gaseous phase of reaction medium 36 in reactor 20 consists mainly of evaporated solvent and undissolved, unreacted part of the oxidant stream. Some prior art oxidation reactors use heat exchange tubes / fins to heat or cool the reaction medium. However, such heat exchange elements may be undesirable in the reactor according to the invention and in the method described herein. Therefore, it is preferred that the bubble column reactor 20 substantially does not contain surfaces in contact with reaction medium 36 and having a time-averaged heat flux above 30,000 watts per square meter.

[0041] The concentration of dissolved oxygen in the liquid phase in reaction medium 36 remains in dynamic balance between the mass transfer rate from the gas phase and the consumption rate in the liquid phase reaction (i.e. it does not simply result from the partial pressure of molecular oxygen in the supplied gas phase , although it is one of the factors related to the rate of dissolved oxygen delivery and significantly reduces the upper concentration of dissolved oxygen). The amount of dissolved oxygen varies locally; is greater near the bubble interface. In general, the amount of dissolved oxygen depends on the balance of supply and demand factors in various areas of the reaction medium 36. In time, the amount of dissolved oxygen depends on the homogeneity of gas and liquid mixing relative to the rate of chemical consumption. When designing to properly adjust the dissolved oxygen supply and demand for it in the liquid phase of the reaction medium 36, it is preferred that the time-averaged and volume-averaged oxygen concentration in the liquid phase of the reaction medium 36 is maintained above about 1 molar ppm, more preferably in the range of from about 4 up to about 1000 molar ppm, even more preferably in the range of from about 8 to about 500 molar ppm, and most preferably in the range of from 12 to 120 molar ppm.

[0042] The liquid phase oxidation reaction taking place in the bubble column reactor 20 is preferably a precipitation reaction that forms solids. More preferably, the liquid phase oxidation reaction in the bubble column reactor 20 causes at least about 10 weight percent of the oxidizable compound (e.g. para-xylene) introduced into reaction zone 28 to form a solid compound (e.g. crude terephthalic acid particles) ) in the reaction medium 36. Even more preferably, the liquid phase oxidation reaction causes that at least about 50 percent by weight of the compound that can be oxidized forms a solid compound in reaction medium 36. Most preferably the liquid phase oxidation reaction causes at least 90 percent by weight of the compound that can be subjected to oxidation. oxidation, forms a solid compound in the reaction medium 36. It is preferred that the total solids content of reaction medium 36 is greater than about 3 weight percent based on time average and volume average. More preferably, the total solids content of reaction medium 36 is maintained in the range of from about 5 to about 40 weight percent, even more preferably in the range of about 10 to about 35 weight percent, and most preferably in the range of 15 to 30 weight percent. It is preferred that a significant portion of the oxidation product (e.g. terephthalic acid) produced in the bubble column reactor 20 is present in the reaction medium 36 in solid form as opposed to the residue dissolved in the liquid phase of reaction medium 36. The content of the solid phase oxidation product in reaction medium 36 is preferably at least about 25 weight percent of the total oxidation product (solid and liquid phase) in reaction medium 36, more preferably at least about 75 weight percent of the total oxidation product in reaction medium 36, and most preferably at least 95 weight percent of total oxidation product in reaction medium 36. The abovementioned numerical ranges of solids content in reaction medium 36 relate essentially to the operation of the bubbling column 20 in a steady state, rather than starting, shutting down or suboptimal operation of the bubble column reactor 20. The solid content in the reaction medium 36 is determined by gravimetric method. In this gravimetric method, a representative part of the suspension is taken from the reaction medium and weighed. Under conditions that allow effective maintenance of the overall separation of solids and liquids in the reaction medium, the free liquid is removed from the solids by sedimentation or filtration, in practice without loss of precipitated solids, so that less than about 10 percent of the initial liquid mass remains in the solids. The liquid remaining in solids is evaporated to dryness, in practice without sublimation of solids. The remaining solids are weighed. The ratio of the weight of the solid to the weight of the original suspension is a fraction of solids, usually expressed as a percentage.

[0043] The precipitation reaction taking place in the bubble column reactor 20 can cause contamination (i.e. accumulation of solids) on the surface of some rigid elements in contact with the reaction medium 36. Therefore, in one embodiment of the present invention, it is preferred that the bubble column reactor 20 does not substantially contain internal heat exchange, mixing or baffle components in reaction zone 28, since such components would be exposed to contamination. If internal structures are present in reaction zone 28, it is preferable not to allow these internal structures to have external surfaces containing a significant amount of flat upward surface, because such flat upward surfaces would be highly exposed to contamination. Therefore, if internal structures are present in reaction zone 28, it is preferred that less than about 20 percent of the total upwardly exposed outer surface of such inner structures is formed of substantially flat surfaces with an inclination of less than about 15 degrees relative to horizontal. [0044] With reference again to FIG. 1 the physical configuration of the bubble column reactor 20 helps to achieve optimized oxidation of the compound that can be oxidized (para-xylene) with minimal formation of impurities. It is preferred that the elongated reaction interval 24 of the tank jacket 22 comprises a substantially cylindrical main body 46 and a lower head 48. The upper end of the reaction zone 28 is defined by a horizontal plane 50 extending through the upper surface of the cylindrical main body 46. The bottom end 52 of reaction zone 28 is defined by the lowest inner surface of the bottom head 48. Usually, the bottom end 52 of reaction zone 28 is near the slurry outlet opening 38. Therefore, the elongated reaction zone 28 in the bubble column reactor 20 has a maximum length "L" measured from the upper end 50 to the lower end 52 of the reaction zone 28 along the extension axis of the cylindrical main body 46. The length "L" of reaction zone 28 is preferably in the range of about 10 to about 100 meters, more preferably in the range of about 20 to about 75 meters, and most preferably in the range of 25 to 50 meters. The reaction zone 28 has a maximum diameter (width) "D" usually equal to the maximum internal diameter of the cylindrical main body 46. The maximum diameter "D" of reaction zone 28 is preferably in the range of about 1 to about 12 meters, more preferably in the range of about 2 to about 10 meters, even more preferably in the range of about 3.1 to about 9 meters and most preferably in the range of 4 up to 8 meters. 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 a L: D ratio in the range of from about 8: 1 to about 20: 1. Most preferably, reaction zone 28 has a L: D ratio in the range from 9: 1 to 15: 1.

[0045] As mentioned above, the reaction zone 28 of the bubble column reactor 20 receives a multi-phase reaction medium 36. The lower end of the reaction medium 36 coincides with the lower end 52 of reaction zone 28 and the upper end located on the upper surface 44. The upper surface 44 of the reaction medium 36 is defined by a horizontal plane crossing the reaction zone 28 at a point on the vertical line where the content of the reaction zone 28 passes from the continuous gas phase to the continuous liquid phase. The top surface 44 is preferably at a point on the vertical line where the local gas retention averaged over a thin horizontal slice of the content of reaction zone 28 is 0.9.

[0046] Reaction medium 36 has a maximum height "H" measured between its upper and lower ends. The maximum width "W" of the reaction medium 36 is usually equal to the maximum diameter "D" of the cylindrical main body 46. During liquid phase oxidation in bubble column reactor 20 it is preferred to maintain an H value of about 60 to about 120 percent L, more preferably about 80 to about 110 percent L and most preferably 85 to 100 percent L. In a preferred embodiment of the present invention, reaction medium 36 has a height / width ratio "H: W" 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, the reaction medium has an H: W ratio in the range of from about 8: 1 to about 20: 1. Most preferably, reaction medium 36 has a H: W ratio in the range of 9: 1 to 15: 1. In one embodiment, L = H and D = W, so that the individual dimensions or ratios given herein in the case of L and D also apply to H and W, and vice versa.

[0047] The relatively high L: D and H: W ratios according to the embodiment can provide a number of significant advantages of the system according to the invention. As discussed in detail below, it has been found that higher L: D and H: W ratios, as well as some other features discussed below, can facilitate the formation of a favorable vertical gradient of molecular oxygen concentration and / or a compound that can be oxidized (e.g., para- xylene) in reaction medium 36. In contrast to conventional knowledge, according to which a strongly mixed reaction medium with a relatively homogeneous concentration throughout the volume is preferred, it has been found that a vertical gradient of oxygen and / or compound that can be oxidized promotes a more efficient and more economical oxidation reaction. Reducing to a minimum the concentration of oxygen and an oxidizable compound near the top surface of reaction medium 36 may help to avoid the loss of unreacted oxygen and unreacted oxidizable compound by the upper gas outlet 40. However, if the concentration of the compound that can be oxidized , and unreacted oxygen is low throughout the entire reaction medium 36, oxidation rate and / or selectivity is reduced. It is therefore preferred that the concentration of molecular oxygen and / or the compound that can be oxidized is significantly higher near the bottom of reaction medium 36 than near the upper surface of reaction medium 36.

[0048] In addition, high L: D and H: W ratios can cause the pressure at the bottom of reaction medium 36 to be significantly higher than the pressure at the upper surface of reaction medium 36. Such a vertical pressure gradient results from the height and density of reaction medium 36. One of the advantages of this vertical pressure gradient is that greater pressure at the bottom of the tank increases oxygen solubility and mass transfer than would otherwise be possible at comparable temperature and pressure over the medium in shallow reactors. As a result, the oxidation reaction can be carried out at a lower temperature than would be required for a shallower tank. If the bubble column reactor 20 is used to partially oxidize para-xylene to crude terephthalic acid (CTA), the ability to operate at a lower reaction temperature with the same or higher oxygen mass transfer rate has many advantages. For example, low temperature oxidation of para-xylene reduces the amount of solvent burned during the reaction. As further discussed below, low temperature oxidation further promotes the formation of small, loosely bound, easily soluble CTA particles with a large surface area that can be subjected to more economical purification methods than large, dense CTA particles with a low surface area, produced by conventional oxidation processes in high temperature.

[0049] During oxidation in reactor 20, it is preferred that the time-averaged and volume-averaged temperature of the reaction medium 36 be maintained 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 to 170C. The pressure over reaction medium 36 is preferably in the range of about 1 to about 20 bar relative pressure (barg), more preferably in the range of about 2 to about 12 barg and most preferably in the range of 4 to 8 barg. The pressure difference between the upper surface of reaction medium 36 and the bottom of reaction medium 36 is preferably in the range of about 0.4 to about 5 bar, more preferably the pressure difference is in the range of about 0.7 to about 3 bar, and most preferably the pressure difference is 1 to 2 bar. Although it is generally preferred that the pressure over reaction medium 36 be maintained so that its value is relatively constant, one embodiment of the present invention provides for pulsed pressure over the medium that facilitates better mixing and / or mass transfer in reaction medium 36. If the pressure over the medium is pulsed, it is preferred that the pulse pressure is in the range of about 60 to about 140 percent of the pressure over the steady state medium given herein, more preferably from about 85 to about 115 percent of the pressure over the steady state medium given in herein, and most preferably from 95 to 105 percent of the pressure over the steady state medium given herein.

[0050] Another advantage of the high L: D ratio in reaction zone 28 is that it can contribute to an increase in the average speed calculated per cross-section of the empty reactor in reaction medium 36. The terms "velocity calculated on the cross section of an empty reactor" and "gas velocity calculated on the cross section of an empty reactor" used herein with reference to the reaction medium 36 mean the volumetric flow rate of the gas phase in the reaction medium 36 at a given height of the reactor divided by the cross-sectional area of ​​the horizontal reactor at this height. Increasing the speed calculated per cross-section of the empty reactor due to the high L: D ratio in reaction zone 28 can facilitate local mixing and increase gas retention in reaction medium 36. The time-averaged velocity based on the cross-section of the empty reactor in reaction medium 36 in one quarter, half and / or three quarters of the height of reaction medium 36 is preferably greater than about 0.3 meter per second, more preferably it is in the range of about 0.8 to about 5 meters per second, even more preferably in the range of about 0.9 to about 4 meters per second and most preferably in the range of 1 to 3 meters per second.

[0051] With reference again to FIG. 1 separation compartment in the bubble column reactor 20 is simply an expanded portion of the tank jacket 22 directly above the reaction compartment 24. The separation interval 26 reduces the velocity of the upwardly moving gas phase in the bubble column reactor 20 when the gas phase rises above the upper surface 44 reaction medium 36 and approaching gas outlet 40. This reduction in the upward velocity of the gas phase helps to improve the removal of liquids and / or solids entrained by the upwardly moving gas phase, thereby reducing unwanted losses of some of the components present in the liquid phase of reaction medium 36.

[0052] The separation compartment 26 preferably comprises a transition wall 54, generally in the shape of a truncated cone, a generally cylindrical wide sidewall 56 and an upper head 58. The narrow lower end of the transition wall 54 is connected to the top of the cylindrical main body 46 in the reaction interval 24. Wide end the upper transition wall is connected to the bottom of the wide side wall 56. It is preferred that the transition wall 54 extends up and out from the narrow bottom end at an angle in the range of about 10 to about 70 degrees from vertical, more preferably in the range of about 15 to about 50 degrees from vertical, and most preferably in the range of 15 to 45 degrees from the vertical. The wide side wall 56 has a maximum diameter "X" substantially greater than the maximum diameter "D" of the reaction compartment 24, however, if the upper portion of the reaction compartment 24 has a diameter smaller than the total maximum diameter of the reaction compartment 24, then X may be smaller than D. In a preferred embodiment of the present invention, the ratio of the width of the wide side wall 56 to the maximum diameter of the 24 "X: D" reaction interval is in the range of from about 0.8: 1 to about 4: 1, most preferably in the range of from 1.1: 1 to 2: 1. The upper head 58 is connected to the upper surface of the wide side wall 56. The upper head 58 is preferably a substantially elliptical head element defining a central opening to allow gas to escape from the separation zone 30 through the gas outlet 40. Alternatively, the upper head 58 may have any shape, including conical. The separation zone 30 has a maximum height "Y" measured from the top surface 50 of the reaction zone 28 to the top of the separation zone 30. The ratio of the length of reaction zone 28 to the height of the "L: Y" separation zone 30 is preferably in the range of from about 2: 1 to about 24: 1, more preferably in the range of from about 3: 1 to about 20: 1, and most preferably in the range of 4: 1 to 16: 1.

[0053] With reference to FIG. 1-5, the location and configuration of the oxidant sprinkler 34 will be discussed in more detail below. FIG. 2 and 3 show that the oxidant sprinkler device 34 may include an annular member 60, a cross member 62, and a pair of inlet conduits for oxidizer 64a, b. These inlet lines to the oxidant 64a, b may conveniently enter the tank at a height above the annular member 60 and then twist down as shown in FIG. 2 and 3. Optionally, the inlet line to the oxidizer 64a, b may enter the reservoir below the annular member 60 or approximately in the same horizontal plane as the annular member 60. Each inlet conduit for oxidant 64a, b includes a first end connected to a corresponding inlet of oxidant 66a, b formed in the shell of the tank 22 and a second end connected in a fluid permeable manner to the annular member 60. The annular member 60 is preferably formed from conduits, more preferably a plurality of straight sections of conductor and most preferably of many straight sections of conductors connected rigidly to each other, thanks to which they form a polygonal tubular ring. Preferably, the annular element 60 is formed of at least 3 straight pipe sections, more preferably 6 to 10 pipe sections and most preferably 8 pipe sections. Accordingly, if the annular element 60 is formed of 8 wire sections, it has a substantially octagonal configuration. The cross member 62 is preferably formed of a substantially straight pipe section connected in fluid manner with opposite sections of the annular member 60 and runs diagonally between them. The pipe section used as cross member 62 preferably has substantially the same diameter as the pipe sections used to form the annular member 60. It is preferred that the pipe sections forming inlet conduits for the oxidant

64a, b, annular member 60 and cross member 62 had a nominal diameter greater than about 0.1 meters, more preferably in the range of about 0.2 to about 2 meters, and most preferably in the range of 0.25 to 1 meter. As most likely shown best in FIG. 3, each of the annular member 60 and cross member 62 has a plurality of upper oxidant apertures 68 for leading the oxidant stream up into reaction zone 28. As most likely shown in FIG. 4, ring member 60 and / or cross member 62 may have one or more lower oxidant openings 70 for discharging the oxidant stream down to reaction zone 28. Lower oxidant openings 70 may also be used to discharge liquids and / or solids that could get stuck in the ring element 60 and / or in the cross member 62. To prevent the accumulation of solids inside the spraying device into the oxidant 34, the liquid stream can be passed continuously or periodically through the spraying device 34 to flush out the accumulated solids.

[0054] With reference again to FIG. 1-4, during oxidation in the bubble column reactor 20 oxidant streams are injected through oxidant inlets 66a, b and into oxidizer inlet pipes 64a, b, respectively. The oxidant streams are then transported through the inlet lines to the oxidant 64a, b to the annular element 60. When the oxidant stream reaches ring element 60, the oxidant stream is distributed throughout the entire internal volume of ring element 60 and cross member 62. The oxidant stream is then extruded from the oxidant spraying device 34 into reaction zone 28 through the upper and lower oxidant openings 68.70 in the annular and transverse element 62.

[0055] The outlets of the upper oxidant openings 68 are arranged in a horizontal line relative to each other and are located substantially at the same height of the reaction zone 28. Therefore, the outlets of the upper oxidant openings 68 are located along a substantially horizontal plane defined by the upper surface of the spraying device for oxidiser 34. The outlets of the bottom holes for oxidant 70 are arranged in a horizontal line relative to each other and are located substantially at the same height of the reaction zone 28. Therefore, the outlets of the bottom holes for oxidizer 70 are located along a substantially horizontal plane defined by the bottom of the spraying device for oxidant 34.

[0056] In one embodiment, the oxidant sprinkler device 34 comprises at least about 20 of the top oxidant apertures 68 therein. More preferably, the oxidant sprinkler device 34 includes in the range of about 40 to about 800 top oxidant apertures therein . Most preferably, the oxidant sprinkler 34 comprises in the range of 60 to 400 the oxidant top openings 68 therein. The oxidant sprinkler device 34 preferably includes at least about 1 lower oxidant opening 70 therein. More preferably, the oxidant sprinkler device 34 has a range of about 2 to about 40 lower oxidant openings 70 therein. oxidizer 34 has in the range of 8 to 20 bottom openings for oxidizer 70 therein. The ratio of the number of upper holes to the oxidant 68 to the lower holes to the oxidant 70 in the spraying device to the oxidant 34 is preferably in the range from about 2: 1 to about 100: 1, more preferably in the range from about 5: 1 to about 25: 1 and most preferably in range from 8: 1 to 15: 1. The diameters of essentially all upper and lower openings to the oxidant 68.70 are preferably substantially the same, therefore the ratio of the volumetric flow rate of the oxidant stream from the upper and lower openings to the oxidant 68.70 is essentially the same as the ratios of the relative number of upper and lower openings to the oxidant

68.70 given above.

[0057] FIG. 5 shows the direction of the oxidant from the upper and lower openings to the oxidant 68.70. With respect to the upper oxidant openings 68, it is preferred that at least a portion of the upper oxidant openings 68 discharge the oxidant stream at an "A" angle off the vertical. It is preferred that the percentage of upper oxidant openings 68 deviated from the vertical by the "A" angle is in the range of about 30 to about 90 percent, more preferably in the range of about 50 to about 80 percent, even more preferably in the range of 60 to 75 percent and most preferably it was about 67 percent. The angle "A" preferably remains in the range of about 5 to about 60 degrees, more preferably in the range of about 10 to about 45 degrees, and most preferably in the range of 15 to 30 degrees. With regard to the bottom holes for oxidizer 70, it is preferred that substantially all bottom holes for oxidizer 70 are located at the lowest part of the annular member 60 and / or cross member 62. Therefore, any liquids and / or solids that may accidentally enter to the oxidant sprinkler 34, it can easily be removed from the oxidant sprinkler 34 through the bottom oxidant openings 70. Preferably, the bottom openings for the oxidant 70 draw the oxidant stream downward in a substantially vertical direction. For the purposes of this specification, the top hole for the oxidant may be any hole that discharges the oxidant stream in a substantially upward direction (i.e. at an angle above level), and the bottom hole for the oxidant can be any hole that discharges the oxidant stream in a substantially downward direction (i.e. at an angle below level).

[0058] In many conventional bubble column reactors containing a multi-phase reaction medium substantially all of the reaction medium downstream of the oxidant spraying device (or other mechanism for introducing the oxidant stream into the reaction zone) has a very low gas retention value. It is known in the art that "gas retention" is simply a fraction of the volume of the multiphase medium present in the gaseous state. Zones with low gas retention can also be referred to as "non-aerated" zones. In many conventional bubble column reactors, a significant portion of the total volume of the reaction medium is below the oxidant spray device (or other mechanism for introducing the oxidant stream into the reaction zone). Therefore, a significant portion of the reaction medium located at the bottom of conventional bubble column reactors is not aerated.

[0059] It has been found that reducing to a minimum the size of non-aerated zones in the oxidation reaction medium in the bubble column reactor allows to minimize the formation of some types of undesirable impurities. Non-aerated zones of the reaction medium contain relatively few oxidant bubbles. Such a small volume of oxidant bubbles reduces the amount of molecular oxygen available for dissolution in the liquid phase in the reaction medium. Therefore, the liquid phase in the non-aerated zone of the reaction medium has a relatively low concentration of molecular oxygen. In such non-aerated zones with oxygen deficiency, there is a tendency in the reaction medium favoring undesirable side reactions instead of the desired oxidation reaction. For example, if para-xylene is oxidized partially forming terephthalic acid, the insufficient availability of oxygen in the liquid phase in the reaction medium may result in the formation of undesirably large amounts of benzoic acid and condensed aromatic rings, including in particular the very undesirable colored particles known as fluorenones and anthraquinones .

[0060] According to one embodiment, the liquid phase oxidation is carried out in a bubble column reactor whose configuration and operation are such that a fraction of the volume of the reaction medium with a low gas retention value is kept to a minimum. Minimizing non-aerated zones can be quantified by theoretically dividing the entire volume of the reaction medium into 2000 separate horizontal slices of uniform volume. In addition to the highest and lowest horizontal sections, each horizontal section is a separate volume limited on the sides by the side wall of the reactor and limited on top and bottom by theoretical horizontal planes. The highest horizontal slice is limited below by the theoretical horizontal plane and above by the upper surface of the reaction medium. The lowest horizontal section is limited at the top by the theoretical horizontal plane and at the bottom by the bottom end of the tank. After theoretical division of the reaction center into 2000 separate horizontal slices of equal volume, time-averaged and volume-averaged gas retention in individual horizontal slices can be determined. If this method of quantifying the content of non-aerated zones is used, it is preferred that the number of horizontal slices in which the time-averaged and volume-averaged gas retention is less than 0.1, less than 30, more preferably less than 15, more preferably less than 6 , even more preferably less than 4 and most preferably less than 2. It is preferred that the number of horizontal slices in which gas retention is less than 0.2, less than 80, more preferably less than 40, more preferably less than 20, even more preferably less than 12, and most preferably less than 5. It is preferred that the number of horizontal slices in which gas retention is less than 0.3, is 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.

[0061] With reference again to FIG. 1 and 2, it has been found that placing the spraying device for the oxidant 34 lower in reaction zone 28 provides a number of benefits, including limiting the content of non-aerated zones in reaction medium 36. At a given height "H" of reaction medium 36, length "L" of reaction zone 28 and the maximum diameter "D" of reaction zone 28, it is preferred that the majority (i.e. > 50 weight percent) of the oxidant stream introduced into reaction zone 28 within about 0.025H, 0.022L and / or 0.25D from the lower end 52 of reaction zone 28. More preferably most of the oxidation stream is introduced into reaction zone 28 within about 0, 02H, 0.018L and / or 0.2D from the bottom end of the 52 reaction zone

28. Most preferably, most of the oxidant stream is introduced into reaction zone 28 within 0.015H, 0.013L and / or 0.15D from the lower end 52 of reaction zone 28.

[0062] In the embodiment shown in FIG. The vertical distance "Y1" between the bottom end 52 of the reaction zone 28 and the outlet of the top oxidant apertures 68 in the oxidant sprayer 34 is less than about 0.25H, 0.022L and / or 0.25D, so that substantially all of the oxidant stream enters to reaction zone 28 within about 0.25H, 0.022L and / or 0.25D from the bottom end of reaction zone 28. More preferably, Y1 is less than about 0.02H, 0.018L and / or 0.2D. Most preferably Y1 is less than 0.015H, 0.013L and / or 0.15D but greater than 0.005H, 0.004L and / or 0.06D. FIG. 2 shows the tangent line 72 at this point, where the lower edge of the cylindrical main body 46 of the tank shell 22 joins with the upper edge of the elliptical lower head 48 of the tank shell 22. Alternatively, bottom 48 may have any shape, including conical, and the tangent line is still defined as the lower edge of the cylindrical main body 46. The vertical distance "Y2" between the tangent line 72 and the top surface of the sprinkler device to the oxidant 34 is preferably at least about 0.0012H, 0.001L 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 vertical distance "Y3" between the bottom end 52 of the reaction zone 28 and the outlet of the bottom openings to the oxidant 70 of the spraying device to the oxidant 34 is preferably less than about 0.015H, 0.013L and / or 0.15D, more preferably less than about 0.012H, 0, 01L and / or 0.1D and most preferably less than 0.01H, 0.008L and / or

0.075D, but more than 0.003H, 0.002L and / or 0.025D.

[0063] In a preferred embodiment, the holes for the oxidant stream and the feed stream into the reaction zone are configured such that the amount (by weight) of the oxidant stream or feedstock output from the hole is directly proportional to the surface of the hole. Therefore, for example, if 50 percent of the total area of ​​the holes of all oxidant holes is within about 0.15D from the bottom of the reaction zone, then 50 percent by weight of the oxidant stream enters the reaction zone within 0.15D from the bottom of the reaction zone and vice versa.

[0064] In addition to the benefits of minimizing non-aerated zones (i.e. zones with low gas retention) in reaction medium 36, it has been found that oxidation can be improved by maximizing gas retention throughout the entire reaction medium 36. Time-averaged and averaged reaction medium 36 the gas retention by volume is preferably at least about 0.4, more preferably it is in the range of about 0.6 to about 0.9 and most preferably in the range of 0.65 to 0.85. A series of physical and operational parameters of the bubble column reactor 20 contribute to the large gas retention referred to above. For example, for a given reactor size and oxidant stream flow, the high L: D ratio in reaction zone 28 results in a lower diameter increasing velocity based on the cross section of the empty reactor in reaction medium 36, which in turn increases gas retention. In addition, it is known that the actual diameter of the bubble column and the L: D ratio have an effect on the average gas retention even for a given constant speed calculated per section of the empty reactor. In addition, minimizing non-aerated zones, in particular at the bottom of reaction zone 28, increases the gas retention value. In addition, the pressure above the medium and the mechanical configuration of the bubble column reactor may affect the stability of operation at high speed calculated on the cross-section of the empty reactor and gas retention values ​​disclosed herein.

[0065] In addition, the inventors have discovered the importance of working at optimized pressure over the medium, which results in greater gas retention and increased mass transfer. It would seem that working at a lower pressure over the medium, which reduces the solubility of molecular oxygen according to Henry's law, would reduce the rate of mass transfer of molecular oxygen from gas to liquid. In a mechanically mixed tank this usually happens because the level of aeration and the mass transfer rate are mainly influenced by the design of the stirrer and the pressure above the medium. However, it has been found in the bubble column reactor according to the preferred embodiment of the present invention how to use a lower pressure over the medium so that a given mass of the oxidant stream in the gas phase occupies a larger volume, which increases the speed calculated per cross-section of the empty reactor in the reaction medium 36, and this in turn, it will increase gas retention and the rate of molecular oxygen permeation.

[0066] The balance between the merging and separation of gas bubbles is an extremely complex phenomenon, which on the one hand causes foaming, which reduces the rate of internal circulation of the liquid phase and may require extremely large separation zones, while on the other hand it tends to form fewer very large bubbles leading to lower gas retention and a lower mass transfer rate from the oxidant stream to the liquid phase. With regard to the liquid phase, it is known that, among other things, its composition, density, viscosity and surface tension interact in an extremely complex manner giving very complicated results, even in the absence of a solid phase. For example, researchers have found in laboratory methods that it is useful to determine whether water is tap water, distilled water, or deionized water when reporting results and evaluating observations, even for simple water-air bubble columns. For complex mixtures in the liquid phase and the addition of a solid phase, the degree of complexity is even greater. Surface irregularities of individual solid particles, average solids size, particle size distribution, amount of solids relative to the liquid phase and the ability of the liquid to wet the surface of solids are examples of factors all of which are relevant due to interaction with the liquid phase and oxidant stream for obtaining bubbling properties and natural convection flow patterns.

[0067] Therefore, the ability of the bubble column reactor to be useful at high speed based on the cross-section of the hollow reactor and the high gas retention values ​​disclosed herein depends, for example, on the appropriate choice of: (1) the liquid phase composition of the reaction medium; (2) the amount and type of solids precipitated; both of these factors can be regulated by reaction conditions; (3) the amount of oxidant stream entering the reactor; (4) pressure over the medium, which affects the volumetric flow of the oxidant stream, bubble stability and through the energy balance on the reaction temperature; (5) the reaction temperature itself, which affects the fluid properties, properties of the precipitated solids, and the specific volume of the oxidant stream; and (6) the geometry and mechanical details of the reaction vessel, including the L: D ratio.

[0068] With reference again to FIG. 1, it has been found that better distribution of the compound which can be oxidized (para-xylene) in reaction medium 36 is achieved by introducing the liquid phase feed stream into reaction zone 28 at multiple points distant in the vertical plane. Preferably, the liquid phase feed stream is introduced into reaction zone 28 through at least 3 inlets, more preferably through at least 4 inlets. The term "inlet holes" as used herein means holes through which a liquid phase feed stream is discharged into reaction zone 28 for mixing with reaction medium 36. It is preferred that at least 2 inlets are spaced vertically by at least about 0.5D, more preferably at least about 1.5D and most preferably at least 3D. It is however preferred that the highest inlet is vertically spaced from the lowest oxidant opening by no more than about 0.75H, 0.65L and / or 8D, more preferably not more than about 0.5H, 0.4L and / or 5D and most preferably not more than 0.4H, 0.35L and / or 4D.

[0069] Although it is desirable to introduce the liquid phase feed stream at multiple points in the vertical plane, it has further been found that better distribution of the oxidizable compound is obtained in reaction medium 36 if most of the liquid phase feed stream introduced into the lower half of reaction medium 36 and / or reaction zone 28. Preferably, at least about 75 weight percent of the liquid phase feed stream is introduced into the lower half of reaction medium 36 and / or reaction zone 28. Most preferably, at least 90 weight percent of the liquid stream feed stream is introduced into the lower half of the reaction medium 36 and / or zone reaction 28. Furthermore, it is preferred that at least about 30 weight percent of the liquid phase feed stream is introduced into reaction zone 28 within about 1.5D from the lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28. The lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28, is usually at the bottom of the oxidant sprinkler; however, the preferred embodiment of the present invention provides various alternative configurations for introducing the oxidant stream into reaction zone 28. Preferably, at least about 50 weight percent of the liquid phase feedstock is introduced within about 2.5D from the lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28. Preferably at least about 75 weight percent of the liquid phase feedstock is introduced within 5D of the lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28.

[0070] Each inlet determines the area of ​​the opening through which the raw material is discharged. It is preferred that at least about 30 percent of the total openings of all feed inlets be within about 1.5D from the lowest point 49 in the vertical plane, where the oxidant stream is introduced into reaction zone 28. Preferably, at least about 50 percent of the total openings of all feed inlets are within about 2.5D of the lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28. Preferably at least about 75 percent of the total openings of all feed inlets are approximately 5D from the lowest point in the vertical plane, where the oxidant stream is introduced into reaction zone 28.

[0071] With reference again to FIG. 1 in one embodiment, the raw material inlets 32a, b, c, d are simply a series of vertically arranged holes along one side of the tank shell 22. These inlets preferably have a substantially similar diameter of less than about 7 centimeters, more preferably in the range of from about 0, 25 to about 5 centimeters and most preferably in the range of 0.4 to 2 centimeters. The bubble column reactor 20 is preferably equipped with a system controlling the flow rate of the liquid phase feed stream flowing out of each inlet. Such a flow control system preferably includes a separate flow control valve 74a, b, c, d for each respective feed inlet 32a, b, c, d. Furthermore, it is preferred that the bubble column reactor 20 is equipped with a flow control system allowing at least a portion of the liquid phase feed stream to enter the reaction zone 28 at an elevated inlet velocity based on the hollow reactor section of at least about 2 meters per second, more preferably at at least about 5 meters per second, even more preferably at least about 6 meters per second and most preferably in the range of 8 to 20 meters per second. The term "inlet speed calculated per cross-section of an empty reactor" as used herein means the time-averaged volume flow rate of the raw material stream from the inlet divided by the surface of the inlet. Preferably, at least about 50 weight percent of the feed stream is introduced into reaction zone 28 at an elevated inlet velocity based on the cross section of the empty reactor. Most preferably, substantially all of the feed stream is introduced into reaction zone 28 at an elevated inlet velocity based on the cross section of the empty reactor.

[0072] With reference to FIG. 6 and 7, an alternative system for introducing a liquid phase feed stream into reaction zone 28 is shown. In this embodiment, the feed stream is introduced into reaction zone 28 at four different heights. An appropriate distribution system for raw material 76a, b, c, d is installed at every height. Each feedstock distribution system 76 includes feedstock main 78 and branch 80. Each branch 80 is equipped with at least two outlets 82.84 connected to respective insertion pipes 86.88, which enter the reaction zone 28 of the tank shell 22. Each insertion pipe 86.88 has a corresponding inlet opening 87.89 for discharging the feed stream to reaction zones 28. The inlet openings 87.89 preferably have a substantially similar diameter of less than about 7 centimeters, more preferably in the range of about 0.25 to about 5 centimeters, and most preferably in the range of 0.4 to 2 centimeters. It is preferred that the inlet openings 87.89 in each distribution system for feedstock 76a, b, c, d are arranged on opposite sides of the diameter to feed the feedstock stream to reaction zone 28 from opposite directions. Furthermore, it is preferred that the inlet openings 86.88 located on opposite sides of the diameter in adjacent raw material distribution systems 76 are arranged at 90 degrees of rotation relative to each other. During operation, the feed stream in the liquid phase is fed into the main feed line to feed 78 and then goes to branch 80. Branch 80 separates the feed stream evenly to simultaneously feed on opposite sides of the reactor 20 through the holes 87.89 into the feedstock.

[0073] FIG. 8 illustrates an alternative configuration in which each raw material distribution system 76 is equipped with bayonet tubes 90,92 instead of insertion tubes 86,88 (shown in FIG. 7). Bayonet tubes 90, 92 go deep into reaction zone 28 and include a plurality of small inlet holes 94.96 for discharging the liquid phase feedstock to reaction zone 28. It is preferred that the small inlet holes 94.96 of bayonet tubes 90.92 have substantially the same diameter of less than about 50 millimeters, more preferably about 2 to about 25 millimeters and most preferably 4 to 15 millimeters.

[0074] FIG. 9-11 show an alternative distribution system for feedstock 100. The distribution system for feedstock 100 introduces the liquid phase feed stream through a plurality of holes arranged in a vertical and transverse plane, which does not require a plurality of holes in the side wall of the bubble column reactor 20. The distribution system for the raw material 100 essentially includes one inlet conduit 102, manifold 104, a plurality of vertical separation pipes 106, transverse support mechanism 108 and vertical support mechanism 110. Inlet conduit 102 passes through the side wall of main body 46 of tank shell 22. Inlet conduit 102 is connected so as to allow fluid to flow with the manifold 104. The manifold 104 divides the feed stream coming through the inlet conduit 102 evenly between the vertical separating pipes 106. Each separating pipe 106 has a plurality of vertical inlet openings 112a, b, c, d for discharging the raw material stream to reaction zone 28. With each separation pipe 106 a transverse support mechanism 108 is connected which prevents relative transverse movement of the separation pipes 106. The vertical support mechanism 110 is preferably connected to the transverse support mechanism 108 and the top surface of the sprinkler device to the oxidant 34. The vertical support mechanism 110 substantially prevents the vertical movement of the separation pipes 106 in the reaction zone 28. It is preferred that the inlets 112 have substantially the same diameter of less than about 50 millimeters, more preferably about 2 to about 25 millimeters and most preferably 4 to 15 millimeters. The vertical distance of the inlets 112 of the distribution system 100 to the feedstock shown in FIG. 9-11 may be substantially the same as described above with respect to the raw material distribution system of FIG. 1.

[0075] It has been found that the flow patterns of the reaction medium in many bubble column reactors can allow uneven azimuthal distribution of the oxidizable compound in the reaction medium, in particular if the oxidizable compound is introduced mainly along one side of the medium reaction. The term "azimuth" as used herein means an angle or arrangement about the vertical axis of the extension of the reaction zone. The term "vertical" in this document means within 45 ° of the vertical. In one embodiment of the present invention, a feed stream containing a compound that can be oxidized (e.g., para-xylene) is introduced into the reaction zone through a plurality of inlets arranged in an azimuth plane. Such inlet openings arranged in the azimuthal plane can help prevent areas with excessively high and excessively low concentrations of a compound that can be oxidized in the reaction medium. Various feed introduction systems shown in FIG. 6-11 are examples of arrangements where the inlet holes are positioned appropriately in the azimuth plane.

[0076] With reference again to FIG. 7, to quantify the introduction of the feed stream in the liquid phase into the reaction medium in the azimuth plane, the reaction medium can theoretically be divided into four vertical azimuth quadrants "Qi, Q₂, Q₃, Q₄"Of approximately equal volume. These azimuth quadrants "Q₁, Q₂, Q₃, Q₄"Are defined by a pair of theoretical intersecting perpendicular vertical planes" P₁, P₂"Extending beyond the maximum vertical dimension and the maximum radial dimension of the reaction medium. If the reaction center is in a cylindrical tank, the intersection line of the theoretical intersecting vertical planes P₁P₂ approximately coincides with the vertical cylinder center line and each Q azimuth quadrant₁, Q₂, Q₃, Q₄ it is essentially wedge-shaped vertical volume equal to the height of the reaction medium. It is preferred that a significant portion of the compound that can be oxidized is discharged into the reaction medium through inlet openings located in at least two different azimuth quadrants. [0077] In a preferred embodiment, no more than about 54-80 percent by weight of the compound that can be oxidized is discharged into the reaction medium through inlet openings that may be in a single azimuth quadrant. More preferably, no more than about 60 weight percent of the compound that can be oxidized is discharged into the reaction medium through inlets that may be in one azimuth quadrant. Most preferably, no more than 40 percent by weight of the compound that can be oxidized is discharged into the reaction medium through inlets that may be located in one azimuth quadrant. These parameters of the azimuthal distribution of a compound that can be oxidized are measured when the azimuthal quadrants are oriented in the azimuthal plane so that the maximum possible amount of the compound that can be oxidized is derived to one of the azimuthal quadrants. For example, if the entire feed stream is discharged into the reaction medium through two inlet openings that are spaced 89 degrees azimuthal to each other, to determine azimuthal distribution in four azimuth quadrants, 100 weight percent of the feed stream is discharged into the reaction medium in one the azimuthal quadrant, because the azimuth quadrants can be oriented in the azimuth plane in such a way that both inlets are in one azimuth quadrant.

[0078] In addition to the benefits of an appropriate arrangement of the inlets in the azimuth plane, it has been found that the appropriate radial arrangement of the inlets in the bubble column reactor can also be important. It is preferred that a significant proportion of the oxidizable compound introduced into the reaction medium is discharged through inlet openings arranged radially inward from the side wall of the tank. Therefore, in one embodiment of the present invention, a significant portion of the compound that may be oxidized enters the reaction zone through inlet openings in the "preferred radial insertion zone" arranged in a radial plane inward from the vertical side walls defining the reaction zone.

[0079] With reference again to FIG. 7, the preferred radial insertion zone "FZ" may have the shape of a theoretical vertical cylinder with the center in reaction zone 28, with an outer diameter of "D_ABOUT"of 0.9D, where" D "is the diameter of the reaction zone 28. Therefore, between the preferred radial insertion zone FZ and the interior of the sidewall defining the reaction zone 28, the 0.05D outer ring" OA "is defined. It is preferred that the zone reaction 28, little or no compound was introduced through the inlet openings located in this OA outer ring, which could be oxidized.

[0080] In another embodiment, it is preferred that a small or no amount of oxidizable compound is introduced into the center of reaction zone 28. Therefore, as shown in FIG. 8, the preferred radial insertion zone FZ may have the shape of a theoretical vertical ring with the center in the reaction zone 28, with an outer diameter DO of 0.9D and an inner diameter D1 of 0.2D. Therefore, in this embodiment, an inner cylinder IC 0.2D in diameter is cut from the center of the preferred radial insertion zone FZ. It is preferred that a small or no amount of oxidizable compound is introduced into reaction zone 28 through the inlet openings located in this internal cylinder IC.

[0081] In a preferred embodiment, a significant portion of the compound that can be oxidized is introduced into reaction medium 36 through inlet openings in the preferred radial introduction zone, regardless of whether the preferred radial introduction zone has the cylindrical or annular shape described above. More preferably, at least about 25 weight percent of the compound that can be oxidized is discharged into reaction medium 36 through inlet openings in the preferred radial introduction zone. Even more preferably, at least about 50 weight percent of the compound that can be oxidized is discharged into reaction medium 36 through inlet openings in the preferred radial introduction zone. Most preferably, at least 75 weight percent of the compound that can be oxidized is discharged into reaction medium 36 through inlet openings in the preferred radial introduction zone.

[0082] Although the theoretical azimuthal quadrants and theoretical preferred radial insertion zone shown in FIG. 7 and 8 are described with reference to the distribution of the feed stream in the liquid phase, it has been found that appropriate azimuthal and radial distribution of the oxidant stream in the gas phase may also provide some benefits. Therefore, in one embodiment of the present invention, the description of the azimuthal and radial distribution of the liquid phase feed stream, given above, also relates to the method by which the oxidant stream in the gas phase is introduced into the reaction medium

36.

[0083] With reference to FIG. 12-15 shows an alternative spraying device for oxidizer 200 comprising essentially annular member 202 and a pair of inlet conduits for oxidizer 204, 206. Oxidizer Sprinkler 200 of FIG. 12-15 is similar to the oxidant sprinkler 34 of FIG. 1-11, with the following three fundamental differences: (1) the oxidizing sprinkler 200 does not include a diagonal cross member; (2) the upper portion of the annular member 202 has no openings for leading the oxidant up; and (3) the oxidizer sprinkler 200 has significantly more holes in the bottom of the annular member 202.

[0084] As most likely shown best in FIG. 14 and 15, the bottom portion 202 of the oxidant sprinkler device includes a plurality of holes for oxidizer 208. The holes for oxidizer 208 preferably have a configuration such that at least about 1 percent of the total surface area of ​​the holes defined by the oxidant holes 208 is below center line 210 ( FIG. 15) the annular element 202, the centerline 210 being at the height of the centroid of the bulk annular element 202. More preferably, at least about 5 percent of the total area of ​​the holes defined by all oxidant holes 208 is below center line 210, wherein at least about 2 percent of the total area of ​​the holes is determined by the holes 208 leading the oxidant stream in a substantially downward direction within about 30 degrees from the vertical. Even more preferably, at least about 20 percent of the total area of ​​the holes defined by all of the oxidant holes 208 is below center line 210, wherein at least about 10 percent of the total area of ​​the holes is defined by the holes 208 leading the stream of oxidant in a substantially downward direction within degrees from the vertical. Most preferably, at least about 75 percent of the total area of ​​the holes defined by all holes for oxidant 208 is below center line 210, wherein at least about 40 percent of the total area of ​​the holes is defined by holes 208 leading the stream of oxidant in a substantially downward direction within 30 degrees from the vertical. The fraction of the total area of ​​the holes defined by all oxidant holes 208 above center line 210 is preferably less than about 75 percent, more preferably less than about 50 percent, even more preferably less than about 25 percent, and most preferably less than 5 percent.

[0085] As shown in FIG. 14 and 15, oxidant holes 208 include down holes 208a and oblique holes 208b. The falling holes 208a are configured to lead the oxidant stream downwards substantially down to an angle within approximately 30 degrees from vertical, more preferably within approximately 15 degrees from vertical and most preferably within 5 degrees from vertical. The oblique holes 208b have the configuration to discharge the oxidant stream in a substantially outward and downward direction at an "A" angle in the range of about 15 to about 75 degrees from vertical, more preferably the angle A is in the range of about 30 to about 60 degrees from and most preferably the angle A remains in the range of 40 to 50 degrees from the vertical.

[0086] It is preferred that substantially all oxidant holes 208 have approximately the same diameter. The average diameter of oxidant 208 holes is preferably in the range of about 2 to about 300 millimeters, more preferably in the range of about 4 to about 120 millimeters, and most preferably in the range of 8 to 60 millimeters. The total number of holes for oxidizer 208 in ring member 202 is selected to meet the low pressure drop criteria set out below. Preferably, the total number of oxidant apertures 208 formed in the ring member 202 is at least about 10, more preferably the total number of oxidant apertures 208 is in the range of about 20 to about 200, and most preferably the total number of oxidant apertures 208 is in the range of 40 to 100.

[0087] Although FIG. 12-15 show a very specific configuration of the oxidant sprinkler 200, it should be noted that various configurations of the oxidant sprinkler can be used to achieve the benefits described herein. For example, the oxidizer sprinkler need not necessarily have the octagonal ring member configuration shown in FIG. 12-13. Instead, it is possible for the oxidant sprinkler to have any configuration of the fluid line or conduits in which a plurality of spaced openings are provided to discharge the oxidant stream. The size, number and direction of discharge from the oxidant holes in the fluid conduit are preferably within the ranges given above. In addition, the oxidant sprinkler device preferably has a configuration to provide the azimuthal and radial distribution of molecular oxygen as described above.

[0088] Regardless of the specific configuration of the spraying device for the oxidant, it is preferred that the oxidizing spraying device has a physical configuration and operates in a manner that minimizes pressure drop due to the discharge of the oxidant stream from the conduit or lines to the fluid through the oxidant openings and into reaction zones. This pressure drop is calculated as the time-averaged static pressure of the oxidant stream inside the fluid conduit at the oxidant inlets 66a, b oxidant sprayer minus the time-averaged static pressure in the reaction zone at a given height, where half of the oxidant stream is introduced above this point on the line vertical and half of the oxidant stream is introduced below this point on the vertical line. In a preferred embodiment of the present invention, the time-averaged pressure drop due to withdrawing the oxidant stream from the spraying device to the oxidant is less than about 0.3 megapascal (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 preferred operating conditions of the bubble column reactor described herein, the pressure of the oxidant stream within the conduit or conduits to the fluid of the spraying device to the oxidant is preferably in the range of from about 0.35 to about 1 MPa, more preferably in the range of from about 0.45 to about 0, 85 MPa and most preferably in the range of 0.5 to 0.7 MPa. [0089] As mentioned earlier regarding the configuration of the oxidant sprinkler device shown in FIG. 2-5, it may be desirable to continuously or periodically flush the spray device into the oxidant with a liquid (e.g., acetic acid, water, and / or para-xylene) to prevent solids from spraying the device into the oxidant. If such a liquid flush is used, it is preferred that an effective amount of liquid (i.e. not just a small amount of liquid droplets that may naturally be present in the oxidant stream) to pass through the oxidant spray device and through the oxidant openings for at least one period of more than one minute each day. If the liquid is continuously or periodically discharged from the spraying device to the oxidant, it is preferred that the time-average ratio of the liquid mass flow rate through the spraying device to the oxidant to the mass flow rate of molecular oxygen through the spraying device to the oxidant remains in the range of about 0.05: 1 up to about 30: 1 or in the range of from about 0.1: 1 to about 2: 1 or even in the range of 0.2: 1 to 1: 1.

[0090] In one embodiment, a significant portion of the compound that can be oxidized (para-xylene) can be introduced into the reaction zone through an oxidant misting device. In such a configuration, it is preferred that the oxidizable compound and molecular oxygen are withdrawn from the spraying device to the oxidant through the same openings in the oxidant spraying device. As noted above, a compound that can be oxidized is usually a liquid at normal temperature and pressure (STP). Accordingly, in this embodiment, the two-phase stream may be output from the spraying device to the oxidant, such that the liquid phase contains a compound that can be oxidized and the gas phase contains molecular oxygen. It should be understood, however, that at least a portion of the compound that can be oxidized may be in a gaseous state when it exits the spraying device into the oxidant. In one embodiment, the aqueous phase discharged from the spraying device into the oxidant is predominantly a compound that can be oxidized. In another embodiment, the aqueous phase discharged from the spraying device into the oxidant has substantially the same composition as the feed stream described above. If the aqueous phase exiting the spraying device into the oxidant has essentially the same composition as the feed stream, such liquid phase may contain the solvent and / or catalyst system in the amounts and ratios described above with respect to the composition of the feed stream.

[0091] In one embodiment, it is preferred that at least about 10 weight percent of the total oxidizable compound introduced into the reaction zone is introduced through the spraying device into the oxidant, more preferably at least about 40 weight percent of the oxidizable compound. introduce through the spraying device into the oxidant and most preferably at least 80 percent by weight of the compound that can be oxidized, introduce into the reaction zone through an oxidant sprinkler. If all or part of the compound that may be oxidized is introduced into the reaction zone through an oxidant spraying device, it is preferred that at least about 10 weight percent of all molecular oxygen introduced into the reaction zone is introduced through the same spraying device into the oxidant, more preferably at least about 40 weight percent of a compound that can be oxidized, introduce through the same sprinkler into the oxidant and most preferably at least 80 percent by weight of the compound which may be oxidized, introduce into the reaction zone through the same sprinkler into the oxidant. If a significant proportion of the compound that can be oxidized is introduced into the reaction zone through an oxidant misting device, it is preferred that the oxidant misting device has one or more temperature measuring devices (e.g., thermocouples). Such temperature sensors can be used to make it easier to guarantee that the temperature of the oxidant sprinkler will not be dangerously high.

[0092] Referring to Figs. 16-18, a bubble column reactor 20 showing an internal vent tank 300 located at the bottom of reaction zone 28 near the slurry outlet 38 is shown. Side reactions in which impurities are formed occur with relatively high intensity during the venting of the reaction medium 36. The term "venting" herein means the separation of the gas phase from the multiphase reaction medium. If reaction medium 36 is heavily aerated (> 0.3 gas retention), the formation of impurities is minimal. If reaction medium 36 is heavily unventilated (<0.01 gas retention), the formation of impurities is also minimal. However, if the reaction medium is partly non-aerated (0.01-0.3 gas retention), this promotes undesirable side reactions and an increased amount of impurities is formed. The vent tank 300 solves this and other problems by minimizing the volume of the reaction medium 36 in a partially aerated state and minimizing the venting time of the reaction medium 36. A substantially vented slurry forms at the bottom of the venting tank 300, which exits the reactor 20 through the slurry outlet 38. The substantially deaerated suspension preferably contains less than about 5 volume percent of the gas phase, more preferably less than about 2 volume percent of the gas phase, and most preferably less than 1 volume percent of the gas phase.

[0093] In FIG. 16 shows a bubble column reactor 20 comprising a level sensor 302 and a flow control valve 304. The level sensor 302 and a flow control valve 304 work together to keep reaction center 36 at a substantially constant height in reaction zone 28. The level 302 sensor works to measure (eg. by measuring the pressure difference level or nuclear level measurement) the height of the upper surface 44 of the reaction medium 36 and produces a control signal 306 corresponding to the height of the reaction medium 36. The flow control valve 304 receives the control signal 306 and corrects the suspension flow rate through the discharge line to the suspension 308. In this way, the slurry flow rate through the slurry outlet 38 may vary from the maximum slurry volumetric flow rate (Fmax) if the height of the reaction medium 36 is too high to the minimum slurry volumetric flow rate (Fmin) if the height of the reaction medium 36 is too low .

[0094] To discharge the solid phase oxidation product from reaction zone 28, a portion thereof must first pass through the vent tank 300. The vent tank 300 provides a low turbulence internal volume that allows the gas phase in the reaction medium 36 to naturally rise above the phases liquid and solid reaction medium 36, when liquid and solids flow down towards the slurry outlet 38. The rise of the gas phase over the liquid and solid phases is due to the upward natural buoyancy force of the gas phase in the liquid and solid phases. When a deaeration vessel 300 is used, the transformation of reaction medium 36 from a fully aerated three-phase medium to a fully deaerated two-phase suspension is fast and efficient. [0095] With reference to Fig. 17 and 18, the vent tank 300 includes a substantially vertical sidewall 308 defining an inward venting zone 312. Preferably, sidewall 308 is directed up approximately 30 degrees from vertical, more preferably approximately 10 degrees from vertical. Most preferably sidewall 308 is substantially vertical. The venting zone 312 is separate from the reaction zone 28 and has a height "h" and a diameter "d". The upper end 310 of sidewall 308 is open so that it receives a reaction medium from reaction zone 28 to an internal volume 312. The lower end of sidewall 308 is fluidly connected to the slurry outlet 38 through transition compartment 314. In some cases, for example when the slurry outlet opening 38 is large or the diameter "d" of sidewall 308 is small, the transition interval 314 can be eliminated. As most likely shown best in FIG. 18, the venting tank 300 may also include an anti-vortex device 316 located in the deaeration zone 312. The anti-vortex device 316 may be any structure designed to inhibit the formation of vortices when solid and liquid phases flow down towards the slurry outlet 38.

[0096] To enable proper separation of the gaseous phase from the solid and liquid phases in the venting tank 300, the height "h" and the horizontal cross-sectional area of ​​the internal venting zone 312 must be carefully determined. The height "h" and the horizontal cross-sectional area of ​​the internal venting zone 312 should provide sufficient distance and time so that even when the maximum amount of suspension is discharged (i.e. when the suspension is discharged at Fmax), substantially all of the gas bubble volume could rise above the solid and liquid phases before the gas bubbles reach the outlet of the lower vent tank 300. It is therefore preferred that the cross-sectional area of ​​the vent zone 312 is such that the maximum downward velocity (Vdmax) of the liquid and solid phases through the vent zone 312 is significantly lower than the natural rate of rise (Vu) of the gas phase bubbles through the liquid and solid phases. The maximum downward velocity (Vdmax) of the liquid and solid phases through the venting zone 312 occurs at the maximum slurry flow rate (Fmax) discussed above. The natural rate of rise (Vu) of gas bubbles through the liquid and solid phases varies depending on the size of the bubbles; however, the natural ascent rate (Vu0.5) for gas bubbles with a diameter of 0.5 centimeters moving through the liquid and solid phases can be used as the cut-off value, because the volume of essentially all bubbles initially in reaction medium 36 will be greater than 0.5 centimeter. Preferably, the cross sectional area of ​​the venting zone 312 is such that Vdmax is less than about 75 percent Vu0.5, more preferably Vdmax is less than about 40 percent Vu0.5, most preferably Vdmax is less than 20 percent Vu0.5.

[0097] The downward velocity of liquid and solid phases in deaeration zone 312 of deaeration tank 300 is calculated as the volumetric flow rate of deaerated suspension through the suspension outlet 38 divided by the minimum cross-sectional area of ​​deaeration zone 312. The downward velocity of the liquid and solid phases in deaeration zone 312 of deaeration tank 300 is preferably less than about 50 centimeters per second, more preferably less than about 30 centimeters per second and most preferably less than 10 centimeters per second.

[0098] It should be noted here that although the vertical side wall 308 of the vent tank 300 is shown in a cylindrical configuration, the side wall 308 may include a plurality of side walls forming different configurations (e.g., triangular, square or oval), provided that the walls these determine the internal volume having the appropriate volume, cross-sectional area, width "d" and height "h". In a preferred embodiment of the present invention, "d" remains in the range of about 0.2 to about 2 meters, more preferably in the range of about 0.3 to about 1.5 meters, and most preferably in the range of 0.4 to 1.2 meters . In a preferred embodiment of the present invention, "h" remains in the range of about 0.3 meters to about 5 meters, more preferably in the range of about 0.5 to about 3 meters, and most preferably in the range of 0.75 to 2 meters.

[0099] In a preferred embodiment of the present invention, sidewall 308 is substantially vertical, such that the horizontal cross-sectional area of ​​the vent zone 312 is substantially constant over the entire height "h" of the vent zone 312. Preferably, the maximum horizontal cross-sectional area of ​​the vent zone 312 is less than about 25 percent of the maximum horizontal cross-sectional area of ​​reaction zone 28. More preferably, the maximum horizontal cross-sectional area of ​​the vent zone 312 is in the range of about 0.1 to about 10 percent of the maximum horizontal cross-sectional area of ​​the reaction zone 28. Most preferably, the maximum horizontal cross-sectional area of ​​vent 312 is in the range of 0.25 to 4 percent the maximum horizontal cross-sectional area of ​​the reaction zone 28. Preferably, the maximum horizontal cross-sectional area of ​​the vent zone 312 remains in the range of from about 0.02 to about 3 square meters, more preferably in the range of from about 0.05 to about 2 square meters, most preferably in the range of from 0.1 to 1.2 square meters. The volume of vent zone 312 is preferably less than about 5 percent of the total volume of reaction medium 36 or reaction zone 28. More preferably, the volume of vent zone 312 remains in the range of about 0.01 to about percent of the total volume of reaction medium 36 or reaction zone 28. Most preferably the volume of vent zone 312 remains in the range of 0.05 to about 1 percent of the total volume of reaction medium 36 or reaction zone 28. The volume of vent zone 312 is preferably less than about 2 cubic meters, more preferably it is in the range of from about 0.01 to about 1 cubic meter, and most preferably in the range of from 0.05 to 0.5 cubic meters.

[0100] Turning now to FIG. 19 shows a bubble bubbler reactor 20 comprising an external vent tank 400. In this configuration, the aerated reaction medium 36 is discharged from the reaction zone 28 through an erected side opening in the tank jacket 22. The discharged aerated medium is transported to the external vent tank 400 via outlet pipe 402 to separation of gas phase from solid and liquid phases. The separated gas phase exits deaeration vessel 400 via conduit 404, and substantially deaerated suspension exits deaeration vessel 400 via conduit 406.

[0101] In FIG. 19 shows outlet conduit 402 as approximately straight, horizontal, and perpendicular to the shell of the tank

22. This is just one of the convenient configurations, and the exhaust line 402 can be different in every sense as long as it usefully connects the bubble column reactor 20 to the external vent tank 400. Turning to conduit 404, it is useful for the conduit to be connected to or near the top surface of the vent tank 400, whereby the safety issues associated with a stationary gas bubble containing a compound that can be oxidized and an oxidant are controlled. In addition, conduits 402 and 404 may usefully include flow isolation means, e.g., valves.

[0102] If the reaction medium 36 is discharged from the reactor 20 through the raised outlet, as shown in FIG. 19, it is preferred that the bubble column reactor 20 is equipped with a bottom outlet 408 at the bottom 52 of the reaction zone 28. The bottom outlet 408 and the bottom line 410 connected thereto can be used to remove the contents (i.e. emptying) of the reactor 20 during standstill. Preferably, one or more of the lower holes 408 are in the lower third of the height of the reaction medium 36, more preferably the lower quarter of the reaction medium 36, and most preferably at the lowest point of the reaction zone 28.

[0103] In the case of an elevated suspension drainage and venting system shown in FIG. 19, bottom line 410 and outlet 408 are not used to drain slurry from reaction zone 28 during oxidation. It is known in the art that solids usually settle under gravity in non-aerated and otherwise unmixed parts of the suspension, including in lines with standing fluid. In addition, solids that have settled (e.g. terephthalic acid), they tend to solidify into large agglomerates due to continuous precipitation and / or remodeling of crystals. Therefore, to prevent blockage of the bottom line 410 to the fluid, a portion of the vented slurry from the bottom of the vent tank 400 can be used to continuously or periodically flush the bottom line 410 during normal operation of the reactor 20. The preferred way to provide such a flush with the slurry of the line 410 is to periodically open the valve 412 in the line 410 and allow the vented portion of the slurry to flow through the conduit 410 to the reaction zone 28 through the bottom opening 408. Even if the valve 412 is fully or partially open, only a portion of the vented slurry flows through bottom line 410 and back to reaction zone 28. The remaining portion of the vented slurry not used for flushing the bottom line 410 is discharged through line 414 from the reactor 20 for further processing (e.g. purification).

[0104] During normal operation of the bubble column reactor 20 for a longer time (e.g.> 100 hours), it is preferred that the amount of vented slurry used to flush bottom line 410 is less than 50 weight percent of the total vented slurry formed at the bottom of vent tank 400, more preferably less than about 20 weight percent and most preferably less than 5 weight percent. It is further preferred that, for a longer period of time, the average mass flow rate of the vented slurry used to flush the lower conduit 410 is less than about 4 times the average mass mass flow rate of the compound that can be oxidized to reaction zone 28, more preferably less than about 2- times the average mass flow rate of the compound which may be oxidised to reaction zone 28, even more preferably less than the average mass mass flow of the compound which can be oxidized to reaction zone 28 and most preferably less than 0.5 times the average mass mass flow rate of the compound which can be oxidized to reaction zone 28. [0105] With reference again to Figure 19, the vent tank 400 includes a substantially vertical, preferably cylindrical sidewall 416 defining a vent zone 418. The vent zone 418 has a diameter "d" and a height "h". The height "h" is measured as the vertical distance between the point where the aerated reaction medium enters the vent tank 400 and the bottom of the side wall 416. Height "h", diameter "d", surface and volume of vent zone 418 is preferably substantially the same as described above with respect to vent zone 312 of vent tank 300 shown in FIG. 16-18. Furthermore, the vent tank 400 includes an upper portion 420 formed by sidewall 416 above the vent zone 418. The upper portion 420 of the vent tank 400 may be of any height, however it is preferably directed up or above the level of reaction medium 36 in reaction zone 28. Upper portion 420 provides space for the gas phase to be properly separated from the liquid and solid phases before leaving the vent tank 400 via conduit 404. It should be noted at this point that, although the passageways 404 are shown as returning the separated gas phase to the separation zone in the reactor 20, the conduit 404 can optionally be connected to the shell of the tank 22 at any height above the discharge conduit 402. The conduit 404 can optionally be connected to the gas outlet conduit 40, so that the separated gas phase from the vent tank 400 connects to the vapor stream discharged from above the medium in the conduit 40, sent for further processing.

[0106] Turning now to FIG. 20 shows a bubble bubbler reactor 20 comprising a hybrid internal and external vent tank 500. In this configuration, part of the reaction medium 36 is discharged from the reaction zone 28 through a relatively large, raised opening 502 in the side wall of the tank jacket 22. The discharged reaction medium 36 is then sent through angular conduit 504 with a relatively large diameter and hits the top of the vent tank 500. In FIG. twenty shown is an angular conduit 504 connected perpendicularly to the side wall of the shell of the tank 22, which includes a gentle bend at an angle of about 90 degrees. This is just one of the convenient configurations, and the angled conduit 504 can be different in every sense as long as it usefully connects the bubble column reactor 20 to the external vent tank 500 as described. In addition, the angled conduit 504 may usefully contain flow isolation means, e.g., valves.

[0107] In the vent tank 500, the gas phase moves up, while the solid and liquid phases move down. The upwardly moving gas phase can again go to the angled conduit 504 and then exit through the opening 502 again into the reaction zone 28. In this way, there may be a counter-current flow at the opening 502 of the incoming reaction medium 36 and the outgoing separated gas. The vented suspension leaves the vent tank 500 via conduit 506. The vent tank 500 comprises a substantially vertical, preferably cylindrical sidewall 508 defining a vent zone 510. The vent zone 510 has a height "h" and a diameter "d". It is preferred that the raised opening 502 and the angled conduit 504 have a diameter equal to or larger than the diameter "d" of the venting zone 510. The height "h", diameter "d", surface and volume of vent zone 510 are preferably substantially the same as described above with respect to vent zone 312 of vent tank 300 shown in FIG. 16-18.

[0108] FIG. 19 and 20 show an embodiment of bubble column reactor 20 in which a solid product (e.g., crude terephthalic acid) produced in reaction zone 28 is withdrawn from reaction zone 28 through an elevated outlet. Draining aerated reaction medium 36 from a raised place above the bottom of bubble column reactor 20 avoids accumulation and stagnation of the poorly aerated reaction medium 36 at the bottom of reaction zone 28.

[0109] The concentration of oxygen and the oxidizable compound (para-xylene) in the reaction medium 36 near the upper surface of the reaction medium 36 is preferably lower than at the bottom. Therefore, discharging reaction medium 36 from a raised site can increase efficiency by reducing the amount of unreacted substrates discharged from reactor 20. In addition, the temperature of reaction medium 36 varies significantly in the vertical direction when the bubble column reactor 20 operates under high STR values ​​and chemical composition gradients disclosed herein. Under such conditions, the temperature of reaction medium 36 will typically have local minima at the lower and upper ends of reaction zone 28. At the bottom end, the minimum corresponds to solvent evaporation near where all or part of the oxidant is fed. At the upper end, the minimum is again due to the evaporation of the solvent, however here as a result of the pressure drop in the reaction medium. In addition, other local minima may occur between the upper and lower ends as long as additional feed or oxidant is introduced into the reaction medium. Therefore, there are one or more temperature maxima associated with the heat of exothermic oxidation reactions between the lower end and the upper end of reaction zone 28. The discharge of the reaction medium 36 at a higher temperature site may be particularly advantageous if further processing takes place at a higher temperature, because the energy costs associated with heating the discharged medium before further processing are reduced.

[0110] Therefore, in a preferred embodiment and in particular, if further processing takes place at a higher temperature, the reaction medium 36 is discharged from the bubble column reactor 20 through an elevated outlet or outlets above the site or places where at least 50 weight percent of the stream the liquid phase feedstock and / or the gas phase oxidant stream goes to reaction zone 28. More preferably, the reaction medium 36 is discharged from the bubble column reactor 20 through an elevated outlet or outlets above the point or places where substantially all of the liquid stream feedstock and / or the gas phase oxidant stream goes to reaction zone 28. Preferably, at least 50 percent by weight of the solid and liquid phase components are discharged from the bubble column reactor through the erected outlet or outlets. More preferably, substantially all solid phase and liquid phase components are discharged from the bubble column reactor 20 through the erected outlet or outlets. Preferably, the raised outlet or outlets are at least about 1D above the bottom end 52 of reaction zone 28. More preferably, the raised outlet or outlets are at least about 2D above the bottom end 52 of reaction zone 28. Most preferably, the raised outlet or outlets are at least 3D above the bottom end 52 of reaction zone 28. At a given height "H" of the reaction medium 36, it is preferred that the raised outlet or outlets are in a vertical plane from about 0.2H to about 0.8H , more preferably from about 0.3H to about 0.7H and most preferably from 0.4H to 0.6 H. Furthermore, it is preferred that the temperature of reaction medium 36 at the raised outlet of reaction zone 28 is at least 1 ° C higher than the temperature of reaction medium 36 at the lower end 52 of reaction zone 28. More preferably, the temperature of reaction medium 36 at the raised outlet of reaction zone 28 is in the range of about 1.5 to about 16 ° C above the temperature of reaction medium 36 at the lower end 52 of reaction zone 28. Most preferably, the temperature of reaction medium 36 at the raised outlet of reaction zone 28 is in the range of 2 to 12 ° C higher than the temperature of reaction medium 36 at the lower end 52 of reaction zone 28.

[0111] Further referring to FIG. 21 shows a bubble column reactor 20 containing an alternative hybrid vent tank 600 located at the bottom of the reactor

twenty. In this configuration, the aerated reaction medium 36 is discharged from the reaction zone 28 through a relatively large opening 602 in the lower end 52 of the tank jacket 22. The opening 602 defines the open upper end of the vent tank 600. In the vent tank 600, the gas phase moves up and the solid phases and liquid are moving down. The upwardly moving gas phase can again enter the reaction zone 28 through the opening 602. In this way, countercurrent flow of the incoming reaction medium 36 and the outgoing separated gas can occur at the opening 602. The vented suspension exits the vent tank 600 through conduit 604. The vent tank 600 includes a substantially vertical, preferably cylindrical sidewall 606 defining vent zone 608. Vent zone 608 has a height "h" and a diameter "d". It is preferred that the opening 602 has a diameter equal to or larger than the diameter "d" of the venting zone 608. The height "h", diameter "d", the surface and volume of the venting zone 608 are preferably substantially the same as described above in relation to the venting zone 312 deaeration tank 300 shown in FIG. 16-18. [0112] Further referring to FIG. 22 shows the bubble column reactor 20 of FIG. 21 containing an alternative oxidant sprinkler 620. The oxidant sprinkler 620 includes an annular member 622 and a pair of inlet lines 624.626. Ring element 622 preferably has substantially the same configuration as ring element 202 described above with reference to FIG. 12-15. The inlet ducts 624.626 extend up through the holes in the bottom head 48 in the shell of the tank 22 and supply the oxidant stream to the annular member 622.

[0113] Further referring to FIG. 23 shows the bubble column reactor 20 of FIG. 21 comprising an element for introducing the oxidant stream into the reaction zone 28 without the spraying device. In the configuration of FIG. 23 the oxidant stream is fed to the reactor 20 via oxidant lines 630,632. Oxidizer lines 630.632 are connected to respective oxidant openings 634.636 in the bottom head 48 of the tank jacket 22. The oxidant stream is fed directly to reaction zone 28 through the oxidant openings 634.636. Optional 638,640 reflective plates can be used which deflect the flow of the oxidant stream when it initially goes to reaction zone 28.

[0114] As mentioned above, it is preferred that the configuration and operation of the oxidation reactor are such that no high concentration zones of the compound which can be oxidized are formed in the reaction medium, since such zones can lead to the formation of impurities. One way to increase the initial dispersion of an oxidizable compound (para-xylene) in the reaction medium is to oxidize the oxidizable compound by dilution with a liquid. The liquid used to dilute the oxidizable compound may be from a part of the reaction medium located at a considerable distance from the place or places where the oxidizable compound is introduced into the reaction zone. Such a liquid from a distant part of the reaction medium may circulate to a location close to the place of introduction of the compound, which may be oxidized, through a fluid conduit located inside and / or outside the main reaction vessel.

[0115] FIG. 24 and 25 show two preferred methods for circulating liquids from a distant portion of the reaction medium to a location near the inlet of a compound that can be oxidized by means of an internal (FIG. 24) or external (FIG. 25) conduit. Preferably, the length of the fluid conduit from its inlet (i.e., the opening or openings through which the liquid enters the conduit) to its outlet (i.e. or holes through which the liquid is discharged from the conduit) is greater than about 1 meter, more preferably greater than about 3 meters, even more preferably larger than about 6 meters, and most preferably greater than 9 meters. However, the actual length of the conduit is less important if the liquid is obtained from a separate tank, which may be located directly above or next to the tank, which is initially fed with a stream of compound that may be oxidized. The liquid from any separate reservoir containing at least a portion of the reaction medium is a preferred source for the initial dilution of the compound that may be oxidized.

[0116] It is preferred that the liquid flowing through the conduit, regardless of its origin, has a lower constant concentration of the compound that can be oxidized than the reaction medium directly at at least one outlet of the conduit. Furthermore, it is preferred that the liquid flowing through the conduit has a concentration of oxidizable compound in the liquid phase below about 100,000 ppmw, more preferably below about 10,000 ppmw, even more preferably below about 1000 ppmw and most preferably below 100 ppmw, the concentration being measured before introducing into the conduit portions of a stream of compound that can be oxidized, and any optional separate solvent stream. If the measurement is carried out after introducing a portion of the compound stream that can be oxidized and any solvent stream, it is preferred that the combined liquid stream entering the reaction medium has a concentration of the compound which can be oxidized in the liquid phase below about 300,000 ppmw, more preferably below about 50,000 ppmw and most preferably below 10,000 ppmw. [0117] It is desirable to keep the flow through the conduit at a low enough rate that the circulating liquid significantly reduces the desired overall gradient of the compound that can be oxidized in the reaction medium. In this regard, it is preferred that the mass ratio of the liquid phase in the reaction zone into which a portion of the oxidizable compound is initially introduced to the mass flow rate of liquid flowing through the conduit is greater than about 0.3 minutes, more preferably greater than about 1 minutes, even more preferably from about 2 minutes to about 120 minutes and most preferably from 3 minutes to 60 minutes.

[0118] There are many ways to cause a liquid to flow through a conduit. Preferred methods include gravity, ejectors of all types in which gas or liquid are used as propellant, or both and mechanical pumps of all types. If an ejector is used, in one embodiment of the invention at least one fluid selected from the group consisting of the following may be used as the driving fluid: stream of a compound that can be oxidized (liquid or gas), stream of oxidant (gas), stream of solvent (liquid ) and the pumped source of the reaction medium (suspension). In another embodiment, at least two fluids selected from the group consisting of a oxidizable compound stream, an oxidant stream, and a solvent stream are used as the propelling fluid. In yet another embodiment, a combined stream of the oxidizable compound, oxidant stream, and solvent stream is used as the propelling fluid.

[0119] The appropriate diameter or diameters of the circulation line may vary depending on the amount and properties of the material being transferred, the energy available to induce fluid movement, and after capital costs. It is preferred that the minimum diameter of such a duct is greater than about 0.02 meters, more preferably from about 0.06 meters to about 2 meters, and most preferably from 0.12 to 0.8 meters. [0120] As mentioned above, it is desirable to control flow through the conduit in certain preferred ranges. Many methods are known in the art to accomplish this control by determining the appropriate geometry constant when designing the fluid line. Another preferred embodiment is the use of variable geometry during operation, in particular including valves of all types and forms, including manually operated and driven by any method, including control loops with or without feedback from the sensor. Another preferred way to control the flow of dilution liquid is to change the energy supplied between the inlet and outlet of the conduit. Preferred methods include changing the flow rate of one or more driving fluids into the ejector, changing the energy supplied to the pump motor, and changing the density difference or elevation difference if gravity is used. These preferred methods can also be used in all combinations.

[0121] The conduit used to circulate the liquid from the reaction medium may be of any type known in the art.

In one embodiment, a conduit constructed wholly or partly from conventional pipe materials is used. In another embodiment, a conduit constructed wholly or partly using the wall of the reaction vessel as one conduit is used. The conduit may be constructed in such a way that it is completely enclosed by the boundaries of the reaction vessel (FIG. 24) or it may be constructed entirely outside the reaction vessel (FIG. 25) or it may contain parts located both inside and outside the reaction vessel.

[0122] The inventors envisage that in particular in larger reactors it may be desirable to have a plurality of lines with different designs for flowing liquid through the line. In addition, it may be desirable to provide multiple outlets at multiple locations on one or all of the ducts. Design details must balance the desired overall gradient of the constant oxidizable concentration of the compound with the desired initial dilution of the oxidizable compound stream, according to other aspects of the present invention.

[0123] Both of FIGS. 24 and 25 show structures in which a venting tank connected to the conduit is used. Due to this deaeration vessel, part of the reaction medium used to dilute the oxidationable compound being introduced is essentially a vented suspension. It should be noted here, however, that the liquid or suspension used to dilute the introduced compound, which may be oxidized, may be in aerated form as well as in deaerated form.

[0124] The use of liquid flowing through the conduit to obtain dilution of the oxidised compound stream is particularly useful in bubble column reactors. In addition, a suitable initial dilution of a compound stream that can be oxidized, even without introducing a stream of compound which can be oxidized, directly into the conduit can be obtained in bubble bubbler reactors, provided that the conduit outlet is close enough to the place of introduction of the compound oxidation. In such an embodiment, it is preferred that the outlet of the conduit is within about 27 outlet pipe diameters from the nearest oxidation site of the compound which may be oxidized, more preferably within about 9 outlet pipe diameters, even more preferably within about 3 outlet pipe diameters and most preferably within 1 outlet duct diameter. [0125] It has further been found that fluid ejectors can be useful for initial dilution of a stream of compound that can be oxidized in bubble columns for oxidation according to an embodiment of the present invention, even without the use of liquid supply lines for dilution from a distant part of the reaction medium. In such cases, the ejector is in the reaction medium, and the path from the reaction medium to the ejector throat is open, so that pressure draws in the nearby reaction medium. Examples of two possible configurations of ejectors are shown in FIG. 26 and 27. In a preferred embodiment of these ejectors, the closest oxidizing compound introduction site is within about 4 meters, more preferably within about 1 meter, and most preferably 0.3 meter from the ejector throat. In another embodiment, the compound that can be oxidized is introduced under pressure as a propelling fluid. In yet another embodiment, the solvent or oxidant is introduced under pressure as an additional propelling fluid together with the compound that can be oxidized. In yet another embodiment, both the solvent and the oxidant are introduced under pressure as an additional propelling fluid together with the compound that can be oxidized.

[0126] The inventors predict that in particular in larger reactors it may be desirable to have multiple ejectors of various designs located at different locations in the reaction medium. Design details must balance the desired overall gradient of the constant oxidizable concentration of the compound with the desired initial dilution of the oxidizable compound stream, according to other aspects of the present invention. In addition, the developers predict that fluid ejector streaks may be directed in either direction. If multiple ejectors are used, each of them may have a different orientation, again in any direction.

[0127] As mentioned above, certain physical and operational features of the bubble column reactor 20 described above with reference to FIG. 1-27 cause vertical gradients of pressure, temperature and concentration of substrates (i.e. oxygen and a compound that can be oxidized) in reaction medium 36. As mentioned above, these vertical gradients can provide a more efficient and economical oxidation process compared to conventional oxidation processes that prefer a highly mixed reaction medium with relatively uniform pressure, temperature and concentration of substrates throughout the volume. The vertical gradients of oxygen, a compound that can be oxidized (para-xylene), and temperatures that can be obtained by using the oxidation system of the embodiment of the present invention will be discussed in more detail below.

[0128] Further referring to FIG. 28 in order to quantify the concentration gradients of substrates present in reaction medium 36 during oxidation in bubble column reactor 20, the entire volume of reaction medium 36 can theoretically be divided into 30 separate horizontal sections of equal volume. FIG. 28 presents the concept of dividing the reaction medium 36 into 30 separate horizontal slices of equal volume. In addition to the highest and lowest horizontal slices, each horizontal slice constitutes a separate volume limited on top and bottom by theoretical horizontal planes and limited on the sides by the wall of the reactor 20. The highest horizontal slice is bounded at the bottom by the theoretical horizontal plane and at the top by the upper surface of the reaction medium 36. The lowest the horizontal section is limited at the top by the theoretical horizontal plane and at the bottom by the bottom of the tank jacket. After dividing the theoretical reaction medium 36 into 30 separate horizontal slices of equal volume, time-averaged and volume-averaged concentration in individual horizontal slices can be determined. A single horizontal slice having a maximum concentration among all 30 horizontal slices can be referred to as "C-max horizontal slice". A single horizontal slice above the horizontal slice with C-max. and having a minimum concentration among all horizontal slices above the horizontal slice with C-max. can be defined as "horizontal slice from Cmin." The vertical concentration gradient can then be calculated as the concentration ratio of the horizontal slice from C-max. to the concentration in the horizontal slice with C-min.

[0129] With respect to the quantification of the oxygen concentration gradient, if the reaction center 36 is theoretically divided into 30 separate horizontal slices of equal volume, the horizontal slice from O2-max is determined. having the maximum oxygen concentration among all 30 horizontal slices and the horizontal slice with O2-min is determined. having a minimum oxygen concentration among all horizontal slices above the horizontal slice with O2-max. The oxygen concentration in horizontal sections is measured in the gas phase in the reaction medium 36 in terms of moles of moist substances averaged by time and by volume averaged. It is preferred that the ratio of oxygen concentration in the horizontal slice with O2-max. for oxygen concentration in the horizontal slice with O2-min. it was in the range of from about 2: 1 to about 25: 1, more preferably in the range of from about 3: 1 to about 15: 1, and most preferably in the range of from 4: 1 to

10:1.

[0130] Usually the horizontal slice with O₂-Max. is near the bottom of reaction medium 36, and the horizontal slice from O₂-min. located near the upper surface of the reaction medium. Preferably horizontal slice with O₂-min. is one of the 5 most horizontal horizontal slices out of 30 separate horizontal slices. Most preferably horizontal slice with O₂-min. is the topmost slice of the 30 separate horizontal slices, as shown in FIG. 28. Preferably, horizontal slice with O₂-Max. is one of the 10 most horizontal sections at the bottom of 30 separate horizontal sections. Most preferably horizontal slice with O₂-Max. is one of the 5 most horizontal horizontal sections out of 30 separate horizontal sections. For example, FIG. 28 shows the horizontal slice with O₂-Max. as a third horizontal slice from the bottom of reactor 20. It is preferred that the vertical distance between the horizontal slices with O₂-min. and o₂-Max. at least about 2W, more preferably at least about 4W and most preferably at least 6W. It is preferred that the vertical distance between the horizontal slices with O₂-min. and o₂-Max. at least about 0.2H, more preferably at least about 0.4H and most preferably at least 0.6H.

[0131] Time averaged and volume averaged oxygen concentration based on wet matter in the O2-min horizontal slice. preferably it is in the range of from about 0.1 to about 3 mole percent, more preferably in the range of from about 0.3 to about 2 mole percent, and most preferably in the range of from 0.5 to 1.5 mole percent. Time-averaged and volume-averaged oxygen concentration in the O2-max horizontal slice preferably it is in the range of from about 4 to about 20 mole percent, more preferably in the range of from about 5 to about 15 mole percent, and most preferably in the range of from 6 to 12 mole percent. The time-average oxygen concentration based on dry matter in the gas stream discharged from the reactor 20 through the gas outlet 40 is preferably in the range of about 0.5 to about 9 mole percent, more preferably in the range of about 1 to about 7 mole percent, and most preferably in the range from 1.5 to 5 mole percent.

[0132] Since the oxygen concentration decreases so significantly towards the top surface of the reaction medium 36, it is desirable to reduce the oxygen demand at the top of the reaction medium

36. Reduced oxygen demand near the upper surface of the reaction medium 36 can be obtained by creating a vertical gradient of the concentration of the compound that can be oxidized (para-xylene), with the minimum concentration of the compound that can be oxidized is near the upper surface of the reaction medium 36 .

[0133] With respect to the quantification of the concentration gradient of the compound that can be oxidized (para-xylene), if the reaction center 36 is theoretically divided into 30 separate horizontal slices of equal volume, the horizontal slice from OC-max is determined. having the maximum concentration of the compound that can be oxidized among all 30 horizontal slices and the horizontal slice with OC-min is determined. having a minimum concentration of a compound that can be oxidized among all horizontal slices above the horizontal slice with OC-max. The concentration of the compound that can be oxidized in horizontal slices is measured in the liquid phase as a mass fraction averaged over time and volume averaged. It is preferred that the ratio of the concentration of the compound that can be oxidized in the horizontal slice with OC-max. to the concentration of the compound which may be oxidized in the horizontal slice with OC-min. was greater than about 5: 1, more preferably greater than about 10: 1, even more preferably greater than about 20: 1, and most preferably it was in the range of 40: 1 to 1000: 1.

[0134] Typically horizontal slice with OC-max. located near the bottom of reaction medium 36, and the horizontal slice with OC-min. located near the upper surface of reaction medium 36. Preferably horizontal slice with OC-min. is one of the 5 most horizontal horizontal slices out of 30 separate horizontal slices. Most preferably horizontal OC-min slice. is the topmost slice of the 30 discrete horizontal slices, as shown in FIG. 28. Preferably, horizontal slice with OC-max. is one of the 10 most horizontal sections at the bottom of 30 separate horizontal sections. Most preferably horizontal slice with OC-max. is one of the 5 most horizontal horizontal sections out of 30 separate horizontal sections. For example, FIG. 28 shows the horizontal slice with OC-max. as the fifth horizontal slice from the bottom of reactor 20. It is preferred that the vertical distance between horizontal slices from OC-min. and OC-max. was at least about 2W, where "W" is the maximum width of the reaction medium 36. More preferably the vertical distance between horizontal slices with 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 preferred that the vertical distance between the horizontal slices from OC-min. and OC-max. at least about 0.2H, more preferably at least about 0.4H and most preferably at least 0.6H.

[0135] Time-averaged and volume-averaged concentration of the compound that can be oxidized (para-xylene) in the liquid phase in the horizontal slice with OC-min. preferably less than about 5000 ppmw, more preferably less than about 2000 ppmw, even more preferably less than about 400 ppmw, and most preferably it ranges from 1 ppmw to 100 ppmw. Time-averaged and volume-averaged concentration of a compound that may be oxidized in the liquid phase in a horizontal slice with OCmax. preferably remains in the range of from about 100 ppmw to about 10,000 ppmw, more preferably in the range of from about 200 ppmw to about 5000 ppmw, and most preferably in the range of from 500 ppmw to 3000 ppmw.

[0136] Although it is preferred that the columnar column reactor 20 provides vertical gradients for the concentration of the oxidizable compound, it is also preferred that the volume percentage of reaction medium 36 in which the concentration of the oxidizable compound in the liquid phase exceed 1000 ppmw, reduce to a minimum. Preferably, the time averaged volume percent of reaction medium 36 in which the concentration of the compound that can be oxidized in the liquid phase is above 1000 ppmw, is less than about 9 percent, more preferably less than about 6 percent, and most preferably less than 3 percent. Preferably, the time averaged volume percent of reaction medium 36 in which the concentration of the compound that can be oxidized in the liquid phase is above 2500 ppmw, is less than about 1.5 percent, more preferably less than about 1 percent, and most preferably less than 0.5 percent . Preferably, the time averaged volume percent of reaction medium 36 in which the concentration of the compound that can be oxidized in the liquid phase is above 10,000 ppmw, is less than about 0.3 percent, more preferably less than about 0.1 percent, and most preferably less than 0, 03 percent. Preferably, the time averaged volume percent of reaction medium 36 in which the concentration of the compound that can be oxidized in the liquid phase is above 25,000 ppmw, is less than about 0.03 percent, more preferably less than about 0.015 percent, and most preferably less than 0.007 percent. The inventors note that the volume of reaction medium 36 in which there is an elevated level of compound that can be oxidized does not need to be in a single continuous volume. Chaotic flow patterns in the bubble column reactor repeatedly cause that simultaneously two or more continuous but separate parts of reaction medium 36 have an elevated level of compound that can be oxidized. Whenever averaging occurs, such continuous but separate volumes greater than 0.0001 volume percent of the total reaction medium are added together to determine the total volume having an elevated concentration of the compound which can be oxidized in the liquid phase.

[0137] In addition to the oxygen concentration and oxidizable compound gradients discussed above, it is preferred that a temperature gradient occurs in reaction medium 36. With reference again to FIG. This temperature gradient can be quantified in a manner similar to concentration gradients by theoretically dividing the reaction medium 36 into 30 separate horizontal slices of equal volume and measuring time-averaged and volume-averaged temperature in each slice. The horizontal slice having the lowest temperature among the lowest 15 horizontal slices can then be referred to as the horizontal slice with T-min, and the horizontal slice above the horizontal slice with T-min. and having the maximum temperature of all horizontal slices above the horizontal slice with T-min. can then be described as "horizontal slice from T-max". It is preferred that the temperature in the horizontal slice with T-max. was at least about 1 ° C higher than the temperature in the horizontal slice with T-min. More preferably the temperature in the horizontal slice with T-max. is in the range of about 1.25 to about 12 ° C higher than the temperature in the horizontal slice with T-min. Most preferably the temperature in the horizontal slice with T-max. is in the range of 2 to 8 ° C higher than the temperature in the horizontal slice with T-min. Temperature in the horizontal slice with Tmax. preferably it is 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 from 150 to 170 ° C.

[0138] Typically horizontal slice with T-max. located near the center of reaction medium 36, and the horizontal section with T-min. is near the bottom of reaction medium 36. Preferably, horizontal slice with T-min. is one of the 10 lowest sections at the bottom of the 15 lowest horizontal sections. Most preferably horizontal slice with T-min. is one of the 5 most horizontal horizontal slices among the 15 lowest horizontal slices. For example, FIG. 28 shows horizontal slice from T-min. as a second horizontal slice from the bottom of reactor 20. Preferably, a horizontal slice with T-max. is one of the 20 middle horizontal slices out of 30 separate horizontal slices. Most preferably horizontal slice with T-min. is one of the 14 middle horizontal slices out of 30 separate horizontal slices. For example, FIG. 28 shows the horizontal slice from T-max. as the twentieth horizontal slice from the bottom of reactor 20 (i.e. one of the 10 middle horizontal slices). It is preferred that the vertical distance between horizontal slices from T-min. and T-max. at least about 2W, more preferably at least about 4W and most preferably at least 6W. It is preferred that the vertical distance between horizontal slices from T-min. and T-max. at least about 0.2H, more preferably at least about 0.4H and most preferably at least 0.6H.

[0139] As discussed above, if there is a vertical temperature gradient in the reaction medium, it may be advantageous that the reaction medium 36 is discharged at an elevated site, where the temperature of the reaction medium is highest, in particular if the discharged product undergoes further further processing in a higher temperature. Thus, if reaction medium 36 is discharged from reaction zone 28 through one or more elevated outlets, as shown in FIG. 19 and 20, it is preferred that the elevated outlet or outlets are near the horizontal section with T-max. Preferably, the elevated outlet is within 10 horizontal slices with a T-max horizontal slice, more preferably within 5 horizontal slices with a T-max horizontal slice. and most preferably within 2 horizontal slices with a horizontal slice from T-max.

[0140] It should now be noted that many of the features of the invention described herein can be used in many oxidation reactor systems, not just systems that use a single oxidation reactor. In addition, some of the features of the invention described herein can be used in mechanical and / or flow mixing oxidation reactors, and not only in bubble mixing reactors (i.e. bubble column reactors). For example, the inventors have discovered some benefits associated with the degree / variation of oxygen concentration and / or the rate of oxygen consumption over the entire volume of the reaction medium. The benefits obtained by grading the concentration / consumption of oxygen in the reaction medium can be realized regardless of whether the total volume of the reaction medium is in a single tank or in multiple tanks. In addition, the benefits obtained by the degree of oxygen concentration / consumption in the reaction medium can be realized irrespective of whether the reaction vessel or vessels are mechanically mixed, flow-through and / or bubbling.

[0141] One way to quantify the degree of concentration and / or oxygen consumption rate in a re93 action medium is to compare two or more separate 20% continuous volumes of reaction medium. These 20% continuous volumes need not be defined by any particular shape. However, each 20% continuous volume must be formed from the continuous volume of the reaction medium (i.e. each volume is "continuous") and 20% continuous volumes cannot overlap (ie, the volumes are "separate"). FIG. 29-31 show that these separate 20% continuous volumes can be in the same reactor (FIG. 29) or in multiple reactors (FIGS. 30 and 31). It should be noted that the reactors shown in FIG. 29-31 can be mechanically mixed, flow-through and / or bubbler reactors. In one embodiment, it is preferred that the reactors shown in FIG. 29-31 were bubble-mixed reactors (i.e. bubble column reactors).

[0142] Further referring to FIG. 29 shows a reactor 20 containing a reaction medium 36. The reaction medium 36 contains a first separate 20% continuous volume 37 and a second separate 20% continuous volume 39.

[0143] Further referring to FIG. 30, a multiple reactor system including the first reactor 720a and the second reactor 720b is shown. Reactors 720a, b together contain the total volume of reaction medium 36. The first reactor 720a contains the first part of the reaction medium 736a and the second reactor 720b contains the second part of the reaction medium 736b. A first separate 20 percent continuous volume 737 of reaction medium 736 is shown so that it is in the first reactor 720a, while a second separate 20 percent continuous volume 739 of reaction medium 736 is shown so that it is in the second reactor

720b.

> [0144] Further referring to FIG. 31, a multiple reactor system including the first reactor 820a, the second reactor 820b and the third reactor 820c is shown. Reactors 820a, b, c together contain the total volume of reaction medium 836. The first reactor 820a contains the first portion of reaction medium 836a; the second reactor 820b comprises a second part of the reaction medium 836b; and the third reactor 820c contains a third portion of the reaction medium 836c. A first separate 20 percent continuous volume 837 of reaction medium 836 is shown so that it is in the first reactor 820a; a second separate 20 percent continuous volume 839 of reaction medium 836 is shown so that it is in the second reactor 820b, while a third separate 20 percent continuous volume 841 of reaction medium 836 is shown so that it is in the third reactor 820c.

[0145] The degree of oxygen availability in the reaction medium can be quantified in relation to the 20% volume of the continuous reaction medium in which the molar fraction of oxygen in the gas phase is largest and in relation to the 20% volume of the continuous reaction medium in which the molar fraction of oxygen in the gas the gas phase is the smallest. In the gas phase, a separate 20% volume of the continuous reaction medium containing the highest concentration of oxygen in the gas phase, time-averaged and volume-averaged oxygen concentration on wet basis remains preferably in the range of from about 3 to about 18 mole percent, more preferably in the range of from about

3.5 to about 14 mole percent and most preferably in the range of 4 to 10 mole percent. In the gas phase of a separate 20% volume of continuous reaction medium containing the lowest oxygen concentration in the gas phase, the time-averaged and volume-averaged oxygen concentration based on sub95 wet conditions is preferably in the range from about 0.3 to about 5 mole percent, more preferably in the range from about 0 , 6 to about 4 mole percent and most preferably in the range of 0.9 to 3 mole percent. In addition, the ratio of the time-averaged and volume-averaged oxygen concentration in terms of wet matter in the 20% volume of the continuous concentration medium with the highest concentration compared to the 20% volume of the continuous concentration medium with the lowest concentration is preferably in the range from about 1.5: 1 to about 20 : 1, more preferably in the range of about 2: 1 to about 12: 1 and most preferably in the range of 3: 1 to 9: 1.

[0146] The degree of concentration and / or rate of oxygen consumption in the reaction medium can be quantified using the oxygen STR value described above. The oxygen STR value was previously described in general terms (i.e. with respect to the average oxygen STR value throughout the reaction medium); however, the STR value for oxygen can also be considered locally (i.e. part of the reaction medium) to quantify the degree of oxygen consumption rate throughout the entire reaction medium.

[0147] The inventors have found that it is very useful to cause the oxygen STR value to change in the reaction medium substantially in line with the desired gradients disclosed herein with respect to the pressure in the reaction medium and the mole fraction of molecular oxygen in the gas phase in the medium reaction. It is therefore preferred that the ratio of the oxygen STR value in the first separate 20% continuous volume of the reaction medium to the oxygen STR value in the second separate 20% continuous volume of the reaction medium remains in the range from about 1.5: 1 to about

20: 1, more preferably in the range of about 2: 1 to about 12: 1 and most preferably in the range of 3: 1 to 9: 1. In one embodiment, the "first separate 20% continuous volume" is closer than the "second separate 20% continuous volume" of the place where the molecular oxygen is initially introduced into the reaction medium. Such large oxygen STR gradients are desirable regardless of whether the partial oxidation reaction medium is in a bubble oxidation column reactor or in any other type of reaction vessel in which pressure gradients and / or a mole fraction of molecular oxygen are formed in the gas phase in the reaction medium (e.g. in a mechanically mixed tank having a plurality of vertically arranged mixing zones, which is obtained by using multiple impellers with high radial flow, possibly supported by substantially horizontal baffle modules, so that the oxidant flow generally rises from the inlet near the bottom of the reaction tank, regardless of that there may be significant backmixing of the oxidant flow in each vertically arranged mixing zone, and that some backmixing of the oxidant flow may occur between adjacent vertically arranged mixing zones). That is, if there is a pressure gradient and / or mole fraction of molecular oxygen in the gas phase in the reaction medium, the inventors have found that it is desirable to produce a similar chemical demand gradient for dissolved oxygen by the methods disclosed herein.

[0148] A preferred way to induce variability of local STR values ​​for oxygen is to control the places of introduction of a compound that can be oxidized by controlling the liquid phase mixing of the reaction medium to control the concentration gradients of the compound which can be oxidized, according to other disclosures of the present invention. Another preferred way to induce variation in local oxygen STR values ​​is to induce variation in reaction activity by inducing local temperature changes and by changing the local mixture of catalyst and solvent components (e.g. by introducing additional gas, which results in cooling by evaporation in a given part of the reaction medium, and by adding a solvent stream containing more water to reduce activity in that part of the reaction medium.

[0149] As stated above with reference to FIG. thirty and 31, the partial oxidation reaction can be usefully performed in a plurality of reaction vessels, wherein at least a portion, preferably at least 25 percent, more preferably at least 50 percent and most preferably at least 75 percent of the molecular oxygen leaving the first reaction vessel is sent to one or more further reaction tanks to consume the next portion, preferably over 10 percent, more preferably more than 20 percent and most preferably more than 40 percent, of molecular oxygen leaving the first / previous reaction tank. If such a serial flow of molecular oxygen from one reactor to the other is used, it is desirable for the first reaction vessel to operate at a higher reaction rate than at least one of the subsequent reaction vessels, preferably such that the ratio of the average oxygen STR value in the first reaction vessel to the average STR value for oxygen in the next reaction vessel ranges from about 1.5: 1 to about

20: 1, more preferably in the range of about 2: 1 to about 12: 1 and most preferably in the range of 3: 1 to 9: 1.

[0150] As mentioned above, the bubble column reactor constituting the first reaction vessel and all types of subsequent reaction vessels, e.g. bubble column, mechanically agitated, backmixed, internally divided, with piston flow and so on, which may be of a different type than the first reaction vessel or not of a different type, are useful for flowing serial molecular oxygen into subsequent reaction vessels according to the present invention . Methods for obtaining a reduction in the mean STR value for oxygen in the tank for subsequent reaction tanks include useful reduction of temperature, reduction of the concentration of the compound that may be oxidized, and reduction of the reaction activity of a given mixture of catalytic components and solvent (e.g., reduced cobalt concentration, increased water concentration and the addition of a slowing down agent (e.g. small amounts of copper ions).

[0151] The flush of oxidant flowing from the first reaction vessel to the next reaction vessel can be treated by any means known in the art, such as compression or reduction of pressure, cooling or heating, and removal of mass or introduction of mass in any amount or type. However, the use of reducing the mean STR value for oxygen in the tank for subsequent reaction tanks is especially useful if the absolute pressure in the upper part of the first reaction tank is less than about 2.0 megapascals, more preferably less than about 1.6 megapascals and most preferably less than 1.2 megapascals. In addition, the use of reducing the mean STR value for oxygen in the tank for subsequent reaction tanks is particularly useful if the ratio of the absolute pressure in the upper part of the first reaction tank to the absolute pressure in the top of at least one subsequent reaction tank remains in the range from about 0.5 : 1 to 6: 1, more preferably in the range of about 0.6: 1 to about 4: 1 and most preferably in the range of 0.7: 1 to 2: 1. Decreasing the pressure in subsequent tanks below these lower limits causes an excessive reduction in the availability of molecular oxygen, while increasing the pressure above these upper limits becomes excessively expensive compared to using a fresh oxidant stream.

[0152] If a serial flow of molecular oxygen is used for subsequent reaction tanks with a reduced mean STR value for oxygen in the tank, fresh feed streams of the compound that can be oxidized, solvent and oxidant can be fed to the next reaction tanks and / or to the first reaction tank . The flow of liquid phase and solid phase, if present, in the reaction medium can take place in any direction between the reaction vessels. All or part of the gas phase leaving the first reaction tank and going to the next reaction tanks can flow separately or together with parts of the liquid phase or solid phase, if present, in the reaction medium from the first reaction tank. The flow of the product stream containing the liquid phase and the solid phase, if present, can be withdrawn from the reaction medium in any reaction vessel in the system.

100 [0153] Referring again to FIG. 1-29 oxidation is preferably carried out in a bubble column reactor under significantly different conditions, according to the preferred embodiments disclosed herein, from conventional oxidation reactors. If the bubble column reactor 20 is used to conduct partial liquid phase oxidation of para-xylene to crude terephthalic acid (CTA) according to preferred embodiments disclosed herein, spatial profiles of local reaction rate, local evaporation rate and local temperature combined with flow diagrams liquids in the reaction medium, and also beneficial relatively low oxidation temperatures contribute to the formation of CTA particles having unique and beneficial properties.

[0154] FIG. 32A and 32B show basic CTA particles produced in accordance with one embodiment of the present invention. FIG. 32A shows CTA base particles at a magnification of 500 times, and FIG. 32B shows the approximation of one basic CTA particle and represents this particle at a magnification of 2000 times. As most likely shown best in FIG. 32B, each CTA base particle is typically formed by a large number of small, aggregated smaller CTA particles, resulting in a CTA base particle with relatively large surface area, high porosity, low density, and good solubility. Unless otherwise stated, the various CTA properties of the invention described below are measured with a representative CTA sample, where the representative sample weighs at least 1 gram and / or is formed from at least 10,000 individual CTA particles. Primary CTA particles usually have an average particle size in the range of

101 about 20 to about 150 microns, more preferably in the range of about 30 to about 120 microns, and most preferably in the range of 40 to 90 microns. Smaller CTA particles typically have an average particle 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 in the range of 2 to 5 microns. The relatively large surface area of ​​the primary CTA particles shown in FIG. 32A and 32B can be quantified using the Braunauer-Emmett-Teller method (BET). Preferably, the CTA base particles have an average BET surface area of ​​at least about 0.6 square meters per gram (m²/ G). More preferably, the CTA base particles have an average BET surface area in the range of about 0.8 to about 4 m²/ G. Most preferably the basic CTA particles have an average BET surface area in the range of 0.9 to 2 m²/ G. The physical properties (e.g. particle size, BET surface area, porosity and solubility) of the basic CTA particles formed in the optimized oxidation process according to the preferred embodiment of the present invention allow the purification of CTA particles by more efficient and / or more economical methods, described in more detail below with reference to FIG.

35.

[0155] The mean particle size values ​​given above were determined by polarized light microscopy and image analysis. The equipment used for particle size analysis included a Nikon E800 optical microscope with 4x Plan Fluor NA 0.13 lens, a Spot RT ™ digital camera and a personal computer with Image Pro Plus ™ V4.5.0.19 image analysis software installed. The particle size analysis method involved the following main steps: (1) dispersion

102 CTA powders in mineral oil; (2) preparing a dispersion preparation on a microscope slide / cover slide; (3) examination of the slide by polarized light microscopy (cross polarization conditions - the particles are visible as bright objects on a black background); (4) recording of different images for each sample preparation (image size = 3 x 2.25 mm; pixel size = 1.84 micron / pixel); (5) performing image analysis using Image Pro Plus ™ software; (6) exporting particle measurements to a spreadsheet; and (7) performing statistical calculations in a spreadsheet. Step (5): "performing image analysis using Image Pro Plus ™" included the following sub-steps: (a) setting the image threshold to detect white particles against a dark background; (b) creating a binary image; (c) using a single pass open filter to remove pixel noise; (d) measuring all particles in the image; and (e) providing an average measured diameter of each particle. Image Pro Plus ™ software determines the average diameter of individual particles as the number average length of particle diameters measured at 2 degree intervals and passing through the centroid of the particle. Step 7: "performing statistical calculations in a spreadsheet" includes calculating the volume-weighted average particle size as follows. The volume of each of the n particles in the sample is calculated as if spherical using the formula pi / 6 * d_and^AND3; the volume of each particle is multiplied by its diameter to give pi / 6 * di^AND4; pi / 6 * di values^AND4 for all particles in the sample are added together; the volumes of all particles in the sample are added and the volume average particle diameter is calculated as the sum of all n particles in the sample (pi / 6 * d_and^AND4) divided by the sum of all n particles in the sample (pi / 6 * d_and^AND3). The term "average particle size" as used herein means the volume average particle size determined according to the test method above; it is also referred to as D (4.3).

<img file="PL1784378T3_D0001.tif" />

[0156] In addition, step 7 includes finding particle sizes for which different fractions of the total sample volume are smaller. For example, D (v, 0.1) is a particle size such that 10 percent in the total sample volume is smaller and 90 percent is larger; D (v, 0.5) is the particle size such that half in the sample volume is larger and half is smaller; D (v, 0.9) is a particle size such that 90 percent of the total sample volume is smaller and so on. In addition, step 7 includes calculating the value of D (v, 0.9) minus D (v, 0.1), which is defined herein as "particle size distribution", and step 7 includes calculating the value of particle size distribution divided by D ( 4,3), which is defined herein as "relative particle size distribution".

[0157] It is further preferred that the D (v, 0.1) value of the CTA particles measured in the above manner remain 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. It is preferred that the D (v, 0.5) value for the CTA particles measured in the above manner remains in the range from about 10 to about 90 microns, more preferably in the range

104 from about 20 to about 80 microns, and most preferably in the range of 30 to 70 microns. It is preferred that the D (v, 0.9) value of the CTA particles measured in the above manner remains in the range of about 30 to about 150 microns, more preferably in the range of about 40 to about 130 microns, and most preferably in the range of 50 to 110 microns . It is preferred that the relative particle size distribution is in the range of from about 0.5 to about 2.0, more preferably in the range of from about 0.6 to about

1.5 and most preferably in the range of 0.7 to 1.3.

[0158] The BET surface area values ​​given above were measured using a Micromeritics ASAP2000 instrument (available from Micromeritics Instrument Corporation, Norcross, GA). In the first stage of the measurement process, 2 to 4 grams of sample particles were weighed and dried under vacuum at 50 ° C. The sample was then placed in a gas analysis collector and cooled to 77 ° K. At least 5 equilibrium pressure values ​​measured the nitrogen adsorption isotherm by exposing the sample to known volumes of nitrogen gas and measuring the pressure drop. The equilibrium pressure values ​​remained in the range P / P0 = 0.01-0.20, respectively, where P is the equilibrium pressure and P0 is the vapor pressure of liquid nitrogen at 77 ° K. Then the resulting isotherm was plotted according to the following BET equation:

_ ^p v_and(P.-P) vc kc {pJ where V_and is the volume of gas adsorbed on the sample under pressure P, V_m is the volume of gas necessary to cover the entire surface of the sample with a single layer of gas, and C is a constant. From this graph V values ​​were determined_m and C. Then V_m calculated on the surface area using the nitrogen cross-sectional area at 77 ° K from the formula:

105

<img file="PL1784378T3_D0002.tif" />

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

[0159] As suggested above, CTA prepared according to one embodiment has excellent dissolution properties compared to conventional CTA produced by other methods. This increased dissolution rate allows the purification of the CTA according to the invention by more efficient and / or more effective purification methods. The following description relates to a method by which the CTA dissolution rate can be quantified.

[0160] The dissolution rate of a known amount of solids in a known amount of solvent in a mixed mixture can be measured by various methods. This document defines the measurement method referred to as "dissolution time testing" as follows. An ambient pressure of about 0.1 megapascal was used throughout the entire dissolution study. In the entire dissolution test over time, the ambient temperature was around 22 ° C. In addition, solids, solvent and the entire dissolution measurement device were fully thermally balanced at this temperature before the start of the test and there was no significant heating or cooling of the beaker or its contents during dissolution. The volume of fresh solvent, HPLC grade tetrahydrofuran (purity&gt; 99.9 percent), hereinafter referred to as THF, in an amount of 250 grams was placed in a clean, high 400 ml KIMAX glass beaker (Kimble^®, catalog number 14020, Kimble / Kontes, Vineland, NJ), constituting an uninsulated vessel, with smooth

106 walls and generally cylindrical. A teflon coated magnetic stirrer (VWR, catalog number 58948-230, about 1 inch long and 3/8 inch diameter, with octagonal cross section, VWR International, West Chester, PA 19380) was placed in a beaker; it settled naturally at the bottom of the beaker. The sample was mixed using a Variomag magnetic stirrer^® Multipoint 15 (H&amp;P Labortechnik AG, Oberschleissheim, Germany) at 800 rpm. Mixing starts no more than 5 minutes before adding solids and runs continuously for at least 30 minutes after adding solids. A solid sample of crude or purified TPA particles weighing 250 milligrams was weighed into a weigh boat for a non-stick sample. At the initial time t = 0, all of the weighed solids were introduced into the mixed THF at once and the stopwatch was started at the same time. If the procedure is carried out correctly, THF moistens the solids very quickly and forms a diluted, strongly mixed suspension within 5 seconds. Successively 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, samples of the mixture were taken. Each small sample was taken from the diluted, well mixed mixture using a fresh disposable syringe (Becton, Dickinson and Co, 5 milliliters, REF 30163, Franklin Lakes, NJ 07417). Immediately after sampling from the beaker through a fresh, unused syringe filter (25 mm diameter, 0.45 micron, Gelman GHP Acrodisc GF^®, Pall Corporation, East Hills, NY 11548) about 2 milliliters of transparent liquid sample was introduced into a new, labeled glass sample vial. The duration of filling each syringe, fitting the filter and introducing it into the sample vial is less than about 5 seconds under normal conditions, this time starting under normal conditions 3 seconds before and ending 3 seconds after the target sampling time. Within about five minutes of each filling, the sample vials were capped and kept at approximately constant temperature until the following chemical analysis was performed. After the last sample was taken at 30 minutes after t = 0, all sixteen samples were analyzed for dissolved TPA content using the HPLC-DAD method generally described elsewhere in this disclosure. In this study, however, values ​​for calibration standards and results obtained are given in milligrams of dissolved TPA per gram of THF solvent (referred to below as "ppm in THF"). For example, if all of 250 milligrams of solids were very pure TPA and all this was completely dissolved in 250 grams of THF solvent before sampling, a properly measured concentration would be about 1000 ppm in THF. [0161] If the CTA of the present method is subjected to the dissolution test at the time described above, it is preferred that the sample taken within one minute after t = 0 dissolve to give a concentration of at least about 500 ppm in THF, more preferably at least 600 ppm in THF . For the sample taken two minutes after t = 0, it is preferred that the CTA according to the present invention dissolves to give a concentration of at least about 700 ppm in THF, more preferably at least 750 ppm in THF. For the sample taken within four minutes after t = 0, it is preferred that the CTA according to the present invention dissolves to give a concentration of at least about 840 ppm in THF, more preferably at least 880 ppm in THF.

108 [0162] The inventors have found that a relatively simple negative exponential growth model is useful to describe the relationship of an entire dataset from a complete dissolution test over time, regardless of the complexity of the particle samples and the dissolution process. 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; units: minutes;

S = solubility; units: ppm in THF, at t;

exp = exponential function based on the natural logarithm of

2;

A, B = regression constants in units of ppm in THF, where A is most associated with the rapid dissolution of smaller particles in a very short time, and the sum of A + B is most associated with the total dissolution at the end of the specified test period; and

C = time regression constant; units: one minute to the power minus one.

[0163] Regression constants were adjusted to minimize error squares between actual data points and corresponding model values; this is usually referred to as the least squares fit. A preferred software package for calculating such regression for data is JMP version 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary, NC 27513).

[0164] If the CTA of the present method is subjected to a dissolution time study and matched to the dissolution model over time described above, it is preferred that the CTA time constant for CTA is greater than about 0.5 minutes to the first fifth, more preferably greater than about 0.6 minutes to the minus one power and most preferably greater than 0.7 minutes to the minus one power.

[0165] FIG. 33A and 33B show a conventional CTA particle produced by a conventional high temperature oxidation process in a continuous stirred tank reactor (CSTR). FIG. 33A shows a conventional CTA particle at a magnification of 500 times, and FIG. 33B is an approximation and shows a CTA particle at 2000 times magnification. Visual comparison of the CTA particles according to the invention shown in FIG. 32A and 32B with the conventional CTA particle shown in FIG. 33A and 33B show that a conventional CTA particle has a higher density, smaller surface area, lower porosity, and a larger particle size than the CTA particles of the invention. Conventional CTAs shown in FIG. 33A and 33B have an average particle size of about 205 microns and a BET surface area of ​​about 0.57 m²/ G.

[0166] FIG. 34 illustrates a conventional method of producing purified terephthalic acid (PTA). In the conventional PTA production process, para-xylene is partially oxidized in a high temperature oxidation reactor with mechanical agitation 700. The suspension containing CTA is withdrawn from the reactor 700 and then purified in a 702 purification system. The PTA product from the purification system 702 is introduced into the separation system 706 to separate and dry the PTA particles. The 702 purification system consumes a significant portion of the cost of producing PTA particles by conventional methods. The purification system 702 generally includes a water addition / exchange system 708, dissolution system 710, hydrogenation system 712, and three separate crystallization tanks 704a, b, c. In the 708 water addition / exchange system, much of the mother liquor is replaced with water. After adding water, the water / CTA suspension is introduced into the dissolution system 710, where the water / CTA mixture is heated until the CTA particles are completely dissolved in the water. After dissolving the CTA, the CTA solution in water is hydrogenated in a 712 hydrogenation system. The hydrogenated stream withdrawn from the hydrogenation system 712 is then subjected to three crystallization steps in the crystallization tanks 704a, b, c, followed by separation of the PTA in the separation system 706.

[0167] FIG. 35 depicts an improved method of producing PTA using bubbling oxidation column reactor 800 in a configuration in accordance with an embodiment of the present invention. The initial suspension containing the solid CTA particles and the liquid mother liquor is withdrawn from the reactor 800. Typically, the initial suspension may contain from about 10 to about 50 weight percent of the CTA solid particles, the remainder being liquid mother liquor. The solid CTA particles present in the initial suspension usually contain at least about 400 ppmw 4-carboxybenzaldehyde (4-CBA), more usually at least about 800 ppmw 4-CBA and most often in the range from 1000 to 15,000 ppmw 4-CBA. The initial slurry discharged from the reactor 800 is introduced into the 802 purification system to reduce the concentration of 4-CBA and other impurities found in CTA. Produces a more pure / purified slurry in the 802 purification system, which is subjected to separation and drying in the 804 separating system, resulting in higher purity terephthalic acid solids containing less than about 400 ppmw 4-CBA, more preferably less than about 250 ppmw 4-CBA and most preferably in the range of 10 to 200 ppm in 4-CBA.

[0168] The purification system 802 in the PTA production system shown in FIG. 35 provides a number of advantages over the prior art purification system 802 shown in FIG. 34. The purification system 802 generally includes preferably a lye exchange system 806, a reoxidation chamber 808 and a single crystallizer 810. In the lye exchange system 806, at least about 50 weight percent of the mother liquor present in the initial suspension is replaced with fresh replaced solvent to give a suspension with said solvent containing CTA particles and said solvent. The slurry with said solvent leaving the lye exchange system 806 is introduced into the secondary oxidation chamber (or secondary oxidation reactor) 808. In the secondary oxidation chamber 808, an additional oxidation reaction occurs at a temperature slightly higher than that used in the initial / principal oxidation reaction carried out in the bubble column reactor 800. As stated above, the large surface area, low particle size, and low density of CTA particles produced in the 800 reactor make some of the impurities trapped in the CTA particles become available for oxidation in the secondary oxidation chamber 808, so that it is not necessary to completely dissolve the CTA particles in the secondary oxidation chamber 808. Therefore, the temperature in the secondary oxidation chamber 808 may be lower than in many similar prior art methods. The additional oxidation carried out in the secondary oxidation chamber 808 preferably reduces the concentration of 4-CBA in the CTA by at least 200 ppmw, more preferably at least about 400 ppmw and most preferably in the range from 600 to 6000 ppmw. Preferably, the post oxidation temperature in the secondary oxidation chamber 808 is at least about 10 ° C higher than the basic oxidation temperature in bubble column reactor 800, more preferably about 20 to about 80 ° C higher than the basic oxidation temperature in reactor 800 and most preferably 30 to 50 ° C higher than the basic oxidation temperature in the 800 reactor. The additional 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 most preferably in the range of 190 to 210 ° C. The purified product from the secondary oxidation chamber 808 requires only one crystallization step in the crystallizer 810 before separation in the separation system 804. Suitable methods for additional oxidation / reoxidation are discussed in more detail in US Patent Application Publication No. 2005/0065373.

[0169] Terephthalic acid (e.g., PTA) produced in the system shown in FIG. 35 is preferably formed of PTA particles having an average particle size of at least about 40 microns, more preferably in the range of about 50 to about 2000 microns, and most preferably in the range of 60 to 200 microns. The PTA particles preferably have an average BET surface area of ​​less than about 0.25 m²/ g, more preferably in the range of from about 0.005 to about 0.2 m²/ g and most preferably in the range of 0.01 to 0.18 m²/ G. PTA produced in the system shown in FIG. 35 is suitable for use as a raw material for PET. PET is usually produced by esterification of terephthalic acid with ethylene glycol followed by polycondensation. Preferably, terephthalic acid prepared according to the embodiment of the present invention is used as a raw material in the process of obtaining PET in re113 tubular actuator described in US Patent Application No. 10 / 013,318 made on December 7, 2001.

[0170] CTA particles having the preferred morphology disclosed herein are particularly useful in the secondary oxidation method described above to reduce the content of 4-CBA. In addition, these preferred CTA particles provide benefits in a wide range of other subsequent processes involving dissolution and / or chemical reaction of the particles. Such additional subsequent processes include, but are not limited to, reaction with at least one hydroxyl-containing compound to form ester compounds, in particular the reaction of CTA with methanol to form dimethyl terephthalate and esters; reaction with at least one diol to form compounds that are monomers and / or polymers of esters, in particular the reaction of CTA with ethylene glycol to form polyethylene terephthalate (PET) and complete or partial dissolution in solvents, including but not limited to water, acid acetic acid and N-methyl-2-pyrrolidone, which may include further processing, including but not limited to re-precipitation of terephthalic acid with greater purity and / or selective chemical reduction of carbonyl groups other than carboxylic acid groups. Particular consideration is given to the significant dissolution of CTA in a water-containing solvent in combination with partial hydrogenation, as a result of which the content of aldehydes, in particular 4-CBA, fluorenones, phenones and / or anthraquinones is reduced.

[0171] The inventors further envisage that particles having advantageous properties disclosed herein may be made from CTA particles having no beneficial properties disclosed.

114 in this document (non-compliant CTA particles) by, inter alia, mechanical grinding of non-compliant CTA particles and complete or partial dissolution of non-compliant CTA particles, followed by complete or partial re-precipitation.

[0172] According to one embodiment, a method is disclosed for partially oxidizing an aromatic compound that can be oxidized to one or two types of aromatic carboxylic acid in which the purity of the solvent portion of the feed stream (ie, "solvent stream") and the purity of the portion of the feed stream which is a compound that can be oxidized (i.e. "Oxidative compound stream") is controlled within the specific ranges given below. Together with other embodiments, this makes it possible to control the purity of the liquid phase and, if present, the solid phase and the combined suspension phase (i.e. solids and liquids) in the reaction medium in certain preferred ranges given below.

[0173] With respect to the solvent stream, it is known to oxidize an aromatic compound or compounds that can be oxidized to form an aromatic carboxylic acid, the solvent stream introduced into the reaction medium being a mixture of pure acetic acid for analysis and water, which is often used on a laboratory and semi-technical scale. Similarly, it is known to oxidize an aromatic compound that can be oxidized to an aromatic carboxylic acid, wherein the solvent leaving the reaction medium is separated from the aromatic carboxylic acid produced and then recycled back to the reaction medium as a solvent stream, mainly

115 due to production costs. Recycling of the solvent accumulates certain impurities from the raw material and by-products of the process in the recycled solvent over time. Various methods are known in the art to help purify recycled solvent before it is reintroduced into the reaction medium. Generally, a higher degree of purification of recycled solvent results in a significant increase in production costs compared to a lower degree of purification by similar methods. One embodiment of the present invention relates to the knowledge of preferred ranges and their determination in the event of a large number of impurities in the solvent stream, many of which have until now been considered relatively harmless, to find an optimal balance between overall production costs and overall product purity.

[0174] "Stirred erythema recycle" is defined herein as a solvent stream containing at least about 5 weight percent mass that has previously passed through a reaction medium containing one or more oxidizable aromatic compounds that undergo partial oxidation. Due to the solvent reserves and the period of the production plant being in operation, it is preferred that the parts of recycled solvent pass through the reaction center at least once a day of operation, more preferably at least once a day for at least seven consecutive days of operation and most preferably at least once per day for at least 30 consecutive days of operation. For economic reasons, it is preferred that at least about 20 weight percent of the solvent stream in the reaction medium of the present invention is recycled solvent, more preferably at least about 40 weight percent, even more preferably at least about 80 weight percent, and most preferably at least 90 weight percent.

[0175] The inventors have found that, due to the activity of the reaction and taking into account the metallic impurities remaining in the oxidation product, the concentration of selected polyvalent metals in the recycled solvent stream is preferably in the ranges given immediately below. The concentration of iron in recycled solvent is preferably below about 150 ppmw, more preferably below about 40 ppmw and most preferably from 0 to 8 ppmw. The concentration of nickel in recycled solvent is preferably below about 150 ppmw, more preferably below about 40 ppmw and most preferably from 0 to 8 ppmw. The concentration of chromium in recycled solvent is preferably below about 150 ppmw, more preferably below about 40 ppmw and most preferably from 0 to 8 ppmw. The concentration of molybdenum in recycled solvent is preferably below about 75 ppmw, more preferably below about 20 ppmw and most preferably from 0 to 4 ppmw. The concentration of titanium in recycled solvent is preferably below about 75 ppmw, more preferably below about 20 ppmw and most preferably from 0 to 4 ppmw. The concentration of copper in recycled solvent is preferably below about 20 ppmw, more preferably below about 4 ppmw and most preferably from 0 to 1 ppmw. Other metallic impurities are also typically present in the recycled solvent, generally at a lower level varying relative to one or more of the metals listed above. Controlling the above mentioned metals within the preferred ranges maintains an appropriate level of other metallic impurities.

117 [0176] These metals may be impurities in any of the process streams introduced (e.g., in the feed that may be oxidized, the solvent, the oxidant, and the catalytic compounds). Alternatively, the metals may be corrosion products of any of the technological devices in contact with the reaction medium and / or in contact with recycled solvent. Methods for controlling metals in the disclosed concentration ranges include appropriate specification and monitoring of the purity of various streams, and the appropriate use of structural materials, including, among others, many industrial grades of titanium and stainless steel, including grades known as duplex stainless steels and high molybdenum stainless steels.

[0177] The inventors have also developed preferred ranges for selected aromatics in recycled solvent. These include both precipitated and dissolved aromatics in recycled solvent.

[0178] Surprisingly, even a precipitated product (e.g. TPA) from partial oxidation of para-xylene is an impurity that needs to be controlled in recycled solvent. Because there are surprisingly favorable ranges of levels of solids in the reaction medium, any precipitated product in the solvent stream directly reduces the amount of oxidizable compound that can be introduced simultaneously. In addition, it has been found that the incorporation of large amounts of precipitated TPA solids in the recycled solvent has an adverse effect on the nature of the particles formed in the oxidation reaction medium during precipitation, which has an undesirable effect on further processing (e.g. product filtration, solvent washing, secondary oxidation of crude

118 product, dissolving the crude product for further processing, etc.). Another undesirable property of precipitated solids present in the recycled solvent stream is that they often contain a significant level of precipitated impurities compared to the concentration of the impurities in the volume of solids in the TPA suspension from which most of the recycled solvent is obtained. Increased levels of impurities found in solids suspended in the recycled filtrate may be likely related to the nucleation time for precipitation of some impurities from recycled solvent and / or cooling of recycled solvent, either intentionally or due to a decrease in ambient temperature. For example, the concentration of the undesirable 2,6-dicarboxyfluorenone with a high coloration at a much higher level was observed in solids present in recycled solvent at 80 ° C compared to solid TPA components after separation from recycled solvent at 160 ° C. Similarly, the concentration of isophthalic acid was observed at a much higher level in solids found in recycled solvent compared to the level observed in solid TPA components from the reaction medium. It seems that the properties of specific precipitated impurities trapped in the recycled solvent after reintroduction into the reaction medium are variable. It depends perhaps on the relative solubility of the liquid phase impurity in the reaction medium, perhaps on the layered distribution of the precipitated impurity in the precipitated solids, and perhaps on the local TPA precipitation rate at the point where the solid first re-enters the reaction medium. Therefore, the inventors have found that it is useful to control the level of certain impurities in the recycled solvent, as discussed below, regardless of whether these impurities are present in the recycled solvent in dissolved form, or whether they are trapped in solid particles.

[0179] The content of precipitated solids present in the recycled filtrate is determined by the gravimetric method as follows. A representative sample is taken from the solvent fed to the reaction medium as the solvent flows through the line to the reaction medium. A useful sample size is about 100 grams collected in a glass container with a total capacity of about 250 milliliters. Before opening to atmospheric pressure, but with continuous flow to the sample container, the recycled filtrate is cooled to below 100 ° C; cooling is intended to limit solvent evaporation in the short term before sealing in a glass container. After sampling at atmospheric pressure, the glass container is sealed immediately. The sample is then allowed to cool to about 20 ° C, in air at about 20 ° C and without forced circulation. After reaching a temperature of about 20 ° C, the sample is stored under these conditions for at least about 2 hours. The closed container is then shaken vigorously until a visually uniform distribution of solids is obtained. Immediately afterwards, a magnetic stirrer is introduced into the sample container and mixed at a speed sufficient to maintain a virtually homogeneous distribution of solids. A volume of 10 milliliters of mixed liquid with suspended solids is taken with a pipette and weighed. Then the superna120 liquid phase from this volume is separated by vacuum filtration, still at a temperature of about 20 ° C and with virtually no loss of solids. Most of the solids filtered off from this volume are then dried, with virtually no sublimation of solids, and the dried solids are weighed. The ratio of the weight of dried solids to the weight of the original volume of suspension is a fraction of solids, usually expressed as a percentage and referred to herein as the amount of "precipitated solids at 20 ° C" in the solvent stream.

[0180] The inventors have found that aromatic compounds dissolved in the liquid phase in the reaction medium and containing aromatic carboxylic acids not containing non-aromatic hydrocarbyl groups (e.g., isophthalic acid, benzoic acid, 2,5,4'-tricarboxybiphenyl) are surprisingly disadvantageous components. Although the chemical activity of these compounds in the current reaction medium is significantly reduced compared to compounds that can be oxidized, having non-aromatic hydrocarbyl groups, the inventors have found that these compounds, however, undergo many adverse reactions. Therefore, it is preferable to control the content of these compounds in preferred ranges in the liquid phase in the reaction medium. As a result, there are preferred ranges of selected compounds in the recycled solvent stream, as well as preferred ranges of selected precursors in the aromatic compound stream that can be oxidized.

[0181] The inventors have found that during partial oxidation in the liquid phase of para-xylene to terephthalic acid (TPA), the undesirable impurity 2,7-dicarboxyfluorenon (2.7DCF) with a strong color is practically undetectable in

121 in the reaction medium and in the discharged product if there is a very low level of meta-substituted aromatic compounds in the reaction medium. The inventors have found that if there is an increasingly higher level of isophthalic acid contamination in the solvent stream, the formation of 2,7-DCF increases almost directly in proportion. The inventors also found that if there is impurity (meta-xylene) in the para-xylene stream, the formation of 2,7-DCF again increases almost directly proportionally. Furthermore, the inventors have found that even if the solvent stream and the compound stream which may be oxidized do not contain meta-substituted aromatic compounds, during the typical partial oxidation of very pure para-xylene some isophthalic acid is formed, in particular if the liquid phase in benzoic acid is present in the reaction medium. Such spontaneous isophthalic acid may, due to the greater solubility than TPA in a solvent containing acetic acid and water, accumulate over time in industrial installations using recycled solvent. Therefore, all factors among the amount of isophthalic acid in the solvent stream, the content of meta-xylene in the aromatic compound stream that can be oxidized, and the rate of spontaneous formation of isophthalic acid in the reaction medium is taken into account in an appropriate manner relative to each other and in relation to any reactions in which it is consumed isophthalic acid. Isophthalic acid has been found to undergo other reactions in which it is consumed, in addition to the formation of 2,7-DCF, as discussed below. In addition, the inventors have found that other factors should also be considered in determining the appropriate ranges for meta-substituted aromatics during the partial oxidation of para-xylene to TPA. Other undesirable, highly colored impurities, such as 2,6-dicarboxyfluorenon (2,6-DCF), appear to be dependent to a large extent on the dissolved, para-substituted aromatics that always occur if liquid phase is introduced during oxidation in the liquid phase para-xylene. Therefore, reducing the 2.7-DCF content is best considered considering the level of other colored impurities produced.

[0182] The inventors have found that during the partial oxidation of para-xylene to TPA in the liquid phase, the formation of trimellitic acid increases with increasing levels of isophthalic acid and phthalic acid in the reaction medium. Trimellitic acid is a trifunctional carboxylic acid that causes branching of polymer chains during the production of PET from TPA. In many PET applications, the branching level should be controlled to have a low value, therefore trimellitic acid should be controlled to have a low level in purified TPA. In addition to the formation of trimellitic acid, the presence of meta-substituted and ortho-substituted compounds in the reaction medium causes the formation of other tricarboxylic acids (e.g.

1,3,5-trikarboksybenzenu). In addition, the increased content of tricarboxylic acids in the reaction medium increases the level of tetracarboxylic acid formation (e.g. 1,2,4,5-tetracarboxybenzene). Controlling the overall production of all aromatic carboxylic acids having more than two carboxylic acid groups is one factor in determining the preferred levels of metasubstituted and ortho-substituted compounds in the recycled solvent stream, in the oxidizable compound stream, and in the reaction medium of the present disclosure.

123 [0183] The inventors have found that during partial oxidation of para-xylene to TPA in the liquid phase, increasing the level in the liquid phase in the reaction medium of a series of dissolved aromatic carboxylic acids not containing non-aromatic hydrocarbyl groups directly leads to increased formation of carbon monoxide and carbon dioxide. This increased production of carbon oxides results in a reduction in yield for both the oxidant and the oxidizable compound of the latter because many of the simultaneously produced aromatic carboxylic acids can be considered as impurities on the one hand, but on the other hand they also have commercial value . Therefore, proper removal of relatively soluble carboxylic acids not containing non-aromatic hydrocarbyl groups from recycled solvent has economic value in preventing loss of oxidation-related compound and oxidant performance, in addition to the limited formation of very undesirable impurities such as various fluorenones and trimellitic acid. [0184] The inventors have found that during the partial oxidation of para-xylene to TPA in the liquid phase, apparently 2,5,4'-tricarboxybiphenyl formation cannot be avoided. 2,5,4'-tricarboxybiphenyl is an aromatic tricarboxylic acid formed by condensing two aromatic rings, optionally by condensing dissolved para-substituted aromatic compounds with an aryl radical, optionally with an aryl radical formed by decarboxylation or decarbonylation of the para-substituted aromatic compound. Fortunately, 2,5,4'-tricarboxybiphenyl is usually less than trimellitic acid and usually does not lead to significant problems with the branching of polymer particles during PET production. However, the inventors have found that an increased level of 2,5,4'-tricarboxybiphenyl in a reaction medium involving the oxidation of alkylaromatic compounds according to preferred embodiments of the present invention leads to an increase in the level of undesirable 2,6-DCF with a high coloration. Increased levels of 2.6-DCF are probably formed as a result of ring closure in 2,5,4'-tricarboxybiphenyl with the loss of a water molecule, but the exact and reliable mechanism of the reaction is unknown. If too much accumulation of 2,5,4'-tricarboxybiphenyl, which is better than TPA, is dissolved in a solvent containing acetic acid and water in recycled solvent, the conversion to 2,6-DCF may become unacceptably high.

[0185] For example, the inventors have found that during the partial oxidation of para-xylene to TPA in the liquid phase, aromatic carboxylic acids not containing non-aromatic hydrocarbyl groups (e.g., isophthalic acid) generally cause a slight reduction in the chemical activity of the reaction medium if present in the liquid phase in the appropriate concentration.

[0186] For example, the inventors have found that during the partial oxidation of para-xylene to TPA in the liquid phase, precipitation is very often imperfect (i.e., no equilibrium) with respect to the relative concentration of various solid phase and liquid phase chemical species. This may be due to the high precipitation rate at the reaction rates over time and space preferred herein, which leads to imperfect co-precipitation of the impurities or even occlusion. Therefore, when it is desired to limit the concentration of certain impurities (e.g. trimellitic acid and 2.6125)

DCF) in raw TPA due to the configuration of processes in the further part of the installation, it is beneficial to control their concentration in the solvent stream, as well as the speed of their formation in the reaction medium.

[0187] For example, the inventors have found that benzophenone derivatives (e.g., 4,4'-dicarboxybenzophenone and 2,5,4'tricarboxybenzophenone) formed during the partial oxidation of para-xylene have an undesirable effect on the PET reaction medium, although benzophenone derivatives are not as strongly colored in TPA as fluorenones and anthraquinones. Accordingly, it is desirable to limit the occurrence of benzophenones and selected precursors in the recycled solvent and in the stream of the compound that can be oxidized. In addition, the inventors have found that the presence of elevated levels of benzoic acid, provided in recycled solvent or generated in the reaction medium, leads to an increased formation rate of 4,4'-dicarboxybenzophenone.

[0188] In total, the inventors have found and quantified sufficiently a surprising reaction group of aromatic compounds not containing non-aromatic hydrocarbyl groups that occur during partial oxidation of paraxylene to TPA in the liquid phase. Returning to only one case of benzoic acid, the inventors have found that the increased level of benzoic acid in the reaction medium in some embodiments of the present invention leads to a significant increase in the production of undesirable 9-fluorenone-2-carboxylic acid with a strong color, to a significant increase in level 4, 4'-dicarboxybiphenyl, to increase the level of 4,4'-dicarboxybenzophenone, to slightly reduce the chemical activity of the desired oxidation of para-xylene and to increase the level of carbon oxides and associated yield losses. The inventors have found that the increased level of benzoic acid in the reaction medium also leads to an increase in the production of isophthalic acid and phthalic acid, the levels of which are desirably controlled to be in a low range according to similar aspects of the present invention. The number and significance of the reactions involving benzoic acid are probably even more surprising given that some inventors now envisage using benzoic acid instead of acetic acid as the main solvent component (see, e.g., US Pat. No. 6,562,997). In addition, the present inventors have found that benzoic acid forms spontaneously during oxidation of para-xylene at a rate that is relatively important compared to its formation from impurities such as toluene and ethylbenzene, often found in the stream of an oxidizable compound containing steam commercial grade xylene.

[0189] On the other hand, the inventors have found little relevance to the additional regulation of recycled solvent composition in relation to the presence of an aromatic compound that can be oxidized, and in relation to aromatic reaction intermediates in which both non-aromatic hydrocarbyl groups are retained and are relatively soluble in recycled solvent. Generally, these compounds are introduced into the reaction medium or are formed in it at a rate much higher than their presence in the recycled solvent; in addition, the consumption rate of these compounds, which retain one or more non-aromatic hydrocarbyl groups in the reaction medium is high enough,

127 to properly reduce their accumulation in recycled solvent. For example, during partial oxidation of para-xylene in a multiphase para-xylene reaction medium, it evaporates to a limited extent along with large amounts of solvent. When such a vaporized solvent leaves the off-gas reactor and condenses to be recovered as recycled solvent, much of the vaporized para-xylene also condenses therein. It is not necessary to limit the concentration of this para-xylene in recycled solvent. For example, if the solvent separates from the solids when the suspension leaves the para-xylene oxidation reaction center, the recovered solvent contains a similar concentration of dissolved para-toluic acid as present at the time of removal from the reaction center. Although limiting the constant concentration of para-toluic acid in the liquid phase of the reaction medium may be important, see below, it is not necessary to separately regulate para-toluic acid in this part of the recycled solvent due to its relatively good solubility and low mass flow rate relative to formation para-toluic acid in the reaction medium. Similarly, the inventors have found little importance in limiting the concentration of recycled aromatic compounds with methyl substituents (e.g. toluic acids), aromatic aldehydes (e.g. terephthaldehyde), aromatic compounds with hydroxymethyl substituents (e.g. 4-hydroxymethylbenzoic acid) and brominated aromatics with at least one a non-aromatic hydrocarbyl group (e.g. alpha-bromo-para-toluic acid) below the concentration naturally occurring in the liquid phase leaving the reaction medium during partial oxidation of xylene

128 according to preferred embodiments of the present invention. The inventors have also surprisingly found that it is not necessary to regulate in the recycled solvent the concentration of selected phenols produced naturally during partial oxidation of xylene, since these compounds are formed and degrade in the reaction medium at a rate much higher than their occurrence in the recycled solvent. For example, the inventors have found that 4-hydroxybenzoic acid has relatively little effect on chemical activity in preferred embodiments of the present invention if it is additionally added in an amount of more than 2 grams of 4-hydroxybenzoic acid per kilogram of para-xylene, much greater than natural occurrence in recycled solvent, although other authors have determined that this is a significant impurity in a similar reaction medium (see e.g. Partenheimer, Catalysis Today 23 (1995) p. 81).

[0190] This means that there are many reactions and many factors that should be considered when determining the preferred ranges of various aromatic impurities in the solvent stream, as disclosed below. These findings are given in relation to the cumulative weighted average composition of all solvent streams introduced into the reaction medium over a given period of time, preferably one day, preferably one hour and most preferably one minute. For example, if one solvent stream flows substantially continuously and the composition contains 40 ppmw of isophthalic acid with a flow rate of 7 kilograms per minute, the other solvent stream flows essentially continuously and the composition contains 2000 ppmw of isophthalic acid with a flow rate of 10 kilograms per minute

129 and there are no other solvent streams introduced into the reaction medium, then the cumulated weighted average composition of the solvent stream is calculated as follows: (40 * 7 + 2000 * 10) / (7 + 10) = 1193 ppmw of isophthalic acid. It should be noted that when calculating the cumulative weighted average composition of the solvent stream, the weight of any stream of compound that can be oxidized, or any oxidant stream that can be combined with the solvent stream before entering the reaction medium is not taken into account.

[0191] Table 1 below gives preferred values ​​for some of the components of the solvent stream introduced into the reaction medium. The components of the solvent stream given in Table 1 are as follows: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxystilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenon (2,6-DCF), 2,7-dicarboxyfluorenon (2,7-DCF), 3,5-dicarboxyfluorenon (3,5DCF), 9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), sum of fluorenones , including other fluorenones not specifically mentioned (sum of fluorenones), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA ), isophthalic acid (IPA), benzoic acid (BA), trimellitic acid (TMA), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4'-dicarboxybenzil (4,4'-DCBZ), 4,4'-dicarboxybenzophenone (4,4'- DCBP), 2,5,4'-tricarboxybenzophenone (2,5,4'-TCBP), terephthalic acid (TPA), precipitated solids at 20 ° C and the sum of aromatic carboxylic acids not containing non-aromatic hydrocarbyl groups. Table 1 below gives the preferred content of these impurities in the CTA prepared according to the embodiment of the present invention.

TABLE 1 - Components of the solvent stream introduced into the reaction medium

<td>Ingredient</td><td>Content favorable (Ppmw)</td><td>Content favorable (Ppmw)</td><td>Content most (Ppmw)</td>

<td>4-CBA</td><td> &lt; 1200</td><td> 30 - 600</td><td> 60 - 300</td>

<td>4,4'-DCS</td><td> &lt; 3</td><td> &lt; 2</td><td> &lt; 1</td>

<td>2,6-DCA</td><td> &lt; 6</td><td> 0,1 - 3</td><td> 0,2 - 1</td>

<td>2,6-DCF</td><td> &lt; 20</td><td> 0,1 - 10</td><td> 0,5 - 5</td>

<td>2,7-DCF</td><td> &lt; 10</td><td> 0,1 - 5</td><td> 0,5 - 2</td>

<td>3,5-DCF</td><td> &lt; 10</td><td> &lt; 5</td><td> &lt; 2</td>

<td>9F-2CA</td><td> &lt; 10</td><td> 0,1-5</td><td> 0,5 - 2</td>

<td>9F-4CA</td><td> &lt;5</td><td> &lt; 3</td><td> &lt; 1</td>

<td>Sum of fluorenones</td><td> &lt; 40</td><td> &lt; 20</td><td> 1 - 8</td>

<td>4,4'-DCB</td><td> &lt; 45</td><td> &lt; 15</td><td> 0,5 - 5</td>

<td>2,5,4'-TCB</td><td> &lt; 45</td><td> 0,1 - 15</td><td> 0,5 - 5</td>

<td>PA</td><td> &lt; 1000</td><td> 15 - 400</td><td> 40 - 150</td>

<td>IPA</td><td> 2500</td><td> 40 - 1200</td><td> 120 - 400</td>

<td>BA</td><td> &lt; 4500</td><td> 50 - 1500</td><td> 150 - 500</td>

<td>TMA</td><td> &lt; 1000</td><td> 15 - 400</td><td> 40 - 150</td>

<td>2,6-DCBC</td><td> &lt; 40</td><td> &lt; 20</td><td> &lt; 5</td>

<td>4,4'-DCBZ</td><td> &lt; 40</td><td> &lt; 20</td><td> &lt; 5</td>

<td>4,4'-DCBP</td><td> &lt; 40</td><td> &lt; 20</td><td> &lt; 5</td>

<td>2,5,4'-TCBP</td><td> &lt; 40</td><td> &lt; 20</td><td> 05 - 5</td>

<td>TPA</td><td> &lt; 9000</td><td> 200 - 6000</td><td> 400 - 2000</td>

<td>Precipitated solids at 20 ° C</td><td> &lt; 9000</td><td> 200 - 6000</td><td> 600 - 2000</td>

<td>The sum of aromatic carboxylic acids not containing non-aromatic hydrocarbyl groups</td><td> &lt; 18000</td><td> 300 - 9000</td><td> 450 - 3000</td>

[0192] Many other aromatic impurities are also typically present in the recycled solvent, generally at varying even lower levels and / or proportions with respect to one or more aromatic compounds disclosed. Control methods for disclosed aromatic compounds within preferred ranges typically allow an appropriate level of other aromatic impurities to be maintained.

[0193] If bromine is used in the reaction medium, it is known that a large number of ionic and organic bromine forms are present in dynamic equilibrium. These different forms of bromine have different stability parameters after leaving the reaction medium and carrying out various unit processes to which the recycled solvent is subjected. For example, alphabromo-para-toluic acid may persist unchanged under some conditions, while in others it may rapidly hydrolyze to 4-hydroxymethylbenzoic acid and hydrogen bromide. In the present invention, it is preferred that at least about 40 weight percent, more preferably at least about 60 weight percent, and most preferably at least about 80 weight percent of the total weight of bromine present in the cumulative solvent stream introduced into the reaction medium has one or more of the following chemical forms: ionic bromine, alpha-bromo-para-toluylic acid and bromoacetic acid.

[0194] Although the importance and value of the control of the cumulative weighted average purity of the solvent stream within the disclosed desirable ranges of the present invention have not yet been discovered and / or disclosed, an appropriate method of controlling the purity of the stream can be developed based on various methods already known in the art. solvent. First, each solvent evaporated from the reaction medium is usually of sufficient purity so that the liquid or solids from the reaction medium are not trapped in the evaporated solvent. The introduction of reflux solvent droplets into the off-gas separation space above the reaction medium as disclosed herein appropriately limits such entrapment; in addition, such a waste gas may be condensed into a solvent of suitable purity with respect to the aromatic compound. Secondly, much more difficult and costly purification of the recycled solvent stream usually refers to the solvent withdrawn from the reaction medium in liquid form and the solvent, which then contacts the liquid and / or solid phases in the reaction medium discharged from the reaction tank (e.g. recycled solvent obtained from the filter on which solids are concentrated and / or washed, recycled solvent obtained from the centrifuge on which the solids are concentrated and / or washed, recycled solvent obtained from the crystallization process, etc.). However, methods are also known in the art for carrying out the appropriate purification of such recycled solvent streams by one or more of the previous disclosures. Regarding the control of precipitated solids in recycled solvent to keep them within the given ranges, suitable control methods include, among others, gravimetric sedimentation, mechanical filtration with filter cloth on rotary belt filters and rotary drum filters, mechanical filtration with stationary filtration material in pressure tanks, hydrocyclones and centrifuges. Regarding the control of dissolved aromatics in recycled solvent to remain within the ranges, suitable control methods include, but are not limited to those disclosed in US Pat. U.S. Patent No. 4,939,297 and U.S. Patent Application Publication No. 2005-0038288. However, none of these earlier inventions identified and disclosed preferred levels of purity in the cumulative solvent stream as disclosed in this document. These prior art inventions merely provide methods for purifying selected and partial recycled solvent streams without specifying the present optimal values ​​of the cumulative weighted average composition of the solvent stream introduced into the reaction medium. [0195] Then, regarding the purity of the oxidizable compound stream, it is known that some amounts of isophthalic acid, phthalic acid and benzoic acid are present and can be tolerated at low levels in purified TPA used in the production of polymers. It is further known that these compounds are relatively more soluble in many solvents and can advantageously be removed from purified TPA in crystallization processes. However, it is now known from the embodiment of the present invention disclosed herein that controlling the level of a number of relatively soluble aromatic compounds, in particular isophthalic acid, phthalic acid and benzoic acid, in the liquid phase in the reaction medium is unexpectedly important for controlling the level of polycyclic and colorful aromatic compounds formed in the reaction medium, for the control of compounds having more than 2 carboxylic acid functional groups in the molecule, for the control of the activity of the reaction in the partial oxidation reaction medium and for the control of loss of efficiency relative to the oxidant and the aromatic compound.

[0196] It is known in the art that isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium as follows. The introduced meta-xylene impurity is oxidized at a high degree of conversion and yield to

IPA (isophthalic acid). The impurity introduced ortho 134 xylene oxidizes at a high degree of conversion and yield to phthalic acid. The impurities introduced in ethylbenzene and toluene are oxidized at a high degree of conversion and yield to benzoic acid. However, the inventors have found that significant amounts of isophthalic acid, phthalic acid and benzoic acid are also formed in a reaction medium containing para-xylene in processes other than oxidation of meta-xylene, ortho-xylene, ethylbenzene and toluene. These other natural chemical reactions may include decarbonylation, decarboxylation, transition state reorganization, and the addition of methyl and carbonyl radicals to aromatic rings.

[0197] When determining the preferred ranges of impurities in a stream of a compound that can be oxidized, many factors are important. Any contamination in the stream can cause direct losses in efficiency and product purification costs if the purity requirements of the oxidized product are high enough (e.g. in the partial oxidation reaction medium of para-xylene, toluene and ethylbenzene usually found in commercially pure para-xylene give benzoic acid, which is mostly removed from most commercial forms of TPA). If the partial oxidation product of a feed impurity participates in additional reactions, factors other than simple yield losses and removal (e.g. in the reaction medium for partial oxidation of para-xylene, ethylbenzene gives benzoic acid, which is then transformed into, inter alia, highly colored 9-fluorenone-2-carboxylic acid, isophthalic acid and phthalic acid and an increased amount of carbon oxides). If the re135 action medium naturally produces additional amounts of contaminants through chemical mechanisms not directly related to the contaminants of the raw material, the analysis becomes even more complex (e.g. in the medium for partial oxidation of paraxylene, benzoic acid is also spontaneously produced from para-xylene alone). In addition, further processing of the raw oxidation product may affect the analysis of favorable raw material purity. For example, the cost of removing and obtaining an appropriate level of direct impurity (benzoic acid) and secondary impurities (isophthalic acid, phthalic acid, 9-fluorenone-2-carboxylic acid and others) may be one and the same, may be different and may be different from requirements for the removal of substantially unrelated impurities (e.g., the product of incomplete oxidation of 4-CBA during oxidation of para-xylene to TPA).

[0198] The raw material purity ranges disclosed below for para-xylene are preferred, with para-xylene being introduced with the solvent and oxidant into the partial oxidation reaction medium in the production of TPA. These ranges are more advantageous in the TPA production process, which uses post-oxidation steps to remove non-oxidant and solvent contaminants (e.g., catalytic metals) from the reaction medium. These ranges are even more advantageous for TPA production processes in which additional 4-CBA is removed from CTA (e.g., by converting CTA into dimethyl terephthalate with ester impurities and subsequent separation of 4-CBA methyl ester by distillation, by secondary oxidation methods, in which 4-CBA is converted to TPA, by hydrogenation methods, in which 4-CBA is converted to paratoluic acid, which is then separated by partial methods

136 crystallization). These ranges are most preferred for TPA production processes in which additional 4-CBA is removed from CTA by secondary oxidation methods, in which 4-CBA is converted to TPA.

[0199] Thanks to new knowledge of beneficial ranges for recycled aromatics and related amounts of aromatics produced directly as a result of oxidation of raw material impurities compared to other natural chemical reactions, improved impurity ranges were found for impure para-xylene introduced into partial oxidation processes in production TPA. Table 2 below gives preferred values ​​for meta-xylene, ortho-xylene and ethylbenzene + toluene in the para-xylene stream expressed in parts per million by weight para-xylene.

TABLE 2 - Components of the impure paraxylene stream

<td>Ingredient</td><td>Content favorable (Ppmw)</td><td>More favorable content (ppmw)</td><td>Most favorable content (ppmw)</td>

<td>metaxylene</td><td> 20 - 800</td><td> 50 - 600</td><td> 100 - 400</td>

<td>orthoxylene</td><td> 10 - 300</td><td> 20 - 200</td><td> 30 - 100</td>

<td>ethylbenzene + toluene *</td><td> 2 - 700</td><td> 50 - 500</td><td> 100 - 300</td>

<td>sum</td><td> 50 - 900</td><td> 100 - 800</td><td> 200 - 700</td>

<td colspan="4">* Requirements for ethylbenzene + toluene are given separately and as a sum</td>

[0200] Those skilled in the art will now understand that the above impurities of impure para-xylene may have the greatest effect on the reaction medium after accumulating their partial oxidation products in recycled solvent. For example, entering the upper value of the most preferred range for meta-xylene, 400 ppmw, will immediately produce about 200 ppmw of liquid phase isophthalic acid in the reaction medium if the reaction proceeds at about 33 weight percent solids in the reaction medium. This can be compared with the introduction of the upper value of the most favorable range for isophthalic acid in the recycled solvent of 400 ppmw, which, after taking into account the typical evaporation of the solvent to cool the reaction medium, means about 1200 ppmw of liquid phase isophthalic acid in the reaction medium. This means that the accumulation of partial oxidation products over time in recycled solvent will cause the greatest impact of impurities of meta-xylene, ortho-xylene, ethylbenzene and toluene in the stream of impure para-xylene. Therefore, it is preferable to maintain the above ranges of impurities in the stream of contaminated para-xylene for at least half of each day of use of each partial oxidation reaction medium in a given production facility, more preferably for at least three quarters of each day for at least seven consecutive days of use and, if the mass-average averages of the contaminated para-xylene stream remain within the preferred ranges, for at least 30 consecutive days of use.

[0201] Methods for obtaining impure para-xylene with favorable purity are already known in the art and include, but are not limited to, distillation, partial crystallization methods below ambient temperature, and molecular sieve methods using selective adsorption depending on on pore size. However, the preferred purity ranges given herein are with them

138 upper bound more demanding and costly than typically provided by commercial suppliers of para-xylene; on the other hand, at the lower limit, in the preferred ranges, excessively expensive purification of para-xylene is avoided in order to introduce partial oxidation into the reaction medium by finding and disclosing that the combined effects of the spontaneous formation of impurities from the para-xymelene alone and the reaction of consuming impurities in the reaction medium become more important on the rate of introduction of impurities in the contaminated para-xylene.

[0202] If the xylene-containing feed stream contains selected impurities, such as ethylbenzene and / or toluene, benzoic acid may form as a result of oxidation of these impurities. In this document, "benzoic acid made from impurities" means benzoic acid derived from any source other than xylene during oxidation of xylene.

[0203] As disclosed herein, part of the benzoic acid produced during xylene oxidation is derived from xylene alone. Importantly, the formation of benzoic acid from xylene is added to any amount of benzoic acid produced, which may be benzoic acid produced from impurities. Without theoretical restrictions, it is believed that benzoic acid is formed from xylene in the reaction medium when various intermediate xylene oxidation products undergo spontaneous decarbonylation (carbon monoxide loss) or decarboxylation (carbon dioxide loss), resulting in the formation of aryl radicals. These aryl radicals can cleave a hydrogen atom from any of the many sources available in the reaction medium and produce spontaneously generated benzoic acid. Independently

139 from the chemical mechanism the term "spontaneously produced benzoic acid" as used herein means benzoic acid derived from xylene during oxidation of xylene.

[0204] As further disclosed herein, if para-xylene is oxidized to produce terephthalic acid (TPA), the production of spontaneously produced benzoic acid causes losses of para-xylene yield and loss of oxidant efficiency. In addition, the presence of spontaneously produced benzoic acid in the liquid phase in the reaction medium is correlated with the severity of many undesirable side reactions, in particular the formation of highly colored compounds called monocarboxyfluorenones. The spontaneously generated benzoic acid also contributes to the undesirable accumulation of benzoic acid in the recycled filtrate, which further increases the concentration of benzoic acid in the liquid phase in the reaction medium. In this way, the formation of spontaneously produced benzoic acid is preferably kept to a minimum, but it should also be considered simultaneously with benzoic acid formed from impurities and factors affecting the consumption of benzoic acid, factors related to other issues of reaction selectivity and general economic aspects.

[0205] The inventors have found that the spontaneous production of benzoic acid can be controlled to remain low due to the appropriate selection, for example, of temperature, xylene distribution and oxygen availability in the reaction medium during oxidation. Without being limited to theory, it appears that lower temperature and increased oxygen availability reduce the rate of decarbonylation and / or decarboxylation,

140 thanks to which there is no loss of efficiency due to spontaneously produced benzoic acid. Sufficient oxygen availability seems to direct aryl radicals towards the formation of other, less harmful products, in particular hydroxybenzoic acids. The distribution of xylene in the reaction medium can also affect the balance between the conversion of aryl radicals to benzoic acid or hydroxybenzoic acids. Regardless of the chemical mechanisms, the inventors have set reaction conditions that are mild enough to reduce benzoic acid production, but strong enough to oxidize a significant portion of the hydroxybenzoic acid produced to carbon monoxide and / or carbon dioxide, which is easily removed from the oxidation product.

[0206] In a preferred embodiment of the present invention, the configuration and operation of the reactor are such that the formation of spontaneously produced benzoic acid is minimized and the oxidation of hydroxybenzoic acids to carbon monoxide and / or carbon dioxide is maximum. If an oxidation reactor is used to oxidize para-xylene to terephthalic acid, it is preferred that the para-xylene accounts for at least about 50 weight percent of total xylene in the feed stream entering the reactor. More preferably, para-xylene makes up at least about 75 weight percent of total xylene in the feed stream. Even more preferably, para-xylene makes up at least about 95 weight percent of total xylene in the feed stream. Most preferably, para-xylene is substantially all of the total xylene in the feed stream.

[0207] If a reactor is used to oxidize para-xylene to terephthalic acid, it is preferred that the rate of terephthalic acid production is maximized and the rate of formation of spontaneously generated benzoic acid is kept to a minimum. Preferably, the ratio of the production rate (by weight) of terephthalic acid to the production rate (by weight) of self-produced benzoic acid is at least about 500: 1, more preferably at least about 1000: 1 and most preferably at least 1500: 1. It will be explained below that the rate of self-produced benzoic acid formation is preferably measured if the concentration of benzoic acid in the liquid phase in the reaction medium is below 2000 ppmw, more preferably below 1000 ppmw and most preferably below 500 ppmw, since such a low concentration limits the conversion of benzoic acid to other compounds at a low enough speed.

[0208] After the spontaneous benzoic acid and benzoic acid produced from impurities have been combined, the ratio of the production rate (by weight) of terephthalic acid to the production rate (by weight) of the sum (spontaneously produced and made from impurities) of benzoic acid is preferably at least about 400: 1, more preferably at least about 700: 1 and most preferably at least 1100: 1. According to the present invention, the sum of the rate of self-produced benzoic acid formation together with benzoic acid produced from impurities is measured if the concentration of benzoic acid in the liquid phase in the reaction medium is below 500 ppmw, since such a low concentration limits the reactions of converting benzoic acid to other compounds, respectively low speed.

[0209] As disclosed herein, an increased concentration of benzoic acid in the liquid phase in the reaction medium

142 leads to increased formation of many other aromatics, many of which are harmful TPA contaminants; in addition, as disclosed herein, the increased concentration of benzoic acid in the liquid phase in the reaction medium leads to increased formation of gaseous carbon oxides, the formation of which reduces the yield relative to the oxidant, as well as aromatic compounds and / or solvent. Furthermore, it is disclosed below that the inventors have found that a significant portion of this increased concentration of other aromatics and carbon oxides results from reactions in which some benzoic acid molecules are transformed, unlike benzoic acid that catalyzes other reactions in which it does not consume. Accordingly, "accidental benzoic acid formation" is defined herein as the time-weighted average weight of all benzoic acid leaving the reaction medium minus the time-weighted average weight of all benzoic acid entering the reaction medium over the same period. This resultant benzoic acid formation is often positive, which is due to the formation rate of benzoic acid formed from impurities and spontaneously generated benzoic acid. The inventors have found, however, that it appears that the degree of conversion of benzoic acid to carbon oxides and to a number of other compounds increases approximately linearly with increasing concentration of benzoic acid in the liquid phase in the reaction medium, measured when other reaction conditions including temperature, oxygen availability, the STR value and reaction activity are kept constant, respectively. Therefore, if the concentration of benzoic acid in the liquid phase in the reaction medium is high enough, probably in

143 As a result of the increased concentration of benzoic acid in recycled solvent, the conversion of benzoic acid molecules to other compounds, including carbon oxides, may even or outweigh the chemical formation of new benzoic acid molecules. In this case, the resultant benzoic acid formation may balance near zero or even be negative. The inventors have found that if the resultant benzoic acid formation is positive, the ratio of the production rate (by weight) of terephthalic acid in the reaction medium to the resultant resultant benzoic acid formation in the reaction medium is preferably above about 700: 1, more preferably above about 1100: 1 and most preferably above 4000 1. The inventors have found that if the incident benzoic acid formation is negative, the ratio of the production rate (by weight) of terephthalic acid in the reaction medium to the resultant benzoic acid formation in the reaction medium is preferably above about 200: (- 1), more preferably above about 1000: (- 1) and most preferably above 5000: (- 1).

[0210] The inventors have also found advantageous ranges for the composition of the suspension (liquid + solids) discharged from the reaction medium and the CTA solid in the suspension. The preferred suspension composition and the preferred CTA composition are surprisingly better and useful. For example, purified TPA made from such beneficial CTA by secondary oxidation has a sufficiently low level of the sum of impurities and colored impurities, making purified TPA suitable, without hydrogenation of additional 4-CBA and / or colored impurities, for a wide range of applications in PET and use in PET packaging. For example, the preferred composition of the suspension gives a liquid phase in the reaction medium which has a relatively low concentration

144 significant impurities, which significantly reduces the formation of other, even more undesirable impurities, as disclosed herein. In addition, the preferred composition of the suspension significantly facilitates the subsequent processing of the liquid from the suspension to obtain a suitably pure recycled solvent according to other embodiments of the present invention.

[0211] CTA prepared according to one embodiment of the present invention contains less impurities of selected types than CTA produced in conventional methods and installations, in particular when recycled solvent is used. Pollutants that may be present in CTA include the following: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxystilbene (4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenon (2,6-DCF), 2, 7-dicarboxyfluorenon (2,7-DCF), 3,5-dicarboxyfluorenon (3,5-DCF), 9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F4CA), sum of fluorenones, in including other fluorenones not specifically mentioned (sum of fluorenones), 4,4'-dicarboxybiphenyl (4,4'DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TCB), phthalic acid (PA), acid isophthalic (IPA), benzoic acid (BA), trimellitic acid (TMA), para-toluic acid (PTAC), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4'-dicarboxybenzil (4,4'-DCBZ), 4,4 ' -dicarboxybenzophenone (4,4'-DCBP), 2,5,4'-tricarboxybenzophenone (2,5,4'-TCBP). Table 3 below gives the preferred content of these impurities in the CTA prepared according to the embodiment of the present invention.

TABLE 3 - CTA pollution

<td>Contamination</td><td>Content favorable</td><td>More favorable content</td><td>Most favorable content</td>

145

<td></td><td>(Ppmw)</td><td>(Ppmw)</td><td>(Ppmw)</td>

<td>4-CBA</td><td> &lt; 15000</td><td> 100 - 8000</td><td> 400 - 2000</td>

<td>4,4'-DCS</td><td> &lt; 12</td><td> &lt; 6</td><td> &lt; 3</td>

<td>2,6-DCA</td><td> &lt; 9</td><td> &lt; 6</td><td> &lt; 2</td>

<td>2,6-DCF</td><td> &lt; 100</td><td> 2-50</td><td> 5-25</td>

<td>2,7-DCF</td><td> &lt; 30</td><td> &lt; 15</td><td> &lt; 5</td>

<td>3,5-DCF</td><td> &lt; 16</td><td> &lt; 8</td><td> &lt; 2</td>

<td>9F-2CA</td><td> &lt; 16</td><td> &lt; 8</td><td> &lt; 4</td>

<td>9F-4CA</td><td> &lt; 8</td><td> &lt; 4</td><td> &lt; 2</td>

<td>Sum of fluorenones</td><td> &lt; 100</td><td> 2 - 60</td><td> 4 - 35</td>

<td>4,4'-DCB</td><td> &lt; 64</td><td> 1 - 32</td><td> 2 - 8</td>

<td>2,5,4'-TCB</td><td> &lt; 24</td><td> &lt; 12</td><td> &lt; 8</td>

<td>PA</td><td> &lt; 200</td><td> 3 - 100</td><td> 5 - 50</td>

<td>IPA</td><td> &lt; 800</td><td> 10 - 400</td><td> 20 - 200</td>

<td>BA</td><td> &lt; 600</td><td> 5 - 300</td><td> 15 - 100</td>

<td>TMA</td><td> &lt; 800</td><td> 10 - 400</td><td> 20 - 200</td>

<td>PTAC</td><td> &lt; 2000</td><td> 10 - 1000</td><td> 50 - 500</td>

<td>2,6-DCBC</td><td> &lt; 64</td><td> &lt; 32</td><td> &lt; 8</td>

<td>4,4'-DCBZ</td><td> &lt; 12</td><td> &lt; 8</td><td> &lt; 4</td>

<td>4,4'-DCBP</td><td> &lt; 40</td><td> &lt; 30</td><td> &lt; 20</td>

<td>2,5,4'-TCBP</td><td> &lt; 32</td><td> &lt; 16</td><td> &lt; 4</td>

[0212] In addition, it is preferred that the CTA prepared according to the present invention has a reduced content of colored compounds relative to the CTA produced in conventional methods and installations, in particular when recycled solvent is used. Therefore, it is preferred that the CTA produced according to one embodiment of the present invention has a transmittance percentage at 340 nanometers (nm) of at least about 25 percent, more preferably at least about 50 percent and most preferably at least 60 percent. It is further preferred that the CTA produced according to one embodiment of the present invention has a transmittance percentage at 400 nanometers (nm) of at least about 88 percent, more preferably at least about 90 percent and most preferably at least 92 percent.

[0213] The percentage transmittance test is a measure of the colored, light-absorbing impurities found in

146

TPA or CTA. The test used in this document concerns measurements made on a portion of the solution obtained by dissolving 2.00 grams of dry solid TPA or CTA in 20.0 milliliters of dimethyl sulfoxide (DMSO) with a purity for analysis or greater. The volume of this solution is then placed in the Hellma flow micro-cuvette, catalog number 176,700, made of quartz, with an optical path length of 1.0 cm and a volume of 0.39 milliliters (Hellma USA, 80 Skyline Drive, Plainview, NY 11803). The Agilent 8453 diode spectrophotometer (Agilent Technologies, 395 Page Mill Road, Palo Alto, CA 94303) is used to measure the transmittance of different wavelengths of light passing through a filled flow cell. After appropriate correction for background absorbance, including but not limited to, associated with the cuvette and solvent used, the device directly reports the percentage transmittance measurement describing the fraction of incident light that passes through the solution. The percent transmittance values ​​at 340 nanometers and 400 nanometers are especially useful for distinguishing pure TPA from many of the impurities usually found in it.

[0214] The preferred ranges of the various aromatic impurities in the suspension phase (solids and liquid) in the reaction medium are given in Table 4 below.

TABLE 4 - Impurities in suspension

<td>Contamination</td><td>Content favorable (Ppmw)</td><td>More favorable content (ppmw)</td><td>Most favorable content (ppmw)</td>

<td>4-CBA</td><td> &lt; 8000</td><td> &lt; 5000</td><td> &lt; 2500</td>

<td>4,4'-DCS</td><td> &lt; 4</td><td> &lt; 2</td><td> &lt; 1</td>

<td>2,6-DCA</td><td> &lt; 6</td><td> &lt; 3</td><td> &lt; 1</td>

<td>2,6-DCF</td><td> &lt; 70</td><td> 2 - 40</td><td> 4 - 20</td>

<td>2,7-DCF</td><td> &lt; 12</td><td> &lt; 8</td><td> &lt; 4</td>

<td>3,5-DCF</td><td> &lt; 12</td><td> &lt; 8</td><td> &lt; 4</td>

147

<td>9F-2CA</td><td> &lt; 12</td><td> &lt; 8</td><td> &lt; 4</td>

<td>9F-4CA</td><td> &lt; 8</td><td> &lt; 4</td><td> &lt; 2</td>

<td>Sum of fluorenones</td><td> &lt; 90</td><td> 2 - 60</td><td> 5 - 30</td>

<td>4,4'-DCB</td><td> &lt; 64</td><td> 1 - 16</td><td> 2 - 4</td>

<td>2,5,4'-TCB</td><td> &lt; 60</td><td> 2 - 40</td><td> 4 - 20</td>

<td>PA</td><td> &lt; 3000</td><td> 25 - 1500</td><td> 75 - 500</td>

<td>IPA</td><td> 9000</td><td> 75 - 4500</td><td> 225 - 1500</td>

<td>BA</td><td> &lt; 15000</td><td> 100 - 6000</td><td> 300 - 2000</td>

<td>TMA</td><td> &lt; 3000</td><td> 25 - 1500</td><td> 75 - 500</td>

<td>PTAC</td><td> &lt; 8000</td><td> 100 - 4000</td><td> 200 - 2000</td>

<td>4,4'-DCBZ</td><td> &lt; 5</td><td> &lt; 4</td><td> &lt; 3</td>

<td>4,4'-DCBP</td><td> &lt; 240</td><td> &lt; 160</td><td> &lt; 80</td>

<td>2,5,4'-TCBP</td><td> &lt; 120</td><td> &lt; 80</td><td> &lt; 40</td>

[0215] These preferred slurry compositions are the preferred liquid phase composition of the reaction medium, such that the experimental difficulties associated with precipitation of additional liquid phase components from the reaction medium into the solid phase components are usefully avoided during sampling from the reaction medium, liquid and solid separation, and shifts of analysis conditions.

[0216] In the suspension phase in the reaction medium and in CTA in the reaction medium there are usually also many other aromatic impurities, generally at a variable, even lower level and / or proportion relative to one or more aromatic compounds disclosed. Controlling disclosed aromatic compounds within preferred ranges maintains an appropriate level of other aromatic impurities. These preferred compositions of the slurry phase in the reaction medium and solid CTA taken directly from the slurry are made possible by using the embodiments of the invention disclosed herein for the partial oxidation of para-xylene to TPA.

[0217] Measurement of the concentration of components present at low levels in the solvent, recycled solvent, CTA, suspension from the reaction medium and PTA is performed by methods

148 liquid chromatography.

[0218] The method referred to in claim 1 is HPLC-DAD, which includes high pressure liquid chromatography (HPLC) connected to a diode detector (DAD) enabling the separation and quantification of various molecules in a given sample. The device used in these measurements is a model 1100 HPLC apparatus equipped with a DAD detector, supplied by Agilent Technologies (Palo Alto, CA), but other suitable instruments are available for purchase, as well as from other suppliers. It is known in the art that both the elution time and the detector response are calibrated with known compounds present in known amounts, these compounds and amounts being adapted to those found in real, unknown samples.

[0219] Another method is HPLC-MS, which involves high pressure liquid chromatography (HPLC) coupled with mass spectrometry (MS) that allows separation, identification and quantification of different molecules in a given sample. The device used in these measurements is the HPLC Alliance apparatus and the ZQ mass spectrometer supplied by Waters Corp. (Milford, MA), but other suitable cameras are available for purchase, as well as from other suppliers. It is known in the art that both the elution time and the mass spectrometer response are calibrated using known compounds present in known amounts, these compounds and amounts being adapted to those found in real, unknown samples.

[0220] Another embodiment relates to the partial oxidation of an aromatic compound that can be oxidized with an appropriate balance of harmful aromatic impurities

149 on the one hand, and the production of carbon dioxide and carbon monoxide, i.e. combined carbon oxides (COx), on the other. These carbon oxides typically leave the reaction tank in the waste gas and correspond to the destructive losses of the solvent and the compound that can be oxidized, including ultimately preferred oxidized derivatives (e.g. acetic acid, paraxylene and TPA). The inventors have found the lower limits of carbon monoxide production, below which it appears that the production of the significant amounts of harmful aromatic impurities described below, and the low overall conversion are inevitably too low to be economically useful. The inventors have also found upper limits for carbon oxides, above which the production of carbon oxides continues to increase with little added value obtained by reducing the formation of harmful aromatic impurities.

[0221] The inventors have found that the reduction in the liquid phase concentration of the oxidizable feed that may be oxidized and the aromatic intermediates in the reaction medium leads to a lower rate of formation of harmful impurities during the partial oxidation of the aromatic that can be oxidized. These harmful impurities include fused aromatic rings and / or aromatic molecules containing more than the desired number of carboxylic acid groups (e.g., during oxidation of para-xylene, harmful impurities include 2.6-dicarboxyanthraquinone, 2,6-dicarboxyfluorenon, trimellitic acid, 2.5, 4'-tricarboxybiphenyl and 2,5,4'-benzophenone). Aromatic intermediates include aromatic compounds derived from an introduced aromatic compound, which may

150 undergo oxidation and still retaining non-aromatic hydrocarbyl groups (e.g., when oxidizing para-xylene to aromatic intermediates include para-tolualdehyde, terephthaldehyde, para-toluic acid, 4-CBA, 4-hydroxymethylbenzoic acid and alpha-bromo-para-toluylic acid). It seems that the introduced aromatic compound and aromatic intermediates retaining non-aromatic hydrocarbyl groups, if present in the liquid phase in the reaction medium, lead to the formation of harmful impurities in a manner similar to that disclosed earlier in the case of dissolved aromatic compounds not containing non-aromatic hydrocarbyl groups (e.g. isophthalic acid).

[0222] Due to the desire for higher reaction activity to reduce the formation of harmful aromatic impurities during the partial oxidation of an aromatic compound that may be oxidized, the inventors have found that an undesirable associated result is increased production of carbon oxides. It is important to remember that these carbon oxides cause performance losses relative to the oxidizable compound and the oxidant, not just the solvent. Namely, the significant and sometimes major part of the carbon oxides comes from the oxidizable compound and its derivatives, not from the solvent; in addition, often a compound that can be oxidized costs more per carbon than a solvent. In addition, it should be remembered that the desired product, carboxylic acid (e.g. TPA) also undergoes excessive oxidation to carbon oxides if it occurs in the liquid phase of the reaction medium.

[0223] It should further be remembered that the present invention relates to the reaction of the liquid phase in the reaction medium and the concentration of substrates therein. This sets it apart from some prior art inventions that directly relate to the production of an aromatic compound that retains non-aromatic hydrocarbyl groups in the form of a precipitated solid. In particular, in the case of partial oxidation of para-xylene to TPA, some prior art inventions relate to the amount of 4-CBA precipitated in the CTA solid phase. However, current inventors have found a variation of more than two to one in the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase using the same requirements for temperature, pressure, catalysis, solvent composition and reaction rate over time and space of para-xylene in depending on whether the partial oxidation is carried out in a strongly mixed autoclave or in a reaction medium with a degree of oxygen and paraxylene concentration according to the present invention. In addition, the inventors have found that the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase can vary by more than two to one in a strongly mixed or graduated reaction medium depending on the reaction rate over time and space of para-xylene outside similar requirements for temperature, pressure, catalysis and solvent composition. In addition, it appears that 4-CBA in CTA in the solid phase does not contribute to the formation of harmful impurities, moreover, 4-CBA in the solid phase can be recovered and oxidized to TPA in a simple manner and with high efficiency (e.g. by secondary oxidation of the CTA suspension described in this document); in turn, the removal of harmful impurities is much more difficult and expensive than the removal of 4-CBA in the solid phase, and the formation of carbon oxides is a permanent loss of efficiency. That is why it is important to

152 to distinguish that this aspect of the present invention relates to the liquid phase composition of the reaction medium.

[0224] Regardless of whether the carbon oxides are derived from a solvent or from an oxidizable compound, the inventors have found that, with economically useful conversion, their production is strongly associated with the overall activity of the reaction despite the large variability of specific temperature combinations of metals, halogen, temperature, acidity of the reaction medium measured on the basis of pH, water concentration used to obtain the level of general reaction activity. The inventors have found that in the case of partial oxidation of xylene, it is useful to assess the level of overall reaction activity by means of the concentration of toluic acids in the liquid phase at half the height of the reaction medium, at the bottom of the reaction medium and on the upper surface of the reaction medium.

[0225] Therefore, there is a significant need to achieve a balance at the same time to minimize the formation of harmful impurities by increasing the reaction activity and at the same time minimizing the formation of carbon oxides by reducing the reaction activity. This means that if the overall carbon monoxide production is reduced to too low, excessive amounts of harmful impurities will form, and vice versa.

[0226] Furthermore, the inventors have found that the solubility and relative reactivity of the desired carboxylic acid (e.g. TPA) and the presence of other dissolved aromatic compounds not containing non-aromatic hydrocarbyl groups are very important factors in balancing carbon oxides against harmful impurities. The carboxylic acid (desired product) usually dissolves in the liquid phase in

153 reaction medium, even if it is also in solid form. For example, at temperatures in preferred ranges, TPA is soluble in a reaction medium containing acetic acid and water at levels ranging from about one thousand ppmw to over 1 weight percent, with solubility increasing with increasing temperature. However, there are differences in the reaction rate towards the formation of various harmful impurities from the introduced aromatic compound, which can be oxidized (para-xylene), from aromatic reaction intermediates (para-toluic acid), from the desired product - aromatic carboxylic acid (TPA) and from aromatic compounds not containing non-aromatic hydrocarbyl groups (isophthalic acid), wherein the occurrence and reactivity of the last two groups determines the area of ​​reduced profits with respect to further reducing the content of the first two groups, the oxidizable aromatic compound introduced, and the aromatic reaction intermediates. For example, during partial oxidation of para-xylene to TPA, if dissolved TPA is present in the amount of 7000 ppmw in the liquid phase in the reaction medium under given conditions, dissolved benzoic acid is present in the amount of 8000 ppmw, dissolved isophthalic acid is present in the amount of 6000 ppmw and dissolved phthalic acid is present in the amount of 2000 ppmw, then the value of further reduction of the sum of harmful compounds begins to decrease, because the reaction activity increases, causing the liquid phase concentration of para-toluic acid and 4-CBA to fall below similar levels. This means that the occurrence and concentration in the liquid phase of the reaction medium of aromatic compounds not containing non-aromatic hydrocarbyl groups is very low

154 changes due to the increase in the activity of the reaction, their presence causing an upward expansion of the area of ​​reduced profits while reducing the concentration of reaction intermediates to reduce the formation of harmful impurities. [0227] Therefore, one embodiment reveals preferred ranges of carbon oxides limited at the lower end by low reaction activity and excessive formation of harmful impurities, and at the upper end by excessive carbon loss, but at a lower level than previously discovered and disclosed as being economically useful. Therefore, the formation of carbon oxides is preferably controlled in the following manner. The ratio of the number of moles of the sum of carbon oxides produced to the number of moles of aromatic compound introduced which may be oxidized is preferably greater than about 0.02: 1, more preferably greater than about 0.04: 1, even more preferably greater than about 0.05: 1 and most preferably greater than 0.06: 1. At the same time, the ratio of the number of moles of the sum of carbon oxides produced to the number of moles of aromatic compound introduced which may be oxidized 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 than 0.15: 1. The ratio of moles of carbon dioxide produced to the number of moles of aromatic compound introduced which may be oxidized is preferably greater than about 0.01: 1, more preferably greater than about 0.03: 1, even more preferably greater than about 0.04: 1, and most preferably greater than 0.05: 1. At the same time, the ratio of the number of moles of carbon dioxide produced to the number of moles of aromatic compound introduced which may be oxidized is preferably less than about 0.21: 1, more preferably less than about 0.19: 1, more preferably less than about 0.16: 1 and most preferably less than 0.11. The ratio of the number of moles of carbon monoxide produced to the number of moles of aromatic compound introduced which may be oxidized is preferably greater than about 0.005: 1, more preferably greater than about 0.010: 1, even more preferably greater than about 0.015: 1 and most preferably greater than 0.020: 1 . At the same time, the ratio of the number of moles of carbon monoxide produced to the number of moles of aromatic compound introduced which may be oxidized 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.

[0228] The carbon dioxide content of the dry off-gas from the oxidation reactor is preferably greater than about 0.10 mol percent, more preferably greater than about 0.20 mol percent, even more preferably greater than about 0.25 mol percent, and most preferably greater than 0, 30 mole percent. At the same time, the carbon dioxide content of the dry off-gas from the oxidation reactor is preferably less than about

1.5 mole percent, more preferably less than about 1.2 mole percent, even more preferably less than about 0.9 mole percent, and most preferably less than 0.8 mole percent. The carbon monoxide content of the dry off-gas from the oxidation reactor is preferably greater than about 0.05 mole percent, more preferably greater than about 0.10 mole percent, even more preferably greater than 0.15 and most preferably greater than 0.18 mole percent. At the same time, the carbon monoxide content of the dry exhaust gas from the oxidation reactor is preferably less than about 0.60 mole percent, more preferably less than about 0.50 mole percent,

156 even more preferably less than about 0.35 mole percent and most preferably less than 0.28 mole percent.

[0229] The inventors have found that an important factor in limiting the formation of carbon oxides to these preferred ranges is to increase the purity of the recycled filtrate and the oxidizable compound introduced to reduce the concentration of aromatic compounds not containing non-aromatic hydrocarbyl groups according to the disclosures of the present invention, which simultaneously reduces the formation of carbon oxides and harmful impurities. Another factor is the improved distribution of para-xylene and oxidant in the reaction vessel according to the disclosures of the present invention. Other factors that enable the above preferred carbon monoxide levels to be achieved are work under the gradient conditions in the reaction medium disclosed herein with respect to pressure, temperature, oxidizable compound concentration, liquid phase and gas phase oxidant. Other factors for achieving the above preferred carbon monoxide levels are the work of the disclosures herein with respect to reaction time and space time, pressure, temperature, solvent composition, catalyst composition, and mechanical geometry of the reaction vessel.

[0230] An important benefit of working in the preferred ranges of carbon monoxide formation is that molecular oxygen consumption can be reduced, but not to stoichiometric values. Regardless of the proper gradation of the oxidant and the compound that can be oxidized, according to the present invention, excess oxygen above the stoichiometric value should be maintained; it is calculated on the basis of the compound stream itself, which can be oxidized, to account for some losses to carbon oxides and to obtain excess molecular oxygen to control the formation of harmful impurities. In particular, in the case where the oxidizable compound stream is xylene, the ratio of molecular weight introduced to molecular weight 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 the molecular weight of molecular oxygen introduced to the 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. In particular for xylene incorporation, the time-averaged molecular oxygen content of the dry off-gas from the oxidation reactor is preferably greater than about 0.1 mole percent, more preferably greater than about 1 mole percent, and most preferably greater than 1.5 mole percent. At the same time, the time-average molecular oxygen content of the dry off-gas from the oxidation reactor is preferably less than about 6 mole percent, more preferably less than about 4 mole percent, and most preferably less than 3 mole percent.

[0231] Another significant benefit of working in the preferred ranges of carbon monoxide formation is that less aromatic compound is converted to carbon oxides and other less valuable forms. This benefit is assessed on the basis of the sum of moles of all aromatics leaving the reaction environment divided by the sum of moles of all aromatics entering the reaction environment over a continuous period, preferably one hour, more preferably

158 one day and most preferably 30 consecutive days. This ratio is referred to below as the "molar persistence ratio" for aromatics passing through the reaction medium and is expressed as a percentage. If all incoming aromatics leave the reaction medium as aromatics, but mainly in the oxidized form of the incoming aromatics, then the molar persistence ratio has a maximum value of 100 percent. If exactly 1 in 100 aromatic molecules entering is converted to carbon oxides and / or other non-aromatic molecules (e.g. acetic acid) passing through the reaction medium, then the molar persistence ratio is 99 percent. Particularly in the case where xylene is the main component of the aromatic compound stream that can be oxidized, the molar ratio of the persistence of aromatics passing through the reaction medium is preferably greater than about 98 percent, more preferably greater than about 98.5 percent, and most preferably less than 99.0 percent. At the same time and in an order such that there is sufficient overall reaction activity, the molar ratio of persistence of aromatics passing through the reaction medium is preferably less than about 99.9 percent, more preferably less than about 99.8 percent, and most preferably less than 99.7 percent if xylene is the main component of the aromatic stream that can be oxidized.

[0232] Another aspect of the present disclosure relates to the formation of methyl acetate in a reaction medium containing acetic acid and one or more aromatic compounds that may be oxidized. Methyl acetate is relatively volatile in comparison

159 to water and acetic acid, therefore it accompanies waste gases, unless additional cooling or other processes in the installation are used to recover and / or destroy them before the waste gases are released into the environment. Therefore, the formation of methyl acetate constitutes operating and capital costs. Methyl acetate is probably formed first by combining a methyl radical, probably derived from the decomposition of acetic acid, with oxygen to give methyl hydroperoxide, then by decomposition into methanol and finally by reacting the produced methanol with the remaining acetic acid to give methyl acetate. Regardless of the chemical reaction path, the inventors have found that when the rate of methyl acetate formation is too low, the production of carbon oxides is also too low, and the formation of harmful aromatic impurities is too high. When the rate of methyl acetate formation is too high, then the production of carbon oxides is also unnecessarily high, which leads to a loss of yield relative to the solvent, the oxidizable compound, and the oxidant. If the preferred embodiments disclosed herein are used, the ratio of the number of moles of methyl acetate produced to the number of moles of aromatic compound introduced which may be oxidized is preferably greater than about 0.005: 1, more preferably greater than about 0.010: 1, and most preferably greater than 0.020 1. At the same time, the ratio of the production of moles of methyl acetate to the moles of introduced aromatic compound that may be oxidized 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.

160 [0233] The inventors reserve that for all numerical ranges given herein, the upper and lower limits of the ranges may be independent of each other. For example, a number range of 10 to 100 means more than 10 and / or less than 100. Therefore, the range of 10 to 100 is a literal confirmation of the reservation, in which the expression above 10 (without upper limit) and the reservation, in which the expression less than 100 (without lower limit) is given, as well as the full range of 10 to 100 ( together with the lower and upper limits).

EXAMPLE (references) [0234] In this example, some properties of new CTA particles (shown in FIGS. 32A and 32B) and comparative CTA (shown in FIGS. 33A and 33B) were evaluated. Particularly for CTA particles, particle size, BET surface area and dissolution rate were tested. The study of CTA particles was based on the test descriptions given above in the Detailed description of this document.

[0235] New CTA particles were produced by partial oxidation of commercial grade para-xylene by recycling the filtrate in a bubble column reactor operating according to many embodiments of the present invention containing a solvent containing acetic acid and water, cobalt, manganese and bromine containing catalyst components , at a temperature close to 160 ° C at half height of the reaction medium, with favorable gas phase composition gradients, liquid phase composition and reaction medium temperature, preferred overall STR values ​​and preferred oxygen STR gradients. Comparative CTA particles were produced in the tank reactor with high stirring again as a result

161 partial oxidation of commercially pure para-xylene using recycled filtrate with a solvent containing acetic acid and water, catalytic components containing cobalt, manganese and bromine, but with a more spatially homogeneous reaction rate, solvent evaporation rate, liquid phase composition and gas phase composition and under higher pressure and at a temperature close to 200 ° C. Most of the test results are given in tabular form in Table 5, and the complete data sets for the time dissolution test are shown in the graph in FIG. 36.

TABLE 5 - CTA properties

<td></td><td>New CTA</td><td>Comparative CTA</td>

<td>photomicrographs</td><td>FIG. 32A and</td><td>FIG. 33A and</td>

<td></td><td>32B</td><td>33B</td>

<td>D (4.3) (microns)</td><td> 62</td><td> 205</td>

<td>D (v, 0,1) (microns)</td><td> 35</td><td> 27</td>

<td>D (v, 0.5) (microns)</td><td> 58</td><td> 181</td>

<td>D (v, 0.9) (microns)</td><td> 92</td><td> 369</td>

<td>Relative size distribution</td><td> 0,92</td><td> 1,67</td>

<td>particles</td><td></td><td></td>

<td>BET surface area (m²/ G)</td><td> 1,18</td><td> 0,57</td>

<td>Dissolving in time</td><td> 621</td><td> 458</td>

<td>After 1 minute (ppm in THF)</td><td></td><td></td>

<td>After 2 minutes (ppm in THF)</td><td> 765</td><td> 579</td>

<td>After 4 minutes (ppm in THF)</td><td> 886</td><td> 695</td>

<td>Dissolution model constants in</td><td> 198</td><td> 243</td>

<td>time A (ppm in THF)</td><td></td><td></td>

<td>B (ppm in THF)</td><td> 748</td><td> 695</td>

<td>C (minutes^-1)</td><td> 0,77</td><td> 0,29</td>

162 [0236] When comparing two CTAs, similar amounts of the smallest particles should be noted, as indicated by D (v, 0.1) values, however the new CTA had a significantly smaller mean particle size (D (4.3)) and a median particle size (D ( v, 0.5)). In addition, the new CTA had a much narrower relative particle size distribution and fewer very large particles (D (v, 0.9)), which usually dissolve and / or react in later technological processes the most difficult. In addition, and despite the similarity with respect to the number of very small particles with a very large surface area, the new CTA had a much higher overall surface area measured on the basis of the BET value, which was expected from the visual inspection of photomicrography. The combination of all physical factors gave a significantly and useful higher dissolution rate for the new CTA, although the final equilibrium solubility was similar. It was found that the new CTA was surprisingly useful in many subsequent technological processes involving dissolution and / or chemical reaction, as described in the disclosure, but the new CTA was at the same time more difficult and less desirable in other technological processes such as filtration and liquid washing in to remove cobalt, manganese and bromine residues.