CN105001073A - PTA production method with low energy consumption, low material consumption andlow refuse discharge - Google Patents

PTA production method with low energy consumption, low material consumption andlow refuse discharge Download PDF

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CN105001073A
CN105001073A CN201510380472.9A CN201510380472A CN105001073A CN 105001073 A CN105001073 A CN 105001073A CN 201510380472 A CN201510380472 A CN 201510380472A CN 105001073 A CN105001073 A CN 105001073A
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tower
acetic acid
extraction agent
outlet
extraction
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CN105001073B (en
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周华堂
许贤文
李利军
王新兰
刘凤
孙爱军
谢萍
刘晓玲
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China National Petroleum Corp
China Kunlun Contracting and Engineering Corp
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China National Petroleum Corp
China Kunlun Contracting and Engineering Corp
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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00—Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/16—Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation
    • C07C51/21—Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen
    • C07C51/255—Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen of compounds containing six-membered aromatic rings without ring-splitting
    • C07C51/265—Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen of compounds containing six-membered aromatic rings without ring-splitting having alkyl side chains which are oxidised to carboxyl groups
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00—Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/42—Separation; Purification; Stabilisation; Use of additives

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  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Engineering & Computer Science (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

The invention relates to a Purified Terephthalic Acidproduction method with low energy consumption, low material consumption and low refuse discharge. The PTA production method comprises the following steps: tail gas of an oxidation reactor is outputted and passes through by-product steam of a first stage steam generator and is sent to an acetic acid regeneration tower for recycling acetic acid; a refined mother liquor is returned to the acetic acid regeneration tower, the original vacuum filtration filter is replaced by a pressure filter, and filter cakes are directly pulped and supplied to a refining section; through recovered acetic acid recovery extraction, an oxidized mother liquor is processed by extraction, and an acetic acid methyl ester removed tower top backflow liquid of the acetic acid regeneration tower is used for oxidizing, washing filter cakes of the refining section and blending slurries of the refining section. The method and system provided by the invention can effectively control acetic acid consumption in tail gas and waste water, recycle organic matters of the refined mother liquid, realize reuse of water, reduce consumption of steam and electric energy, reduce refuse discharge, and also can supply benzoic acid, mixed phthalic acids and other by-products. The invention has the advantages of short process flow, good operationality, low production cost and low apparatus investment.

Description

The PTA production method of less energy-consumption material consumption, few waste discharge and system
Technical field
The present invention relates to a kind of PTA production method, compare traditional its raw material of PTA production method, process solvent, catalyzer and steam consumption and all significantly reduce, is the novel PTA production technique of the low material consumption of a kind of less energy-consumption.
Background technology
Pure terephthalic acid (Purified Terephthalic Acid, be called for short PTA) be the main raw material producing trevira and resin, being nonpoisonous and tasteless white crystal or powder under normal temperature, is important large Chemicals, closely bound up with the daily life of the people.
PTA mainly adopts p-Xylol (p-Xylene at present, be called for short PX) air oxidation process production, technological process mainly comprises oxidation and refining Liang Ge workshop section, and forming a whole set of process system with process matched therewiths such as oxidation mother liquor process and oxidized tail gas process, Fig. 2 is the schema of an embodiment of conventional P TA production technique.
(1) be oxidized
Raw material PX300 is that solvent generates terephthalic acid TA with oxygen generation oxidizing reaction in air 301 under cobalt manganese bromine catalyst 302 acts on acetic acid in oxidation reactor 201, and most of terephthalic acid crystallizes out formation slurry in the reactor; Slurry is sent into oxidation filter 205(and is generally rotary vaccum filter after oxidizing and crystallizing device 202-204 step by step decrease temperature and pressure flash distillation and crystallization), slurry completes filtration, filter cake washing, blow dried totally three processes successively in filter: first filter, by most of liquid phase and TA crystal separation, rate of cake moisture content controls at 28wt%, filtrated stock is the acetic acid of 91%, except small part delivers to mother liquid disposal subsystem, most of mother liquor is back to oxidation reactor; Then wash filter cake, washing process adopts and washs containing the acetum of 10% water together; Finally carry out blow dried, dry gas is dried oxidized tail gas, intermediates crude terephthalic acid (being called for short CTA) is obtained after filter filter cake drying machine 206 drying out, deliver to intermediate bunker 207 with blast system, filter 205 filter obtain oxidation mother liquor major part Returning oxidation reactor, small part extract out go to carry out oxidation mother liquor process.
(2) refining
Refining slurry blend tank 208, carry out making beating from intermediate bunker 207 out crude terephthalic acid powder and hot water and obtain slurry, slurry is through six grades of refining paste heater 209-214 stepped heatings, slurry through heating carries out hydrogenation reaction with hydrogen 303 in hydrogenator 215, catalyst for hydrogenation is palladium charcoal solid particulate, hydrogenation reaction makes the impurity p-carboxybenzaldehyde in CTA (being called for short 4-CBA) react with hydrogen and is reduced to PT acid (i.e. p-methylbenzoic acid more soluble in water, also be another kind of impurity component in CTA), the discharging of hydrogenator 215 carries out decrease temperature and pressure flash distillation through refining crystallization device 216-220, crystallization, refining crystallization device 220 bottom discharge is delivered to refining filter-press 221 and is carried out solid-liquid separation, filtration successively in refining filter-press 221, filter cake washing and gas drying obtain filter cake and refinement mother liquor (being made up of filter liquor and wash filtrate), filter cake becomes PTA product 304 after PTA drying machine 222 drying, refinement mother liquor is successively through flash distillation, cooling, filtration treatment, filtrate is directly disposed to sewage plant with the form of refined wastewater.
(3) solvent recuperation
From oxidation reactor 201 top, oxidized tail gas is out successively through vapour generator 231-234 condensation and byproduct water steam, and then dereaction device return tank 236 carries out gas-liquid separation after de-mineralized water water cooler 235 continuation cooling, return tank outlet gas phase sends to oxidized tail gas process (entering high-pressure absorber 238) after water cooler 237 cools, return tank 236 exports liquid phase (water-containing acetic acid) part and goes dehydration tower 251(entrainer to be n-butyl acetate) carry out azeotropic distillation dehydration, all the other return oxidation reactor 201; Dehydration tower 251 outlet at bottom liquid phase is back to oxidation reactor 201 on a small quantity, main as device spray acid, and a part is as the flushing acetic acid 306 of whole system, and all the other deliver to high-pressure absorber 238 as washing acetic acid through condenser 252; Dehydration tower 251 top gaseous phase realizes being separated of oil phase and aqueous phase after water cooler 253 cools in decanting vessel 254, oil phase is back to dehydration tower and recycles, aqueous phase sends into entrainer recovery tower 255, ritalin is recycled to oxidation reactor 201 from the discharge of entrainer recovery tower 255 tower top through supercooler 256 cooling by stripping, in the middle part of tower, decanting vessel 254 is delivered in entrainer extraction, and tower bottoms delivers to sewage works 257 as waste water discharging device.
(4) oxidized tail gas process
The object of oxidized tail gas process carries out purification utilization to oxidized tail gas, treating processes is: the oxidized tail gas from water cooler 237 enters at the bottom of high-pressure absorber 238 tower, adopt in the middle part of high-pressure absorber bottom dehydration tower 251 through the cooled dense acetic acid spray of water cooler 252, top adopts de-mineralized water 305 to spray, the high-pressure absorber top exit gas phase gas incineration system 239 that truncates purifies, tail gas after purification goes off-gas expander 240 to reduce pressure acting, and the tail gas gas wash tower 241 that truncates after decompression acting is finally disposed in air after alkali cleaning.
(5) oxidation mother liquor process
Oxidation mother liquor is first through peroxyacetic acid stripping tower 261 concentrate recovery of acetic acid simultaneously, the discharging of acetic acid stripping tower bottom is mother liquor residue slurry, mother liquor residue slurry enters thin-film evaporator 262, recovery of acetic acid is evaporated further in thin-film evaporator, device bottom is discharged mother liquor residue and is entered catalyzer water extractor 263, in catalyzer water extractor 263, add water to stir simultaneously, water carrys out fresh food frozen water 307, the discharging of catalyzer water extractor goes residue whizzer 264 to isolate oxidation residua 308 to transport outward, liquid phase enters after concentrated solution well heater 265 heats and delivers to catalyst recovery tank 266, in catalyst recovery tank, be back to oxidation reactor 201 after making beating recycle.
Traditional PTA technological process of production length, energy consumption and material consumption are high.To produce 1200000 tons of PTA devices per year, PTA products material PX unit consumption 655-657kg/tPTA per ton, acetic acid unit consumption 38-40kg/tPTA, steam mono-consumption 0.7-1.0t/tPTA, de-mineralized water unit consumption 2.2-3.0t/tPTA, especially the pond waste material that detention tank is pulled out just has 2000-4000 ton/year, amounts to PX unit consumption 1.2-2.4kg/tPTA, and full factory sewage quantity reaches 3,600,000 tons/year, COD discharge reaches 1.5 ten thousand tons/year, and amounting to ton products C OD is 12.5kg/tPTA.
The disadvantage of conventional P TA technique is in particular in following five aspects:
(1) oxidation section adopts vacuum filtration+drying mode, and energy consumption is high.Oxidation subsystem adopts rotary vaccum filter to filter and acetic acid mode of washing, and filter discharging adopts drying machine dry with recovery of acetic acid, and the CTA obtained after drying delivers to intermediate bunker by wind and stores.Like this, for the vacuum system that filter is supporting needs power consumption 8-12KW/tPTA, drying machine consumes more than 5barG pressure steam 0.1-0.2t/tPTA, and blast system consumes high pressure exhaust gas thus reduces tail gas expansion work and is about 5KW/tPTA.
(2) dewatering system energy consumption is high.Add water for what remove the generation water of oxidizing reaction and high-pressure absorber washing tail gas, need to arrange acetic acid dehydration system.But because the relative volatility of acetic acid and water is close to 1, therefore the general azeotropic distillation that adopts dewaters, and remove water per ton and need consume water vapor 1.4-2.0t, steam energy consumption is high, causes plant energy consumption higher.
(3) oxidation mother liquor dedoping step energy consumption is high.For control PTA quality product, extraction section oxidation mother liquor is needed to go mother liquid disposal to remove the impurity in mother liquor, this mother liquor is first through acetic acid gas stripping column concentrate, again through the further recovery of acetic acid of thin-film evaporator, thin-film evaporator need consume extraneous high pressure steam to heat, and adds the steam consumption 0.03t/tPTA of device.
(4) oxidation residua does not effectively utilize.Acetic acid in thin-film evaporator blanking residue and catalyzer do not reclaim, and cause acetic acid consumption 0.5-1.0kg/tPTA, cobalt to consume 0.01kg/tPTA, manganese consumes 0.02kg/tPTA; Thin-film evaporator blanking residue is directly outer after water making beating and cooling to be arranged, phenylformic acid wherein, terephthalic acid, m-phthalic acid, cobalt manganese are not all recycled, conversion PX unit consumption is about 2-3kg/tPTA, cobalt unit consumption 0.03kg//tPTA, manganese unit consumption 0.03kg//tPTA, acetic acid unit consumption 2kg/tPTA.
(5) refinement mother liquor does not obtain effective recycling.Refinement mother liquor usually adopt first flash distillation be cooled to 101 DEG C, be again cooled to about 45 DEG C, the scheme of then filtering processes, filtrate is directly disposed to sewage plant with the form of refined wastewater, the refined wastewater of discharge amounts to 2-3t/tPTA, simultaneously in water also containing 300-600ppm(mass concentration) acetic acid, the PT acid of 300-600ppm, 4-CBA and 200-500ppmw of 10ppmw TA; In addition, the impurity in refinement mother liquor is if PT acid is in mother liquor process of cooling, and meeting fouling in equipment and pipeline, has a strong impact on the operate continuously of device, meanwhile, can cause the loss of PT acid in cleaning process, amounts to PX consumption and is about 0.5kg/tPTA.Also have Chinese patent CN201210382308.8, CN201210428330.1 to propose respectively to adopt adsorbed ion exchange and extraction mode to reclaim organism, the cost of waste water recycling is higher, is difficult to actual motion.
To sum up, in view of conventional P TA energy consumption in production process, material consumption are all higher, be necessary to find brand-new process program to reduce the consumption such as raw material PX, process solvent acetic acid and water, steam energy consumption and catalyzer, promote cleaner production and save energy and reduce the cost.
Summary of the invention
In order to overcome the above-mentioned defect under prior art, the object of the present invention is to provide PTA production method and the system of a kind of less energy-consumption material consumption, less waste discharge, the method and system not only effectively can control the acetic acid consumption in tail gas, waste water, the organism of high efficiente callback refinement mother liquor also realizes the recycling of water, reduce steam, power consumption, reduce waste discharge, but also add the byproduct such as phenylformic acid, mixed phthalic acid, and technical process is short, workable, production cost is low, plant investment is low.
Technical scheme of the present invention is:
A kind of less energy-consumption material consumption, the PTA production method of few waste discharge, comprise oxidation section and refinement, the oxidation reactor top oxidized tail gas out of described oxidation section is introduced into first step vapour generator byproduct steam, the thermal medium outlet gas phase of first step vapour generator enters Acetic Acid Recovering Column at the bottom of tower, the vapour generator group that the discharging of described Acetic Acid Recovering Column top gaseous phase is composed in series through multiple vapour generator successively carries out multi-stage condensing and byproduct steam, at the bottom of most of tower of described Acetic Acid Recovering Column, liquid-phase reflux is to the tower top dewatering period of described oxidation reactor, the refinement mother liquor of described refinement enters described Acetic Acid Recovering Column on one piece, the top tower tray of described Acetic Acid Recovering Column.
Described less energy-consumption material consumption, the PTA production method of few waste discharge also comprises the solvent recuperation of liquid phase at the bottom of the small portion tower to described Acetic Acid Recovering Column, described solvent recuperation can comprise extraction, extraction agent reclaims and extraction agent rectification step, described extraction step carries out extraction to the pending liquid through cooling to be extracted oil phase and aqueous phase extracted, extraction agent preferably adopts methyl acetate, isopropyl acetate, n-butyl acetate, the mixture of any one or a few the arbitrary proportion in isobutyl acetate, described extraction agent recycling step is separated further described aqueous phase extracted to obtain gaseous state extraction agent and water, wherein be back to use described extraction step after the cooling of gaseous state extraction agent, described extraction agent rectification step obtains gas phase mixture and dense acetic acid to the rectifying of described extraction oil phase, wherein gas phase mixture carries out oily water separation after cooling, obtain oil phase extraction agent and be back to use described extraction step, obtain aqueous phase substance and be back to described extraction agent recycling step, described dense acetic acid is back to use in oxidizing reaction and/or is used as the flushing acetic acid of whole system.
The terephthalic acid crystals that oxidizing reaction described in described oxidation section generates separates out formation slurry, described slurry is sent into filter-press successively and is filtered successively after some crystallizer crystallizations, filtration cakes torrefaction and the multistage washing of filter cake obtain crude terephthalic acid wet cake, oxidation mother liquor and wash filtrate, described crude terephthalic acid wet cake is directly sent to refinement as main raw material and is carried out slurry preparation, described for major part oxidation mother liquor is back to described oxidizing reaction, oxidation mother liquor described in rest part sends to oxidation mother liquor process, drying gas preferably adopts dried oxidized tail gas, one or more in fresh nitrogen or water vapor, described wash filtrate is back to described Acetic Acid Recovering Column and carries out acetic acid recovery, in the multistage washing of described filter cake, last step washing and described slurry preparation liquid all adopt water, the liquid phase that the gas phase that described water preferably adopts described vapour generator group to export obtains through gas-liquid separation.
The liquid phase that the gas phase that in the multistage washing of described filter cake of oxidation section, the described slurry preparation liquid of last step washing liquid, refinement and pure terephthalic acid's filter cake washing liquid of refinement all can export for described vapour generator group obtains through the liquid phase that gas-liquid separation goes out after removing ritalin again, preferred employing gas stripping process removes ritalin, and steam stripped thermal source is from the partial gas phase of the lime set after the crystallizer flash steam heated slurry preheater of refinement and last step crystallizer.
The step of described oxidation mother liquor process can be: oxidation mother liquor is first by acetic acid stripping tower concentrate recovery of acetic acid, the mother liquor residue slurry thin up obtained after concentrate also drops into extraction agent and extracts, bottom slurry after extraction obtains filter cake and filtrate through filtration, filter cake drying recovery of acetic acid and extraction agent obtain mixed phthalic acid coarse fodder, filtrate obtains oil phase and aqueous phase through oily water separation, first reclaim extraction agent to oil phase to carry out rectification and purification again and cool obtaining phenylformic acid product, the extraction of oxidation mother liquor is returned after the extraction agent cooling reclaimed, aqueous phase is back to described oxidizing reaction.
Further, in the mother liquor residue slurry obtained after concentrate, acetate concentration preferably reaches 60-85wt%, extraction agent used is preferably the mixture of one or more in methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate, during extraction, the mass ratio of described mother liquor residue slurry, water and extraction agent is preferably 1:0.5-2.0:1.5-4.0, and extraction temperature is preferably 40-90 DEG C.
A kind of less energy-consumption material consumption, the PTA production system of few waste discharge, comprise oxidation subsystem, refining subsystem and oxidized tail gas processing subsystem, top gas phase outlet connection one vapour generator of the oxidation reactor of described oxidation subsystem, the thermal medium outlet of described vapour generator connects the bottom gas phase import of an Acetic Acid Recovering Column, the top gas phase outlet of described Acetic Acid Recovering Column connects the thermal medium inlet of the vapour generator group be composed in series by multiple vapour generator, the thermal medium outlet of vapour generator group connects the import of Acetic Acid Recovering Column return tank, the top gas phase outlet of described Acetic Acid Recovering Column return tank connects described oxidized tail gas processing subsystem, liquid phase primary outlet at the bottom of the tower of described Acetic Acid Recovering Column (outlet for liquid phase at the bottom of the most of tower of outer row) connects the tower top dewatering period import of described oxidation reactor, the refinement mother liquor outlet of described refining subsystem connects tower top first fluid inlet of described Acetic Acid Recovering Column, described first fluid inlet is positioned at the top of the one piece of tower tray in the top of described Acetic Acid Recovering Column.
Described solvent recuperation subsystem can comprise extraction dehydration tower, extraction agent recovery tower and extraction agent rectifying tower, liquid phase sub-outlet at the bottom of the tower of described Acetic Acid Recovering Column (outlet for liquid phase at the bottom of outer row's small portion tower) connects the tower top import of described extraction dehydration tower through water cooler, the tower top oil phase outlet of described extraction dehydration tower connects the opening for feed in the middle part of described extraction agent rectifying tower, aqueous phase outlet at the bottom of the tower of described extraction dehydration tower connects tower top first import of described extraction agent recovery tower, the top gaseous phase outlet of described extraction agent recovery tower connects extraction agent import bottom described extraction dehydration tower through water cooler, aqueous phase outlet at the bottom of the tower of described extraction agent recovery tower connects sewage works, the top gas phase outlet of described extraction agent rectifying tower connects the import of decanting vessel through water cooler, the oil phase outlet of described decanting vessel connects the described extraction agent import bottom described extraction dehydration tower, first aqueous phase outlet of described decanting vessel connects tower top second import of described extraction agent recovery tower, second aqueous phase outlet of described decanting vessel connects the refluxing opening of described extraction agent rectifier, described extraction agent rectifying tower is provided with acetic acid discharge port at the bottom of some towers, extraction agent used is preferably methyl acetate, isopropyl acetate, n-butyl acetate, the mixture of any one or a few the arbitrary proportion in isobutyl acetate.
The multiple oxidizing and crystallizing devices of bottom discharge mouth through connecting successively of described oxidation reactor connect the opening for feed of a rotary pressure filter, described rotary pressure filter is provided with filtration successively, blow dried and filter cake washing three districts, the outlet of filtrating area is connected with the mother liquor reflux mouth bottom described oxidation reactor, the dry gas import in blow dried district connects dried oxidized tail gas, one or more in the air duct of fresh nitrogen or water vapor, filter cake washing district is provided with multistage washing subregion, the washings import of last step washing subregion connects the first bottom liquid phases outlet of described acetic acid reflux tower return tank, in all the other washing subregions at different levels, the washings import of previous stage washing subregion connects the outlet of the wash filtrate of rear stage washing subregion, the filtered solid outlet in described filter cake washing district connects the solid material import of the slurry preparation tank of described refining subsystem, the water inlet of described slurry preparation tank connects the first bottom liquid phases outlet of described acetic acid reflux tower return tank.
Described less energy-consumption material consumption, the PTA production system of few waste discharge is also preferably provided with ritalin stripping tower, the water inlet of washings import and described slurry preparation tank that the first bottom liquid phases outlet of described Acetic Acid Recovering Column return tank washs subregion through described ritalin stripping tower and described last step is connected, concrete mode of connection is: the first bottom liquid phases outlet of described Acetic Acid Recovering Column return tank is connected to the top inlet of described ritalin stripping tower, the outlet at bottom of described ritalin stripping tower connects the washing washings import of subregion of described last step and the water inlet of described slurry preparation tank, the top gaseous phase outlet connection one of described ritalin stripping tower is for import at the bottom of the tower of the ordinary-pressure absorption tower of recovery of acetic acid methyl esters.
Described less energy-consumption material consumption, the PTA production system of few waste discharge also comprises oxidation mother liquor processing subsystem, described oxidation mother liquor processing subsystem can comprise acetic acid stripping tower, residue extractor, residue filter machine, residue drying machine, decanting vessel, residue extraction agent recovery tower and phenylformic acid rectifying tower, the bottom inlet of described acetic acid stripping tower is as the pending liquid import of described oxidation mother liquor processing subsystem, the outlet at bottom of described acetic acid stripping tower connects the bottom inlet of described residue extractor, described residue extractor is provided with upper water import and top extraction agent import, described upper water import connects the second bottom liquid phases outlet of described Acetic Acid Recovering Column return tank, the outlet at bottom of described residue extractor connects the import of described residue filter machine, the filtered solid outlet of described residue filter machine connects the import of described residue drying machine, the filtrate outlet of described residue filter machine connects the import of described decanting vessel, the aqueous phase outlet of described decanting vessel connects the mother liquor reflux mouth bottom described oxidation reactor, the oil phase outlet of described decanting vessel connects the bottom inlet of described residue extraction agent recovery tower, the top extraction agent outlet of described residue extraction agent recovery tower connects the top extraction agent import of described residue extractor through water cooler, the outlet at bottom of described residue extraction agent recovery tower connects the import of described phenylformic acid rectifying tower.
Enter acetate concentration in the mother liquor residue slurry of described residue extractor from described acetic acid stripping tower and be preferably 60-85wt%, the extraction agent entering described residue extractor is preferably the mixture of any one or a few the arbitrary proportion in methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate, described in described residue extractor, the mass ratio of mother liquor residue slurry, water and extraction agent is preferably 1:0.5-2.0:1.5-4.0, and extraction temperature is preferably 40-90 DEG C.
Further, the crude terephthalic acid moisture percentage 5-20% in oxidation section after filtration drying, liquid phase is the acetum of 85-95%; The crude terephthalic acid moisture percentage of described filter outlet is 25-35%, acetic acid content is 1-5%, and the acetic acid content of the wash filtrate of last step is 25-45%.
Further, described vapour generator (the 1st grade of vapour generator 21) byproduct steam grade is pressure 0.3-0.6MPaG, and steam generator outlet liquid phase temperature is 155-185 DEG C.
Further, described Acetic Acid Recovering Column column bottom temperature is 155-185 DEG C, water-content is 45-75wt% in dilute acetic acid at the bottom of tower, the temperature that return tank of top of the tower goes out saliva is 30-50 DEG C, acetic acid content is 0.5-2wt% in water, and in return tank outlet liquid phase, 97 ~ 99% go ritalin stripping tower, the residue extractor of 3-1% deoxidation mother liquid disposal subsystem, rest part to be back to Acetic Acid Recovering Column.
Further, described ritalin stripping (carrying out in ritalin stripping tower) working pressure is 0-0.3MPaG, tower top outlet gas phase temperature is 90-130 DEG C, drop temperature at the bottom of tower is 102-145 DEG C, ritalin content is 0-0.1wt%, and the mass fraction removing methyl acetate recovery tower last step crystallizer top flash-off steam is 20-60%.20-35% deoxidation filter in discharging at the bottom of tower, 15-25% removes refining filtering machine, and 45-65% removes refining slurry preparation.
Further, the mixture that dehydration tower extraction agent used is preferably any one or a few the arbitrary proportion in methyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate is extracted described in solvent recuperation, extraction dehydration tower operation preferable temperature is 30 ~ 50 DEG C, and the mass ratio of feeding extraction agent at the bottom of tower and tower top charging dilute acetic acid is preferably 1.5-4:1; Acetate concentration in water outlet at the bottom of described extraction agent recovery tower is 0-0.2wt%, extractant concentration is 0-1.5wt%; In discharging acetic acid at the bottom of described extraction agent rectifying tower, extractant concentration is 0-0.1wt%, water concentration is 5-15wt%, and the acetate concentration in decanting vessel aqueous phase is 0-0.1wt%.
Beneficial effect of the present invention is:
(1) by arranging Acetic Acid Recovering Column 22 and refinement mother liquor being back to Acetic Acid Recovering Column, eliminate the treatment system that conventional P TA technique is filtered refinement mother liquor cooling, not only simple flow but also can the organism such as water, cobalt manganese and acetic acid, PT acid, 4-CBA, TA in efficient recovery refinement mother liquor, reduce PX unit consumption 2.5kg, the acetic acid unit consumption 1.0kg of PTA product per ton, de-mineralized water unit consumption 2.0t, catalyst cobalt and manganese unit consumption 0.01kg separately, also reduce sewage discharge 2.0t/tPTA, the heat of refining subsystem mother liquor is also recycled simultaneously; Visible the consumption that Acetic Acid Recovering Column reduces raw material PX, acetic acid, catalyzer is set, greatly reduces de-mineralized water consumption and quantity of wastewater effluent, facilitate energy-conserving and environment-protective and cleaner production effectively; Significantly, by the operation of Acetic Acid Recovering Column, oxidation section and refinement is made to control to become to jointly control from independent, the mother liquor refined is made to save the process of cooling, remain the higher water of temperature and enter oxidation subsystem, not only do not consume waste heat, effectively reclaim the heat in refinement mother liquor on the contrary; The reuse of refinement mother liquor, eliminates the scale formation of former refinement mother liquor treatment unit, makes whole plant running more steady;
(2) first step vapour generator 21 is set between oxidation reactor top gas phase and Acetic Acid Recovering Column, the water yield of discharging from acetic acid recovery tower top is regulated, with the water consumption of the water loss and oxidation mother liquor processing subsystem that meet refining subsystem by the temperature out of this vapour generator;
(3) arranged by foregoing first step vapour generator and Acetic Acid Recovering Column, oxidizing reaction is made to produce pennies from heaven: in the oxidation reaction, not only there is the oxidizing reaction of PX, simultaneously with the oxidizing reaction of acetic acid, because water is in oxidation subsystem and refining subsystem recycle, water-content in oxidation reactor is increased, but when the content of water increases, obviously can suppress the oxidation of acetic acid, the consumption of acetic acid is reduced, and the water yield increased can be discharged from oxidation reactor top by flash distillation, to follow-up crystallization, hydrofining but can not have an impact,
(4) adopt extraction dehydration tower 32 to remove water that oxidizing reaction produces and the water that adds of washing tail gas, the charging of extraction dehydration tower is from the extraction of acetic acid recovery tower bottom, then draining after extraction dehydration tower and extraction agent recovery tower 33, draining acetic acid content is 0-0.2wt%; Compared with traditional technology adopts azeotropic distillation to dewater, about the steam consumption that extraction process of the present invention dewaters only has the half of azeotropic distillation;
(5) be oxidized the filter-press that subsystem adopts water washing, eliminate the oxidizing and drying machine of conventional P TA technique, gas conveying and intermediate bunker, significantly simplify and shorten technical process, significantly reducing one-time investment expense; Filter adopt first filter, gas drying, the technical process of finally washing again, significantly can improve the washing effect of filter cake and reduce washing water consumption; Cancel the steam consumption that drying machine then decreases device;
(6) in multistage washing, last step washing washings adopts the water containing a small amount of acetic acid of Acetic Acid Recovering Column overhead reflux, and the water of this backflow is recycled fully, the consumption of the outer water of the system that simultaneously significantly reduces;
(7) water of Acetic Acid Recovering Column overhead reflux is first sent to as the filter cake washing liquid of oxidation section, the making beating liquid of refinement and filter cake washing liquid through ritalin stripping tower 29 stripping removing ritalin again, (existence of organism ritalin can cause expensive hydrogenation palladium carbon catalyst inactivation on the impact of palladium carbon catalyst in hydrogenation reaction effectively can to avoid ritalin, blocking catalyst active site, affect hydrogenation reaction, cause the finished product PTA defective); The ritalin that stripping tower reclaims simultaneously is all back to oxidation subsystem, reduces the discharge of ritalin, thus reduces the consumption of acetic acid;
(8) oxidation mother liquor removal of impurities process adopts extraction process, eliminate thin-film evaporator, both the phenylformic acid in oxidation residua, terephthalic acid, m-phthalic acid and catalyst cobalt manganese had been reclaimed, eliminate again the high pressure steam consumption of former thin-film evaporator, high pressure steam can be reduced and consume 0.03t/tPTA, reduce PX unit consumption 3.0kg/tPTA, eliminate oxidation residua discharge; Particularly part is, though the acetate concentration of the slurry after acetic acid stripping tower 50 significantly reduces, but still up to 60-85wt%, and when the content of water in acetic acid higher with the mode of extraction all can not be adopted time lower to be separated, therefore usually can not expect under prior art adopt extraction process.The present invention is the dilution that will add a part in acetic acid lower for water content and carry out to a certain degree from the water of acetic acid recovery top of tower, make originally cannot adopt and extract the solution that the solution carrying out being separated becomes applicable extracting and separating, and achieve above-mentioned good result thus;
After comprehensive employing every innovative technology described above, great effect is created to PTA technological innovation, compared with conventional P TA technique, present invention process flow process is simple, material consumption energy consumption is low, waste water and dregs is few, change PTA device and be regarded as energy consumption rich and influential family for a long time, water consumption rich and influential family, sewage discharge rich and influential family, the features such as waste residue not easily processes, then become resource-conserving, environmentally friendly PTA novel process, not only do not need power consumption, also electric power can be sent, while meeting own device use, also can transmit electric power outside PTA device, become a part for extra earning, it is energy-conservation typical case, there is good environmental benefit, economic benefit and operability.The technology of the present invention is not only applicable to newly-built PTA device, is applicable to the transformation and upgrade of existing PTA device too, makes the more energy-saving and cost-reducing operation of device.
Accompanying drawing explanation
Fig. 1 is the process flow sheet of PTA production method of the present invention;
Fig. 2 is the process flow sheet of traditional PTA production method.
Reference numeral is:
1, oxidation reactor, 2, be oxidized the first crystallizer, 3, be oxidized the second crystallizer, 4, be oxidized the 3rd crystallizer, 5, oxidative pressure filter, 6, refining slurry blend tank, 7 ~ 12, refining paste heater, 13, hydrogenator, 14 ~ 18, first ~ level V refining crystallization device, 19, refining filter-press, 20, the dry machine of PTA, 21, first step vapour generator, 22, Acetic Acid Recovering Column, 23, second stage vapour generator, 24, third stage vapour generator, 25, fourth stage vapour generator, 26, de-mineralized water well heater, 27, water recirculator, 28, Acetic Acid Recovering Column return tank, 29, ritalin stripping tower, 30, ordinary-pressure absorption tower, 31, water cooler, 32, extraction dehydration tower, 33, extraction agent recovery tower, 34, extraction agent reclaims overhead condenser, and 35, extraction agent rectifying tower, 36, washing acetic acid water cooler, 37, extraction agent rectifying tower overhead condenser, 38, decanting vessel, 39, sewage works, 40, high-pressure absorber, 41, tail gas incineration system, 42, off-gas expander, 43, tail gas washing tower, 50, acetic acid stripping tower, 51, residue extractor, 52, residue filter machine, 53, residue drying machine, 54, decanting vessel, 55, residue extraction agent recovery tower, 56, residue extraction agent condenser, 57, phenylformic acid rectifying tower, 58, phenylformic acid water cooler, 100, raw material p-Xylol, 101, air, 102, catalyzer, 103, hydrogen, 104, PTA product, 105, water, 106, rinse acetic acid, 107, mixed phthalic acid, 108, phenylformic acid,
201, oxidation reactor, 202, be oxidized the first crystallizer, 203, be oxidized the second crystallizer, 204, be oxidized the 3rd crystallizer, 205, oxidation filter, 206, oxidizing and drying machine, 207, intermediate bunker, 208, refining slurry blend tank, 209 ~ 214, refining paste heater, 215, hydrogenator, 216 ~ 220, first ~ level V refining crystallization device, 221, refining filter-press, 222, the dry machine of PTA, 231, first step vapour generator, 232, second stage vapour generator, 233, third stage vapour generator, 234, fourth stage vapour generator, 235, water cooler, 236, reactor return tank, 237, water cooler, 238, high-pressure absorber, 239, tail gas incineration system, 240, off-gas expander, 241, tail gas washing tower, 251, dehydration tower, 252, washing acetic acid water cooler, 253, dehydration overhead condenser, 254, decanting vessel, 255, entrainer recovery tower, 256, entrainer reclaims overhead condenser, and 257, sewage works, 261, acetic acid stripping tower, 262, thin-film evaporator, 263, catalyzer water extractor, 264, residue whizzer, 265, concentrated solution well heater, 266, catalyst recovery tank, 300, raw material p-Xylol, 301, air, 302, catalyzer, 303, hydrogen, 304, PTA product, 305, de-mineralized water, 306, rinse acetic acid, 307, water, 308, oxidation residua.
Embodiment
Fig. 1 gives and implements technical process of the present invention, comprises following five integral parts:
(1) p xylene oxidation subsystem: in oxidation reactor 1 take acetic acid as solvent, and with Cobaltous diacetate, manganese acetate, Hydrogen bromide for catalyzer, raw material p-Xylol 100 and air 101 react and generate CTA under catalyzer 102 acts on.Reactor is the non-stirring reactor of band dewatering period, also can for any type of reactor applied in other existing PTA devices, service temperature is 180-200 DEG C, and working pressure is 1-1.5MPa(absolute pressure), at this, patent ZL03142246.2 and ZL200910076703.1 is referred to this.The dewatering period of tower top plays acetic acid and is separated and reduces the effect that solid phase carries secretly with water, the height of dewatering period can regulate according to water-content, device quantity of steam and acid balance in oxidation reactor, and heat of reaction is taken away from tower top by acetic acid, water and the tail gas evaporated.Because reaction product TA solubleness in solvent acetic acid is not high, the TA crystallization formation slurry in the reactor of more than 90%, follow-up oxidation first crystallizer 2 sent into successively by slurry, be oxidized the second crystallizer 3, be oxidized the 3rd crystallizer 4 with decrease temperature and pressure flash distillation and crystallization step by step, being oxidized the 3rd crystallizer 4 temperature controls at 90-140 DEG C, bottom it, oxidative pressure filter 5 delivered to by slurry, filter-press 5 is rotary pressure filter, slurry completes filtration successively in filter, blow dried, filter cake washing is totally three processes: first filter, by most of liquid phase and TA crystal separation, rate of cake moisture content controls at 15-35%, filtrated stock is the acetic acid of 85-95%, except small part delivers to mother liquid disposal subsystem (entering mother liquid disposal processing subsystem from acetic acid stripping tower 50), major part mother liquor is back to oxidation reactor, then carry out blow dried, dry gas can be dried oxidized tail gas, fresh nitrogen or water vapor, and after dry, rate of cake moisture content can be reduced to 10 ~ 12% even less than 8% from 15-35%, finally filter cake is washed, washing process adopts 2 ~ 4 grades of washings, the washings of last step is ritalin stripping tower 29 bottom discharge, the washings of other grade is the wash filtrate after next stage washing, in wash filtrate after last step washing, acetic acid content is less than 10wt%, in wash filtrate after 1st grade of washing, acetic acid content is 25-45%, this wash filtrate send Acetic Acid Recovering Column 22, after washing, gained filter cake is crude terephthalic acid wet feed, the moisture percentage of this wet feed is 20-40wt%, acetic acid content is 1-5%, wet feed removes refinement of crude terephthalic acid subsystem (entering refining slurry blend tank 6).
(2) refinement of crude terephthalic acid subsystem: pull an oar refining slurry blend tank 6 from oxidative pressure filter 5 crude terephthalic acid wet feed out and the water from ritalin stripping tower 29, slurry solid content controls at 25-35wt%, temperature controls at 95-140 DEG C, and pressure-controlling is at 0-0.5MPaG; From refining slurry blend tank 6 slurry out through six grades of refining paste heater 7-12 stepped heatings to 280-290 DEG C, 1st ~ 5 grades of well heaters above and 1st ~ 5 grades of refining crystallization device one_to_one corresponding, use the flash steam of corresponding crystallizer to heat, last step adopts high pressure steam heating; Slurry through heating carries out hydrogenation reaction with hydrogen 103 in hydrogenator 13, hydrogenator is fixed-bed reactor, catalyzer is palladium charcoal solid particulate, and the object of hydrogenation process the impurity 4-CBA in crude terephthalic acid and hydrogen is reacted make 4-CBA be reduced to PT acid more soluble in water; The discharging of hydrogenator 13 enters 1st ~ 5 grades of refining crystallization devices 14 ~ 18 and carries out decrease temperature and pressure flash crystallization, and the temperature of the 5th grade of crystallizer 18 is 140-150 DEG C; The flash steam of crystallizer at different levels is for heating paste heater above to reclaim heat, and the 5th grade of crystalliser feet discharging is delivered to refining filter-press 19 and carried out solid-liquid separation.Feed slurry carries out successively filtering in refining filter-press 19, filter cake washing and gas drying totally three processes: first filter, gained rate of cake moisture content is 25-30wt%; Then wash, washing process adopts multistage washing, the washings of every grade of washing is the wash filtrate after next stage washing, the washing water of last step are from ritalin stripping tower 29, washing sum of series washing water consumption reaches product requirement to control according to foreign matter content in filter cake, filter liquor and wash filtrate composition refinement mother liquor, refinement mother liquor enters Acetic Acid Recovering Column 22 from top, thus has reclaimed the organism such as PT acid, 4-CBA and TA in refinement mother liquor; Finally carry out gas drying, desiccant gas is dried oxidized tail gas, nitrogen, water vapor.From the wet cake of refining pressure filter blanking, moisture percentage is the preferred 8-12wt% of 5-15wt%(), this wet cake obtains PTA product 104 after finally entering PTA drying machine 20 drying.
(3) solvent recovery cycle subsystem: from oxidation reactor 1 top oxidized tail gas out by steam needed for first step vapour generator 21 by-product device, the grade of byproduct steam adjusts at 0.3-0.6MPaG according to device balance of steam, the temperature out of vapour generator controls at 150-180 DEG C, and the liquid phase that in tail gas, condensation is got off is back to oxidation reactor, and gas phase delivers to Acetic Acid Recovering Column 22, by the water yield regulating the temperature out of first step vapour generator 21 to regulate Acetic Acid Recovering Column return tank 28 to discharge, to meet the demand of Purified Water, at the bottom of tower, Acetic Acid Recovering Column 22 is entered from first step vapour generator 21 oxidized tail gas out, wash filtrate from oxidative pressure filter 5 enters the bottom of Acetic Acid Recovering Column 22, the refinement mother liquor carrying out autofining filter-press 19 enters Acetic Acid Recovering Column 22 from tower top and is back to first piece of tower tray, each feed stream realizes being separated of acetic acid and water in Acetic Acid Recovering Column, tower reactor is acetic acid-aqueous solution (water containing a small amount of acetic acid), the PT acid of reclaiming from refinement mother liquor has been dissolved in solution, the organism such as 4-CBA and TA, tower reactor discharging is extracted small portion out and is dewatered to extraction dehydration tower 31, all the other are back to oxidation reactor 1, at the bottom of the tower of Acetic Acid Recovering Column 22, service temperature is 155-185 DEG C, is preferably 160-180 DEG C, and the rate of recovery of acetic acid is greater than 98.5%, and the water-content of tower reactor is 45-75wt%, Acetic Acid Recovering Column ejects the oxidized tail gas that gas mainly contains acetic acid and water, this oxidized tail gas is successively through the 23-25 condensation of 2-4 level vapour generator and byproduct water steam, and then after de-mineralized water well heater 26 and water recirculator 27 continue to be cooled to 30-50 DEG C, go Acetic Acid Recovering Column return tank 28 to carry out gas-liquid separation, return tank outlet gas phase truncates gas disposal subsystem (entering high-pressure absorber 40), return tank outlet liquid phase is water, the 97-99% of this water yield goes ritalin stripping tower 29 to remove ritalin, the residue extractor 51 of 1-3% deoxidation mother liquid disposal subsystem, rest part is back to the tower top second liquid phase import that Acetic Acid Recovering Column 22(enters the top of the first piece of tower tray being positioned at described Acetic Acid Recovering Column), oxidation section (not shown) is back to after the ritalin that the stripping of ritalin stripping tower 29 goes out enters ordinary-pressure absorption tower 30, water after stripping flows out from tower reactor, 20-35% deoxidation filter, 15-25% removes refining filtering machine, 45-65% sends into refining slurry blend tank 6 as making beating water, thus achieve recycling of water, ritalin stripping tower working pressure is 0-0.3MPaG, tower top outlet gas phase temperature is 90-130 DEG C, and leaving water temperature(LWT) at the bottom of tower is 102-140 DEG C, ritalin content is 0-0.1wt%, after cooling, extraction dehydration tower 32 is entered from tower top from liquid phase at the bottom of the tower of Acetic Acid Recovering Column 22, extraction agent enters extraction dehydration tower 32 at the bottom of tower, extraction dehydration tower bottom discharge is aqueous phase extracted, this aqueous phase enters extraction agent recovery tower 33 from tower top, extraction agent recovery tower bottom discharge is water, this water is disposed to sewage works 39, extraction agent recovery tower top exit gas phase is gaseous state extraction agent, after condensation, be back to extraction dehydration tower 32 recycle, the top discharge of extraction dehydration tower is extraction oil phase, oil phase removes extraction agent rectifying tower 35, discharging at the bottom of extraction agent rectifying tower is dense acetic acid, this dense acetic acid part goes high-pressure absorber 40 to use as washing acetic acid after water cooler 36 cools, a part uses as the flushing acetic acid 106 of whole device, rest part is back to oxidation reactor 1, the tower top outlet gas phase of extraction agent rectifying 35 goes decanting vessel 38 to be divided into water-oil phase after condenser 37 condensation, oil phase returns extraction dehydration tower 32 and recycles, aqueous portion enters extraction agent rectifying tower 35 as backflow from tower top, rest part enters extraction agent recovery tower 33 from tower top, the extraction agent that extraction dehydration tower 32 uses is methyl acetate, the mixture of one or more in isopropyl acetate, n-butyl acetate, isobutyl acetate, the service temperature of tower is 30-50 DEG C, the mass ratio of feeding extraction agent at the bottom of tower and tower top charging dilute acetic acid is 1.5-4:1, acetate concentration in water outlet at the bottom of the tower of extraction agent recovery tower 33 is 0-0.2wt%, extractant concentration is 0-0.2wt%, and in discharging acetic acid at the bottom of the tower of extraction agent rectifying tower 35, extractant concentration is 0-0.2wt%, water concentration is 5-15wt%.
(4) vent gas treatment subsystem: from Acetic Acid Recovering Column return tank 28 oxidized tail gas at the bottom of tower, enter high-pressure absorber 40, the dense acetic acid spray from extraction agent rectifier bottoms water cooler 36 is adopted in the middle part of high-pressure absorber, top adopts de-mineralized water 105 to spray, the high-pressure absorber top exit gas phase gas catalytic burning system 41 that truncates purifies, tail gas after purification goes off-gas expander 42 to reduce pressure acting, and the tail gas gas wash tower 43 that truncates after decompression acting is finally disposed in air after alkali cleaning.
(5) mother liquid disposal: the oxidation mother liquor extracting 5-30% from oxidative pressure filter 5 out carries out mother liquid disposal, first through peroxyacetic acid stripping tower 50 concentrate recovery of acetic acid simultaneously, the discharging of acetic acid stripping tower bottom is mother liquor residue slurry, mother liquor residue slurry enters residue extractor 51, simultaneously in residue extractor 51, add extraction agent and water stirs, extraction, water exports liquid phase from Acetic Acid Recovering Column return tank 28, the discharging of residue extractor obtains filter cake and filtrate after going residue filter machine 52 to filter, filter cake is through drying machine 53 dried recovered acetic acid, extraction agent also obtains mixed phthalic acid coarse fodder 107, filtrate lease making decanting vessel 54 is divided into water-oil phase, oil phase first sends to residue extraction agent recovery tower 55 to reclaim extraction agent, enter phenylformic acid rectifying tower 57 again carry out rectification and purification and obtain phenylformic acid product 108 after condenser 58 condensation, the extraction agent reclaimed returns residue extractor 51 after condenser 56 condensation, aqueous phase is back to oxidation reactor 1, acetate concentration in mother liquor residue slurry is 60-85wt%, the extraction agent added in residue extractor 51 is the mixture of one or more in methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate, mother liquor residue slurry: water: the mass ratio of extraction agent is 1:0.5-2.0:1.5-4.0, the extraction temperature of residue extractor is 40-90 DEG C.
Be below a concrete PTA production technique operating process embodiment of the present invention, designed annual production capacity is the PTA production equipment of 1,200,000 tons, and PTA output is 150t/h, comprises following five subsystems.
(1) p xylene oxidation subsystem: in oxidation reactor 1, take acetic acid as solvent, with Cobaltous diacetate, manganese acetate, Hydrogen bromide for catalyzer, the raw material p-Xylol 100 of 98.25t/h reacts with the air 101 of 470t/h and generates crude terephthalic acid CTA under catalyzer 102 acts on.Reactor is the non-stirring reactor of band dewatering period, and service temperature is 187 DEG C, and working pressure is 1.1MPaG, and in reactor liquid phase, catalyst concn is cobalt 400ppm, manganese 400, bromine 950pppm, and the residence time is 65min.The dewatering period of tower top plays acetic acid and is separated and reduces the effect that solid phase carries secretly with water, and dewatering period is be highly 500mm, and heat of reaction is taken away from tower top by acetic acid, water and the tail gas evaporated.Follow-up oxidation first crystallizer 2 sent into successively by reactor discharging slurry, be oxidized the second crystallizer 3, be oxidized the 3rd crystallizer 4 with decrease temperature and pressure flash distillation and crystallization step by step, being oxidized the 3rd crystallizer 4 temperature controls at 90 DEG C, bottom it, oxidative pressure filter 5 delivered to by slurry, filter-press 5 is rotary pressure filter, slurry completes filtration successively in filter, blow dried, filter cake washing is totally three processes: first filter, by most of liquid phase and TA crystal separation, rate of cake moisture content controls at 28wt%, filtrated stock is the acetic acid of 91%, except small part delivers to mother liquid disposal subsystem (entering mother liquid disposal processing subsystem from acetic acid stripping tower 50), major part mother liquor is back to oxidation reactor, then carry out blow dried, dry gas is dried oxidized tail gas, and after dry, rate of cake moisture content can be reduced to less than 15% from 28%, finally filter cake is washed, washing process adopts 2 grades of washings, the washings of last step is refining ritalin stripping tower bottom discharge 100t/h, the washings of other grade is the wash filtrate after next stage washing, in wash filtrate after last step washing, acetic acid content is less than 10wt%, in wash filtrate after 1st grade of washing, acetic acid content is 40%, this wash filtrate send Acetic Acid Recovering Column 22, after washing, gained filter cake is crude terephthalic acid wet feed, the moisture percentage of this wet feed is 35wt%, acetic acid content is 3%, wet feed removes refinement of crude terephthalic acid subsystem (entering refining slurry blend tank 6).
(2) refinement of crude terephthalic acid subsystem: pull an oar refining slurry blend tank 6 from oxidative pressure filter 5 crude terephthalic acid wet feed out and the water from ritalin stripping tower 29, slurry solid content controls at 31wt%, temperature controls at 135 DEG C, and pressure-controlling is at 0.2MPaG; From refining slurry blend tank 6 slurry out through seven grades of refining paste heater 7-12 stepped heatings to 286 DEG C, 1-5 level well heater above and 1-5 level refining crystallization device one_to_one corresponding, use the flash steam of corresponding crystallizer to heat, last two-stage adopts high pressure steam heating; Slurry through heating carries out hydrogenation reaction with hydrogen 103 in hydrogenator 13, hydrogenator is fixed-bed reactor, catalyzer is palladium charcoal solid particulate, and the object of hydrogenation process the impurity 4-CBA in crude terephthalic acid and hydrogen is reacted make 4-CBA be reduced to PT acid more soluble in water; The discharging of hydrogenator 13 enters 1-5 level refining crystallization device 14-18 and carries out decrease temperature and pressure flash crystallization, and the temperature of the 5th grade of crystallizer 18 is 148 DEG C; The flash steam of crystallizer at different levels is for heating paste heater above to reclaim heat, 5th grade of crystalliser feet discharging is delivered to refining filter-press 19 and is carried out solid-liquid separation, feed slurry carries out successively filtering in refining filter-press 19, filter cake washing and gas drying totally three processes: first filter, gained rate of cake moisture content is 28wt%; Then wash, washing process adopts two-stage washing, the washings of one-level washing is the wash filtrate after secondary washing, the washing water of secondary are from Acetic Acid Recovering Column return tank 28(consumption 70t/h), filter liquor and wash filtrate composition refinement mother liquor, a refinement mother liquor part delivers to the washings of oxidative pressure filter 5 as its last step, and residue refinement mother liquor enters Acetic Acid Recovering Column 22 from top, thus has reclaimed the organism such as PT acid, 4-CBA and TA in refinement mother liquor; Finally carry out gas drying, desiccant gas is dried oxidized tail gas, nitrogen, water vapor.From the wet cake of refining pressure filter blanking, moisture percentage is 12wt%, and this wet cake obtains the PTA product 104 of 150t/h after finally entering PTA drying machine 20 drying.
(3) solvent recovery cycle subsystem: from oxidation reactor 1 top oxidized tail gas out by steam needed for first step vapour generator 21 by-product device, the grade of byproduct steam is 0.35MPaG, the temperature out of vapour generator controls at 176 DEG C, and the liquid phase that in tail gas, condensation is got off is back to oxidation reactor, and gas phase delivers to Acetic Acid Recovering Column 22, by the water yield regulating the temperature out of first step vapour generator 21 to regulate Acetic Acid Recovering Column return tank 28 to discharge, to meet the demand of water, at the bottom of tower, Acetic Acid Recovering Column 22 is entered from first step vapour generator 21 oxidized tail gas out, wash filtrate from oxidative pressure filter 5 enters the bottom of Acetic Acid Recovering Column 22, the refinement mother liquor carrying out autofining filter-press 19 enters Acetic Acid Recovering Column 22 from tower top and is back to first piece of tower tray, each feed stream realizes being separated of acetic acid and water in Acetic Acid Recovering Column, tower reactor is acetic acid-aqueous solution, the PT acid of reclaiming from refinement mother liquor has been dissolved in solution, the organism such as 4-CBA and TA, tower reactor discharging is extracted 110t/h out and is dewatered to extraction dehydration tower 31, all the other are back to oxidation reactor 1, service temperature at the bottom of the tower of Acetic Acid Recovering Column 22 is 160 DEG C, and the rate of recovery of acetic acid reaches 99.0%, and the water-content of tower reactor is 55wt%, Acetic Acid Recovering Column ejects the oxidized tail gas that gas mainly contains acetic acid and water, this oxidized tail gas is successively through the 23-26 condensation of 2-4 level vapour generator and byproduct water steam, and then after de-mineralized water well heater 27 and water recirculator 28 continue to be cooled to 40 DEG C, go Acetic Acid Recovering Column return tank 28 to carry out gas-liquid separation, return tank outlet gas phase truncates gas disposal subsystem (entering high-pressure absorber 40), return tank outlet liquid phase is water, the 300t/h of this water yield goes ritalin stripping tower 29 to remove ritalin, the residue extractor 51 of 5t/h deoxidation mother liquid disposal subsystem, all the other 70t/h are back to Acetic Acid Recovering Column 22, oxidation section is back to after the ritalin that the stripping of ritalin stripping tower 29 goes out enters ordinary-pressure absorption tower 30, water after stripping, 100t/h deoxidation pressure filter 5 goes refining pressure filter 19 to flow out from tower reactor as washings, 130t/h as washings, 70t/h and sends into refining slurry blend tank 6 as making beating water, thus achieve recycling of water, ritalin stripping tower working pressure is 0.2MPaG, tower top outlet gas phase temperature is 90 DEG C, and leaving water temperature(LWT) at the bottom of tower is 140 DEG C, ritalin content is 0.05wt%, after cooling, send into tower top from liquid phase at the bottom of the tower of Acetic Acid Recovering Column 22 and enter extraction dehydration tower 32, extraction agent enters extraction dehydration tower 32 at the bottom of tower, extraction dehydration tower bottom discharge is aqueous phase extracted, this aqueous phase enters extraction agent recovery tower 33 from tower top, extraction agent recovery tower bottom discharge is water, this water is disposed to sewage works 39, extraction agent recovery tower top exit gas phase is gaseous state extraction agent, after condensation, be back to extraction dehydration tower 32 recycle, the top discharge of extraction dehydration tower is extraction oil phase, oil phase removes extraction agent rectifying tower 35, discharging at the bottom of extraction agent rectifying tower is dense acetic acid, this dense acetic acid part goes high-pressure absorber 40 to use as washing acetic acid after water cooler 36 cools, a part uses as the flushing acetic acid 106 of whole device, rest part is back to oxidation reactor 1, the tower top outlet gas phase of extraction agent rectifying 35 goes decanting vessel 38 to be divided into water-oil phase after condenser 37 condensation, oil phase returns extraction dehydration tower 32 and recycles, aqueous portion enters extraction agent rectifying tower 35 as backflow from tower top, rest part enters extraction agent recovery tower 33 from tower top, the extraction agent that extraction dehydration tower 32 uses is n-butyl acetate, the service temperature of tower is 40 DEG C, the mass ratio of feeding extraction agent at the bottom of tower and tower top charging dilute acetic acid is 2:1, acetate concentration in water outlet at the bottom of the tower of extraction agent recovery tower 33 is 0.05wt%, extractant concentration is 0.02wt%, and in discharging acetic acid at the bottom of the tower of extraction agent rectifying tower 35, extractant concentration is 0.05wt%, water concentration is 5wt%.
(4) vent gas treatment subsystem: from Acetic Acid Recovering Column return tank 28 oxidized tail gas at the bottom of tower, enter high-pressure absorber 40, the dense acetic acid spray from extraction agent rectifier bottoms water cooler 36 is adopted in the middle part of high-pressure absorber, top adopts de-mineralized water 105 to spray, the high-pressure absorber top exit gas phase gas catalytic burning system 41 that truncates purifies, tail gas after purification goes off-gas expander 42 to reduce pressure acting, and the tail gas gas wash tower 43 that truncates after decompression acting is finally disposed in air after alkali cleaning.
(5) mother liquid disposal: the oxidation mother liquor extracting 10% from oxidative pressure filter 5 out carries out mother liquid disposal, first through peroxyacetic acid stripping tower 50 concentrate recovery of acetic acid simultaneously, the discharging of acetic acid stripping tower bottom is mother liquor residue slurry, mother liquor residue slurry enters residue extractor 51, simultaneously in residue extractor 51, add extraction agent and water stirs, extraction, water exports liquid phase from Acetic Acid Recovering Column return tank 28, the discharging of residue extractor obtains filter cake and filtrate after going residue filter machine 52 to filter, filter cake is through drying machine 53 dried recovered acetic acid, extraction agent also obtains mixed phthalic acid coarse fodder 107, filtrate lease making decanting vessel 54 is divided into water-oil phase, oil phase first sends to residue extraction agent recovery tower 55 to reclaim extraction agent, the extraction agent reclaimed returns residue extractor 51 and then enters phenylformic acid rectifying tower 57 and carries out rectification and purification and obtain phenylformic acid product 108 after condenser 58 condensation after condenser 56 condensation, aqueous phase is back to oxidation reactor 1, acetate concentration in mother liquor residue slurry is 72wt%, the extraction agent added in residue extractor 51 is the mixture of one or more in methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate, mother liquor residue slurry: water: the mass ratio of extraction agent is 1:1:2, the extraction temperature of residue extractor is 52 DEG C.
The result of implementation of the present embodiment is: the PX unit consumption of the PTA product per ton of PTA full device is 650kg, acetic acid unit consumption 35kg, de-mineralized water unit consumption 0.2t, catalyst cobalt unit consumption 0.01kg, steam mono-consumption 0.5t, generating 70KW.
And when adopting conventional P TA technique (same apparatus place adopts identical Operating parameters) shown in Fig. 2, the PX unit consumption of the PTA product per ton of PTA full device is 655kg, acetic acid unit consumption 38kg, de-mineralized water unit consumption 2.2t, catalyst cobalt unit consumption 0.04kg, steam mono-consumption 0.70t, generating-15KW.
Visible, adopt method of the present invention to reduce 5kg, acetic acid unit consumption reduction 3kg, de-mineralized water unit consumption reduction 2.0t, catalyst cobalt unit consumption reduction 0.03kg, steam mono-consumption reduction 0.2t, generating increase 55KW than the PX unit consumption of conventional P TA technique.
Wherein, due to the existence of dewatering system of the present invention, the water that reaction is generated and the fresh water being entered system have all shifted out system, the usage quantity of the steam that therefore can decline to a great extent.For the PTA device of the annual output 1200000 tons/year of the present embodiment, the quantity of steam adopting the present invention to consume 0.35MPaG is 20t/h, amount to 0.13t/tPTA, than adopting traditional technology, (quantity of steam consuming 0.35MPaG is 50t/h, amount to 0.33t/tPTA) time quantity of steam consumption the range of decrease can reach 60%, with 180 yuan/t vaporimeter, year cost saving is about 3,0*8,000,*18,0/1,000,0=4,320 ten thousand yuan.
The present invention adopts new oxidation mother liquor treatment process, not only reduces the consumption of steam, acetic acid and catalyzer, avoids the Hazardous wastes processing cost to oxidation residua, has also reclaimed other materials, energy conservation and consumption reduction effects highly significant.Still for the PTA device of the annual output 1200000 tons/year of the present embodiment, adopt oxidation mother liquor treatment process of the present invention, the quantity of steam consuming 9.5MPaG is 1.3t/h, low-pressure steam consumption is 1t/h, catalyst cobalt consumption is 1.5kg/h, acetic acid all reclaims, simultaneously recyclable phenylformic acid 300kg/h, mixed phthalic acid 300kg/h.And the quantity of steam adopting traditional technology to consume 9.5MPaG is 2.5t/h, low-pressure steam consumption is 5t/h, acetic acid consumption 300kg/h; Catalyst cobalt consumption is that 4.5kg/h, 1000kg/h oxidation residua send Hazardous wastes processing enter.Visible, adopt the present invention to carry out oxidation mother liquor process, the usage quantity 1.2t/h of high pressure steam can be reduced, the usage quantity 4t/h of low-pressure steam, acetic acid consumption amount 300kg/h, catalyst cobalt consumption 3kg/h, recovery of benzoic acid 300kg/h simultaneously, year cost saving is about 4,050 ten thousand yuan.

Claims (10)

1. a less energy-consumption material consumption, the PTA production method of few waste discharge, comprise oxidation section and refinement, it is characterized in that the oxidation reactor top oxidized tail gas out of described oxidation section is introduced into first step vapour generator byproduct steam, the thermal medium outlet gas phase of first step vapour generator enters Acetic Acid Recovering Column at the bottom of tower, the vapour generator group that the discharging of described Acetic Acid Recovering Column top gaseous phase is composed in series through multiple vapour generator successively carries out multi-stage condensing and byproduct steam, at the bottom of most of tower of described Acetic Acid Recovering Column, liquid-phase reflux is to the tower top dewatering period of described oxidation reactor, the refinement mother liquor of described refinement enters described Acetic Acid Recovering Column on one piece, the top tower tray of described Acetic Acid Recovering Column.
2. less energy-consumption material consumption as claimed in claim 1, the PTA production method of few waste discharge, characterized by further comprising the solvent recuperation of liquid phase at the bottom of the small portion tower to described Acetic Acid Recovering Column, described solvent recuperation comprises extraction, extraction agent reclaims and extraction agent rectification step, described extraction step carries out extraction to the pending liquid through cooling to be extracted oil phase and aqueous phase extracted, extraction agent preferably adopts methyl acetate, isopropyl acetate, n-butyl acetate, the mixture of any one or a few the arbitrary proportion in isobutyl acetate, described extraction agent recycling step is separated further described aqueous phase extracted to obtain gaseous state extraction agent and water, wherein be back to use described extraction step after the cooling of gaseous state extraction agent, described extraction agent rectification step obtains gas phase mixture and dense acetic acid to the rectifying of described extraction oil phase, wherein gas phase mixture carries out oily water separation after cooling, obtain oil phase extraction agent and be back to use described extraction step, obtain aqueous phase substance and be back to described extraction agent recycling step, described dense acetic acid is back to use in oxidizing reaction and/or is used as the flushing acetic acid of whole system.
3. less energy-consumption material consumption as claimed in claim 2, the PTA production method of few waste discharge, it is characterized in that the terephthalic acid crystals that oxidizing reaction described in described oxidation section generates separates out formation slurry, described slurry is sent into filter-press successively and is filtered successively after some crystallizer crystallizations, filtration cakes torrefaction and the multistage washing of filter cake obtain crude terephthalic acid wet cake, oxidation mother liquor and wash filtrate, described crude terephthalic acid wet cake is directly sent to refinement as main raw material and is carried out slurry preparation, described for major part oxidation mother liquor is back to described oxidizing reaction, oxidation mother liquor described in rest part sends to oxidation mother liquor process, drying gas preferably adopts dried oxidized tail gas, one or more in fresh nitrogen or water vapor, described wash filtrate is back to described Acetic Acid Recovering Column and carries out acetic acid recovery, in the multistage washing of described filter cake, last step washing and described slurry preparation liquid all adopt water, the liquid phase that the gas phase that described water preferably adopts described vapour generator group to export obtains through gas-liquid separation.
4. less energy-consumption material consumption as claimed in claim 3, the PTA production method of few waste discharge, it is characterized in that last step washing liquid in the multistage washing of described filter cake of oxidation section, the liquid phase that the gas phase that the described slurry preparation liquid of refinement and pure terephthalic acid's filter cake washing liquid of refinement are the output of described vapour generator group obtains through the liquid phase that gas-liquid separation goes out after removing ritalin again, preferred employing gas stripping process removes ritalin, and steam stripped thermal source is from the partial gas phase of the lime set after the crystallizer flash steam heated slurry preheater of refinement and last step crystallizer.
5. less energy-consumption material consumption as claimed in claim 4, the PTA production method of few waste discharge, it is characterized in that the step of described oxidation mother liquor process is: oxidation mother liquor is first by acetic acid stripping tower concentrate recovery of acetic acid, the mother liquor residue slurry thin up obtained after concentrate also drops into extraction agent and extracts, bottom slurry after extraction obtains filter cake and filtrate through filtration, filter cake drying recovery of acetic acid and extraction agent obtain mixed phthalic acid coarse fodder, filtrate obtains oil phase and aqueous phase through oily water separation, first reclaim extraction agent to oil phase to carry out rectification and purification again and cool obtaining phenylformic acid product, the extraction of oxidation mother liquor is returned after the extraction agent cooling reclaimed, aqueous phase is back to described oxidizing reaction, in the mother liquor residue slurry obtained after concentrate, acetate concentration preferably reaches 60-85wt%, extraction agent used is preferably the mixture of one or more in methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, isobutyl acetate, during extraction, the mass ratio of described mother liquor residue slurry, water and extraction agent is preferably 1:0.5-2.0:1.5-4.0, and extraction temperature is preferably 40-90 DEG C.
6. a less energy-consumption material consumption, the PTA production system of few waste discharge, comprise oxidation subsystem, refining subsystem and oxidized tail gas processing subsystem, it is characterized in that top gas phase outlet connection one vapour generator of the oxidation reactor of described oxidation subsystem, the thermal medium outlet of described vapour generator connects the bottom gas phase import of an Acetic Acid Recovering Column, the top gas phase outlet of described Acetic Acid Recovering Column connects the thermal medium inlet of the vapour generator group be composed in series by multiple vapour generator, the thermal medium outlet of vapour generator group connects the import of Acetic Acid Recovering Column return tank, the top gas phase outlet of described Acetic Acid Recovering Column return tank connects described oxidized tail gas processing subsystem, liquid phase primary outlet at the bottom of the tower of described Acetic Acid Recovering Column connects the tower top dewatering period import of described oxidation reactor, the refinement mother liquor outlet of described refining subsystem connects tower top first fluid inlet of described Acetic Acid Recovering Column, described first fluid inlet is positioned at the top of the one piece of tower tray in the top of described Acetic Acid Recovering Column.
7. less energy-consumption material consumption as claimed in claim 6, the PTA production system of few waste discharge, characterized by further comprising solvent recuperation subsystem, described solvent recuperation subsystem comprises extraction dehydration tower, extraction agent recovery tower and extraction agent rectifying tower, liquid phase sub-outlet at the bottom of the tower of described Acetic Acid Recovering Column connects the tower top import of described extraction dehydration tower through water cooler, the tower top oil phase outlet of described extraction dehydration tower connects the opening for feed in the middle part of described extraction agent rectifying tower, aqueous phase outlet at the bottom of the tower of described extraction dehydration tower connects tower top first import of described extraction agent recovery tower, the top gaseous phase outlet of described extraction agent recovery tower connects extraction agent import bottom described extraction dehydration tower through water cooler, aqueous phase outlet at the bottom of the tower of described extraction agent recovery tower connects sewage works, the top gas phase outlet of described extraction agent rectifying tower connects the import of decanting vessel through water cooler, the oil phase outlet of described decanting vessel connects the described extraction agent import bottom described extraction dehydration tower, first aqueous phase outlet of described decanting vessel connects tower top second import of described extraction agent recovery tower, second aqueous phase outlet of described decanting vessel connects the refluxing opening of described extraction agent rectifier, described extraction agent rectifying tower is provided with acetic acid discharge port at the bottom of some towers, extraction agent used is preferably methyl acetate, isopropyl acetate, n-butyl acetate, the mixture of any one or a few the arbitrary proportion in isobutyl acetate.
8. less energy-consumption material consumption as claimed in claim 7, the PTA production system of few waste discharge, it is characterized in that the multiple oxidizing and crystallizing devices of the bottom discharge mouth of described oxidation reactor through connecting successively connect the opening for feed of a rotary pressure filter, described rotary pressure filter is provided with filtration successively, blow dried and filter cake washing three districts, the outlet of filtrating area is connected with the mother liquor reflux mouth bottom described oxidation reactor, the dry gas import in blow dried district connects dried oxidized tail gas, one or more in the air duct of fresh nitrogen or water vapor, filter cake washing district is provided with multistage washing subregion, the washings import of last step washing subregion connects the first bottom liquid phases outlet of described acetic acid reflux tower return tank, in all the other washing subregions at different levels, the washings import of previous stage washing subregion connects the outlet of the wash filtrate of rear stage washing subregion, the filtered solid outlet in described filter cake washing district connects the solid material import of the slurry preparation tank of described refining subsystem, the water inlet of described slurry preparation tank connects the first bottom liquid phases outlet of described acetic acid reflux tower return tank.
9. less energy-consumption material consumption as claimed in claim 8, the PTA production system of few waste discharge, it is characterized in that also being provided with ritalin stripping tower, the water inlet of washings import and described slurry preparation tank that the first bottom liquid phases outlet of described Acetic Acid Recovering Column return tank washs subregion through described ritalin stripping tower and described last step is connected, concrete mode of connection is: the first bottom liquid phases outlet of described Acetic Acid Recovering Column return tank is connected to the top inlet of described ritalin stripping tower, the outlet at bottom of described ritalin stripping tower connects the washing washings import of subregion of described last step and the water inlet of described slurry preparation tank.
10. less energy-consumption material consumption as claimed in claim 9, the PTA production system of few waste discharge, characterized by further comprising oxidation mother liquor processing subsystem, described oxidation mother liquor processing subsystem comprises acetic acid stripping tower, residue extractor, residue filter machine, residue drying machine, decanting vessel, residue extraction agent recovery tower and phenylformic acid rectifying tower, the bottom inlet of described acetic acid stripping tower is as the pending liquid import of described oxidation mother liquor processing subsystem, the outlet at bottom of described acetic acid stripping tower connects the bottom inlet of described residue extractor, described residue extractor is provided with upper water import and top extraction agent import, described upper water import connects the second bottom liquid phases outlet of described Acetic Acid Recovering Column return tank, the outlet at bottom of described residue extractor connects the import of described residue filter machine, the filtered solid outlet of described residue filter machine connects the import of described residue drying machine, the filtrate outlet of described residue filter machine connects the import of described decanting vessel, the aqueous phase outlet of described decanting vessel connects the mother liquor reflux mouth bottom described oxidation reactor, the oil phase outlet of described decanting vessel connects the bottom inlet of described residue extraction agent recovery tower, the top extraction agent outlet of described residue extraction agent recovery tower connects the top extraction agent import of described residue extractor through water cooler, the outlet at bottom of described residue extraction agent recovery tower connects the import of described phenylformic acid rectifying tower, enter acetate concentration in the mother liquor residue slurry of described residue extractor from described acetic acid stripping tower and be preferably 60-85wt%, the extraction agent entering described residue extractor is preferably methyl acetate, n-propyl acetate, isopropyl acetate, n-butyl acetate, the mixture of any one or a few the arbitrary proportion in isobutyl acetate, mother liquor residue slurry described in described residue extractor, the mass ratio of water and extraction agent is preferably 1:0.5-2.0:1.5-4.0, extraction temperature is preferably
40-90℃。
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