CN115260496A

CN115260496A - Process for producing polyarylene sulfide resin, product thereof and use thereof

Info

Publication number: CN115260496A
Application number: CN202210888486.1A
Authority: CN
Inventors: 尹红; 贾艳宇; 张雄伟; 陈志荣; 连明; 陈兴; 蒋杰; 李沃源
Original assignee: Zhejiang University ZJU; Zhejiang NHU Co Ltd; Zhejiang NHU Special Materials Co Ltd
Current assignee: Zhejiang University ZJU; Zhejiang NHU Co Ltd; Zhejiang NHU Special Materials Co Ltd
Priority date: 2022-07-27
Filing date: 2022-07-27
Publication date: 2022-11-01
Anticipated expiration: 2042-07-27
Also published as: WO2024021816A1; CN115260496B

Abstract

The invention discloses a process for preparing polyarylene sulfide, which takes a sulfur source and a dichloro aromatic compound as raw materials, adopts an organic amide solvent, and is prepared by a series of processes such as dehydration, prepolymerization, polymerization, flash evaporation, post-treatment and the like. The salt content in the prepared crude product of the polyarylene sulfide resin can be greatly reduced by the manufacturing process, the generation of high-salt-content wastewater is greatly reduced, the water consumption and the energy consumption are reduced, and the high-molecular-weight polyarylene sulfide resin can be efficiently prepared; more importantly, the process is not only suitable for an intermittent process, but also suitable for continuous production, is expected to greatly improve the production efficiency of the production system, reduces the production cost, and is more environment-friendly and green.

Description

Process for producing polyarylene sulfide resin, product thereof and use thereof

Technical Field

The invention relates to the field of synthesis of polyarylene sulfide resin, in particular to a continuous and batch type manufacturing process of polyarylene sulfide resin, a product and application thereof.

Background

A typical production method of polyarylene sulfide (PAS) is a polymerization reaction of sodium sulfide/sodium hydrosulfide and a dihalo aromatic compound in an organic amide solvent in the presence of an alkali metal hydroxide, and the solvent is preferably N-methylpyrrolidone (hereinafter, abbreviated as NMP) in general, and the alkali metal hydroxide is preferably sodium hydroxide in general. At present, the industrial production device for synthesizing polyphenylene sulfide (PPS) by the method is operated intermittently, partial reaction raw materials are firstly subjected to dehydration reaction in a reaction kettle, after the rest reaction raw materials are added after being cooled, the reaction kettle is sealed, the temperature is raised for high-temperature polycondensation, a crude product is obtained after the reaction is finished through cooling filtration or flash evaporation, and the crude product is purified to obtain a final product. In the intermittent production process, the reaction kettle needs to be repeatedly charged, heated and cooled, discharged and cleaned, the consumed time is long, the production efficiency is greatly reduced, and the process stability and the reliability are poor. Therefore, researchers have been working on developing continuous production processes for PAS to improve production efficiency.

The patents US4060520, US4056515 and US4066632 propose a continuous production process with multiple kettles connected in series, wherein the reaction temperature of each kettle is different, the kettle pressure is continuously reduced from front to back, and the materials are pushed to flow from front to back by the pressure difference between the adjacent kettles. Chinese patent publication No. CN 108779253A provides a continuous production apparatus and a continuous production method for PAS, which includes a housing chamber housing a plurality of reaction tanks, supplies an organic amide solvent, a sulfur source, and a dihalo aromatic compound into the housing chamber, and forms a reaction mixture by performing a polymerization reaction of the sulfur source and the dihalo aromatic compound in the organic amide solvent; the reaction tanks are connected in series and communicate with each other via a gas phase, and the reaction mixture moves in series to each reaction tank. Chinese patent publication No. CN 113667122a provides a gradient temperature-controlled continuous condensation method of polyarylene sulfide resin, which performs step control (gradual temperature rise) on the temperature of each reactor to adapt to different progress of polymerization reaction, and the reaction raw materials are continuously fed into a series-connected tank-type or tubular reactor through a metering pump.

The production of PPS by the sodium sulfide method described above generates a large amount of sodium chloride as a by-product, and 1.08 tons of sodium chloride are generated per one ton of PPS produced, and sodium chloride starts to be generated and precipitated in a large amount in the early stage of the polymerization reaction, so that the reaction slurry contains a large amount of salt particles, and the by-product salt particles are contained in the polymerization reaction liquid at the end of the reaction, in a typical value where the molar ratio of NMP/sulfur source in the reaction liquid is 3.5%. Large quantities of salt particles in the slurry tend to settle and plug in long distance transport or reaction lines, especially in tubular reactors of relatively fine size, resulting in production shutdowns. In both batch and continuous processes, the increase of the concentration of the reaction substrate is beneficial to the improvement of the production efficiency, and the high concentration of the reaction substrate is also beneficial to the efficient production of the polyphenylene sulfide with high molecular weight, but the content of the byproduct salt particles is also obviously increased, the viscosity of the reaction solution is also sharply increased, and the problems of conveying blockage and stirring are further worsened. The presence of a large amount of by-product salt particles in the PPS reaction slurry severely limits the research and development of the continuous PPS production process and limits the efficient production of higher molecular weight polyphenylene sulfide.

In both batch and continuous manufacturing processes, the byproduct sodium chloride is wrapped in the product after the reaction is finished, a large amount of pure water is needed to be used for washing and removing for many times, and a large amount of washing wastewater with high salt content is generated. In addition, the high corrosiveness of the brine also puts high demands on the material of the polyphenylene sulfide manufacturing equipment. The manufacturing process of the polyphenylene sulfide has the characteristics of high water consumption, high energy consumption and high corrosivity, greatly improves the manufacturing cost and the environmental protection cost of the polyphenylene sulfide, seriously limits the greenness and the sustainability of the production process of the polyphenylene sulfide, and limits the competitiveness and the vitality of the polyphenylene sulfide resin.

The aforementioned problems can be solved by removing precipitated salt particles before solid-liquid separation of the product to reduce the salt particle content in the subsequent equipment and in the crude product. For example, chinese patent publication No. CN104371103A discloses a method for desalting polyphenylene sulfide, which comprises, after polymerization reaction, maintaining temperature and pressure in a reaction kettle, standing, after inorganic salts and inorganic additives settle to the bottom of the kettle, pressing the upper polyphenylene sulfide solution out of the reaction kettle through a bottom internal extension pipe by using pressure difference, cooling the solution, and purifying the precipitated polyphenylene sulfide to obtain a high-purity polyphenylene sulfide product. However, the process described in this patent is only applicable to batch processes, and the inorganic salts, the large amount of reaction solvent and the products in the settled layer in the reactor are further processed, which is not further described in this patent document, and in fact the settled inorganic salt layer makes it impossible to start the polymerization vessel agitation again before it is not completely cleaned.

Also, as disclosed in chinese patent publication No. CN107964098a, it relates to a method and apparatus for desalting in the process of synthesizing polyphenylene sulfide, during the polymerization process, the material in the reaction kettle passes through a filter equipped with a filter screen by a feed transfer pump, the salt is retained in the filter, the residual polyphenylene sulfide material returns to the reaction kettle, the process is circulated until the reaction is completed, and the polyphenylene sulfide resin is obtained by discharging, washing and drying. The method is also only suitable for batch processes, and in the later stage of the reaction, the filtration becomes more difficult as the thickness of the salt cake layer and the viscosity of the reaction solution increase.

Disclosure of Invention

Aiming at the problems in the prior art, the invention discloses a manufacturing process of polyarylene sulfide resin, which can greatly reduce the salt content in the prepared crude polyarylene sulfide resin, greatly reduce the generation of high-salt wastewater, reduce water consumption and energy consumption, and is beneficial to preparing high-molecular-weight polyarylene sulfide resin; more importantly, the process is not only suitable for an intermittent process, but also suitable for continuous production, is expected to greatly improve the production efficiency of the production system, reduces the production cost, and is more environment-friendly and green.

The specific technical scheme is as follows:

a method for continuously producing a polyarylene sulfide, comprising:

(1) Putting a sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an optionally added auxiliary agent and an organic amide solvent into a reaction kettle, continuously heating to 200-220 ℃ for dehydration and saponification reaction, cooling to 150-180 ℃ after dehydration until the molar ratio of water to the sulfur source is 1.5-2.0, and keeping the obtained dehydrated liquid in a storage kettle;

(2) Injecting organic amide crystal slurry containing sodium chloride crystal seeds, dichloro aromatic compound and the dehydration solution into a prepolymerization reactor, controlling the reaction temperature to be 180-240 ℃, injecting the obtained reaction solution into a filtering device after the reaction is carried out until the monomer conversion rate reaches 90% or more, and filtering to remove sodium chloride particles;

(3) And injecting the filtrate obtained after the filtration into a polymerization reactor, controlling the reaction temperature to be 235-280 ℃, reacting until the monomer conversion rate reaches 95% or above, carrying out flash evaporation treatment on the obtained reaction liquid, and finally carrying out post-treatment to obtain the polyarylene sulfide.

The inventor carries out intensive research on the growth process and the particle morphology of salt particles in PPS reaction liquid, after all raw materials are added and heated to a certain temperature, the polymerization reaction starts, sodium chloride is continuously generated, sodium chloride as a byproduct can be dissolved in a reaction solvent in a small amount due to the existence of water in a polyphenylene sulfide reaction system, when the solution is saturated, the generated sodium chloride starts to nucleate and separate out, the reaction continues to be carried out rapidly, and a large amount of generated sodium chloride is rapidly nucleated and separated out. The research finds that: (1) In the previous typical batch PPS (polyphenylene sulfide) manufacturing process, the particle size of salt particles generated in a reaction liquid is small and about 20 mu m, which is not beneficial to filtration, and the fact that the particle size is not good is reported; (2) In the pre-polymerization reaction section, the reaction conversion rate is over 90 percent, the growth of sodium chloride crystals is basically completed at the moment, but the molecular weight of the product is still lower and the viscosity of the reaction solution is not high at the moment, and the product needs to enter a high-temperature polymerization section for continuous reaction to improve the molecular weight, so that the filtration and removal of byproduct salt are more suitable when the growth of the sodium chloride crystals is basically completed in the pre-polymerization reaction section, and the viscosity of the reaction solution is higher after the high-temperature polymerization section, so that the filtration is more difficult; (3) The water content in the system is controlled in a proper range, so that the precipitation nucleation speed of sodium chloride can be reduced, and the growth speed of sodium chloride crystals can be increased; (4) The introduction of a small amount of sodium chloride seed crystals can effectively inhibit the primary nucleation of sodium chloride.

The technical scheme of the invention is provided based on the research results, and experiments show that in the step (1), in order to control the particle size of sodium chloride particles to be more than 100 mu m so as to be beneficial to filtration, the water content needs to be controlled under a proper proportion (the molar ratio of water to a sulfur source is controlled to be 1.5-2.0) after dehydration is finished; because the sodium chloride can be filtered and removed, the generation of high-salt-content wastewater is greatly reduced, the water consumption and the energy consumption are reduced, and meanwhile, the using amount of the solvent in the reaction liquid can be further reduced, so that the high-molecular-weight PPS can be prepared efficiently.

In the step (1):

the concentration of the sulfur source water solution is 28-48 wt%, and the sulfur source is selected from sodium hydrosulfide and/or sodium sulfide;

the alkali metal hydroxide aqueous solution is sodium hydroxide aqueous solution, and the concentration is 30-70 wt%;

the dichloro aromatic compound is selected from one or more of p-dichlorobenzene, dichloronaphthalene, dichlorofluorene and dichlorocarbazole, and preferably p-dichlorobenzene;

the molar ratio of the sulfur source to the sodium hydroxide is 1:0.95 to 1.1;

the auxiliary agent is selected from one or more of sodium acetate, sodium benzoate and C5-C6 sodium aliphatate;

the molar ratio of the sulfur source to the auxiliary agent is 1:0 to 0.5;

the organic amide solvent is selected from one or more of N-methyl pyrrolidone, hexamethyl phosphoric triamide, N-methyl-epsilon-caprolactam and N, N-dimethylformamide, and is preferably N-methyl pyrrolidone;

the molar ratio of the sulfur source to the organic amide solvent is 1:1.5 to 2.5.

The dehydration saponification reaction can be carried out according to the method of the step (1) in batch mode, and also can adopt continuous dehydration schemes reported by other patents, and the key is to control the water content in the proper range after the dehydration is finished.

In the step (2):

the concentration of the sodium chloride crystal seeds in the organic amide crystal slurry is 0.1-1 mol/L, and the particle size of the sodium chloride crystal seeds is 1-20 mu m;

the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydrated liquid is 1.0-1.1: 1; preferably 1.0 to 1.08:1.

after the dehydrated liquid, the dichloro aromatic compound and the organic amide crystal slurry containing the sodium chloride crystal seeds are mixed, the adding amount of the sodium chloride crystal seeds is controlled to be 0.05-0.4 mol% based on 1mol of a sulfur source in a prepolymerization reactor system, and the total amount of an organic amide solvent in the system is controlled to be 2.0-4.0 mol.

Preferably, the total amount of the organic amide solvent is 2.0 to 2.5mol based on 1mol of the sulfur source in the prepolymerization reactor system at this time (i.e., NMP/S =2.0 to 2.5 in the system is controlled).

Experiments show that the optimization not only can reduce the flash energy consumption generated by the production process and the amount of waste liquid to be recovered, but also can properly improve the weight average molecular weight of the PAS prepared.

More preferably, the molar ratio of the dichloroaromatic compound to the sulfur source is controlled to be 1.0 to 1.025:1.

it was found through experiments that when the molar ratio of the dichloroaromatic compound to the sulfur source and the NMP/S ratio are both controlled within the above-mentioned further preferable range, the weight average molecular weight of PAS produced can be significantly increased.

The inventors have conducted intensive studies on the polymerization kinetics of polyphenylene sulfide and found that: (1) The main polymerization reaction of the polyarylene sulfide is a strong exothermic reaction, the temperature runaway in the prepolymerization process can seriously affect the reaction safety and the product quality, and the polymerization reaction is a secondary kinetic reaction, so that in order to meet the requirements of reaction temperature control and reaction efficiency, for a continuous manufacturing process with a plurality of reactors connected in series, the first reactor is preferably a continuous stirred tank reactor and the conversion rate is controlled in a proper range; (2) Increasing the water content in the reaction system can reduce the reaction rate, i.e., the generation rate of sodium chloride, but too high water content can cause problems of too low reaction efficiency, high reaction pressure, higher energy consumption in the flash evaporation process, and the like.

By combining the research results, the method aims at various problems caused by a large amount of byproduct salt in the existing production process of the PAS intermittent manufacturing process and the development process of the continuous manufacturing process: (1) Washing to remove byproduct salt, which results in high water consumption and high salt-containing wastewater discharge; (2) The problem of easy blockage limits the development of the PAS continuous manufacturing process; (3) limiting the increase in substrate concentration; and the main polymerization reaction is a secondary dynamic strong exothermic reaction, so that the temperature control in the reaction process is high, and the invention provides a PAS continuous manufacturing method which is more environment-friendly, green, low in cost and efficient.

Preferably, the following components:

in the step (2), the pre-polymerization reactor is a two-stage reactor connected in series;

in order to control the reaction temperature and the reaction efficiency, the first-stage reactor is selected from a continuous stirring tank type reactor, and the reaction temperature in the continuous stirring tank type reactor is controlled to be 180-220 ℃, preferably 190-210 ℃.

After the monomer conversion rate in the first reactor reaches 60 to 80%, more preferably 70 to 80%, the reaction solution is injected into the second reactor.

The second stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled to be 220-240 ℃.

Preferably, the heating medium adopted by the jacketed tubular reactor is heat conducting oil, the flow direction of the heat conducting oil is consistent with the flow direction of the reactant, the temperature of the inlet oil is controlled to be 220-230 ℃, and the temperature of the outlet oil is controlled to be 230-240 ℃; preferably, the monomer conversion at the outlet is more than 90%, more preferably more than 93%, at this time, the growth of sodium chloride crystals is substantially completed, the particle size is more than 60 μm, preferably more than 100 μm, and the size is uniform.

The filtering device is a continuous filtering device, is specifically selected from a closed type pressure filtering device, is preferably a pressure rotary drum filter, and is suitable for continuous filtering of high-temperature slurry containing volatile components and easily separated crystals.

Preferably, the pore size of the filter medium in the filter device is 60 to 80 μm. Sodium chloride particles of this size can be sufficiently removed without the pressure drop of the filtration being too high.

In the step (3), the step (c),

the polymerization reactor is selected from a tubular reactor, preferably, the tubular reactor is selected from a coil tubular reactor immersed in a high-temperature oil bath, the reaction temperature is controlled to be 235-280 ℃, further molecular weight increase is completed, and the monomer conversion rate is ensured to be more than 95%, preferably more than 98%.

Preferably, the flash evaporation treatment uses superheated steam with the temperature of 260-300 ℃ for auxiliary heat supply, the using amount of the superheated steam is 0.1-1 kg/mol of the sulfur source, the steam and the vaporized solvent are extracted from the upper part of the flash evaporator, the steam and the vaporized solvent enter a solvent recovery system after being condensed, and the crude products of the salt-containing PAS are continuously discharged from the lower part.

The post-treatment comprises drying, washing, filtering and re-drying, and specifically comprises the following steps:

the crude salt-containing PAS product is continuously dried by a dryer to further remove residual solvent.

Adding water into the dried salt-containing PAS crude product, pulping, and then washing with water at high temperature and high pressure continuously for one time, wherein the mass ratio of the water addition amount to the crude product is 3-10: and 1, continuously filtering and drying to obtain a final product.

The continuous high-temperature high-pressure water washing equipment can be continuous kettle type washing equipment or continuous tube type washing equipment; preferably continuous tubular washing equipment with a heat conducting oil jacket, wherein the retention time of the water washing slurry is preferably 1-30 min, and the water washing temperature is preferably 150-220 ℃.

The continuous filtration equipment is selected from continuous filters, such as plate-frame filters, decanter centrifuges, belt filters and the like, and is preferably a continuous vacuum filtration belt filter.

The dryer is selected from a continuous feeding and discharging dryer.

By adopting the preparation process, because large-particle salt is filtered and removed in advance, the salt concentration in the subsequent crude product is greatly reduced, and the pressure in the washing process in the post-treatment is obviously reduced, so that only one-time continuous high-temperature high-pressure washing is needed; if salt particles are not filtered in advance, sodium chloride in the product can be sufficiently washed off only by a plurality of times of normal-pressure water washing and high-pressure water washing, so that the water consumption for washing and the amount of the generated salt-containing wastewater are greatly increased; if no seed crystal is added and the water content in the reaction liquid is controlled at a proper level, the sodium chloride particles are small in size and difficult to effectively intercept by adopting large-aperture filter materials, and the filter is easy to block due to overlarge filter pressure drop by adopting small-aperture filter materials.

The above process is also applicable to the batch production of polyarylene sulfide, and the influence of various process parameters on the properties of PAS prepared by batch production is similar to that in continuous production, and will not be described in detail below.

Accordingly, the present invention also discloses a method for batchwise producing a polyarylene sulfide, comprising:

(a) Putting a sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an optionally added auxiliary agent and an organic amide solvent into a reaction kettle I, continuously heating to 200-220 ℃ for dehydration and saponification reaction, wherein the molar ratio of water to the sulfur source is 1.5-2.0 after dehydration, and then cooling to 150-180 ℃;

(b) Putting organic amide crystal slurry containing sodium chloride crystal seeds, dichloro aromatic compound and supplemented organic amide solvent into a reaction kettle I, firstly heating to 205-220 ℃, preserving heat for 0.5-3 h, then heating to 215-240 ℃, preserving heat for 0.5-5 h, then injecting the obtained reaction liquid into a filtering device, and filtering to remove sodium chloride particles;

(c) And (3) injecting the filtrate obtained after the filtration into a preheated reaction kettle II, heating to 240-280 ℃, preserving the temperature for 0.5-5 h, carrying out flash evaporation treatment on the obtained reaction liquid, and finally carrying out post-treatment to obtain the polyarylene sulfide.

In step (a):

the dichloro aromatic compound is selected from one or more of p-dichlorobenzene, dichloronaphthalene, dichlorofluorene and dichlorocarbazole, and preferably p-dichlorobenzene; the molar ratio of the sulfur source to the sodium hydroxide is 1:0.95 to 1.1;

the molar ratio of the sulfur source to the auxiliary agent is 1:0 to 0.5;

the organic amide solvent is selected from one or more of N-methyl pyrrolidone, hexamethyl phosphoric triamide, N-methyl-epsilon-caprolactam and N, N-dimethylformamide;

the molar ratio of the sulfur source to the organic amide solvent is 1:1.5 to 3.

In step (b):

the concentration of the sodium chloride crystal seeds in the organic amide crystal slurry is 0.1-1 mol/L, and the particle size of the sodium chloride crystal seeds is 1-20 mu m; the addition amount of the sodium chloride seed crystal is 0.05-0.4 mol% of the sulfur source in the system;

the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration solution is 1.0-1.1: 1;

after the organic amide solvent is supplemented, in the reaction kettle I, the molar ratio of the total molar amount of the organic amide solvent to the sulfur source is 2.5-3.5;

preferably:

the temperature is raised to 205-220 ℃ at the heating rate of 0.2-0.5 ℃/min, and then is raised to 215-240 ℃ at the heating rate of 0.2-1 ℃/min.

The reaction kettle used in the steps (a) and (b) is selected from a stirring type reaction kettle with a heat-conducting oil jacket and a coil pipe.

In step (c):

preferably, the temperature is raised to 240-280 ℃ at a rate of 0.2-1 ℃/min.

Preferably, after the flash evaporation is finished, continuously drying in the flash evaporator for 10 min-1 h, then cooling to room temperature, and discharging.

The post-treatment comprises drying, washing, filtering and re-drying.

The present invention also discloses polyarylene sulfide prepared according to the above two processes, respectively.

The invention also discloses application of the polyarylene sulfide in preparing crosslinked polyarylene sulfide, in particular to the preparation of crosslinked polyarylene sulfide by carrying out thermal oxidation treatment on the polyarylene sulfide.

Compared with the prior art, the invention has the following beneficial effects:

the invention discloses a process for preparing polyarylene sulfide resin, which is characterized in that sodium chloride seed crystal is added into PAS reaction liquid, the water content in the system is controlled within a proper range, the reaction temperature of a prepolymerization section is controlled at a lower temperature, so that the nucleation speed of sodium chloride is reduced, the growth speed of crystals is increased, the purposes of inhibiting primary nucleation and promoting secondary nucleation are achieved, and the size of sodium chloride crystals as byproducts is increased finally; the other key point of the process is the time for filtering and removing the salt particles, and the filtering selection of the salt particles is carried out at the end of the prepolymerization reaction stage; the manufacturing process disclosed by the invention can greatly reduce the salt content in the prepared polyarylene sulfide resin crude product, greatly reduce the generation of high-salt-content washing wastewater, reduce water consumption and energy consumption, and is beneficial to efficiently preparing the polyarylene sulfide resin with high molecular weight; more importantly, the process is not only suitable for an intermittent process, but also suitable for continuous production, is expected to greatly reduce the production efficiency and the production cost of the production system, and is more environment-friendly and green. When a continuous production process is adopted, in order to take account of the temperature control and the production efficiency of a prepolymerization section, the prepolymerization section is divided into two sections, the first section adopts a kettle type reactor suitable for stable heat transfer, the conversion rate is controlled to be a proper value so as to remove most of reaction heat, but the reaction rate is not too low, the second section adopts a jacketed pipe type reactor, and the residual heat release is utilized to push the reaction to continue to heat up so as to maintain the reaction efficiency; the high-temperature polymerization section is more suitable for a coil type reactor with high conversion rate and high viscosity state (high molecular weight and high concentration) after salt particles are removed.

Drawings

FIG. 1 is a schematic view showing a process flow of the continuous production of polyarylene sulfide according to the present invention, in which a post-treatment process is omitted;

in the figure, 1-dichlorobenzene, 2-dehydration liquid, 3-NMP crystal slurry containing sodium chloride crystal seeds, 4-a first polymerization reaction kettle, 5-a first reaction mixed liquid, 6-a first heat exchanger, 7-a second tubular reactor, 8-a closed pressure filter, 9-salt particles, 10-a second reaction mixed liquid, 11-a second heat exchanger, 12-a third tubular reactor, 13-superheated steam, 14-an atmospheric flash evaporator, 15-a non-dried crude product, 16-a dryer feeding bin, 17-a continuous dryer, 18-a dried crude product, 19-vaporized solvent and steam, 20-a condenser and 21-waste liquid to be recovered.

Detailed Description

The present invention is described in further detail below with reference to examples and comparative examples, but the embodiments of the present invention are not limited thereto.

Ash content test: and (3) placing the ceramic crucible into a muffle furnace with the constant temperature of 750 ℃ to be burnt to constant weight, taking out the crucible, placing the crucible in a dryer to be cooled, and weighing the crucible as M0. 3g of the sample was weighed into a crucible, 10mL of nitric acid was added, and the mixture was placed on an alcohol burner and burned until no smoke was emitted. And finally, placing the crucible in a muffle furnace with the constant temperature of 750 ℃ for burning for 1h, taking out the crucible, placing the crucible in a dryer for cooling, and weighing the crucible as M1. The calculation formula of the resin ash is as follows: (M1-M0)/3 x 100%.

And (3) testing molecular weight: the molecular weight of PPS is measured by using a gel permeation chromatograph and polystyrene as a standard sample, the mobile phase is 1-chloronaphthalene, the column temperature is 220 ℃, the flow speed is 1mL/mL, and the detector is a refractive index detector.

Example 1

19.82kg (200.0 mol) of N-methyl pyrrolidone, 8.33kg (100.0 mol of NaOH) of 48wt% sodium hydroxide aqueous solution and 11.93kg (100.0 mol of NaHS) of 47wt% sodium hydrosulfide aqueous solution are added into a 100L reaction kettle I, the temperature is slowly raised to 200 ℃ under the protection of nitrogen, water is continuously removed in the process, 10.00kg (98.0% of water content) of distillate is discharged, and the temperature is lowered to 170 ℃ after dehydration is finished. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 147.0mol.

Adding 15.14kg (103.0 mol) of p-dichlorobenzene, 10kg of NMP crystal slurry containing 11.7g of sodium chloride crystal seeds (0.2 mol, the D50 of the sodium chloride crystal seeds is 15 mu m) into a reaction kettle I, then adding 4.46kg of NMP (45.0 mol), sealing the reaction kettle I, heating to 215 ℃ at the speed of 0.3 ℃/min, preserving heat for 3h, heating to 230 ℃ at the speed of 0.2 ℃/min, preserving heat for 1h, filtering reaction liquid through a filter (the diameter of a filter screen is 80 mu m) to remove sodium chloride particles, pumping the reaction liquid into a 100L reaction kettle II which is preheated to 230 ℃, heating the reaction kettle II to 260 ℃ at the speed of 0.5 ℃/min, preserving heat for 1h, then slowly discharging to a normal-pressure flash evaporator for flash evaporation treatment, finally drying to further remove the solvent, obtaining 10.4kg of a crude product, adding 60kg of water for pulping, heating to 180 ℃, preserving heat for 0.5h, reducing the temperature, filtering, and drying to obtain a final product.

In this example, 10.9kg of sodium chloride was obtained when the reaction solution was transferred and filtered, and the D50 of the sodium chloride particles was 130 μm as measured by a laser particle size analyzer; 60.0kg of water is used in the washing process, 50.0kg of salt-containing wastewater is generated, and the salt content is 0.6wt%; the final product was prepared at 10.1kg, weight average molecular weight 22500 and ash at 0.53wt%.

Comparative example 1

Adding 15.14kg (103.0 mol) of p-dichlorobenzene and 14.46kg of NMP into a reaction kettle I, sealing the reaction kettle I, heating to 215 ℃ at the speed of 0.3 ℃/min, preserving heat for 3h, heating to 230 ℃ at the speed of 0.2 ℃/min, preserving heat for 1h, filtering reaction liquid by a filter (the diameter of a filter screen is 20 mu m), removing sodium chloride particles, pumping into a 100L reaction kettle II which is preheated to 230 ℃, heating to 260 ℃ at the speed of 0.5 ℃/min, preserving heat for 1h, slowly discharging into a normal pressure flash evaporator for flash evaporation treatment, drying to further remove the solvent to obtain 10.8kg of a crude product, adding 60kg of water, pulping, heating to 180 ℃, preserving heat for 0.5h, cooling, filtering and drying to obtain a final product

In this comparative example, 10.5kg of sodium chloride was obtained when the reaction solution was transferred and filtered, and the D50 of the sodium chloride particles was 30 μm as measured by a laser particle size analyzer; 60.0kg of water is used in the washing process, 51.0kg of salt-containing wastewater is generated, and the salt content is 1.3wt%; the final product was prepared at 10.1kg, weight average molecular weight 22500 and ash at 0.56wt%.

It should be noted that, because the particle size of the salt and the aperture of the filter screen matched with the particle size are smaller, the pressure drop of the filter is high, the flux is small, a large amount of time is spent in the filtering step, the production efficiency is greatly reduced, and the production cost is obviously increased.

Comparative example 2

Adding 15.14kg (103.0 mol) of p-dichlorobenzene and 10kg of NMP crystal slurry containing 11.7g of sodium chloride crystal seeds (0.2 mol, the D50 of the sodium chloride crystal seeds is 15 mu m) into a reaction kettle I, supplementing 4.46kg of NMP (45.0 mol), sealing the reaction kettle, heating to 215 ℃ at the rate of 0.3 ℃/min, preserving heat for 3h, heating to 230 ℃ at the rate of 0.15 ℃/min, preserving heat for 1h, heating to 260 ℃ at the rate of 0.5 ℃/min, preserving heat for 1h, slowly discharging to a normal-pressure flash evaporator for flash evaporation treatment, drying, and further removing the solvent to obtain 21.2kg of a crude product. The crude product is washed and filtered under normal pressure twice, the water addition amount is 60.0kg each time, 60.0kg of water is added for pulping, the temperature is raised to 180 ℃, the temperature is kept for 0.5h, the temperature is lowered, the filtration is carried out, and the final product is obtained after the filtration and the drying.

In this comparative example, 180.0kg of water was used in the washing process, yielding 182.0kg of salt-containing wastewater with a salt content of 6.0wt%; the final product was prepared at 10.0kg, weight average molecular weight 21500, ash 0.50wt%.

Comparison of example 1 with comparative example 2 shows that removal of salt particles by filtration can significantly reduce the amount of water used for washing and the amount of waste salt water generated.

Example 2

19.82kg (200.0 mol) of N-methyl pyrrolidone, 8.33kg (100.0 mol of NaOH) of 48wt% sodium hydroxide aqueous solution and 11.93kg (100.0 mol of NaHS) of 47wt% sodium hydrosulfide aqueous solution are added into a 100L reaction kettle, the temperature is slowly raised to 200 ℃ under the protection of nitrogen, water is continuously removed in the process, 10.00kg (98.0% of water content) of distillate is discharged, and the temperature is lowered to 170 ℃. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 147.0mol. The dehydration liquid prepared by the method in batches is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

The PAS continuous manufacturing process flow diagram shown in fig. 1 is used. Pumping paradichlorobenzene 1 (PDCB) preheated to 200 ℃, a dehydration solution 2 (with the content of sodium sulfide being 25.48 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the D50 of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.136 mol/L) to a first polymerization reaction kettle 4 (the molar ratio of PDCB/S =1.050 in the first polymerization reaction kettle 4 is controlled, the same holds hereinafter), NMP/S =3.50 (the molar ratio is also held hereinafter), the temperature of the first polymerization reaction kettle 4 is 215 ℃, the retention time is 5h, the obtained first reaction mixed solution 5 is heated to 225 ℃ through a first heat exchanger 6, enters a second tubular reactor 7 (the temperature of the reactor is 225 ℃), the reaction mixed solution is fed into a pressure filter 8 after the retention time is 1h, salt particles are removed through a closed pressure filter 8 (the aperture of the filter is 80 mu m), and the D50 of the filtered salt particles 9 is 140 mu m; heating the obtained second reaction mixed solution 10 to 265 ℃ through a second heat exchanger 11, then entering a third tubular reactor 12 (the temperature of the reactor is 265 ℃), keeping the time for 1.5h, then exchanging heat to 280 ℃ through a heat exchanger (not shown), entering an atmospheric flash evaporator 14 for flash evaporation, using 265 ℃ superheated steam 13 for auxiliary flash evaporation during flash evaporation, wherein the flow of the superheated steam 13 is 5.0kg/h, condensing and liquefying vaporized solvent and steam 19 through a condenser 20 to generate waste liquid 21 to be recovered, entering an undried crude product 15 into a continuous dryer 17 through a dryer feeding bin 16, continuously discharging the crude product 18 obtained after further drying into a pulping kettle, adding water for pulping, wherein the water amount is 6.0kg/h, exchanging heat to 180 ℃, entering a tubular washer with a heat-conducting oil jacket, controlling the temperature of the tubular washer at 180 ℃, keeping the tubular washer for 10min, then cooling to 60 ℃ through heat exchange, and continuously filtering and drying (not shown) to obtain a final product.

The production rate of this example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.6wt%, 5.0kg/h of superheated steam is consumed, and 8.7kg/h of waste liquid to be recovered is generated. The final product was prepared with a weight average molecular weight of about 23000 and ash content of less than 0.55wt%.

Example 3

The dehydrated liquid was prepared in batches as in example 2 and placed in a holding tank and held at 170 ℃ for subsequent continuous feeding. The PAS continuous production apparatus shown in FIG. 1 was used. Pumping paradichlorobenzene 1 (PDCB) preheated to 200 ℃, a dehydration solution 2 (with the content of sodium sulfide being 25.48 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the D50 of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.4 mol/L) to a first polymerization reaction kettle 4 (PDCB/S =1.050, NMP/S = 2.5) by flow rates of 25.21g/min, 50g/min and 8.1g/min in sequence, wherein the temperature of the first polymerization reaction kettle 4 is 215 ℃, the retention time is 3.5h, the obtained mixed solution of a first reaction 5 is heated to 230 ℃ through a first heat exchanger 6, the mixed solution enters a second tubular reactor 7 (the temperature of the reactor is 230 ℃), the mixed solution is introduced into a closed type pressure filter 8 after the retention time is 1h, salt particles are removed through the closed type pressure filter 8 (the aperture of a filter cloth is 80 mu m), and the D50 of the filtered salt particles 9 is 130 mu m; the obtained second reaction mixed solution 10 is heated to 265 ℃ through a second heat exchanger 11, enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), stays for 1h, then is subjected to heat exchange through a heat exchanger (not shown) to 280 ℃, enters an atmospheric flash evaporator 14 for flash evaporation, auxiliary flash evaporation is carried out by using 265 ℃ superheated steam 13 during flash evaporation, the flow of the superheated steam 13 is 3.5kg/h, and the vaporized solvent and steam 19 are condensed and liquefied through a condenser 20 to generate a waste liquid 21 to be recovered. The undried crude product 15 firstly enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, added with water for pulping, the added water amount is 6.0kg/h, enters a tubular washer with a heat conduction oil jacket after heat exchange to 180 ℃, the temperature of the tubular washer is controlled to be 180 ℃, the retention time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and the final product is obtained through continuous filtration and drying (not shown).

The production rate of this example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.6wt%, 3.5kg/h of superheated steam is consumed, and 6.2kg/h of waste liquid to be recovered is generated. The final product was prepared with a weight average molecular weight of about 26000 and ash content of less than 0.55wt%.

The records of the comparative example 2 and the example 3 show that the molar ratio of NMP/S in the prepolymerization stage, specifically in the first polymerization reaction kettle, can be further reduced, so that the solvent dosage and the dosage of superheated steam in the flash evaporation stage can be reduced, the energy consumption and the waste liquid amount can be reduced, and the production cost can be remarkably reduced; the reaction efficiency can be improved, and the retention time of the reaction liquid in each section of reactor can be reduced; the weight average molecular weight of the final product can also be suitably increased.

However, the decrease in the NMP/S molar ratio means that the concentrations of PPS and sodium chloride in the system are increased, and the viscosity of the reaction solution rapidly increases with the concentrations of PPS and sodium chloride in the system and the increase in the molecular weight of PPS, and it is difficult to smoothly carry out the production at a low NMP/S ratio by the production process of the prior art, either batch-type or continuous-type. This indicates that the production process disclosed in the present invention is highly adaptable and particularly suitable for the production conditions of low NMP/S molar ratio, and thus more advantageous for the efficient preparation of high molecular weight PAS.

Example 4

19.82kg (200.0 mol) of N-methyl pyrrolidone, 8.33kg (the mol number of sodium hydroxide is 100.0 mol) of 48wt% sodium hydroxide aqueous solution, 11.93kg (the mol number of sodium hydrosulfide is 100.0 mol) of 47wt% sodium hydrosulfide aqueous solution and 1.6kg of anhydrous sodium acetate are added into a 100L reaction kettle, then the temperature is continuously raised under the protection of nitrogen for dehydration, 10.00kg of aqueous solution (the water content is 98.0 wt%) is removed, and the temperature is reduced to 170 ℃ after the dehydration is finished. At this time, the amount of sulfur in the dehydrated liquid system was 98.0mol, and the water content was 147.0mol. The dehydration liquid prepared by the method in batches is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

The PAS continuous production apparatus shown in FIG. 1 was used. Preheated p-dichlorobenzene 1 (PDCB), dehydrated liquid 2 (with the sodium sulfide content of 24.22 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the particle size of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.4 mol/L) are pumped to a first polymerization reaction kettle 4 (PDCB/S =1.025, NMP/S =2.5, CH) at flow rates of 24.61g/min, 52.6g/min and 8.1g/min in sequence₃COONa/S = 0.2), the temperature of the first polymerization reactor is 215 ℃, the retention time is 3.5 hours, the obtained first reaction mixed solution 5 is heated to 230 ℃ through the first heat exchanger 6, enters the second tubular reactor 7 (the temperature of the reactor is 230 ℃), is led into the closed pressure filter 8 after the retention time is 1 hour, salt particles are removed through the closed pressure filter 8 (the aperture of the filter screen is 80 mu m), and the D50 of the filtered salt particles 9 is 120 mu m; the obtained second reaction mixed solution 10 is heated to 265 ℃ by a second heat exchanger 11, enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), stays for 1h, is subjected to heat exchange by a heat exchanger (not shown) to 280 ℃, enters an atmospheric flash evaporator 14 for flash evaporation, and is subjected to auxiliary flash evaporation by using 265 ℃ superheated steam 13 during flash evaporation, the flow of the superheated steam 13 is 3.3kg/h, and the vaporized solvent and steam 19 are condensed and liquefied by a condenser 20 to generate a waste liquid 21 to be recovered. The undried crude product 15 firstly enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water addition amount is 6.0kg/h, the crude product enters a tubular washer with a heat conduction oil jacket after heat exchange to 180 ℃, the temperature of the tubular washer is controlled to be 180 ℃, the retention time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and continuous filtration and drying (not shown) are carried out, so that the final product is obtained.

The production rate of this example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.65wt%, 3.3kg/h of superheated steam is consumed, and 6.0kg/h of waste liquid to be recovered is generated. The final product was prepared with a weight average molecular weight of about 68000 and an ash content of less than 0.5wt%.

Example 5

In a 100L reaction kettle, 19.82kg (200.0 mol) of N-methylpyrrolidone, 8.33kg (100.0 mol of NaOH) of 48wt% sodium hydroxide aqueous solution, 11.93kg (100.0 mol of NaHS) of 47wt% sodium hydrosulfide aqueous solution are added, the temperature is slowly raised to 195 ℃ under the protection of nitrogen, water is continuously removed in the process, 9.07kg (98.0% of water content) of distillate is obtained, and the temperature is lowered to 170 ℃ after the dehydration is finished. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 196.0mol. The dehydration liquid prepared by the method in batches is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

The PAS continuous manufacturing process flow diagram shown in fig. 1 is used. Pumping p-dichlorobenzene 1 (PDCB) preheated to 200 ℃, a dehydration solution 2 (with the sodium sulfide content of 24.66 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the D50 of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.136 mol/L) to a first polymerization reaction kettle 4 (PDCB/S =1.050, NMP/S = 3.50) at the temperature of 215 ℃ for 6h by using flow rates of 25.19g/min, 51.66g/min and 24.5g/min in sequence, heating the obtained mixed solution of the first reaction kettle 5 to 225 ℃ through a first heat exchanger 6, feeding the mixed solution into a second tubular reactor 7 (the reactor temperature is 225 ℃) for 1.5h, introducing into a closed type pressure filter 8, removing salt particles through the closed type pressure filter 8 (the aperture of a filter cloth is 80 mu m), and the D50 of the filtered salt particles 9 is 150 mu m; the obtained second reaction mixed solution 10 is heated to 265 ℃ through a second heat exchanger 11, then enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), stays for 1.5h, then enters an atmospheric flash evaporator 14 for flash evaporation after being subjected to heat exchange to 280 ℃ through a heat exchanger (not shown), auxiliary flash evaporation is carried out by using 265 ℃ superheated steam 13 during flash evaporation, the flow of the superheated steam 13 is 5.0kg/h, and the vaporized solvent and steam 19 are condensed and liquefied through a condenser 20 to generate waste liquid 21 to be recovered. The crude product 15 which is not dried enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, added with water for pulping, the added water amount is 6.0kg/h, enters a tubular washer with a heat conduction oil jacket after heat exchange to 180 ℃, the temperature of the tubular washer is controlled to be 180 ℃, the retention time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and the final product is obtained through continuous filtration and drying (not shown).

The production rate of this example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.6wt%, 5.0kg/h of superheated steam is consumed, and 8.78kg/h of waste liquid to be recovered is generated. The final product was prepared with a weight average molecular weight of about 21000 and an ash content of less than 0.55wt%.

Example 6

In a 100L reaction kettle, 14.85kg (150.0 mol) of N-methylpyrrolidone, 8.33kg (100.0 mol of NaOH) of 48.0wt% of sodium hydroxide aqueous solution, 11.93kg (100.0 mol of NaHS) of 47.0wt% of sodium hydrosulfide aqueous solution are added, the temperature is slowly raised to 203 ℃ under the protection of nitrogen, water is continuously removed in the process, 9.47kg (98.0% of water content) of distillate is obtained, and the temperature is lowered to 170 ℃ after dehydration is finished. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 176.4mol. The dehydration liquid prepared by the method in batches is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

The PAS continuous manufacturing process flow diagram shown in fig. 1 is used. Pumping p-dichlorobenzene 1 (PDCB) preheated to 200 ℃, dehydration liquid 2 (with the sodium sulfide content of 29.83 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the D50 of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.2 mol/L) to a first polymerization reaction kettle 4 (PDCB/S =1.025, NMP/S = 2.0) at the flow rates of 24.60g/min, 42.71g/min and 8.1g/min in sequence, wherein the temperature of the first polymerization reaction kettle 4 is 215 ℃, the retention time is 3.5h, the obtained first reaction mixed liquid 5 is heated to 225 ℃ through a first heat exchanger 6, enters a second tubular reactor 7 (the reactor temperature is 225 ℃), is introduced into a closed type pressure filter 8 after the retention time is 1h, salt particles are removed through the closed type pressure filter 8 (the aperture of a filter cloth is 80 mu m), and the D50 of the filtered salt particles 9 is 180 mu m; the obtained second reaction mixed solution 10 is heated to 265 ℃ through a second heat exchanger 11, then enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), stays for 1h, then enters an atmospheric flash evaporator 14 for flash evaporation after being subjected to heat exchange to 280 ℃ through a heat exchanger (not shown), auxiliary flash evaporation is carried out by using 265 ℃ superheated steam 13 during flash evaporation, the flow of the superheated steam 13 is 2.8kg/h, and the vaporized solvent and steam 19 are condensed and liquefied through a condenser 20 to generate waste liquid 21 to be recovered. The crude product 15 which is not dried enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water addition amount is 6.0kg/h, the crude product enters a tubular washer with a heat conduction oil jacket after heat exchange to 180 ℃, the temperature of the tubular washer is controlled to be 180 ℃, the retention time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, continuous filtration and drying are carried out (not shown), and then the final product is obtained after drying.

The production rate of this example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.58wt%, 2.8kg/h of superheated steam is consumed, and 5.1kg/h of waste liquid to be recovered is generated. The final product was prepared with a tested weight average molecular weight of about 32000 and ash below 0.52wt%.

Compared with the conventional NMP/S molar ratio example 2 (the molar ratio is 3.5), the continuous preparation process provided by the invention can be used for successfully preparing the polyphenylene sulfide under the conditions that the NMP/S molar ratio is 2.5 (example 4) and the NMP/S molar ratio is 2.0 (example 6), so that the energy consumption (superheated steam consumption) and the three wastes (waste liquid to be recovered) are remarkably reduced, the retention time (reaction time) is shortened, and the reaction efficiency is remarkably improved.

Comparing example 6 with example 2, the superheated steam consumption is reduced by about half, the amount of waste liquid to be recovered is reduced by more than 1/3, the total residence time of the reaction is reduced by about 20%, and the technical effects can greatly reduce the production cost of the continuous production.

Comparative example 3

19.82kg (200.0 mol) of N-methyl pyrrolidone, 8.33kg (100.0 mol) of 48wt% sodium hydroxide aqueous solution and 11.93kg (100.0 mol) of 47wt% sodium hydrosulfide aqueous solution are added into a 100L reaction kettle, then the temperature is continuously raised under the protection of nitrogen for dehydration, 10.51kg of aqueous solution (the water content is 98.0 wt%) is removed, and the temperature is reduced to 170 ℃ after the dehydration is finished. At this time, the amount of sulfur in the dehydrated liquid system was 98.0mol, and the water content was 117.6mol. The dehydration liquid prepared by the method in batches is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

The PAS continuous manufacturing process flow diagram shown in fig. 1 is used. Pumping p-dichlorobenzene 1 (PDCB) preheated to 200 ℃, a dehydration solution 2 (with the sodium sulfide content of 25.86 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (the D50 of the sodium chloride crystal seeds is 15 mu m, the concentration of the crystal slurry is 0.136 mol/L) to a first polymerization reaction kettle 4 (PDCB/S =1.050, NMP/S = 3.50) at the temperature of 215 ℃ for 5h by using flow rates of 25.21g/min, 49.27g/min and 24.5g/min in sequence, heating the obtained mixed solution of the first reaction kettle 5 to 225 ℃ through a first heat exchanger 6, feeding the mixed solution into a second tubular reactor 7 (the reactor temperature is 225 ℃) for 1h, feeding the mixed solution into a closed type pressure filter 8, removing salt particles through the closed type pressure filter 8 (the aperture of a filter cloth is 30 mu m), and the D50 of the filtered salt particles 9 is 60 mu m; the obtained second reaction mixed solution 10 is heated to 265 ℃ through a second heat exchanger 11, then enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), stays for 1.5h, then enters an atmospheric flash evaporator 14 for flash evaporation after being subjected to heat exchange to 280 ℃ through a heat exchanger (not shown), auxiliary flash evaporation is carried out by using 265 ℃ superheated steam 13 during flash evaporation, the flow of the superheated steam 13 is 5.0kg/h, and the vaporized solvent and steam 19 are condensed and liquefied through a condenser 20 to generate waste liquid 21 to be recovered. The crude product 15 which is not dried enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water addition amount is 6.0kg/h, the crude product enters a tubular washer with a heat conduction oil jacket after heat exchange to 180 ℃, the temperature of the tubular washer is controlled to be 180 ℃, the retention time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and continuous filtration and drying (not shown) are carried out to obtain the final product.

The production rate of this comparative example was 1.01kg/h; 5.0kg/h of salt-containing wastewater is generated, the salt content is 0.65wt%, 5.0kg/h of superheated steam is consumed, and 8.7kg/h of waste liquid to be recovered is generated. The final product was prepared with a weight average molecular weight of about 23000 and ash content of less than 0.6wt%.

In the implementation process of the comparative example, the particle size of the salt and the pore size of the filter cloth matched with the particle size are small, so that the resistance of the filter is large, the filter is occasionally blocked, and the stability of continuous implementation is poor.

Claims

1. A method for continuously producing a polyarylene sulfide, comprising:

(1) Putting a sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an optionally added auxiliary agent and an organic amide solvent into a reaction kettle, continuously heating to 200-220 ℃ for dehydration and saponification reaction, cooling to 150-180 ℃ after the dehydration reaction, and keeping the obtained dehydration liquid in a storage kettle;

(2) Injecting organic amide crystal slurry containing sodium chloride crystal seeds, dichloro aromatic compound and the dehydration solution into a prepolymerization reactor, controlling the reaction temperature at 180-240 ℃, reacting until the monomer conversion rate reaches 90% or more, injecting the obtained reaction solution into a filtering device, and filtering to remove sodium chloride particles;

(3) And injecting the filtrate obtained after the filtration into a polymerization reactor, controlling the reaction temperature at 235-280 ℃, reacting until the monomer conversion rate reaches 95% or above, carrying out flash evaporation treatment on the obtained reaction liquid, and finally carrying out post-treatment to obtain the polyarylene sulfide.

2. The continuous production method of a polyarylene sulfide as recited in claim 1, wherein in the step (1):

the dichloro aromatic compound is selected from one or more of p-dichlorobenzene, dichloronaphthalene, dichlorofluorene and dichlorocarbazole;

the molar ratio of the sulfur source to the sodium hydroxide is 1:0.95 to 1.1;

the molar ratio of the sulfur source to the auxiliary agent is 1:0 to 0.5;

3. The continuous production method of a polyarylene sulfide as recited in claim 1, wherein in the step (2):

the molar ratio of the dichloro aromatic compound to the sulfur source in the dehydration solution is 1.0-1.1: 1;

the dehydrated liquid, the dichloro aromatic compound and the organic amide crystal slurry containing the sodium chloride seed crystal are mixed, and the addition amount of the sodium chloride seed crystal is 0.05 to 0.4mol% based on 1mol of the sulfur source in the system, and the total amount of the organic amide solvent in the system is 2.0 to 4.0mol.

4. The method for continuously producing a polyarylene sulfide as recited in claim 3, wherein:

the molar ratio of the dichloro aromatic compound to the sulfur source in the dehydration solution is 1.0-1.025: 1;

the dehydrated liquid, the dichloroaromatic compound and the organic amide crystal slurry containing the sodium chloride crystal seed are mixed, and the total amount of the organic amide solvent is 2.0 to 2.5mol based on 1mol of the sulfur source in the system.

5. The continuous production method of a polyarylene sulfide as recited in claim 1, wherein in the step (2), the prepolymerization reactor is a two-stage reactor connected in series;

the first-stage reactor is selected from a continuous stirred tank reactor, and the reaction temperature in the tank reactor is controlled to be 180-220 ℃;

the second-stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled to be 220-240 ℃;

and when the monomer conversion rate in the first-stage reactor reaches 60-80%, injecting the reaction liquid into the second-stage reactor.

6. The continuous production method of a polyarylene sulfide as recited in claim 1, wherein in the step (3),

the polymerization reactor is selected from a tubular reactor;

the post-treatment comprises drying, washing, filtering and re-drying.

7. A method for producing a polyarylene sulfide in a batch manner, comprising:

(a) Putting a sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, a selectively added auxiliary agent and an organic amide solvent into a reaction kettle I, continuously heating to 200-220 ℃ for dehydration and saponification reaction, wherein the molar ratio of water to the sulfur source is 1.5-2.0 after dehydration, and then cooling to 150-180 ℃;

(b) Putting organic amide crystal slurry containing sodium chloride crystal seeds, dichloro aromatic compound and supplemented organic amide solvent into a reaction kettle I, firstly heating to 205-220 ℃, preserving heat for 0.5-5 h, then heating to 215-240 ℃, preserving heat for 0.5-5 h, then injecting the obtained reaction liquid into a filtering device, and filtering to remove sodium chloride particles;

8. The batch production method of a polyarylene sulfide as recited in claim 7, wherein in the step (a):

the molar ratio of the sulfur source to the sodium hydroxide is 1:0.95 to 1.1;

the molar ratio of the sulfur source to the auxiliary agent is 1:0 to 0.5;

9. The batch production method of a polyarylene sulfide as recited in claim 7, wherein in the step (b):

firstly, heating to 205-220 ℃ at the heating rate of 0.2-0.5 ℃/min, and then heating to 215-240 ℃ at the heating rate of 0.2-1 ℃/min;

in step (c):

heating to 240-280 ℃ at the heating rate of 0.2-1 ℃/min; the post-treatment comprises drying, washing, filtering and re-drying.

10. A polyarylene sulfide produced by the continuous production method according to any one of claims 1 to 6 or the batch production method according to any one of claims 7 to 9.

11. Use of the polyarylene sulfide according to claim 10 for preparing a crosslinked polyarylene sulfide by subjecting the polyarylene sulfide to thermal oxidation.