CN103183607A - Method for producing mixed aromatic dioctyl phthalate - Google Patents
Method for producing mixed aromatic dioctyl phthalate Download PDFInfo
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Abstract
The invention relates to a method for producing mixed aromatic dioctyl phthalate, particularly to a method for producing mixed aromatic dicarboxylic acid through cooxidation of mixed xylenes, and belongs to the field of production of aromatic carboxylic acid. In the production process of a polyester chip, PTA (purified terephthalic acid) and PIA (isophthalic acid) at a certain ration are prepared, the addition of the PIA is to improve the flexibility and the degree of crosslinking of the polyester chip while a serious process problem existing in the current polyester production process is that the PTA and the PIA are hardly mixed very uniformly. However, when PX (p-xylene) and MX (m-xylene) are used to produce mixed carboxylic acid through a cooxidation method, the PTA and the PIA obtained are generated simultaneously in the oxidation reaction, molecules of different types in mixed components form a perfect crystal form according to a fixed order and are highly dispersed and fully mixed, so that the quality problem of a product brought about by the uneven mixing of the PTA and the PIA is effectively avoided. The mixed aromatic dioctyl phthalate produced through the method in the invention can form an ideal high polymer copolymerization configuration when polymerized with ethylene glycol in the polyester process, so that the performance of the polyester chip is improved.
Description
Technical field
The present invention relates to a kind of production method of mixing fragrant dioctyl phthalate, specifically relate to a kind of method of xylol co-oxidation production mixing aromatic dicarboxilic acid, it belongs to the aromatic carboxylic acid production field.
Background technology
Aromatic carboxylic acid is important organic compound and the raw material of production number of chemical product, and wherein that the market demand maximum is pure terephthalic acid (PTA), is used for the production of terylene, and the output of global terephthalic acid (TA) has reached 4,000 ten thousand tons/year; Next is m-phthalic acid (PIA), is used for polyester slice, and output is 100～2,000,000 tons/year.The production method of PTA and PIA is closely similar, all adopts the production of dimethylbenzene liquid-phase air oxidation, obtains terephthalic acid by p-Xylol (PX) oxidation, obtains m-phthalic acid by m-xylene (MX) oxidation.Its technology source is the patented technology US2833816 of Mid-Century company application in 1958, the main points of this patented technology are as reaction solvent with low-molecular-weight carboxylic acid (as acetic acid), cobalt, manganese, bromine compounds with solubility are catalyst system, make aromatic carboxylic acid with the air liquid-phase oxidation, the dimethylbenzene liquid phase catalytic oxidation is that the yield of carboxylic acid can reach 95%.The reaction conditions of PTA and PIA preparation is also about the same, and general temperature is 150～220 ℃, and pressure is 0.5 ~ 2.5MP, and specific embodiment and condition have detailed introduction in many patents and chemical industry document.
At present, the production of PTA and PIA all adopts highly purified raw material monomer (more than 99.5%) to obtain corresponding benzene dicarboxylic acid by the difference oxidation.The dimethylbenzene aromatic monomer is that carbon 8 aroamtic hydrocarbon raw materials that starting raw material obtains are a kind of mixtures that reach chemical equilibrium mainly by carbon 8 aromatic hydrocarbons productions from oil and coal, and m-xylene is maximum, accounts for 1/2, and p-Xylol and o-Xylol respectively account for 1/4.Because the isomer character of these three kinds of carbon 8 aromatic hydrocarbons approaches, separation difficulty, in order to obtain highly purified single component material, the cost of separation is just than higher.For example, existing carbon 8 aroamtic hydrocarbon raw material production technique all are with BTX aromatics condensation of gas liquefy laggard capable Crystallization Separation or fractionation by adsorption, tell p-Xylol, then will between, adjacent mixture heating up gasification, at high temperature carry out isomerization reaction, its part is converted into p-Xylol (per pass conversion about 20%), separates again, so move in circles.On the other hand, though adopting highly purified m-xylene is that raw material is produced m-phthalic acid, but the main application of a large amount of m-phthalic acids is to carry out copolymerization with terephthalic acid to produce polyester slice, and like this, the product m-phthalic acid mixes with terephthalic acid again during use.If can directly adopt xylol raw material rather than its high purity single component material to carry out oxidation, the mixed dicarboxylic acid that obtains no longer separates and directly copolymerization use, and then production technique just can be simplified, and production cost can significantly reduce.
In addition, from the required angle of polyester slice production, use the mixed carboxylic acid more to be conducive to eliminate because PTA and PIA mix the uneven product quality problem that brings.This be because, adopt PX and MX conjugated oxidation to produce the mixed carboxylic acid, resulting PTA and PIA generate in oxidizing reaction simultaneously, they form cocrystallization, the perfect cystal that constitutes according to fixing order between the dissimilar molecules, mixes fully at high dispersing between the molecule, so it is in the polyester process during with the ethylene glycol polymerization, can arrange out desirable polymer copolymerization configuration, the performance of polyester slice is improved.And tradition uses PTA and the pure material blend of PIA to be the method for making polyester of raw material, because PTA and PIA often can not fully mix, the resulting polyester macromolecule product often is the simple mixing of ethylene glycol terephthalate and ethylene m-phthalate polymkeric substance, and does not reach the degree of copolymerization.Therefore, exploitation PTA and PIA co-oxidation technology, not only lucrative aspect the conservation cost, itself is a kind of type material for preparing the high-performance polyester section especially.
Summary of the invention
In view of prior art exists in problem, the objective of the invention is to propose a kind of product of extract after separating with p-Xylol and C8 aromatic hydrocarbons is raw material, by the parallel-two-stage oxidizing reaction, produces the method for mixing fragrant dioctyl phthalate.
To achieve these goals, the technical solution adopted in the present invention is a kind of production method of mixing fragrant dioctyl phthalate, adopting the product of extract after separating of p-Xylol and C8 aromatic hydrocarbons is reaction raw materials, by two-stage oxidation reaction, makes the mixing aromatic dicarboxilic acid; The energy of shared solvent, catalyzer, tail gas and solvent dehydration, mother liquor purification, catalyst recovery system in the described reaction process, and the slurry treatment scheme of employing level Four crystallization, its step is as follows:
1) oxidizing reaction: p-Xylol feeds first oxidation reactor, contains the higher C8 aromatic hydrocarbons extract of m-xylene concentration and delivers to second oxidation reactor, and dimethylbenzene raw material, acetate solvate and catalyzer are fed in second oxidation reactor by pipeline; Acetate solvate and catalyzer are fed in first oxidation reactor by pipeline, air or oxygen rich gas feed first oxidation reactor and second oxidation reactor through inlet pipe after compression, and dimethylbenzene and air catalyzed oxidation under the effect of catalyzer generates pure mixed phthalic acid; The condition of described first oxidation reactor and second oxidation reactor is: pressure is 0.5～2.5Mpa, temperature is 150～220 ℃, water-content 5 ~ 15%(mass content), tail oxygen concentration is 3% ~ 6%(molar content), reaction times is 30 ~ 100min, and described catalyzer is cobalt manganese bromine catalyst, and wherein the total mass concentration of cobalt, manganese, three kinds of ions of bromine is 500～3000ppm, cobalt manganese atom ratio is 30～0.3, and the atomic ratio of cobalt and manganese total concn and bromide anion is 0.5～2.5; The energy that oxidation produces changes into the latent heat of solvent and takes away by tail gas, and the slurry of product solid is extracted out from the reaction bottom, to the slurry processing unit;
2) slurry is handled: it comprises solid-liquid separation, hydrofining and Crystallization Separation;
Described solid-liquid separation is passed through the crystallization of oxidizing and crystallizing device level Four with the slurry of first oxidation reactor and second oxidation reactor; The oxidizing and crystallizing device is made up of the full crystallizer that mixes continuously of 4 polyphones, the temperature of first step crystallizer is 190～200 ℃, the temperature of second stage crystallizer is 160～170 ℃, and the temperature of third stage crystallizer is 90～120 ℃, and the temperature of fourth stage crystallizer is 50～70 ℃; Described first step mould temperature is higher 10 ~ 15 ℃ than the first oxidizing reaction actuator temperature, and main purpose is to promote liquid phase dissolved in the m-phthalic acid, by the foreign pigment in the secondary oxidation reduction m-phthalic acid; Solvent flashing in the crystallizer of the second stage reclaims the solvent latent heat of vaporization; The control of third stage crystallizer reduces system temperature to normal pressure; Fourth stage crystallizer is reduced to lesser temps by vacuum-evaporation, can separate out from liquid phase crystallization fully to guarantee the higher m-phthalic acid of solubleness; The dissolubility data that the selection of Tc provides with reference to table 1;
The 4th crystallizer outlet slurry is by filtering unit, and a liquid part is drawn to catalyzer and residue treatment unit, and a part is returned first oxidation reactor; Solid is sent into refining workshop section then by drying treatment;
Described hydrofining will mix dioctyl phthalate and be dissolved in the water, and with hydrogen generation catalytic reduction reaction, 85 ~ 90% aldehyde will be converted into tolyl acid soluble in water in fixed-bed reactor; The hydrogenation reaction temperature is 250～300 ℃ of temperature, and 5～10Mpa pressure, catalyzer are palladium carbon, is immersed in fully that the residence time of slurry in reactor is 3～10min in the solution; By hydrogenation process, the carboxyl benzaldehyde of massfraction 99% is removed;
The slurry of described Crystallization Separation after with hydrofining feeds the refining crystallization device, crystallizer outlet slurry separates by solid-liquid separation system, the solid-liquid separation operation is with deionized water rinsing, product after the purification makes the mixing prod that high purity terephthalic acid (TA) and m-phthalic acid (IA) binary are formed by dry;
3) energy recovery: the tail gas of first oxidation reactor, second oxidation reactor, first crystallizer and the tail gas of second crystallizer are passed into energy-recuperation system, and the recovery system of energy comprises: multi-stage condensing cooling apparatus, high-pressure absorber and decompressor; Described multi-stage condensing cooling apparatus produces middle pressure steam, and middle pressure steam is as the thermal source of dehydration column reboiler or stripping tower, drying machine; Described high-pressure absorber utilizes pickling and washing to reclaim organic substance; Described decompressor pressure recovery can drive the compressor compresses air; Part phlegma turns back to first step oxidation reactor and the second oxidation state reactor, and a part of phlegma is drawn out of the solvent dehydration system that is passed into, and the water that reaction produces is extracted out;
4) solvent dehydration and catalyst recovery system: the phlegma of extracting out from energy-recuperation system is passed into the solvent dehydration system, and the acetic acid steam that extract out at the second crystallizer top is passed into the solvent dehydration system; High density acetic acid after the dehydration is as the source of the solvent of second oxidation reactor production or the washings of high-pressure absorber and slurry treatment system; A part of mother liquor that the slurry treatment system obtains turns back to first oxidation reactor as circulating solvent; Part mother liquor is extracted out and is carried out mother liquor purification and catalyst recovery, to reduce the content while recovery part catalyzer of impurity in the circulating mother liquor; The steam stripped method of evaporation is adopted in the recovery of acetic acid in the mother liquor, and recovered solvent and catalyzer turn back to first oxidation reactor.
Described reaction raw materials derives from the product of C8 aromatic hydrocarbons extract after separating, and perhaps passes through massfraction〉99.5% p-Xylol and m-xylene be 20:1～3:1 preparation in proportion; Preferred ratio is 10:1～5:1.
Described catalyzer adopts the mixture of Cobaltous diacetate, manganese acetate and hydrogen bromide; Described oxidation reactor adopts selection to have the tank reactor of stirring rake or has the bubbling column reactor of gas distributor; Be preferably stirred-tank reactor, reaction time is 50～120 minutes.
Described oxidizing and crystallizing device is made up of the stirring tank of four series connection, and first step mould temperature is 180～190 ℃, feeds part oxygen and carries out deep oxidation to first crystallizer; The temperature of second stage crystallizer is 110～150 ℃; The temperature of third stage crystallizer is reduced to 20～50 ℃.
Described hydrogenator is continuous plug flow fixed-bed reactor, and an opening for feed is set, and hydrogen enters first oxidation reactor from spout, and the bottom catalyzer of first oxidation reactor supports by spiral filtering net.
Feasibility of the present invention and advantage are: in the polyester slice production process, need a certain proportion of PTA of preparation and PIA, the adding of PIA mainly is for the pliability of improving polyester slice and degree of crosslinking, and the mixing that the serious technological problems that existing polyester production process exists is PTA and PIA is difficult to reach degree very uniformly.And adopt PX and MX conjugated oxidation to produce the mixed carboxylic acid, resulting PTA and PIA generate in oxidizing reaction simultaneously, constitute perfect crystal formation according to fixing order between the dissimilar molecule in the blending ingredients, high dispersing, mixing fully between the molecule, this has just effectively been avoided because PTA and PIA mix the uneven product quality problem that brings.The mixing aromatic dicarboxilic acid of being produced by the present invention during with the ethylene glycol polymerization, can be arranged out desirable polymer copolymerization configuration in the polyester process, the performance of polyester slice is improved.
Beneficial effect of the present invention is: adopt the inventive method can the cheap xylol raw material direct production of applied cost to be applied to the benzene dicarboxylic acid of polyester industrial, reduced dimethylbenzene raw material separation costs, obtain the mixing aromatic dicarboxilic acid of high dispersive simultaneously, promote the quality of product, improved the benzene dicarboxylic acid competitiveness of product in market.
Description of drawings
Fig. 1 is unit module and the process task figure of pure mixed phthalic acid generative process;
Fig. 2 is process flow sheet of the present invention;
Among the figure: 100, first oxidation reactor, 200, second oxidation reactor, 300, crystallizer and solid-liquid separation system, 400, energy-recuperation system, 500, dehydration and catalyst recovery system, 301, first step crystallizer, 302, second stage crystallizer, 303, third stage crystallizer, the 304, the 4th crystallizer, 305, filter, 401, first condenser, 402, second condenser, 501, dehydration tower, 520, catalyzer and residue treatment unit.
Embodiment
Mix the production method of fragrant dioctyl phthalate in order further to understand this, accompanying drawings is as follows.Need to prove that method provided by the invention is not limited to flow process configuration and the processing condition that provide among embodiment and the embodiment, any local improvement to these methods can not change feature of the present invention yet.
The unit module of the pure mixed phthalic acid generative process that the present invention provides and process task figure, production process comprises five formants: first oxidation reactor 100, second oxidation reactor 200, crystallizer and solid-liquid separation system 300, energy-recuperation system 400, dehydration and catalyst recovery system 500.Mother liquor, catalyzer and air that p-Xylol and slurry treating processes produce add first oxidation reactor 100 by filling tube; The product (mainly contain p-Xylol and m-xylene) of the extract of C8 aromatic hydrocarbons after separating adds second oxidation reactor 200 with fresh acetic acid, catalyzer and air that the dehydration tower bottom obtains by filling tube; First oxidation reactor 100 and second oxidation reactor, 200 tail gas merge, export by offgas duct, through most of reactor that refluxes back after the condenser condenses, quantity of reflux is distributed by the oxidation load of first oxidation reactor 100 and second oxidation reactor 200, condensed tail gas is sent into follow-up cell processing, and a little phlegma is delivered to dehydration tower 501 rectifying separation.The slurry that comes out from first oxidation reactor 100 and second reactor 200 merges outlet line to be introduced in the crystallizer; First step crystallizer 301 temperature are higher 10 ~ 15 ℃ than the first oxidizing reaction actuator temperature, and main purpose is to promote liquid phase dissolved in the m-phthalic acid, by the foreign pigment in the secondary oxidation reduction m-phthalic acid.The temperature and pressure of follow-up level Four crystallizer reduces step by step in order to make separates out, two kinds of dicarboxylic acid sufficient crystallisings.The slurry of crystallizer output makes fragrant dioctyl phthalate by filtration, drying.The mother liquor that filters returns first oxidation reactor 100.M-phthalic acid content is 5%～30% in the mixing dioctyl phthalate that said process obtains, and the content of optimization is 10%～15%.
The technical process that the present invention provides is as follows:
Filling tube is to first oxidation reactor 100 dimethylbenzene of supplying raw materials, and p-Xylol concentration (mass concentration) is higher than 99%.Simultaneously, pipeline provides circulating mother liquor, contains m-phthalic acid that a small amount of not crystallization separates out and a large amount of solvent acetic acid in the mother liquor.Air compressor is compressed to 1.2～2.0MPa(than the high 0.2MPa of pressure in first oxidation reactor 100 with atmospheric air), feed in first oxidation reactor 100 through pipeline, the oxygen in the high-pressure air is as the oxygenant of reaction.Temperature of reaction control is 185～200
oC, pressure are 1.0～1.8MPa.Dimethylbenzene generates solid through liquid-phase oxidation and mixes fragrant phthalic acid in reactor, the slurry residence time is 30～80 minutes.Catalyst cobalt, manganese metal and bromine catalyst are sent into by pipeline, catalyzer adopts the industrial cobalt manganese bromine catalyst system of generally using, the mass concentration of cobalt is 100 ~ 900ppm, and manganese cobalt mass ratio is 0.5 ~ 1.5, and the mass ratio of promotor bromine and cobalt manganese is 0.6 ~ 1.2.The content of tolyl acid is 2000 ~ 4000ppm in the pure mixed phthalic acid solid of gained, and particle diameter is 90 ~ 150 μ m.The slurry that contains mixed carboxylic acid's solid of gained is extracted out from first oxidation reactor, 100 bottoms, and the lot of energy that p xylene oxidation produces changes into the latent heat of solvent and takes away by tail gas, enters energy-recuperation system through pipeline.
The product of the extract of C8 aromatic hydrocarbons after separating enters second oxidation reactor 200 by pipeline, and pipeline is passed into second oxidation reactor 200 with part tail gas phlegma simultaneously.High-pressure air or oxygen rich gas are sent to first oxidation reactor 100 by pipeline, as the oxygenant that mixes aromatic oxidation.Catalyzer adopts the industrial cobalt manganese bromine catalyst system of generally using, and the mass concentration of cobalt is 100 ~ 900ppm, and manganese cobalt mass ratio is 0.5 ~ 1.5, and the mass ratio of promotor bromine and cobalt manganese is 0.6 ~ 1.2.The peak optimization reaction temperature of second oxidation reactor 200 is 190 ~ 200 ℃, and pressure is 1200 ~ 1600kPa, water-content 5 ~ 10%(mass content), tail oxygen concentration is 3% ~ 6%(molar content), the residence time is 50 ~ 100min.Liquid-phase oxidation takes place and generates the pure mixed phthalic acid solid in xylol in first oxidation reactor 100, the content of m-phthalic acid is 60 ~ 70% in the solid product of gained, and particle diameter is 60 ~ 80 μ m.The gained slurry of solids is extracted out from second oxidation reactor, 200 bottoms, and the lot of energy that oxidation produces changes into the latent heat of solvent and takes away by tail gas, and the tail gas merging with first oxidation reactor 100 enters energy-recuperation system 400 through pipeline.
Combine from the slurry of first oxidation reactor 100 and the extraction of second oxidation reactor, 200 bottoms, send into the level Four crystallizer through pipeline.First step crystallizer 301 temperature are higher 10 ~ 15 ℃ than first oxidation reactor, 100 temperature, can provide the high temperature acetic acid steam of heat to be sent to first step crystallizer 301 by pipeline, high temperature acetic acid steam can be obtained through heating by the high density acetic acid of solvent dehydration system, high-pressure air or oxygen rich gas enter first step crystallizer 301 through inlet pipe, mix the aromatic carboxylic acid solid particulate deep oxidation at high temperature takes place.The pressure of first step crystallizer 301 is 1300 ~ 2000kPa, and temperature is 190～220 ℃, and the residence time is 20 ~ 40min.The slurry that comes out from first crystallizer 301 is sent to second stage crystallizer 302 through pipeline.The second further decompression cooling of crystallizer 302, the steam that produces is sent into dehydration tower 501 bottoms by pipeline, provide dehydration required portion of energy, most TA and most IA separate out from liquid phase crystallization in second crystallizer 302, the temperature of second crystallizer 302 is 160 ~ 170 ℃, pressure is 400 ~ 600kPa, and the residence time is 20 ~ 40min.Third stage crystallizer 303 further reduces pressure to normal pressure, and temperature is controlled at 90～120 ℃.Fourth stage crystallizer is by vacuum-evaporation, and pressure is gauge pressure-0.05MPa, is reduced to 50～70 ℃ of temperature, can separate out from liquid phase crystallization fully to guarantee the higher m-phthalic acid of solubleness.
The slurry that the 4th crystallizer 304 flows out is delivered to filter 305 by pipeline and is separated mother liquor and solid, simultaneously, partly is transported to 305 pairs of solids of oxidizing reaction filter by the refining acetic acid in dehydration tower 501 bottoms and washs.Oxidizing reaction filter mother liquor part (90～95%) is back to first oxidation reactor 100 by pipeline, partly (5～10%) are extracted out and are carried out mother liquor purification and catalyst recovery, the acetic acid that reclaims is back to first oxidation reactor 100, solid behind the filtration washing is delivered to drying machine through pipeline, obtains mixing the aromatic dicarboxilic acid product after the drying.
The tail gas of first oxidation reactor 100, second oxidation reactor 200 and first step crystallizer 301 is passed into energy-recuperation system 400, the recycling of energy adopts the multi-stage condensing cooling apparatus to produce middle pressure steam, middle pressure steam is realized recycle as the thermal source of dehydration column reboiler or stripping tower, drying machine.The phlegma of part turns back to first oxidation reactor 100 and second oxidation reactor 200 through pipeline and pipeline, and a part of phlegma is passed into the solvent dehydration system through pipeline, and the acetic acid steam of extracting out from the crystallizer top also is passed into the solvent dehydration system through pipeline.High density acetic acid after the dehydration is as the solvent of second oxidation reactor 200 source, or as high temperature acetic acid steam, filtration procedure washings, and the source of residue treatment unit washings.The water vapour lease making pipeline of extracting out from the solvent dehydration system drains into Sewage treatment systems.
The Extract of first condenser 401 and second condenser 402 is delivered to dehydration tower through pipeline, carries out rectifying and dewatering.Simultaneously, the steam that is produced by second crystallizer 302 enters from dehydration tower 501 bottoms, provides dehydration required portion of energy, and its complementary energy is provided by dehydration tower 501 reboiler interchanger.Dehydration tower 501 adopts azeotropic distillation to strengthen, and entrainer is N-BUTYL ACETATE or propyl acetate.Entrainer and water add from cat head, and the vapor phase product after the rectifying is rich in water and entrainer, flow out from cat head, through condenser condenses, are separated into water and oil phase in water-and-oil separator.Oil phase is back to rectifying tower, and water is then discharged system.Separate via water-and-oil separator, entrainer is recycled.The acetate concentration massfraction of cat head drainage water is less than 0.3%, and the acid concentration of the low acetic acid that concentrates of tower is greater than 90%.
Embodiment 1
The method and apparatus that adopts the present invention to provide is produced 600000 tons mixing fragrance dioctyl phthalate flow process and device design per year, 7600 hours production times of year, the processing condition of each unit equipment determine that according to material consumption and energy consumption minimized principle the energy between each stream thigh is flux matched to be optimized according to the integrated principle of system capacity.Gained related process parameter is enumerated as follows according to sequence of unit:
Under these conditions, the index of oxidation reactor 100 output slurries is listed in table 3.
Annotate: PX---terephthalic acid, MX---m-phthalic acid; 3-CBA---3-carboxyl benzaldehyde, 4-CBA---4-carboxyl benzaldehyde.
Under these conditions, the index of oxidation reactor 200 output slurries is listed in table 5.
The crystallizer parameter is enumerated as follows
Material and the energy balance data of 600,000 tons mixing fragrance dioctyl phthalate Production Flow Chart are listed in table 10.
Claims (9)
1. production method of mixing fragrant dioctyl phthalate, adopting the extract of p-Xylol and C8 aromatic hydrocarbons is reaction raw materials through separating after product, by two-stage oxidation reaction, makes the mixing aromatic dicarboxilic acid; Its step is as follows:
1) oxidizing reaction: p-Xylol feeds first oxidation reactor, contains the higher C8 aromatic hydrocarbons extract of m-xylene concentration and delivers to second oxidation reactor, and dimethylbenzene raw material, acetate solvate and catalyzer are fed in second oxidation reactor by pipeline; Acetate solvate and catalyzer are fed in first oxidation reactor by pipeline, air or oxygen rich gas feed first oxidation reactor and second oxidation reactor through inlet pipe after compression, and dimethylbenzene and air catalyzed oxidation under the effect of catalyzer generates pure mixed phthalic acid; The energy that oxidation produces changes into the latent heat of solvent and takes away by tail gas, and the slurry of product solid is extracted out from the reaction bottom, to the slurry processing unit;
2) slurry is handled: it comprises solid-liquid separation, hydrofining and Crystallization Separation;
Described solid-liquid separation is passed through the crystallization of oxidizing and crystallizing device level Four with the slurry of oxidation reactor; The oxidizing and crystallizing device is made up of the full crystallizer that mixes continuously of 4 polyphones, the temperature of first step crystallizer is 190～200 ℃, the temperature of second stage crystallizer is 160～170 ℃, and the temperature of third stage crystallizer is 90～120 ℃, and the temperature of fourth stage crystallizer is 50～70 ℃; The 4th crystallizer outlet slurry is by filtering unit, and liquid portion is drawn to catalyzer and residue treatment unit, and part is returned first oxidation reactor (100); Solid is sent into refining workshop section then by drying treatment;
Described hydrofining will mix dioctyl phthalate and be dissolved in the water, and with hydrogen generation catalytic reduction reaction, partly the aldehyde of (90 ~ 95%) is converted into tolyl acid soluble in water in fixed-bed reactor; The hydrogenation reaction temperature is 250～300 ℃ of temperature, and 5～10Mpa pressure, catalyzer are palladium carbon, is immersed in fully that the residence time of slurry in reactor is 3～10min in the solution; By hydrogenation process, the carboxyl benzaldehyde of massfraction 99% is removed;
The slurry of described Crystallization Separation after with hydrofining feeds the refining crystallization device, crystallizer is discharged slurry and is separated by solid-liquid separation system, the solid-liquid separation operation is with deionized water rinsing, product after the purification makes the mixing prod that high purity terephthalic acid (TA) and m-phthalic acid (IA) binary are formed by dry;
3) energy recovery: the tail gas of first oxidation reactor, second oxidation reactor, first crystallizer and second crystallizer is passed into energy-recuperation system, and energy-recuperation system comprises: multi-stage condensing cooling apparatus, high-pressure absorber and decompressor; Described multi-stage condensing cooling apparatus produces middle pressure steam, and middle pressure steam is as the thermal source of dehydration column reboiler or stripping tower, drying machine; Described high-pressure absorber utilizes pickling and washing to reclaim organic substance; Described decompressor pressure recovery can drive the compressor compresses air; The phlegma of a part turns back to first oxidation reactor and second oxidation reactor, and a part of phlegma is drawn out of the solvent dehydration system that is passed into, and the water that reaction produces is extracted out;
4) solvent dehydration and catalyst recovery system: the phlegma of extracting out from energy-recuperation system is passed into the solvent dehydration system, and the acetic acid steam that extract out at the second crystallizer top is passed into the solvent dehydration system; High density acetic acid after the dehydration is solvent source or high-pressure absorber and the slurry treatment system washings that reactor is produced; The partial mother liquid that the slurry treatment system obtains turns back to reactor as circulating solvent; Partial mother liquid is extracted out and is carried out mother liquor purification and catalyst recovery; The steam stripped method of evaporation is adopted in the recovery of acetic acid in the mother liquor, and recovered solvent and catalyzer turn back to reactor.
2. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: described reaction raw materials derives from the product of C8 aromatic hydrocarbons extract after separating, and perhaps passes through massfraction〉99.5% p-Xylol and m-xylene be 20:1～3:1 preparation in proportion; Preferred ratio is 10:1～5:1.
3. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: the condition of described first oxidation reactor and second oxidation reactor is: pressure is 0.5～2.5Mpa, temperature is 150～220 ℃, water-content 5 ~ 15%(mass content), tail oxygen concentration is 3% ~ 6%(molar content), the reaction times is 30 ~ 100min.
4. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: described catalyzer is cobalt manganese bromine catalyst, the total mass concentration of catalyst cobalt, manganese, three kinds of ions of bromine is 500～3000ppm, cobalt manganese atom ratio is 30～0.3, and the atomic ratio of cobalt and manganese total concn and bromide anion is 0.5～2.5.
5. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: first step mould temperature is higher 10 ~ 15 ℃ than oxidizing reaction actuator temperature, and the solvent latent heat of vaporization is reclaimed in a large amount of flash distillations of solvent in the crystallizer of the second stage; The control of third stage crystallizer reduces system temperature to normal pressure; Fourth stage crystallizer is by vacuum-evaporation.
6. the production method of the fragrant dioctyl phthalate of mixing according to claim 1 is characterized in that: the mixture of described catalyzer employing Cobaltous diacetate, manganese acetate and hydrogen bromide; Described oxidation reactor adopts selection to have the tank reactor of stirring rake or has the bubbling column reactor of gas distributor; Reaction time is 50～120 minutes.
7. the production method of the fragrant dioctyl phthalate of mixing according to claim 6 is characterized in that: described oxidation reactor employing stirred-tank reactor.
8. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: described oxidizing and crystallizing device is made up of the stirring tank of four series connection, first step mould temperature is 180～190 ℃, feeds part oxygen and carries out deep oxidation to first crystallizer; The temperature of second stage crystallizer is 110～150 ℃; The temperature of third stage crystallizer is reduced to 20～50 ℃.
9. the production method of the fragrant dioctyl phthalate of mixing according to claim 1, it is characterized in that: described hydrogenator is continuous plug flow fixed-bed reactor, individual opening for feed is arranged, and hydrogen enters reactor from a spout, and the bottom catalyzer of reactor is supported by spiral filtering net.
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| CN106565457A (en) * | 2015-10-09 | 2017-04-19 | 中国石化仪征化纤有限责任公司 | Method for preparing two products including polymer grade terephthalic acid and polymer grade isophthalic acid by taking mixed aromatics as raw material |
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| Publication number | Priority date | Publication date | Assignee | Title |
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| CN107963964A (en) * | 2017-11-02 | 2018-04-27 | 华东理工大学 | The preparation method of M-phthalic acid |
| CN108003006A (en) * | 2017-11-17 | 2018-05-08 | 杭州多向流化学科技有限公司 | A kind of method that terephthalic acid (TPA) and M-phthalic acid are produced using C8 aronmatic as raw material |
| CN108299182A (en) * | 2018-02-02 | 2018-07-20 | 中国石油化工股份有限公司 | A kind of method that meta-xylene and paraxylene mixed oxidization prepare phthalic acid |
| CN111389314B (en) * | 2020-03-31 | 2020-12-11 | 南京延长反应技术研究院有限公司 | Built-in micro-interface unit enhanced reaction system and process for producing PTA (purified terephthalic acid) by PX (para-xylene) |
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| US2853514A (en) * | 1955-07-22 | 1958-09-23 | Olin Mathieson | Process for the preparation of aromatic carboxylic acids |
| CN1665770A (en) * | 2002-06-27 | 2005-09-07 | 欧洲技术承包与工程股份公司 | Method for separating water generated in production of aromatic polycarboxylic acid by aromatic hydrocarbon catalytic oxidation |
| CN1962598A (en) * | 2006-11-23 | 2007-05-16 | 浙江大学 | Method for producing benzene dicarboxylic acid by mixed dimethylbenzene cooxidation |
| CN101318890A (en) * | 2008-07-15 | 2008-12-10 | 中国石油化工股份有限公司 | Method for preparation of mixed benzol dimethanoic acid and benzene acetic acid |
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Patent Citations (4)
| Publication number | Priority date | Publication date | Assignee | Title |
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| US2853514A (en) * | 1955-07-22 | 1958-09-23 | Olin Mathieson | Process for the preparation of aromatic carboxylic acids |
| CN1665770A (en) * | 2002-06-27 | 2005-09-07 | 欧洲技术承包与工程股份公司 | Method for separating water generated in production of aromatic polycarboxylic acid by aromatic hydrocarbon catalytic oxidation |
| CN1962598A (en) * | 2006-11-23 | 2007-05-16 | 浙江大学 | Method for producing benzene dicarboxylic acid by mixed dimethylbenzene cooxidation |
| CN101318890A (en) * | 2008-07-15 | 2008-12-10 | 中国石油化工股份有限公司 | Method for preparation of mixed benzol dimethanoic acid and benzene acetic acid |
Cited By (1)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN106565457A (en) * | 2015-10-09 | 2017-04-19 | 中国石化仪征化纤有限责任公司 | Method for preparing two products including polymer grade terephthalic acid and polymer grade isophthalic acid by taking mixed aromatics as raw material |
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| CN102584572A (en) | 2012-07-18 |
| CN103183607B (en) | 2016-06-15 |
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