MXPA06005214A - Staged countercurrent catalytic oxidation of disubstituted bemzene - Google Patents

Abstract

A process for oxidation with oxygen of at least one aromatic hydrocarbon having oxidizable, substituents that maximized the utilization of oxygen without reduction in the quality of the carboxylic acid products produced by means of a stagewise countercurrent oxidation system is disclosed.

Description

CATALYTIC OXIDATION IN COUNTERCORRENT, BY STAGES, OF DISSOLVED BENZENE FIELD OF THE INVENTION This invention relates generally to an improved oxidation process for the conversion of one or more aromatic hydrocarbon materials having oxidizable substituents with respect to their corresponding acid derivatives, and relates more particularly to such process involving the oxidation in stages and the recycling of the residual oxygen from a second oxidation stage to a first oxidation stage. BACKGROUND OF THE INVENTION It is well known that aromatic hydrocarbons having at least one and preferably two or more oxidizable substituent groups can be converted to carboxylic acid products by effecting the oxidation of such groups with molecular oxygen under controlled conditions. Such conditions have generally included the use of a known oxidation catalyst together with a suitable solvent. During the present commercial production of aromatic acids such as terephthalic acid, it is essential that the partial pressure of the reactor oxygen in the oxidation of an alkyl aromatic substance be sufficiently Ref. 172689 raised to prevent underfeeding of oxygen. A high partial pressure of the oxygen reduces the formation of undesirable colored byproducts by suppressing the coupling reactions. Also, a high partial pressure increases the reaction rates of the oxidation, which allows higher reactor performances, and reduces the burning of the reaction solvent. However, in a commercial operation of such an oxidation system, a significant loss of oxidation capacity occurs as a result of insufficient utilization of molecular oxygen. Therefore, it is highly desirable to improve the oxygen utilization and thereby improve the efficiency and debit of the process and increase the nominal capacity of a commercially mentioned oxidation system, and simultaneously maintain the high quality of the carboxylic acid products produced. Spillar et al., In U.S. Pat. No. 2,962,361 (November 29, 1960) describes a continuous, step-wise countercurrent oxidation system that "makes the use of practically quantitative oxygen possible ... without substantial detriment to the performance or quality of the product". The highest concentration of oxygen is introduced in the final stage, and the malodorous gases from each stage are returned to the preceding stage while the partially oxidized products move from the first stage to the final stage. It is described as desirable that the final oxidation step be at a higher temperature, pressure and concentration of oxygen. It is also described with respect to the first oxidation stage 11 and the ventilation line 26 thereof in Figure 2 that "additional air or oxygen can be introduced through line 41 ... to prevent oxygen concentration. in the receiver 24, the capacitor 23 or the line 26 from a volume percentage exceeding 8 (it is preferably zero) ". Baldwin et al., In U.S. Pat. No. 3,092,658 (June 4, 1963) describes a continuous, stepwise, countercurrent oxidation system that is very similar to that of U.S. Pat. No. 2,962,361 mentioned above. Baldwin, in the U.S. patent. No. 3,064,044 (November 13, 1962) also describes a step-wise countercurrent oxidation system. The non-condensed malodorous gases leaving the second oxidation stage are returned to the first oxidation stage, and the patent states that they must contain less than 8 percent oxygen but that they may contain approximately 1 to 8 percent oxygen and therefore , is introduced by the lines 14 and 15 to supply oxygen in the container 11. With respect to the first oxidation stage 11 and the condenser 20 and the receiver 21 through which its malodorous gases pass., the patent also states that "the amount of additional air introduced from line 15 must be controlled so that the oxygen content of the gases in the condenser 20 and the receiver 21 will be less than 8 percent, preferably close to zero" . June et al., In the U.S. patent. No. 6,153,790 (November 28, 2000) describes a process for producing substituted aromatic substances, with a diacid, with a purity of at least 97%. The process comprises contacting in a stirred tank reactor, a dialkyl-substituted aromatic substance in an organic acid solvent with an oxidant containing at least 50 volume% oxygen, at a partial oxygen pressure of at least 0.07 kg. / cm2 (1 psia), at a temperature between about 80 ° C (176 ° F) and about 130 ° C (266 ° F), in the presence of a catalyst system comprising zirconium and cobalt. A vapor stream comprising the organic acid solvent, steam and the unreacted oxidant is removed from the reactor. More than 50% by volume of oxygen in the oxidant is required so that the total pressure of the reaction system can be low enough to allow cooling of the reflux of the reaction system at temperatures between 80 ° C (176 ° F) ) and approximately 130 ° C (266 ° F) as a result of the vaporization of the components of the liquid phase to form the vapor stream mentioned above. The design of the reactor must effectively provide almost full oxygen consumption below the liquid / gas interface. The nitrogen can be introduced in the vicinity of the liquid / gas interface in an amount sufficient to make the gas mixture in the vapor phase, non-flammable. The patent discloses that, if desired, after dilution with nitrogen, the unconsumed oxygen can be contacted with the feed streams in an optional pre-reactor to provide full or near full utilization of the oxygen. Turner and Hously, in the U.S. patent application. No. U.S. 2001/0007910 Al, published on July 12, 2001; PCT / US01 / 20109, published July 18, 2002 as WO 02/055468 Al; PCT / USOl / 00825, published July 19, 2001 as WO 01/51442 A2; and PCT / US01 / 19960, published July 18, 2002 as WO 02/055467, disclose a process for the stepwise catalytic liquid phase, the oxidation with air of a suitable precursor, such as paraxylene, to a carboxylic acid, such as terephthalic acid, comprising the oxygenation of a feed stream comprising acetic acid and an oxidation catalyst at a high pressure, from 200 to 20,000 kPa, continuously and simultaneously feeding the oxygenated feed stream and paraxylene to a first zone of reaction which is placed upstream of a conventional oxidation reactor to form a reaction medium in which the mass ratio of acetic acid to paraxylene is in the range from 10: 1 to 20: 1, and the products of reaction are kept in solution. In this first reactor the absorption of oxygen within the reaction medium in the first reaction zone is limited to less than 50% of the oxygen for the total conversion of the paraxylene present to the terephthalic acid. After this, the reaction medium is fed from the first oxidation zone to the conventional oxidation reactor mentioned above and simultaneously, the pressure of the reaction medium is reduced to a pressure in the range from 1000 to 2,000 kPa in the reactor. conventional oxidation. In WO 01/51442 A2 the process is described as a method for increasing the production capacity of a conventional oxidation reactor, while the other three patent publications describe it as a method to reduce the formation of impurities in the acid product. carboxylic acid and to control the degradation of the solvent and the precursor. Although it is highly desirable to maximize the utilization of oxygen and thereby improve the efficiency of the process and increase the nominal capacity of a commercial oxidation system while maintaining the high quality of the carboxylic acid products produced and without the need to add capacity of additional compressor, this goal has never been achieved and the means to achieve it have never been described. BRIEF DESCRIPTION OF THE INVENTION The present invention is a stepwise, continuous, improved countercurrent process of this invention for the catalytic oxidation of at least one benzene disubstituted with oxidizable substituents selected from the class consisting of the alkyl, hydroxyalkyl groups, aldehyde, carboxy and mixtures thereof, up to their corresponding acid derivative in a solvent system. This process comprises the steps of: a) introducing into a first oxidation zone a feed mixture comprising from about 7 to about 60 weight percent of the total amount of at least one disubstituted benzene introduced in steps (a) and (b), and at least a portion of the total amounts to be introduced in steps (a) and (b) of each of (i) the solvent, (ii) the catalyst components comprising at least one metal , catalytic selected from the class consisting of manganese, cobalt, nickel, zirconium, hafnium, and mixtures thereof, and (iii) bromine at a molar ratio based on the total catalytic metals within the range of from about 1:20 to approximately 5: 1; (b) partially oxidizing the disubstituted benzene in the first oxidation step with an oxygen-containing, molecular gas, which initially contains from about 3 to about 20 volume percent molecular oxygen, at a temperature in the range of about 121.22 ° C (250 ° F) up to 205 ° C (401 ° F), and with the relative amounts of the disubstituted benzene, the catalyst components, and the solvent, and with the temperature that is such that from about 25 to about 99.95 percent in weight of disubstituted benzene fed to the first oxidation stage, be oxidized to form a gaseous mixture comprising unreacted molecular oxygen, and the vaporized solvent and a first mixture of the product comprising disubstituted benzene that did not react, disubstituted partially oxidized benzene, the product derived from the acid, and the solvent, and at a pressure that is sufficient to maintain the disubstituted benzene, disubstituted partially oxidized benzene, the product of the acid derivative, and the solvent substantially as a liquid phase or a solid / liquid suspension, and in such a way that the concentration of the molecular oxygen residual in the remaining gas mixture is from about 0.3 to about 2 volume percent; (c) recovering the first mixture of the resulting product from the first oxidation step, and feeding at least a portion of the first mixture of product recovered to a second oxidation step; (d) feeding to the second oxidation stage the molecular oxygen or a gas containing the molecular oxygen and the remainder, if any, of the predetermined total amounts, mentioned above, introduced in steps (a) and (b) ) of the disubstituted benzene, catalyst components, solvent and bromine; (e) substantially completely oxidizing, in the second oxidation step, the disubstituted partially oxidized benzene and disubstituted benzene that did not react, fed to the second oxidation step in the presence of the molecular oxygen-containing gas at a temperature within the range from about 175 ° C (347 ° F) to about 216.1 ° C (421 ° F) and with the relative amounts of the disubstituted benzene, disubstituted partially oxidized benzene, the components of the catalyst and the solvent, and with the temperature that it is such that from about 96 to about 100 weight percent disubstituted benzene and disubstituted partially oxidized benzene are sanitized to form a second product mixture comprising the above-mentioned acid-derived product and the solvent, and at a sufficiently high pressure to keep the product derived from the acid, partially disubstituted benzene oxidized and disubstituted benzene that did not substantially react as a liquid or solid / liquid suspension, and such that the concentration of residual molecular oxygen in the gas removed from the second oxidation stage is in the range of from about 3 to about 15 percent by volume; (f) recovering from the second oxidation step the second mixture of the product comprising the acid derivative; and (g) extracting from the second oxidation stage and recycling to the first oxidation stage the gas containing the residual molecular oxygen. The present invention is also the solid-liquid solution or suspension produced in step (b). BRIEF DESCRIPTION OF THE FIGURES For a more complete understanding of this invention, reference should now be made to the embodiment illustrated in greater detail in the appended figures and those described below by way of examples of the invention. In the figures, Figure 1 is a schematic illustration of one embodiment of the method of this invention employing the continuous, step-wise, countercurrent oxidation of paraxylene that results in the production of high quality terephthalic acid with maximum oxygen utilization.

Figure 2 contains a series of graphs of the total acetic acid, burned, against the concentration of the 4-carboxybenzaldehyde in the reactor for a number of the illustrative examples and the comparative examples given hereinafter. Figure 3 contains a series of graphs of the optical density of the terephthalic acid product after it has been recovered, washed and dried against the concentration of 4-carboxybenzaldehyde in the reactor for a number of illustrative examples and comparative examples they are given here later. It should be understood that Figure 1 is schematic in nature. In certain cases, details that are not necessary for the understanding of the present invention or which make other details difficult to perceive, may be omitted. It should be understood, of course, that the invention is not necessarily limited to the particular embodiments illustrated herein. DETAILED DESCRIPTION OF THE INVENTION The mixture / components of the feed for the process of this invention includes at least one aromatic hydrocarbon having at least one oxidizable substituent group capable of being oxidized to a corresponding dicarboxylic acid, i.e., the product of the acid derivative . Preferred feedstock components include at least one disubstituted benzene material having any of a variety of substituents selected from the class consisting of the alkyl, hydroxyalkyl, aldehyde and carboxyalkyl groups, together with mixtures thereof. Particularly preferred components of the feed mixture include the paradisubstituted benzene derivatives having alkyl groups as substituents, the acid derivative of which is terephthalic acid, and the partially oxidized forms thereof include mono- and di-alcohols and -aldehydes thereof, and the monoacid thereof, such as p-hydroxymethyl benzoic acid, p-tolualdehyde, and p-toluic acid. The alkyl groups preferably contain 1-4 carbon atoms, and are more preferably methyl groups. Accordingly, a component of the especially preferred feed mixture is paraxylene. At least one disubstituted benzene employed in the present invention is provided in a solution with the solvent which is preferably an organic acid solvent. Preferably, organic acids include organic acids having from one to six carbon atoms plus a carboxyl group, such as benzoic acid. The most preferred solvent is acetic acid due to its vapor pressure at the preferred reactor temperatures and its solvent capacities. These organic acids are solvents for the reasonable concentrations of the feed components, the components of the catalyst system, the intermediate oxidation products and the dicarboxylic acid product. The disubstituted benzene feed is preferably provided in a solution of between about 5 and about 25 weight percent. The components of the feed mixture also comprise at least one catalytic metal selected from the class consisting of manganese, cobalt, nickel, zirconium, hafnium, cerium and mixtures thereof, and a material that produces bromine. Preferably, the catalytic metals are cobalt and manganese. The catalytic metals can be in any form that is soluble in the reaction medium. Examples of such soluble forms include salts of organic acid, basic salts, complex compounds and alcoholates. The catalytic metal component can be added to the reaction mixture with disubstituted benzene or separately. Other metals and promoters may also be present in the catalyst system employed in the present invention. The material that produces the bromine can be molecular bromine, a bromide or bromate salt, hydrobromic acid, an organic compound substituted with bromine, or a mixture of any or all of these.

The oxidant employed in the present invention is oxygen, which for the purpose of this invention means molecular oxygen. The oxygen source employed in this invention is typically pure oxygen, air or improved air containing additional amounts of oxygen. From about 7, preferably from about 15, to about 60, preferably up to about 35 weight percent of the total amount of the disubstituted benzene to be added in steps (a) and (b), is added in the first step of oxidation (a). The comparative example A hereinafter, illustrates that if the total p-xylene to be introduced in steps (a) and (b) is introduced in step (a), then the optical density (1.79) of the product The resulting terephthalic acid is unacceptably high. The optical density at 340 nm (OE 340) measures the concentration of the compounds of undesirably high molecular weight, which causes yellowness and fluorescence. Preferably from about 20, more preferably from about 40, to about 100, percent by weight of the predetermined total amount, mentioned above, to be added in steps (a) and (d) of each of the catalytic metals , preferably cobalt and manganese, is added in the first oxidation stage. Preferably from about 20, more preferably from about 40, to about 100, percent by weight of the predetermined total amount mentioned above to be added in steps (a) and (d) of the bromine, is added in the first step of oxidation. Preferably from about 10, more preferably from about 40, preferably up to about 100, percent by weight of the total amount of the solvent to be added in steps (a) and (d), is added in the first oxidation step. . With respect to the first oxidation stage alone, the molar ratio of total catalytic metals to bromine is from about 1:20, preferably from about 1: 5, more preferably from about 1: 4, to about 5: 1. , preferably up to about 2: 1, more preferably up to about 1: 1. The weight ratio of the catalytic metals to the solvent in the first oxidation step is from about 150, preferably from about 400, to about 10,000, preferably up to about 5,000 parts by weight of the catalytic metals per million parts of the solvent. The atomic ratio of manganese to cobalt in the first oxidation step is in the range of about 1: 100, preferably from about 1: 5, to about 100: 1, preferably to about 5: 1. The temperature of the reaction in the first oxidation step is in the range from about 121 ° C (250 ° F), preferably from about 136.11 ° C (277 ° F), up to about 205 ° C (451 ° F), preferably up to about 177.22 ° C (351 ° F). The reaction pressure in the first oxidation stage is sufficiently high at the temperature used to maintain the solvent, the disubstituted partially oxidized benzene, the product of the acid derivative, and the disubstituted benzene that did not react therein, substantially as a liquid or as a solid-liquid suspension. Typically, the reaction pressure in the first oxidation step is in the range of about 9.1481 kg / cm2 (130 psig) to about 15.13 kg / cm2 (215 psig). The oxygen concentration in the oxygen-containing gas, introduced in the first oxidation step, is in the range from about 3, preferably from about 4, to about 20, preferably up to about 11., more preferably up to about 8, percent by volume of oxygen. The reaction conditions are selected such that the concentration of residual oxygen in the gas removed from the first oxidation step is from about 0.3 to about 2, preferably up to about 1, percent by volume. Under such conditions, from about 25, preferably from about 60, more preferably from about 70 to about 99.95 percent by weight of the disubstituted benzene, fed to the first oxidation step, is partially or completely oxidized in the first oxidation step. Comparative examples A and B illustrate that when the oxygen content of the vent gas from the first oxidation stage is less than 0.3%, then the optical density (1.79 and 1.08) of the resulting terephthalic acid product is unacceptably high. With respect to the second oxidation step alone, the molar ratio of the bromine to the total catalytic metals is from about 1:20, preferably from about 1: 5, more preferably from about 1: 4, to about 5: 1, preferably up to about 2: 1, more preferably up to about 1: 1. The weight ratio of the catalytic metals to the solvent in the second oxidation step is from about 150, preferably from about 400, to about 10,000, preferably up to about 5,000 parts by weight of the catalytic metals per million parts of the solvent. The atomic ratio of manganese to cobalt in the second oxidation step is in the range of from about 1: 100, preferably from about 1: 5, to about 100: 1, preferably to about 5: 1. The reaction temperature in the second oxidation step is in the range from about 175 ° C (347 ° F), preferably from about 182.22 ° C. (360 ° F), up to about 216.11 ° C (421 ° F), preferably up to about 205 ° C (401 ° F). The reaction pressure in the second oxidation stage is sufficiently high at the temperature used to maintain the solvent, the acid-derived product, the disubstituted partially oxidized benzene and the disubstituted benzene which did not react, substantially as a liquid or as a suspension solid-liquid. Typically, the reaction pressure in the second oxidation step is in the range from about 11.96 kg / cm2 (170 psig) to about 16.54 kg / cm2 (235 psig). The temperature in the first oxidation step is preferably at least -14.72 ° C (5.5 ° F), more preferably at least 4.44 ° C (40 ° F), lower than the temperature in the second oxidation step. The pressure in the first oxidation step is at least 0.35 kg / cm2 (5 psig), more preferably at least 1.40 kg / cm2 (20 psig) lower than the pressure in the second oxidation stage. The concentration of oxygen in the oxygen-containing gas, introduced in the second oxidation step, is in the range from about 15, preferably from about 20, to about 50, preferably to about 25, percent by volume of oxygen. The reaction conditions are selected within the above ranges for this, such that the concentration of residual oxygen in the gas removed from the second oxidation stage is from about 3, preferably from about 4, to about 15, preferably up to about 11, more preferably up to about 8 volume percent. Under such conditions, from about 97, preferably from about 99 weight percent, to about 100 weight percent of the disubstituted benzene and partially oxidized disubstituted benzene fed to the second oxidation step, is oxidized in the second oxidation step. The first oxidation stage may be comprised of a single reactor or a plurality of reactors operated in parallel. Similarly, the second oxidation step may be comprised of a single reactor or a plurality of reactors operated in parallel. Accordingly, a first oxidation stage comprising a plurality of reactors, for example, four reactors, can be used with a second oxidation stage comprising a single reactor. In that case, the suspension or product in solution of all the reactors in the first oxidation stage could be fed to the single reactor in the second oxidation stage, and the oxygen in the ventilation gas from the single reactor in the Second stage of oxidation could be recycled up to and divided between the four reactors in the first oxidation stage. In an alternative, a first oxidation stage comprising a single reactor could be used with a second oxidation stage comprising a plurality of reactors, for example, four reactors. In that case, the product in solution or suspension of the single reactor in the first oxidation stage could be divided and fed to each of the four reactors in the second oxidation stage. One embodiment of the method of this invention in which illustrative examples 1-18 and comparative examples A and B are modeled, is illustrated by the schematic representation in figure 1. For the purpose of this illustration, the disubstituted benzene employed is the paraxylene, its acid derivative is terephthalic acid, and its partially oxidized disubstituted benzenes (intermediate oxidation compounds) include para-hydroxymethyl benzoic acid, p-tolualdehyde, 4-carboxybenzaldehyde and p-toluic acid. A solution of paraxylene and the catalyst components described above in an acetic acid solvent is introduced into a first reactor 10 from a feed tank 11 through line 12. The contents of the first reactor 10 are kept well mixed using an agitator 13, and the reactor pressure is maintained at the desired level with a pressure regulator 14 in the vent line 15 from the first reactor 10. The temperature of the contents of the first reactor 10 is regulated using a heating jacket 16 around the first reactor 10. The oxygen is introduced into the first reactor 10 through lines 17 and 18 in a stream of the vent gas from the second reactor 20. If desired, the additional oxygen as compressed air in the line 21 can be combined with the vent gas stream from the second reactor 20 to increase the oxygen concentration in the gaseous stream that is introduced to the first reactor 10 through lines 17 and 18. In the alternative, if desired, the additional nitrogen may be combined through line 22 with the vent gas stream from the second reactor 20 to reduce the concentration of oxygen in the gaseous stream that is int it flows to the first reactor 10 through line 17.

Paraxylene and oxygen react in the first reactor 10 to form a solution or suspension of the unreacted paraxylene-containing product, the intermediate oxidation compounds thereof and the terephthalic acid. The exothermic heat of the reaction is removed by vaporizing some of the solvent. A stream of malodorous gases comprising vaporized solvent, nitrogen (from the air) and unreacted oxygen is removed from the first reactor 10 through line 25 to a condenser 26 where most of the vaporized solvent is condensed and already is returned to the first reactor 10 through lines 27, 28 and 29, or passed to the second reactor 20 through lines 27, 28, and 30, or divided with a first portion that is returned to the first reactor 10 through of lines 27, 28, and 29, and a second portion that is passed to the second reactor 20 through lines 27, 28, and 30. The non-condensable vapor is extracted through line 33 and the vent line 15 while a retrograde stream of the non-condensable vapor is extracted for analysis through line 36 to verify the extent of reaction and the concentration of oxygen that did not react in the vent gas. The level of the solution or suspension of the product in the first reactor 10 is maintained by a valve 41 between the first reactor 10 and the second reactor 20. The pressure in the first reactor 10 can be maintained higher than the pressure in the second reactor. 20 so that when the valve 41 is opened, the solution or suspension of the product is transferred from the first reactor, to the second reactor 20 through the lines 42 and 43. In the alternative, a pump (not shown) in the lines 42 or 43 can be used to pump the solution and suspension from the first reactor 10 to the second reactor 20. If necessary or desired, additional amounts of one or more of the paraxylene, one or more of the catalyst components, the solvent and / or bromine can be introduced to the second reactor 20 from tanks 44, 45, and 46, respectively, through lines 47, 48, and 49, respectively. The compressed air is introduced through the line 50 into the second reactor 20. When the disubstituted benzene being oxidized is p-xylene, the product composition of the first oxidation step on a solvent free base as a percentage of the total product mixture comprises from about 5.0 to about 85.0 weight percent of the terephthalic acid, from about 2.0 to about 20.0 weight percent of the 4-carboxybenzaldehyde, from about 0.0 to about 3.0 weight percent of the hydroxymethylbenzoic acid, from about 5.0 to about 65.0 weight percent of the p-toluic acid, from about 0.0 to about 30.0 weight percent of the p-tolualdehyde, and from about 0.0 to about 35.0 weight percent of p-xylene. As with the first reactor 10, the contents of the second reactor 20 are kept well mixed using an agitator 53, and the reactor pressure is maintained at the desired level with a pressure regulator 54 in the vent line 19 from the second reactor. The temperature of the contents of the second reactor 20 is regulated using a heating jacket 55 around the second reactor 20. The oxygen reacts with the unreacted paraxylene and the oxidation intermediates thereof in the second reactor 20 to form a solution or suspension of the product containing terephthalic acid. Oxidation of the paraxylene and its oxidation of the intermediate compounds to the terephthalic acid is carried out to a high degree of complement in the second reactor 20. The solution or suspension of the resulting product of the terephthalic acid is then extracted from the second reactor 20 through line 56, the valve for the water carried by the steam 57 and the line 58. Inside the second reactor 20, the exothermic heat of the reaction is removed by vaporization of some of the solvent. A stream of malodorous gases comprising the vaporized solvent, nitrogen and unreacted oxygen is withdrawn from the second reactor 20 through "line 61 to a condenser 62 where most of the vaporized solvent is condensed and returned to the second reactor on lines 63 and 64. A retrograde stream of the condensed solvent is extracted for analysis through line 65. The non-condensable vapor is extracted through lines 67 and 19 while a retrograde stream of the non-condensable vapor is extracted for analysis through line 68 to verify the extension of the reaction and the concentration of oxygen that did not react in the vent gas of the second reactor 20. The solution or suspension resulting from the aromatic acid product extracted from the second oxidation reactor, it is then typically subjected to crystallization as described in US Patent No. 3,092,658, mentioned above mind in column 2, lines 45-63; U.S. Patent No. 2,962,361 in column 2, line 43 to column 3, line 2; and the U.S. patent No. 3,064,044 in column 3, lines 47-71. In one embodiment, the solution or suspension of the aromatic acid extracted from the second oxidation reactor is passed to one or more containers where it is contacted with the air under the oxidation conditions to further oxidize the intermediate impurities. The solution or suspension is then subjected to crystallization, as described hereinabove. Typically, the aromatic acid can be recovered by centrifugation or filtration, and then purified using a hydrogenation catalyst and an aqueous solvent, as is well known in the art. This invention is applicable to any aromatic acid purification process, such as those known in the art, examples of which are described in U.S. Pat. Nos. 5,354,898 and 5,362,908, both of which are incorporated herein by reference. In general, a purification process of the aromatic acid comprises the hydrogenation of the dissolved raw aromatic acid within a liquid purification stream comprising the solvent to produce the dissolved purified aromatic acid. The dissolved purified aromatic acid is then crystallized and the resulting solid purified acid separated from the liquid purification stream, typically by filtration. The invention can be suitably used in an aromatic acid purification process, wherein the crude aromatic acid (e.g., crude terephthalic acid) is dissolved in a liquid purification stream comprising the solvent, and treated with hydrogen in a vessel of the pressure reactor in a first reaction zone containing a hydrogenation catalyst. The hydrogenation catalyst of the pressurized reactor vessel typically comprises one or more components of the active hydrogenation catalyst supported on a carrier material. The carrier material is typically in a granular form, although the pills or other particular types of shape may be used. When in a granular form, the granules preferably have an average mesh size of -2 to -12 mesh (series of U.A. sieves), more preferably of -4 mesh to -8 mesh. The carrier material is preferably an activated carbon, and more preferably is derived from coconut charcoal. Such activated carbon typically has a surface area of at least 600 m2 / gram (N2, BET method), preferably 800 m2 / gram to 1500 m2 / gram. Although activated carbon derived from coconut charcoal in the form of granules is preferred as a support material for the hydrogenation catalyst component, other metal, carbonaceous, porous, or other supports or substrates can be used. The hydrogenation catalyst contains at least one active catalytic hydrogenation component. Particularly suitable catalytic hydrogenation components are the metals of group VIII of the periodic table of the elements (IUPAC version), including palladium, platinum, rhodium, osmium, ruthenium, iridium, and mixtures thereof. The catalytic hydrogenation catalyst component can be deposited on, or added to, the carbon or other carrier material by any suitable method, for example, by treatment of the carrier with a solution of one or more soluble group VIII metal compounds, such as palladium chloride, and then drying the resulting material to remove the excess solvent. A preferred charge of the group VIII metal on the carrier is in the range of 0.01 to 2% by weight based on the total weight of the terminal catalyst, i.e., the total weight which is the weight of the dry carbon carrier and the component of active hydrogenation. More preferably, the loading of the group VIII metal on the carbon carrier is from 0.2 to 0.8% by weight. Catalysts and suitable catalyst beds, useful in the embodiment of this invention which relates to the purification of the aromatic acid, are described, for example, in U.S. Pat. Nos. 4,394,299; 4,629,715; 4,728,630 and 4,892,972. A palladium on carbon catalyst suitable, for example, can be obtained from Engelhard Corporation, Edison, N. J. Also, suitable rhodium on carbon catalysts can be obtained from Engelhard Corporation. A suitable reactor for hydrogenation is any reactor vessel that can withstand the temperature and pressure used for the hydrogenation of a crude aromatic acid dissolved in the purification solvent. The preferred configuration of the reactor is a cylindrical reactor, placed with its vertical axis arranged and having the hydrogenation catalyst contained therein in a fixed bed. In the preferred operation mode, the crude aromatic acid dissolved in a purification solvent is added to the reactor vessel in a position at or near the upper portion of the reactor vessel, and the crude aromatic acid dissolved in the liquid stream of purification flows down through the bed of the hydrogenation catalyst contained in the reactor vessel in the presence of hydrogen gas, where the impurities are reacted with the hydrogen gas. In this preferred mode, the crude aromatic acid is purified and the purified product is removed from the reactor vessel to a position at or near the bottom of the reactor. In a suitable reactor vessel apparatus, a hydrogenation catalyst preferably comprising a carbon carrier and an active hydrogenation catalyst component, supported on the carrier, is held within the reactor vessel by a sieve or other means that retains the catalyst particles in the reactor, still allowing the relatively free passage of the crude aromatic acid dissolved in the liquid purification stream. The means used to retain the catalyst particles can be a flat mesh screen or a screen made by parallel wires spaced closely. Other suitable catalyst retention means include, for example, a tubular Johnson screen or a perforated plate. The means used to retain the catalyst particles are constructed of a material that is suitably corrosion resistant and of adequate strength to efficiently retain the catalyst bed. Even more suitably, the means used to retain the catalyst bed, have openings of 1 mm or less and are constructed of a metal, such as stainless steel, titanium, or Hastelloy C. The reactor can be operated in several ways. For example, a predetermined liquid level can be maintained in the reactor and hydrogen can be fed, for any given pressure in the reactor, at a rate sufficient to maintain the predetermined liquid level. The difference between the actual pressure of the reactor and the vapor pressure of the present purification liquid stream is the partial pressure of hydrogen in the reactor vapor space. Alternatively, if the hydrogen is fed in a mixed manner with an inert gas such as nitrogen, the difference between the actual reactor pressure and the vapor pressure of the crude acid solution present, is the combined partial pressure of the hydrogen and the inert gas. mixed with it. In this case, the partial pressure of the hydrogen can be calculated from the known relative amounts of the hydrogen and the inert gas present in the mixture. In yet another operating mode, the reactor can be filled with a liquid purification stream so that no vapor space is provided in the reactor. That is, the reactor can be operated as a fully hydraulic system with the dissolved hydrogen that is fed to the reactor by flow control. In such a case, the concentration of hydrogen in the solution can be modulated by adjusting the speed of hydrogen flow to the reactor. If desired, a partial pressure value of pseudohydrogen can be calculated from the concentration of hydrogen in the solution which, in turn, can be correlated with the flow velocity of hydrogen with respect to the reactor. In the mode of operation wherein the control process is carried out by adjusting the partial pressure of the hydrogen, the partial pressure of hydrogen in the reactor is preferably in the range of 69-1379 kPa (10 pounds per square inch gauge up to 200 pounds per inch square gauge) or higher, depending on the calibration of the reactor service pressure, the degree of contamination of the crude aromatic acid mentioned above, the activity and age of the particular catalyst employed, and other processing considerations known to the skilled person. in art. In the operating mode where the control of the process is carried out by directly adjusting the concentration of the hydrogen in the feed solution, the latter is usually lower than the saturated one with respect to hydrogen and the reactor itself is hydraulically full. Accordingly, an adjustment of the flow velocity of the hydrogen with respect to the reactor will lead to the desired control of the concentration of the hydrogen in the solution. In general, the amount of hydrogen that is to be supplied to the purification reactor under the reaction conditions is, of course, sufficient to effect the desired hydrogenation. The space velocity, reported as the weight of the crude aromatic acid per weight of the catalyst per hour, during the hydrogenation, is typically from 1 hour "1 to 25 hours" 1, preferably from 2 hours "1 to 15 hours" 1. The residence time of the liquid purification stream in the catalyst bed varies, depending on the space velocity. After hydrogenation, the hydrogenation stream, which now comprises the purified aromatic acid and the solvent, is removed from the reactor and cooled to a crystallization temperature. The crystallization temperature is sufficiently low (eg 160 ° C (320 ° F) or below this value) for the crystallization of the purified aromatic acid to occur, whereby crystals are produced within the liquid phase. The crystallization temperature is sufficiently high so that the impurities and their reduction products (the products resulting from the hydrogenation) remain dissolved in the liquid phase. After this, the liquid, which contains the dissolved impurities and their reduction products, is separated (typically by centrifugation or filtration) from the purified, crystallized aromatic acid. A particularly desirable application of the method of the present invention is in conjunction with one or more existing oxidation reactor (s). In this case, the existing reactor (s) could (n) serve as the second stage in the method of this invention, and a pre-oxidation reactor could be installed and used as the first stage in conjunction with the existing reactor (s) as the second stage. The result could be an effective increase in the production capacity of the existing reactor (s) without the use of additional air or additional compressor capacity and without any significant reduction in the quality or yield of the acid derivative. derivative produced or the use of the additional solvent.

The present invention will be understood more clearly from the following examples, which are illustrative, without limitation, of the improved oxidation process of this invention. Illustrative examples 1-18 and comparative examples A and B The feed solution in each of the illustrative examples 1-18 and the comparative examples A and B, was made from measured amounts of the solvent (acetic acid and water), the catalyst (cobalt acetate, manganese acetate, and hydrobromide) and paraxylene (PX), and stored in a solvent / catalyst / PX feed tank. This feed was pumped into the first reactor at a fixed speed. The amounts of nitrogen and compressed air were mixed to produce a current that is equal to the volume and had the same oxygen content of the non-condensable vapor stream that was extracted from the second reactor. Accordingly, the resulting stream of compressed air and nitrogen that is fed to the first reactor, simulated the composition and volume of the vent stream from the second reactor and was introduced into the first reactor. The contents of the first reactor were kept well mixed with an agitator. The reactor pressure was maintained at the desired level with a pressure regulator in a vent line to vent the uncompressed vent gas from the first reactor. The temperature of the reactor was regulated with a heating jacket around the reactor. The level of the reactor was maintained by the opening and closing of a valve between the first and second reactor. For reasons of experimental convenience, the first reactor pressure was maintained slightly higher than the second reactor pressure so that the material was transferred by differential pressure from the first reactor to the second reactor when that valve was opened. The PX and oxygen were reacted in the first reactor to form a solution or suspension of the reaction containing the unreacted PX, the intermediate oxidation compounds thereof, and the terephthalic acid as the final oxidation product. The heat of the reaction was removed by vaporizing some of the solvent. The venting current of the first reactor was sent to a condenser where most of the solvent vapor was condensed. The condensed reflux was either returned to the first reactor or passed to the second reactor or a combination of both. A retrograde stream of the non-condensable vapor was extracted and analyzed to verify the extent of the reaction. The solution or suspension of the reaction was passed from the first reactor to the second reactor. The additional solvent and PX were also fed at specified rates to the second reactor. The amount of compressed air added to the second reactor was such that the oxygen concentration in the ventilation of this reactor was at the desired level. The control of the second reactor was similar to that of the first reactor, except that a small retrograde current was removed from the reflux to control the water content of the reactor content to the desired level. The oxidation of the PX and its intermediate oxidation compounds in the second reactor were brought to a high degree of complement. The terephthalic acid product was removed from the second reactor by means of a valve for the water carried by the steam. The specific reaction conditions employed and the results for the illustrative examples 1-18 and the comparative examples A and B, are presented in Tables 1-7. In Tables 1-7, PX, HAc, HMBA, TA, 4CBA and BA represent p-xylene, acetic acid, hydroxymethylbenzoic acid, terephthalic acid, 4-carboxybenzaldehyde and benzoic acid, respectively. The concentration of H20 in the last column of table 1 under the reactor liquid represents the concentration of water in the first reactor that includes the water formed in the oxidation reaction. SCFH stands for standard cubic feet per hour. Burning means the amount of acetic acid consumed in the example and reported as pounds of HAc per thousand pounds of PX (lbs / MlbPX). The proportion of the solvent in Table 4 is reported as pounds of solvent per pound of total PX feed for both reactors (total solv./lb of the PX feed). OD 340 represents the optical density of TA at 340 nm after it has been recovered, washing and filtering. Comparative Examples 1-9 The same procedure used for Examples 1-18 as for Comparative Examples 1-9 was used, except that Comparative Examples 1-9 did not employ two oxidation reactors. A single oxidation reactor was employed and was controlled substantially in the same manner as the second oxidation reactor employed in Examples 1-18. All of the components of the reaction mixture were introduced directly into the single oxidation reactor employed. The non-condensable vapor from the oxidation reactor was removed from the reactor as described for the first oxidation reactor in Examples 1-18. The specific reaction conditions employed and the results for comparative examples 1-9 are presented in Tables 8-11. The abbreviations and units used in tables 8-11 are the same as those used in tables 1-7. Comparison of the results of the comparative examples, 4, 5, 6 and 7 with those of the comparative examples 1, 2, 3, and 8 illustrates that if the oxygen of the ventilation is reduced to about 1 volume percent of the normal of about 4 volume percent, the color of the terephthalic acid product, as measured by the optical density (OD 340), will be significantly increased and will make the TA product unacceptable. Examples 1 to 18 illustrate that if the reduction of ventilation oxygen to about 1 volume percent is achieved with stepwise countercurrent oxidation, the color of the terephthalic acid product as well as the amount of the acetic acid burned is comparable to that of the common oxidation method, that is, about 4 volume percent of the oxygen of ventilation without the oxidation in step countercurrent. Comparison of the results of Examples 18 and 14 illustrates that the amount of disubstituted benzene introduced into the first reactor can vary from 20 to 55 weight percent of the total predetermined amount of disubstituted benzene introduced into both reactors. Examples 1 to 18 illustrate that the feed gas to the first reactor may contain 4 to 6 volume percent oxygen. Examples 1 and 18 illustrate that the temperature of the first reactor can vary from 136.11 ° C (277 ° F) to 160.55 ° C (321 ° F). Examples 8 and 17 illustrate that about 79 to 99.3 weight percent disubstituted benzene fed to the first oxidation step is oxidized. Examples 13 and 14 illustrate that the residual molecular oxygen concentration in the remaining gas mixture of the first reactor can vary from 0.73 to 1.66 volume percent. Figures 2 and 3 contain graphs illustrating the effect of the concentration of 4-carboxybenzaldehyde (4-CBA) in the oxidation reactor on: (a) the total acetic acid burned (measured in pounds of acetic acid per 1000 pounds of paraxylene ) and, (b) the optical density at 340 nm of the terephthalic acid product after it has been recovered, washed and dried. The graphs in figure 2 that illustrate the total acetic acid burned for runs with stepwise countercurrent oxidation, are comparable with those for runs without stepwise countercurrent oxidation with 4 volume percent ventilation oxygen, when Results are compared based on an equal weight percentage of 4-CBA in the terephthalic acid product. It is further illustrated that if the oxygen of ventilation is reduced to 1 volume percent without the countercurrent oxidation in stages, the resulting acetic acid burned is much higher. Similarly, the graphs in Figure 3 illustrate that the optical density for runs with stepwise countercurrent oxidation is comparable to those for runs without stepwise countercurrent oxidation with 4 volume percent of the ventilated oxygen. It is further illustrated that if the ventilated oxygen is reduced to 1 volume percent without stepwise countercurrent oxidation, the optical density of the resulting terephthalic acid product is much higher. Accordingly, the graphs in Figures 2 and 3 illustrate that the method of the present invention produces an economical method of operation while providing an acid derivative product of a quality that is comparable to or unacceptably different from similar products. They are manufactured commercially and marketed in our days. For the practical implementation of the method of the present invention, the highest permissible temperature, for example 148.88 ° C (300 ° F), for the first oxidation reactor, can be limited by the overall process configuration, in particular, the pressure of the second reactor and the heat balance. Once the temperature limit is set, one or more other variables can be adjusted to achieve a ventilation oxygen that is less than, or equal to about 2%. Examples include the following: the paraxylene in the feed, the amount of the added catalyst, the residence time, and the increased concentration of the catalyst in the reactor achieved by the reduction of the amount of the added solvent. In this regard, it is shown in Examples 8 and 9 that the ratio of oxygen to paraxylene (ox / PX) in the first reactor is too low (ie, too much paraxylene in the feed because the amount of oxygen is fixed by the volume fed thereto from the second reactor, and the desired volume to be ventilated from the first reactor), that the result is less desirable because a significant amount of the paraxylene is unreacted and the The color of the terephthalic acid product is increased (0.58 and 0.56, respectively). Similarly, since all of the effluent from the first reactor is passed directly into the second reactor, all of the catalyst added to the first reactor also ends up in the second reactor. Accordingly, the amount of catalyst that can be added is determined by the requirements for the catalyst in the second reactor as well and therefore, is not really a completely independent variable for the first reactor. In addition, for the same performance, a longer residence time can be achieved by the larger reactor volume, but that may not be desirable because this increases the cost. In addition, after the second reactor, most of the solvent is separated from the product and is recycled. This stream of recycled solvent (mother liquor) also contains the majority of the recycled catalyst. The amount of the free catalyst of the solvent that must be added to the reactor stages is not great. Consequently, it is impossible to reduce by much the amount of the solvent added to the first reactor, without also having to reduce the catalyst at the same time. In contrast, a very convenient to increase the concentration of the catalyst and the residence time in the first reactor at the same time, is to send all or some of the condensate reflux from the first reactor to the second reactor, omitting the passage through the first reactor. As a result, the amount of the solvent in the first reactor effluent is reduced, thus increasing the concentration of the catalyst and the residence time in the first reactor. Examples 2 and 6 indicate that this method is successful in reducing the oxidation temperature (141.66 and 137.22 ° C (287 ° F and 279 ° F)), respectively in the first reactor), and still achieve only about 1 per hundredth volume of oxygen in the vent gas from the first reactor. The color of the product by OD 340 (0.34 and 0.29) and the burned acetic acid (37 and 38 pounds / M pound PX) are both acceptable.

Table 1 Conditions of the First Reactor Table 1 (cont.) Table 2 Analysis of the First Reactor Ventilation Gas Table 3 Analysis of the First Reactor Suspension Table 4 Conditions of the Second Reactor Table 4 (Cont.) Conditions of the Second Reactor Table 5 Analysis of the Ventilation Gas of the Second Reactor Table 6 Analysis of the Second Reactor Suspension Table 7 TA Analysis of the Recovered Product Table 8 Reactor Conditions Table 9 Ventilation Gas Analysis Table 10 Analysis of Reactor Suspension Table 11 Product Analysis of Recovered TA From the foregoing description, it is evident that the goals of the present invention have been achieved although only certain modalities have been described, alternative modalities and various modifications of the foregoing description will be apparent to those skilled in the art. These and other alternatives are considered equivalent and within the spirit and scope of the present invention. It is noted that in relation to this date, the best method known to the applicant to carry out the aforementioned invention is that which is clear from the present description of the invention.

Claims (30)

1. CLAIMS Having described the invention as above, the content of the following claims is claimed as property. 1. A stepwise, continuous countercurrent process for the catalytic oxidation of at least one disubstituted benzene with oxidizable substituents selected from the class consisting of the alkyl, hydroxyalkyl, aldehyde, carboxy groups and mixtures thereof, up to their acid derivatives corresponding in a solvent system, characterized in that it comprises the steps of: (a) introducing into a first oxidation zone a feed mixture comprising at least a portion of the predetermined total amount of each of: (i) the solvent, (ii) catalyst components comprising at least one catalytic metal selected from the class consisting of manganese, cobalt, nickel zirconium, hafnium, cerium and mixtures thereof, and (iii) bromine in a molar ratio based on metals total catalysts within the range from about 1:20 to about 5: 1, and from about 7 to about 60 percent in weight of the predetermined total amount of at least one disubstituted benzene, introduced in steps (a) and (b); (b) partially oxidizing at least one disubstituted benzene in the first oxidation step in the presence of a molecular oxygen containing gas initially containing from about 3 to about 20 volume percent molecular oxygen, at a temperature in the range from about 121.11 ° C (250 ° F) to about 205 ° C (401 ° F) and with the relative amounts of the disubstituted benzene, the catalyst components, and the solvent, and with the temperature being from about 25 to about 99.95 percent by weight of the disubstituted benzene, fed, to the first oxidation stage, is oxidized to form a gaseous mixture comprising the unreacted molecular oxygen and the vaporized solvent and a first product mixture comprising the acid derivative product, disubstituted partially oxidized benzene and disubstituted benzene that did not react and the solvent, and at a pressure that is sufficient to keep the disubstituted benzene, disubstituted partially oxidized benzene, the product of the acid derivative and the solvent substantially as a liquid or a solid-liquid suspension, and in such a way that the concentration of residual molecular oxygen in the remaining gas mixture be from about 0.3 to about 2 volume percent; (c) recovering the first mixture of the product resulting from the first oxidation step, and feeding at least a portion of the first mixture of the recovered product to a second oxidation step; (d) feeding to the second oxidation stage a gas containing molecular oxygen and the remainder, if any, of the predetermined total amounts mentioned above, introduced in steps (a) and (d) of the disubstituted benzene, the components of the catalyst, the solvent and the bromine; (e) substantially completely oxidizing in the second oxidation step disubstituted partially oxidized benzene and unvaccinated disubstituted benzene, fed, to the second oxidation step with a molecular oxygen containing gas containing from about 15 to about 50 volume percent of molecular oxygen at a temperature in the range from about 175 ° C (347 ° F) to about 216.11 ° C (421 ° F), and with the relative amounts of disubstituted benzene, disubstituted partially oxidized benzene, components of the catalyst and the solvent, and with the temperature which is from about 96 to about 100 weight percent of the disubstituted benzene and the disubstituted partially oxidized benzene, are oxidized to form a gaseous mixture comprising the molecular oxygen which did not react and the vaporized solvent, and a second mixture of the product comprising the acid derivative and the solvent, and at a sufficiently high pressure to maintain the product of the acid derivative, disubstituted partially oxidized benzene and disubstituted benzene that did not react, substantially as a liquid or suspension of solid-liquid in such a way that the concentration of the residual molecular oxygen in the remaining gas mixture is in the range of from about 3 to about 15 volume percent; (f) recovering from the second oxidation step the second product mixture comprising the product of the acid derivative; and (g) extracting from the second oxidation stage and recycling it to the first oxidation stage, the gas containing residual molecular oxygen.
2. 2. The process according to claim 1, characterized in that the disubstituted benzene comprises the para-disubstituted benzene and the corresponding acid derivative is the terephthalic acid.
3. 3. The process according to claim 2, characterized in that the substituents in the para-disubstituted benzene are alkyl groups having from one to four carbon atoms.
4. 4. The process according to claim 1, characterized in that the solvent comprises an organic acid.
5. 5. The process according to claim 4, characterized in that the solvent comprises acetic acid.
6. 6. The process according to claim 1, characterized in that the catalytic metals are cobalt and manganese. The process according to claim 1, characterized in that the atomic ratio of manganese to cobalt in the reaction mixture in the first oxidation step is in the range from about 1: 100 to about 100: 1. The process according to claim 1, characterized in that the reaction mixture in the first oxidation step comprises a molar ratio of the total catalytic metals with respect to the bromine therein from about 1: 5 to about 2: 1. 9. The process according to claim 1, characterized in that the temperature in the first oxidation stage is maintained within the range from about 136.11 ° C (277 ° F) to about 177.22 ° C (351 ° F). The process according to claim 1, characterized in that the molecular oxygen-containing gas introduced in the first oxidation step contains from about 3 to about 11 volume percent of the molecular oxygen. The process according to claim 1, characterized in that the residual molecular oxygen concentration in the gas removed from the first oxidation stage is less than about 1 volume percent. 12. The process according to claim 1, characterized in that the degree of conversion of the disubstituted benzene to the disubstituted partially oxidized benzene and the acid derivative thereof, in the first oxidation step, is within the range of from about 60 to about 99.95 percent in weigh. The process according to claim 1, characterized in that the molecular oxygen-containing gas introduced in the second oxidation stage contains from about 20 to about 25 volume percent of the molecular oxygen. 14. The process according to claim 1, characterized in that the residual molecular oxygen concentration in the gas removed from the second oxidation stage is from about 3 to about 11 volume percent. 15. The process according to claim 14, characterized in that the concentration of residual molecular oxygen in the gas removed from the second stage of oxygenation is from about 3 to about 8 volume percent. 16. The process according to claim 1, characterized in that the degree of conversion of the disubstituted benzene and partially oxidized disubstituted benzene to the acid derivative thereof, in the second oxidation step, is within the range of from about 97 to about 100 percent in weigh. The process according to claim 6, characterized in that from about 20 to about 100 weight percent of the predetermined total amount, mentioned above, of the manganese added in steps (a) and (d), is added in the first stage of oxidation. The process according to claim 6, characterized in that from about 20 to about 100 weight percent of the total predetermined amount, mentioned above, of the added cobalt is added in steps (a) and (d) in the first step. oxidation stage. The process according to claim 1, characterized in that from about 20 to about 100 weight percent of the predetermined total amount, mentioned above, of the bromine, added in steps (a) and (d), is added in The first stage of oxidation. The process according to claim 1, characterized in that from about 15 to about 35 weight percent of the predetermined total amount, mentioned above, of the disubstituted benzene added in steps (a) and (d), is added in The first stage of oxidation. 21. The process according to claim 1, characterized in that from about 10 to about 100 weight percent of the aforementioned, total, predetermined amount of the solvent added in steps (a) and (d), is added in the first stage of oxidation. 22. The process according to claim 1, characterized in that the temperature in the first oxidation stage is at least -14.7 ° C (5.5 ° F) lower than the temperature in the second oxidation stage. 23. The process according to claim 1, characterized in that the gas removed from the first oxidation stage is partially condensed to remove the condensable solvent therefrom and at least a portion of the condensed solvent is introduced in the first oxidation stage, the second stage of oxidation, or both of these. 24. The process according to claim 23, characterized in that at least a portion of the condensed solvent is introduced into the first oxidation stage. 25. The process according to claim 23, characterized in that at least a portion of the condensed solvent is introduced into the second oxidation stage. 26. The process according to claim 23, characterized in that substantially all of the condensed solvent is introduced into the second oxidation step. 27. The process according to claim 1, characterized in that the second mixture of the product recovered from the second oxidation stage is subjected to oxidation conditions in a third oxidation step in order to oxidize oxidizable impurities therein. 28. The process according to claim 27, characterized in that the oxidized impurities are separated from the product of the acid derivative to thereby produce a product derived from the purified acid. 29. The first product mixture, characterized in that it is produced according to step (b) according to claim 1. 30. A composition for the product of the first oxidation step on a solvent free base as a percentage of the total product mixture, characterized in that it comprises from about 5.0 to about 85.0 weight percent of the terephthalic acid, from about 2.0 to about 20.0 weight percent of the 4-carboxybenzaldehyde, from about 0.0 to about 3.0 weight percent of the acid hydroxy ethylbenzoic acid, from about 5.0 to about 65 weight percent of p-toluic acid, from about 0.0 to about 30.0 weight percent of p-tolualdehyde, and from about 0.0 to about 35.0 weight percent of p-xylene.