Polyarylene sulfide resin manufacturing process and its products and uses

Abstract

The present invention discloses a process for producing polyarylene sulfide and its product and use, which is produced by using sulfur source, dichloroaromatic compound as raw material, organic amide solvent, and a series of processes such as dehydration, prepolymerization, polymerization, flash evaporation and post-treatment, and the present invention innovatively adds sodium chloride seed crystals to the reaction liquid in the prepolymerization stage, and further adds a filtration step to the reaction liquid after the prepolymerization reaction is completed. This production process can greatly reduce the salt content in the produced polyarylene sulfide resin crude product, greatly reduce the generation of high salt-containing wastewater, reduce water consumption and energy consumption, and help the efficient production of high molecular weight polyarylene sulfide resin. More importantly, this process is not only applicable to batch process, but also to continuous production, which is expected to greatly improve the production efficiency of the production system, reduce production costs and be more environmentally friendly.

Description

[Technical field]

[0001] The present invention relates to the field of polyarylene sulfide resin synthesis, and in particular to continuous and batch production processes for polyarylene sulfide resins and their products and uses. [Background technology]

[0002] The typical method for producing polyarylene sulfide (PAS) is to polymerize sodium sulfide / sodium hydrosulfide and dihalo aromatic compounds in an organic amide solvent in the presence of an alkali metal hydroxide, and the solvent is generally N-methylpyrrolidone (hereinafter referred to as NMP), and the alkali metal hydroxide is generally sodium hydroxide. At present, the industrial production equipment for synthesizing polyphenylene sulfide (PPS) by the above method is all batch operation, and a part of the reaction raw material is first dehydrated in a reaction vessel, cooled, and the remaining reaction raw material is replenished, and then the reaction vessel is sealed and heated to carry out high-temperature polycondensation. After the reaction is completed, the crude product is obtained by cooling filtration or flash evaporation, and the crude product is purified to obtain the final product. In the batch production process, it is necessary to repeatedly add materials, heat and cool them, take them out, and clean the reaction vessel, which is time-consuming, greatly reduces the production efficiency, and the stability and reliability of the process are also low. Therefore, researchers have been making efforts to develop a continuous production method for PAS in order to improve the production efficiency.

[0003] Patents US4060520, US4056515, and US4066632 provide a continuous production process in which multiple tanks are connected in series, the reaction temperature of each tank is different, the pressure of the tank is continuously reduced from front to rear, and the raw material flows from front to rear by the pressure difference between adjacent tanks. A Chinese patent document with publication number CN108779253A provides a continuous PAS production apparatus and method, in which the PAS continuous production apparatus includes a storage chamber that contains multiple reaction tanks, and supplies an organic amide solvent, a sulfur source, and a dihalo aromatic compound to the storage chamber, and performs a polymerization reaction between the sulfur source and the dihalo aromatic compound in the organic amide solvent to form a reaction mixture. A plurality of reaction tanks are connected in sequence and communicate with each other through the gas phase, and the reaction mixture moves to each reaction tank in sequence. A Chinese patent publication with publication number CN113667122A provides a gradient temperature control continuous condensation method for polyarylene sulfide resin, in which the temperature of each reactor is controlled stepwise (gradually increased) to suit the different progress of the polymerization reaction, and the reaction raw materials are continuously fed to the kettle-type or tubular reactors connected in series by metering pumps.

[0004] In the production of PPS by the sodium sulfide method, a large amount of sodium chloride is produced as a by-product, and 1.08 tons of sodium chloride are produced for every ton of PPS produced. A large amount of sodium chloride is produced and begins to precipitate in the early stage of the polymerization reaction, so the reaction slurry contains a large amount of salt particles. When calculated based on a typical value of the molar ratio of NMP / sulfur source in the reaction liquid of 3.5, the content of by-product salt particles in the polymerization reaction liquid after the reaction is completed is about 20.5 wt%. A large amount of salt particles in the slurry settles in the long-distance transportation or reaction pipe, especially in a tubular reactor of a small size, and is prone to clogging, causing production stoppage. Whether in a batch process or a continuous process, increasing the concentration of the reaction substrate is advantageous for improving production efficiency, and a high concentration of the reaction substrate is also advantageous for efficiently producing high molecular weight polyphenylene sulfide, but the content of by-product salt particles also increases significantly accordingly, and the viscosity of the reaction liquid also increases rapidly, aggravating the above-mentioned problems of clogging and stirring in transportation. The presence of a large amount of by-product salt particles in the PPS reaction slurry severely limits research and development of a continuous PPS production process, and also limits the efficient production of higher molecular weight polyphenylene sulfide.

[0005] Whether it is a batch process or a continuous process, the by-product sodium chloride is covered on the product after the reaction is completed, and it is necessary to use a large amount of pure water to wash and remove it multiple times, generating a large amount of washing wastewater containing high salt content. At present, in industrialized production, more than 15 tons of water is consumed for every ton of PPS produced, and the by-product salt dissolved in the water is evaporated and recovered, which further increases energy consumption. In addition, the fact that salt water promotes corrosion imposes high requirements on the materials of polyphenylene sulfide production equipment. The characteristics of the polyphenylene sulfide production process, which consumes a lot of water and energy and promotes corrosion, significantly increase the production cost and environmental protection cost of polyphenylene sulfide, greatly restrict the ecological measures and sustainability of the polyphenylene sulfide production process, and limit the competitiveness and development potential of polyphenylene sulfide resin.

[0006] These problems can be solved by removing the precipitated salt particles before the solid-liquid separation of the product, thereby reducing the content of salt particles in the subsequent equipment and the crude product.For example, the Chinese patent document with publication number CN104371103A discloses a method for desalting polyphenylene sulfide, in which after the polymerization reaction is completed, the reactor is kept warm and pressurized to stand, and inorganic salts and inorganic auxiliary agents settle to the bottom of the reactor, and then the pressure difference is used to push the upper layer of polyphenylene sulfide solution out of the reactor through the inner extension tube at the bottom, and the solution is cooled, and the precipitated polyphenylene sulfide is purified to obtain a high-purity polyphenylene sulfide product.However, the method described in this patent is only applicable to batch processing, and the patent document does not further explain how to further treat the inorganic salts in the precipitated layer in the reactor, the large amount of reaction solvent and the product, and in fact, the stirring of the polymerization reactor cannot be restarted before the precipitated inorganic salt layer is completely treated.

[0007] Furthermore, for example, the Chinese patent publication number CN107964098A relates to a method and device for desalting in the synthesis of polyphenylene sulfide, in which the raw material of the reaction vessel is passed through a filter with a screen by a material conveying pump during polymerization, the salt remains in the filter, and the remaining polyphenylene sulfide raw material returns to the reaction vessel, and this circulation continues until the reaction is completed, and the discharged material is washed and dried to obtain polyphenylene sulfide resin. This method is also only applicable to batch processes, and in the later stage of the reaction, the thickness of the salt filter cake layer and the viscosity of the reaction liquid increase, making filtration more difficult. Summary of the Invention

[0008] In response to the above problems in the prior art, the present invention discloses a process for producing polyarylene sulfide resin, which can greatly reduce the salt content in the crude polyarylene sulfide resin produced, greatly reduce the generation of high salt-containing wastewater, reduce water consumption and energy consumption, and is useful for producing high molecular weight polyarylene sulfide resin.More importantly, this process is not only applicable to batch process, but also to continuous production, which is expected to greatly improve the production efficiency of the production system, reduce production costs, and be more environmentally friendly.

[0009] The technical solutions are specifically as follows: A continuous method for producing polyarylene sulfide, comprising the steps of:

[0010] (1) A sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an auxiliary agent that can be optionally added, and an organic amide solvent are charged into a reaction vessel, and the temperature is continuously raised to 200-220°C to carry out a dehydration and saponification reaction. After the completion of the dehydration, the molar ratio of water to sulfur source is 1.5-2.0. The reaction vessel is cooled to 150-180°C, and the obtained dehydrated liquid is kept warm in a storage vessel.

[0011] (2) The organic amide crystal slurry containing sodium chloride seed crystals, the dichloro aromatic compound and the dehydrated liquid are injected into a prepolymerization reactor, and the reaction temperature is controlled to 180 to 240°C. The reaction is carried out until the monomer conversion rate reaches 90% or more. The resulting reaction liquid is then injected into a filtration device and filtered to remove sodium chloride particles.

[0012] (3) The filtrate obtained through the filtration device is injected into a polymerization reactor, the reaction temperature is controlled at 235-280°C, and the reaction is continued until the monomer conversion rate reaches 95% or more. The resulting reaction liquid is subjected to a flash evaporation treatment, and finally, after post-treatment, polyarylene sulfide is obtained.

[0013] The inventors have conducted intensive research into the growth process and particle morphology of salt particles in the PPS reaction solution. After all the raw materials are added and the temperature is raised to a certain level, the polymerization reaction begins, and sodium chloride is constantly produced. Due to the presence of water in the polyphenylene sulfide reaction system, a small amount of by-product sodium chloride can be dissolved in the reaction solvent. When the dissolution saturation is reached, the sodium chloride that continues to be produced will generate nuclei and begin to precipitate, the reaction continues rapidly, and the large amount of sodium chloride produced will rapidly generate nuclei and precipitate. The research findings are as follows: (1) In the conventional typical batch PPS production process, the particle size of the salt particles produced in the reaction solution is small, about 20 μm, which is disadvantageous for filtration, and there have been few reports on this point. (2) In the preliminary polymerization reaction stage, the reaction conversion rate is more than 90%, and the growth of sodium chloride crystals is almost completed at this time, but the molecular weight of the product at this time is still low, and the viscosity of the reaction liquid is not high, so it is necessary to enter the high-temperature polymerization stage and continue the reaction to improve the molecular weight, so the appropriate time to filter out the by-product salt is when the growth of sodium chloride crystals is basically completed in the preliminary polymerization reaction stage, and after the high-temperature polymerization stage, the viscosity of the reaction liquid is high and filtering is more difficult. (3) By controlling the water content in the system within an appropriate range, on the one hand, it can reduce the rate of precipitation and nucleation of sodium chloride, and on the other hand, it can accelerate the growth rate of sodium chloride crystals. (4) The introduction of a small amount of sodium chloride seed crystals can effectively inhibit the primary nucleation of sodium chloride.

[0014] Based on the above research results, the technical solution of the present invention is put forward, and the test results show that in step (1), in order to control the particle size of sodium chloride particles to 100 μm or more to facilitate filtration, it is necessary to control the water content at an appropriate ratio (control the molar ratio of water to sulfur source to 1.5-2.0) after the dehydration is completed. Sodium chloride can be removed by filtration, which greatly reduces the generation of high-salt wastewater, reduces water consumption and energy consumption, and at the same time further reduces the amount of solvent in the reaction solution, which is conducive to the efficient production of high molecular weight PPS.

[0015] In step (1), The concentration of the sulfur source aqueous solution is 28 to 48 wt %, and the sulfur source is selected from sodium hydrosulfide and / or sodium sulfide; The alkali metal hydroxide aqueous solution is a sodium hydroxide aqueous solution having a concentration of 30 to 70 wt %,

[0016] The dichloroaromatic compound is selected from one or more of paradichlorobenzene, dichloronaphthalene, dichlorofluorene, dichlorocarbazole, preferably paradichlorobenzene; The molar ratio of the sulfur source to sodium oxide is 1:0.95 to 1.1;

[0017] The auxiliary is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; The molar ratio of the sulfur source to the auxiliary is 1:0 to 0.5;

[0018] The organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl-ε-caprolactam, N,N-dimethylformamide, preferably N-methylpyrrolidone; The molar ratio of the sulfur source to the organic amide solvent is 1:1.5-2.5.

[0019] The dehydration and saponification reaction may be carried out batchwise according to the method of step (1), or may be carried out by using a continuous dehydration means reported in other patents, and the key point is to control the water content after the completion of dehydration within the appropriate range.

[0020] In step (2), The organic amide crystal slurry is prepared from sodium chloride seed crystals and the organic amide solvent, wherein the concentration of the sodium chloride seed crystals is 0.1-1 mol / L and the particle size of the sodium chloride seed crystals is 1-20 μm;

[0021] The molar ratio of the dichloroaromatic compound to the sulfur source in the dehydrated liquid is 1.0 to 1.1:1, and preferably 1.0 to 1.08:1.

[0022] After mixing the dehydration liquid, the dichloroaromatic compound, and the organic amide crystal slurry containing the sodium chloride seed crystal, the amount of the sodium chloride seed crystal added is controlled to be 0.05-0.4 mol% relative to 1 mol of the sulfur source in the prepolymerization reactor system, and the total amount of the organic amide solvent in the system is 2.0-4.0 mol.

[0023] Preferably, the total amount of the organic amide solvent is 2.0 to 2.5 mol per mol of the sulfur source in the prepolymerization reactor system (that is, the system is controlled to NMP / S=2.0 to 2.5).

[0024] Test results have shown that this not only advantageously reduces the energy consumption of flash evaporation generated in the production process and the amount of waste liquid to be recovered, but also adequately improves the weight average molecular weight of the PAS produced.

[0025] More preferably, the molar ratio of the dichloroaromatic compound to the sulfur source is controlled to 1.0 to 1.025:1.

[0026] As a result of the tests, it was found that by controlling the molar ratio of the dichloroaromatic compound to the sulfur source and the NMP / S ratio within the above more preferred ranges, the weight average molecular weight of the produced PAS can be significantly improved.

[0027] The inventors further conducted intensive research into the polymerization kinetics of polyphenylene sulfide and found the following: (1) The main polymerization reaction of polyarylene sulfide is a strong exothermic reaction, and temperature runaway during prepolymerization has a significant impact on reaction safety and product quality, and the polymerization reaction is a second-order kinetic reaction, so in order to meet the requirements of reaction temperature control and reaction efficiency, in the continuous production process of connecting multiple reactors in series, the first reactor is preferably a continuous stirred tank reactor, and the conversion rate is controlled within an appropriate range. (2) By increasing the water content in the reaction system, the reaction rate, i.e., the production rate of sodium chloride, can be reduced, but too high a water content will cause problems such as too low reaction efficiency, high reaction pressure, and more energy consumption during flash evaporation.

[0028] To sum up the above research results, a large amount of by-product salts is generated during the production process of PAS by the conventional batch manufacturing process and during the development of a continuous manufacturing process, and the following problems are solved: (1) the removal of the by-product salts by washing consumes a lot of water and produces wastewater with high salt content, (2) the problem of clogging is a limitation on the development of a continuous manufacturing process of PAS, and (3) the improvement of the substrate concentration is limited, and the main polymerization reaction is a second-order kinetic highly exothermic reaction, which requires high temperature control during the reaction process. In response to these problems, the present invention provides an environmentally friendly, low-cost, and highly efficient continuous manufacturing method for PAS.

[0029] Preferably, In step (2), the prepolymerization reactor is a two-stage reactor connected in series, In order to achieve both control of the reaction temperature and reaction efficiency, the first stage reactor is selected from continuous stirred tank reactors, and the reaction temperature in the tank reactor is controlled to 180-220°C, preferably 190-210°C.

[0030] After the monomer conversion in the first-stage reactor reaches 60 to 80%, and more preferably 70 to 80%, the reaction liquid is injected into the second-stage reactor.

[0031] The second stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled at 220-240°C.

[0032] Preferably, the heating medium used in the jacketed tubular reactor is a thermal oil, the flow direction of the thermal oil coincides with the flow direction of the reaction liquid, the oil temperature at the inlet is controlled to 220-230° C., and the oil temperature at the outlet is 230-240° C. Preferably, the monomer conversion rate at the outlet reaches 90% or more, more preferably 93% or more, at which time the growth of sodium chloride crystals is almost complete, the particle size grows to 60 μm or more, preferably 100 μm or more, and the size is uniform.

[0033] The filtration device is a continuous filtration device, specifically selected from closed pressure filtration devices, preferably a pressure rotary drum filter, which is applied to the continuous filtration of high-temperature slurries containing volatile components and crystals that are prone to precipitation.

[0034] Preferably, the pore size of the filter medium in the filtration device is preferably 60 to 80 μm, which allows sufficient removal of sodium chloride particles while at the same time preventing the filtration pressure drop from becoming too large.

[0035] In step (3), The polymerization reactor is selected from a tubular reactor, preferably the tubular reactor is selected from a coil reactor immersed in a high temperature oil bath, to control the reaction temperature at 235-280°C, complete the further increase in molecular weight, and ensure that the monomer conversion reaches 95% or more, preferably 98% or more.

[0036] Preferably, the flash evaporation process uses superheated steam at 260-300°C for auxiliary heat supply, the amount of superheated steam used is 0.1-1 kg / mol of sulfur source, the steam and vaporized solvent are extracted from the top of the flash evaporator and condensed into a solvent recovery system, and the salt-containing PAS crude product is continuously discharged from the bottom.

[0037] The post-treatment includes drying, washing, filtering and redrying, and is specifically as follows:

[0038] The salt-containing PAS crude product is successively dried in a dryer to further remove residual solvent.

[0039] After drying, the salt-containing PAS crude product is mixed with water and beaten, and then washed with high-temperature and high-pressure water once. The mass ratio of the water to the crude product is 3 to 10:1. The final product is obtained after continuous filtration and drying.

[0040] The continuous high-temperature, high-pressure water washing apparatus may be a continuous tank type washing apparatus or a continuous pipe type washing apparatus, and is preferably a continuous pipe type washing apparatus equipped with a heat transfer oil jacket. The residence time of the water-washed slurry is preferably 1 to 30 minutes, and the water-washing temperature is preferably 150 to 220°C.

[0041] The continuous filtration device is selected from a continuous filter such as a multiple filter press, a decanter type centrifuge and a belt filter, preferably a belt filter of continuous vacuum suction filtration.

[0042] The dryer is selected from continuous feed and discharge dryers. The manufacturing process of the present invention is used to filter out large salt particles in advance, so that the salt concentration in the subsequent crude product is greatly reduced, and the pressure of the washing process in the post-treatment is significantly reduced, so that only one continuous high-temperature and high-pressure water washing is required. If the salt particles are not filtered in advance, the sodium chloride in the product needs to be thoroughly washed by normal pressure water washing + high-pressure water washing several times, and the amount of washing water and the amount of salt-containing wastewater generated will increase significantly. If seed crystals are not added and the water content in the reaction solution is not controlled to an appropriate level, the size of the sodium chloride particles is small, and it is difficult to effectively block it when filtered with a filter material with a large pore size, and if a filter material with a small pore size is used, the pressure loss of the filter is too large, and the filter is easily clogged.

[0043] The above process can also be applied to batch production of polyarylene sulfide. The effects of each process parameter on the performance of PAS produced in batch production are similar to those in continuous production, and therefore will not be described below.

[0044] Therefore, the present invention further discloses a batch-wise method for preparing polyarylene sulfide, comprising the steps of:

[0045] (a) A sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an auxiliary that can be optionally added, and an organic amide solvent are charged into a reaction vessel I, and the temperature is continuously raised to 200-220°C to carry out a dehydration and saponification reaction, and the molar ratio of water to sulfur source after the completion of the dehydration is 1.5-2.0, and then the reaction is cooled to 150-180°C.

[0046] (b) The organic amide crystal slurry containing sodium chloride seed crystals, the dichloroaromatic compound and the additional organic amide solvent are charged into the reaction vessel I, and the temperature is first raised to 205-220°C and kept at that temperature for 0.5-3 hours, and then the temperature is further raised to 215-240°C and kept at that temperature for 0.5-5 hours. The resulting reaction liquid is then poured into a filtration device and filtered to remove sodium chloride particles.

[0047] (c) The filtrate obtained through the filtration device is poured into a preheated reaction vessel II, heated to 240-280°C, and kept at that temperature for 0.5-5 hours. The resulting reaction liquid is then subjected to a flash evaporation treatment, and finally, after post-treatment, polyarylene sulfide is obtained.

[0048] In step (a), The concentration of the sulfur source aqueous solution is 28 to 48 wt %, and the sulfur source is selected from sodium hydrosulfide and / or sodium sulfide; The alkali metal hydroxide aqueous solution is a sodium hydroxide aqueous solution having a concentration of 30 to 70 wt %, The dichloroaromatic compound is selected from one or more of paradichlorobenzene, dichloronaphthalene, dichlorofluorene, and dichlorocarbazole, and is preferably paradichlorobenzene; and the molar ratio of the sulfur source to sodium hydroxide is 1:0.95 to 1.1; The auxiliary is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; The molar ratio of the sulfur source to the auxiliary is 1:0 to 0.5; the organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl-ε-caprolactam, N,N-dimethylformamide; The molar ratio of the sulfur source to the organic amide solvent is 1:1.5-3.

[0049] In step (b), The concentration of the sodium chloride seed crystals in the organic amide crystal slurry is 0.1 to 1 mol / L, the particle size of the sodium chloride seed crystals is 1 to 20 μm, and the amount of the sodium chloride seed crystals added is 0.05 to 0.4 mol % of the sulfur source in the system at this time; the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0 to 1.1:1; After adding the organic amide solvent, the molar ratio of the total molar amount of the organic amide solvent to the sulfur source in the reaction vessel I is 2.5 to 3.5; Preferably, First, the temperature is increased to 205 to 220°C at a rate of 0.2 to 0.5°C / min, and then increased to 215 to 240°C at a rate of 0.2 to 1°C / min.

[0050] The reactor used in steps (a) and (b) is selected from stirred reactors equipped with heat transfer oil jackets and coils.

[0051] In step (c), Preferably, the temperature is increased to 240 to 280° C. at a rate of 0.2 to 1° C. / min.

[0052] Preferably, after the flash evaporation is completed, the material is dried in the flash evaporator for 10 min to 1 h, cooled to room temperature, and then discharged.

[0053] The post-treatment includes drying, washing, filtering and redrying.

[0054] The present invention further discloses polyarylene sulfides produced according to the above two processes, respectively.

[0055] The present invention further discloses the use of said polyarylene sulfide in the preparation of crosslinked polyarylene sulfide, specifically by subjecting said polyarylene sulfide to a thermal oxidation treatment.

[0056] Compared with the prior art, the beneficial effects of the present invention are as follows:

[0057] The present invention discloses a process for producing polyarylene sulfide resin, the core of which is to add sodium chloride seed crystals into the PAS reaction solution, while controlling the water content in the system to a suitable range, and controlling the reaction temperature in the prepolymerization stage to a relatively low temperature, thereby achieving the purpose of reducing the nucleation rate of sodium chloride, accelerating the crystal growth rate, inhibiting primary nucleation, promoting secondary nucleation, and finally improving the size of by-product sodium chloride crystals. Another key point of the process is the timing of filtering to remove salt particles, and filtering salt particles is selected at the end of the prepolymerization reaction stage. The production process disclosed in the present invention can greatly reduce the salt content in the produced polyarylene sulfide resin crude product, greatly reduce the generation of high-salt-containing washing wastewater, reduce water consumption and energy consumption, and help the efficient production of high molecular weight polyarylene sulfide resin. More importantly, the process is not only applicable to batch processes, but also to continuous production, which is expected to greatly reduce the production efficiency and production cost of the production system, and is more environmentally friendly. When using a continuous production process, in order to balance the temperature control and production efficiency of the pre-polymerization stage, the pre-polymerization stage is divided into two stages: the first stage selects a tank reactor suitable for stable heat transfer, and controls the conversion rate to an appropriate value to remove most of the reaction heat while preventing the reaction rate from becoming too low; the second stage selects a jacketed tubular reactor, and uses the remaining heat to promote the reaction, and continues to raise the temperature to maintain reaction efficiency. For the high-temperature polymerization stage, a coil reactor is selected, which is more suitable for high conversion and high viscosity conditions (high molecular weight and high concentration) after removing salt particles. [Brief description of the drawings]

[0058] [Figure 1] FIG. 1 is a schematic diagram of a process flow for continuous production of polyarylene sulfide of the present invention, in which post-treatment processes are omitted. DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0059] The present invention will be described in more detail below with reference to examples and comparative examples, but the embodiments of the present invention are not limited to these.

[0060] Ash content test: The porcelain crucible is placed in a constant temperature muffle furnace at 750°C and fired until it reaches a constant weight, then it is taken out and placed in a dryer to cool, then weighed and recorded as M0. 3g of sample is weighed and placed in the crucible, 10mL of nitric acid is added, then it is placed on an alcohol burner and burned until no smoke is produced. Finally, the crucible is placed again in a constant temperature muffle furnace at 750°C and fired for 1h, then it is taken out and placed in a dryer to cool, then weighed and recorded as M1. The formula for calculating resin ash is as follows: (M1-M0) / 3*100%.

[0061] Molecular weight test: The molecular weight of PPS is measured by gel permeation chromatography using polystyrene as the standard sample, the mobile phase is 1-chloronaphthalene, the column temperature is 220°C, the flow rate is 1mL / min, and the detector is a refractive index detector.

[0062] Example 1 In a 100L reactor I, add 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol), and slowly heat up to 200℃ under nitrogen atmosphere, continuously removing water during the process, take out 10.00kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system is 98.0mol, and the water content is 147.0mol.

[0063] To the reactor I, 15.14 kg (103.0 mol) of paradichlorobenzene and 10 kg of NMP crystal slurry containing 11.7 g of sodium chloride seed crystals (0.2 mol, D50 of sodium chloride seed crystals is 15 μm) were added, and 4.46 kg of NMP (45.0 mol) was further added. The reactor I was sealed, and the temperature was increased to 215°C at 0.3°C / min and kept at that temperature for 3 hours, and then the temperature was increased to 230°C at 0.2°C / min and kept at that temperature for 1 hour. After that, the reaction liquid was filtered (screen pore size After filtering through a sieve (80μm) to remove sodium chloride particles, the mixture was pumped into a 100L reactor II preheated to 230℃, the reactor II was then heated to 260℃ at 0.5℃ / min, kept at that temperature for 1 hour, and then slowly released into a normal pressure flash evaporator for flash evaporation treatment, and finally dried to further remove the solvent, obtaining 10.4kg of crude product, which was then mixed with 60kg of water and beaten, heated to 180℃, kept at that temperature for 0.5 hours, cooled, filtered and dried to obtain the final product.

[0064] In this example, 10.9 kg of sodium chloride was obtained when the reaction solution was transferred and filtered, and the D50 of the sodium chloride particles measured by a laser particle size analyzer was 130 μm. 60.0 kg of water was used in the washing process, and 50.0 kg of salt-containing wastewater was generated, with a salt content of 0.6 wt%. 10.1 kg of final product was produced, with a weight average molecular weight of 22500 and an ash content of 0.53 wt%.

[0065] Comparative Example 1 In a 100L reactor I, add 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol), and slowly heat up to 200℃ under nitrogen atmosphere, continuously removing water during the process, take out 10.00kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system is 98.0mol, and the water content is 147.0mol.

[0066] Add 15.14kg (103.0mol) of paradichlorobenzene and 14.46kg of NMP to reactor I, seal reactor I, heat to 215℃ at 0.3℃ / min, keep warm for 3h, then heat to 230℃ at 0.2℃ / min, keep warm for 1h, filter the reaction liquid through a filter (screen pore size 20μm) to remove sodium chloride particles, and pump into 100L reactor II preheated to 230℃. Then heat reactor II to 260℃ at 0.5℃ / min, keep warm for 1h, slowly release into normal pressure flash evaporator for flash evaporation treatment, finally dry to further remove the solvent, obtain 10.8kg of crude product, add 60kg of water, beat, heat to 180℃, keep warm for 0.5h, cool, filter and dry to obtain final product.

[0067] In this comparative example, 10.5 kg of sodium chloride was obtained when the reaction solution was transferred and filtered, and the D50 of the sodium chloride particles measured by a laser particle size analyzer was 30 μm. 60.0 kg of water was used in the washing process, and 51.0 kg of salt-containing wastewater was generated, with a salt content of 1.3 wt%. 10.1 kg of final product was produced, with a weight average molecular weight of 22500 and an ash content of 0.56 wt%.

[0068] Furthermore, due to the small particle size of the salt and the small pore size of the corresponding filter screen, the pressure loss in the filter is high, the flow rate is low, and the filtration step takes a long time, which greatly reduces the production efficiency and significantly increases the production cost.

[0069] Comparative Example 2 In a 100L reactor I, add 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol), and slowly heat up to 200℃ under nitrogen atmosphere, continuously removing water during the process, take out 10.00kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system is 98.0mol, and the water content is 147.0mol.

[0070] Into reactor I, add 10 kg of NMP crystal slurry containing 15.14 kg (103.0 mol) of paradichlorobenzene and 11.7 g of sodium chloride seed crystals (0.2 mol, D50 of sodium chloride seed crystals is 15 μm), then add 4.46 kg of NMP (45.0 mol), seal the reactor, heat to 215 ° C at 0.3 ° C / min, keep warm for 3 hours, then heat to 230 ° C at 0.15 ° C / min, keep warm for 1 hour, then heat to 260 ° C at 0.5 ° C / min, keep warm for 1 hour, then slowly release to normal pressure flash evaporator for flash evaporation treatment, and further dry to further remove the solvent, obtaining 21.2 kg of crude product. The crude product was first washed with water at normal pressure and filtered twice, with 60.0 kg of water added each time. An additional 60.0 kg of water was added and the product was beaten, then heated to 180°C, kept at that temperature for 0.5 hours, cooled, filtered and dried to obtain the final product.

[0071] In this comparative example, the amount of water used in the washing process was 180.0 kg, the amount of salt-containing wastewater produced was 182.0 kg, and the salt content was 6.0 wt%. The final product produced was 10.0 kg, with a weight average molecular weight of 21500 and an ash content of 0.50 wt%.

[0072] A comparison of Example 1 and Comparative Example 2 shows that removing salt particles by filtration significantly reduces the amount of wash water and waste brine produced.

[0073] Example 2 In a 100L reactor, add 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol). Slowly heat to 200℃ under nitrogen atmosphere, continuously remove water during the process, take out 10.00kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system is 98.0mol, and the water content is 147.0mol. The dehydrated liquid produced in the above batch method is put into a storage kettle at 170℃ and stored at a constant temperature for use in subsequent continuous supply.

[0074] The flow chart of the continuous manufacturing process of PAS shown in Figure 1 is used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (content of sodium sulfide is 25.48 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of sodium chloride seed crystals is 15 μm, concentration of crystal slurry is 0.136 mol / L) are fed at flow rates of 25.21 g / min, 50 g / min and 24.5 g / min, respectively, to a first polymerization reactor 4 (in the first polymerization reactor 4, PDCB / S=1.050 (molar ratio, the same applies below). ), NMP / S=3.50 (molar ratio, the same below) was controlled), the temperature of the first polymerization reactor 4 was 215°C, and the residence time was 5h. The obtained first reaction mixture 5 was heated to 225°C through the first heat exchanger 6 and entered the second tubular reactor 7 (reactor temperature was 225°C), and after residence for 1h, entered the closed pressure filter 8, and the salt particles were removed through the closed pressure filter 8 (screen pore size was 80μm), and the D50 of the filtered salt particles 9 was 140μm. The obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (the temperature of the reactor is 265°C) with a residence time of 1.5 hours. Then, it is heat exchanged through a heat exchanger (not shown) to 280°C and enters the atmospheric pressure flash evaporator 14 for flash evaporation. During the flash evaporation, superheated steam 13 at 265°C is used to assist the flash evaporation, and the flow rate of the superheated steam 13 is 5.0 kg / h. The vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, and the waste liquid 21 to be recovered is generated. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 then enters a tubular washer equipped with a heat transfer oil jacket. The temperature of the tubular washer is controlled at 180°C, and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown) to obtain the final product.

[0075] The production rate of this example was 1.01 kg / h. The amount of saline wastewater produced was 5.0 kg / h, with a salt content of 0.6 wt%, the amount of superheated steam consumed was 5.0 kg / h, and the amount of wastewater to be recovered was 8.7 kg / h. The final product produced was measured to have a weight average molecular weight of about 23,000 and an ash content of less than 0.55 wt%.

[0076] Example 3 The dehydration liquid is produced in batches according to the method of Example 2, and stored in a storage tank at 170°C for subsequent use in continuous supply. The PAS continuous production apparatus shown in Figure 1 is used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content is 25.48 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of sodium chloride seed crystals is 15 μm, concentration of crystal slurry is 0.4 mol / L) are fed to a first polymerization reactor 4 (PDCB / S=1.050, NMP) at flow rates of 25.21 g / min, 50 g / min and 8.1 g / min, respectively. The resulting first reaction mixture 5 was heated to 230°C through a first heat exchanger 6 and entered a second tubular reactor 7 (reactor temperature was 230°C), and after 1 hour of residence, entered a sealed pressure filter 8 (screen pore size was 80μm) to remove salt particles. The filtered salt particles 9 had a D50 of 130μm. The obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11 and enters the third tubular reactor 12 (the temperature of the reactor is 265°C) with a residence time of 1 hour, and then passes through a heat exchanger (not shown) for heat exchange up to 280°C before entering the atmospheric pressure flash evaporator 14 for flash evaporation. During the flash evaporation, superheated steam 13 at 265°C is used to assist the flash evaporation, and the flow rate of the superheated steam 13 is 3.5 kg / h. The vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, and the waste liquid 21 to be recovered is generated. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 then enters a tubular washer equipped with a heat transfer oil jacket. The temperature of the tubular washer is controlled at 180°C, and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown) to obtain the final product.

[0077] The production rate of this example was 1.01 kg / h. The amount of saline wastewater produced was 5.0 kg / h, with a salt content of 0.6 wt%, the amount of superheated steam consumed was 3.5 kg / h, and the amount of wastewater to be recovered was 6.2 kg / h. The final product produced was measured to have a weight average molecular weight of about 26,000 and an ash content of less than 0.55 wt%.

[0078] As can be seen from the comparison between the description of Example 2 and Example 3, by further reducing the molar ratio of NMP / S in the prepolymerization reaction step, specifically in the first polymerization reactor, the amount of solvent used and the amount of superheated steam used in the flash evaporation step can be reduced, energy consumption and waste liquid amount can be reduced, and production costs can be significantly reduced. In addition, the reaction efficiency can be improved, the residence time of the reaction liquid in each stage of the reactor can be shortened, and the weight average molecular weight of the final product can be appropriately improved.

[0079] However, a decrease in the molar ratio of NMP / S means that the concentrations of PPS and sodium chloride in the system increase, and the viscosity of the reaction solution increases rapidly with an increase in the concentrations of PPS and sodium chloride in the system and the molecular weight of PPS, making it difficult to smoothly carry out production under low NMP / S in the production process of the prior art, regardless of whether it is a batch or continuous process. This shows that the production process disclosed in the present invention is highly adaptable, and is particularly suitable for production conditions with a low NMP / S molar ratio, and is therefore more advantageous for the efficient production of high molecular weight PAS.

[0080] Example 4 In a 100L reactor, 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of sodium hydroxide is 100.0mol), 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of sodium hydrosulfide is 100.0mol), and 1.6kg of anhydrous sodium acetate are added, and then the temperature is continuously raised under a nitrogen atmosphere to dehydrate, and 10.00kg of aqueous solution (water content 98.0wt%) is removed, and the temperature is lowered to 170℃ after dehydration is completed. At this time, the amount of sulfur in the dehydration liquid system is 98.0mol, and the water content is 147.0mol. The dehydration liquid produced in the above batch method is put into a storage kettle at 170℃ to be stored at a high temperature for subsequent use in continuous supply.

[0081] The continuous PAS manufacturing equipment shown in Figure 1 is used. Preheated paradichlorobenzene (PDCB) 1, dehydration liquid 2 (sodium sulfide content is 24.22 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (particle size of sodium chloride seed crystals is 15 μm, concentration of crystal slurry is 0.4 mol / L) are fed at flow rates of 24.61 g / min, 52.6 g / min and 8.1 g / min, respectively, to the first polymerization reactor 4 (PDCB / S=1.025, NMP / S=2.5, CH 3The first polymerization reactor temperature 4 was 215°C and the residence time was 3.5h. The first reaction mixture 5 obtained was heated to 230°C through the first heat exchanger 6 and entered the second tubular reactor 7 (reactor temperature was 230°C), and after residence for 1h, entered the sealed pressure filter 8 to remove salt particles (screen pore size was 80μm). The D50 of the filtered salt particles 9 was 120μm. The obtained second reaction mixture 10 passes through the second heat exchanger 11 and is heated to 265°C, then enters the third tubular reactor 12 (the temperature of the reactor is 265°C), with a residence time of 1 hour, and then passes through a heat exchanger (not shown) and is heat exchanged up to 280°C, after which it enters the atmospheric pressure flash evaporator 14 for flash evaporation, during which superheated steam 13 at 265°C is used to assist the flash evaporation, with the flow rate of the superheated steam 13 being 3.3 kg / h, and the vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, to generate the waste liquid 21 to be recovered. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 then enters a tubular washer equipped with a heat transfer oil jacket. The temperature of the tubular washer is controlled at 180°C, and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown) to obtain the final product.

[0082] The production rate of this example was 1.01 kg / h. The amount of saline wastewater produced was 5.0 kg / h, with a salt content of 0.65 wt%, the amount of superheated steam consumed was 3.3 kg / h, and the amount of wastewater to be recovered was 6.0 kg / h. The final product produced was measured to have a weight average molecular weight of about 68,000 and an ash content of less than 0.5 wt%.

[0083] Example 5 In a 100L reactor, add 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol). Slowly heat to 195℃ under nitrogen atmosphere, continuously remove water during the process, take out 9.07kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 196.0mol. The dehydrated liquid produced in the above batch method was put into a storage kettle at 170℃ and stored at a constant temperature for use in subsequent continuous supply.

[0084] The continuous manufacturing process of PAS shown in Figure 1 is used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content is 24.66 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of sodium chloride seed crystals is 15 μm, and the concentration of the crystal slurry is 0.136 mol / L) are fed to the first polymerization reactor 4 (PDCB / S=1.050, The first polymerization reactor 4 had a temperature of 215°C and a residence time of 6 hours. The first reaction mixture 5 obtained was heated to 225°C through a first heat exchanger 6 and entered a second tubular reactor 7 (reactor temperature was 225°C). After residence for 1.5 hours, the mixture entered a closed pressure filter 8 (screen pore size was 80μm) to remove salt particles. The filtered salt particles 9 had a D50 of 150μm. The obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (the temperature of the reactor is 265°C) with a residence time of 1.5 hours. Then, it is heat exchanged through a heat exchanger (not shown) to 280°C and enters the atmospheric pressure flash evaporator 14 for flash evaporation. During the flash evaporation, superheated steam 13 at 265°C is used to assist the flash evaporation, and the flow rate of the superheated steam 13 is 5.0 kg / h. The vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, and the waste liquid 21 to be recovered is generated. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 then enters a tubular washer equipped with a heat transfer oil jacket. The temperature of the tubular washer is controlled at 180°C, and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown) to obtain the final product.

[0085] The production rate of this example was 1.01 kg / h. The amount of saline wastewater produced was 5.0 kg / h, with a salt content of 0.6 wt%, the amount of superheated steam consumed was 5.0 kg / h, and the amount of wastewater to be recovered was 8.78 kg / h. The final product produced was measured to have a weight average molecular weight of about 21,000 and an ash content of less than 0.55 wt%.

[0086] Example 6 In a 100L reactor, add 14.85kg (150.0mol) of N-methylpyrrolidone, 8.33kg of 48.0wt% aqueous sodium hydroxide solution (the number of moles of NaOH is 100.0mol), and 11.93kg of 47.0wt% aqueous sodium hydrosulfide solution (the number of moles of NaHS is 100.0mol). Slowly heat to 203℃ under nitrogen atmosphere, continuously remove water during the process, take out 9.47kg of distillate (water content 98.0%), and after dehydration is completed, lower the temperature to 170℃. At this time, the amount of sulfur in the system is 98.0mol, and the water content is 176.4mol. The dehydrated liquid produced in the above batch method is put into a storage kettle at 170℃ and stored at a constant temperature for use in subsequent continuous supply.

[0087] The continuous manufacturing process of PAS shown in Figure 1 is used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content is 29.83 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of sodium chloride seed crystals is 15 μm, concentration of crystal slurry is 0.2 mol / L) are fed to the first polymerization reactor 4 (PDCB / S=1.025, NMP) at flow rates of 24.60 g / min, 42.71 g / min and 8.1 g / min, respectively. The first polymerization reactor 4 had a temperature of 215°C and a residence time of 3.5h. The obtained first reaction mixture 5 was heated to 225°C through a first heat exchanger 6 and entered a second tubular reactor 7 (reactor temperature was 225°C). After residence for 1h, it entered a closed pressure filter 8 (screen pore size was 80μm) to remove salt particles. The filtered salt particles 9 had a D50 of 180μm. The obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (the temperature of the reactor is 265°C) with a residence time of 1 hour. Then, it is heat exchanged through a heat exchanger (not shown) to 280°C and enters the atmospheric pressure flash evaporator 14 for flash evaporation. During the flash evaporation, superheated steam 13 at 265°C is used to assist the flash evaporation, and the flow rate of the superheated steam 13 is 2.8 kg / h. The vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, and the waste liquid 21 to be recovered is generated. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 enters a tubular washer equipped with a heat transfer oil jacket, where the temperature of the tubular washer is controlled at 180°C and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown), after which the final product is obtained.

[0088] The production rate of this example was 1.01 kg / h. The amount of saline wastewater produced was 5.0 kg / h, with a salt content of 0.58 wt%, the amount of superheated steam consumed was 2.8 kg / h, and the amount of wastewater to be recovered was 5.1 kg / h. The final product produced was measured to have a weight average molecular weight of about 32000 and an ash content of less than 0.52 wt%.

[0089] Compared with Example 2 (the molar ratio is 3.5) which is a conventional NMP / S molar ratio, by using the continuous production process of the present invention, polyphenylene sulfide was successfully produced at an NMP / S molar ratio of 2.5 (Example 4) and an NMP / S molar ratio of 2.0 (Example 6), and not only was energy consumption (amount of superheated steam consumption) and the amount of waste liquid, waste gas, and waste material (amount of waste liquid to be recovered) significantly reduced, but the residence time (reaction time) was shortened and the reaction efficiency was significantly improved.

[0090] Taking Example 6 as an example, comparing Example 2, the consumption of superheated steam is reduced by about half, the amount of waste liquid to be recovered is reduced by more than 1 / 3, and the total residence time of the reaction is reduced by about 20%, and the above technical effects can significantly reduce the production cost of the continuous production.

[0091] Comparative Example 3 In a 100L reactor, 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg (100.0mol) of 48wt% aqueous sodium hydroxide solution, and 11.93kg (100.0mol) of 47wt% aqueous sodium hydrosulfide solution are added, and then the temperature is continuously raised under a nitrogen atmosphere to dehydrate, and 10.51kg of aqueous solution (water content 98.0wt%) is removed, and the temperature is lowered to 170℃ after dehydration is completed. At this time, the amount of sulfur in the dehydration liquid system is 98.0mol, and the water content is 117.6mol. The dehydration liquid produced in the above batch method is put into a storage kettle at 170℃ to be stored at a constant temperature, and prepared for use in subsequent continuous supply.

[0092] The flow chart of the continuous manufacturing process of PAS shown in Figure 1 is used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content is 25.86 wt%) and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of sodium chloride seed crystals is 15 μm, and the concentration of the crystal slurry is 0.136 mol / L) are fed to the first polymerization reactor 4 (PDCB / S=1.05) at flow rates of 25.21 g / min, 49.27 g / min and 24.5 g / min, respectively. The first polymerization reactor 4 had a temperature of 215°C and a residence time of 5h. The first reaction mixture 5 obtained was heated to 225°C through a first heat exchanger 6 and entered a second tubular reactor 7 (reactor temperature was 225°C). After residence for 1h, it entered a closed pressure filter 8 (screen pore size was 30μm) to remove salt particles. The D50 of the filtered salt particles 9 was 60μm. The obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (the temperature of the reactor is 265°C) with a residence time of 1.5 hours. Then, it is heat exchanged through a heat exchanger (not shown) to 280°C and enters the atmospheric pressure flash evaporator 14 for flash evaporation. During the flash evaporation, superheated steam 13 at 265°C is used to assist the flash evaporation, and the flow rate of the superheated steam 13 is 5.0 kg / h. The vaporized solvent and steam 19 are condensed and liquefied through the condenser 20, and the waste liquid 21 to be recovered is generated. The undried crude product 15 first passes through the dryer supply chamber 16 and enters the continuous dryer 17. The crude product 18 obtained after further drying continuously enters the beating tank, where water is added at 6.0 kg / h to be beaten, and then heat-exchanged to 180°C. The crude product 18 then enters a tubular washer equipped with a heat transfer oil jacket. The temperature of the tubular washer is controlled at 180°C, and the residence time is 10 minutes. The crude product 18 is then cooled to 60°C through heat exchange, and continuously filtered and dried (not shown) to obtain the final product.

[0093] The production rate of this comparative example was 1.01 kg / h. The amount of salt-containing wastewater produced was 5.0 kg / h, with a salt content of 0.65 wt%, the amount of superheated steam consumed was 5.0 kg / h, and the amount of wastewater to be recovered was 8.7 kg / h. The final product produced was measured to have a weight average molecular weight of about 23,000 and an ash content of less than 0.6 wt%.

[0094] In addition, in the process of carrying out this comparative example, the particle size of the salt and the corresponding pore size of the screen were small, so the resistance of the filter was large and clogging occurred from time to time, making the process less stable when carried out continuously. [Explanation of symbols]

[0095] 1. Paradichlorobenzene 2 Dehydrated liquid 3. NMP crystal slurry containing sodium chloride seed crystals 4. First Polymerization Reactor 5. First reaction mixture 6 1st heat exchanger 7. Second tubular reactor 8. Closed pressure filter 9. Salt particles 10 Second reaction mixture 11 Second heat exchanger 12 Third tubular reactor 13 Superheated steam 14 Atmospheric pressure flash evaporator 15 Undried crude product 16 Dryer supply room 17 Continuous Dryer 18 Crude product after drying 19 Vaporized solvents and vapours 20 Condenser 21 Waste liquid to be collected

Claims

1. A method for continuously producing polyarylene sulfide, comprising the steps of: (1) charging an aqueous sulfur source solution, an aqueous alkali metal hydroxide solution, an optional auxiliary and an organic amide solvent into a reaction vessel, continuously heating the reaction vessel to 200-220°C to carry out a dehydration and saponification reaction, and after completion of the dehydration, the molar ratio of water to sulfur source is 1.5-2.0; and (2) cooling the reaction vessel to 150-180°C, and storing the resulting dehydrated liquid in a storage vessel to keep it warm; (2) injecting the organic amide crystal slurry containing sodium chloride seed crystals, the dichloro aromatic compound and the dehydration liquid into a prepolymerization reactor, controlling the reaction temperature at 180-240°C, and reacting until the monomer conversion rate reaches 90% or more, and then injecting the resulting reaction liquid into a filtration device and filtering to remove sodium chloride particles; and (3) injecting the filtrate obtained through the filtration device into a polymerization reactor, controlling the reaction temperature at 235 to 280°C, reacting until the monomer conversion rate reaches 95% or more, subjecting the resulting reaction liquid to a flash evaporation treatment, and finally obtaining polyarylene sulfide through post-treatment.

2. In step (1), The concentration of the sulfur source aqueous solution is 28 to 48 wt %, and the sulfur source is selected from sodium hydrosulfide and / or sodium sulfide; The alkali metal hydroxide aqueous solution is a sodium hydroxide aqueous solution having a concentration of 30 to 70 wt %, the dichloroaromatic compound is selected from one or more of paradichlorobenzene, dichloronaphthalene, dichlorofluorene, dichlorocarbazole; the molar ratio of the sulfur source to sodium oxide is 1:0.95-1.1; The auxiliary is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; The molar ratio of the sulfur source to the coagent is 1:0-0.5; the organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl-ε-caprolactam, N,N-dimethylformamide; 2. The method for continuously producing polyarylene sulfide according to claim 1, wherein the molar ratio of the sulfur source to the organic amide solvent is 1:1.5-2.

5.

3. In step (2), The concentration of the sodium chloride seed crystals in the organic amide crystal slurry is 0.1 to 1 mol / L, and the particle size of the sodium chloride seed crystals is 1 to 20 μm; the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0 to 1.1:1; The method for continuously producing polyarylene sulfide according to claim 1, characterized in that after mixing the dehydrating liquid, the dichloro aromatic compound, and the organic amide crystal slurry containing sodium chloride seed crystals, the amount of the sodium chloride seed crystals added is 0.05 to 0.4 mol% relative to 1 mol of sulfur source in the system, and the total amount of the organic amide solvent in the system is 2.0 to 4.0 mol.

4. the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0 to 1.025:1; The method for continuously producing polyarylene sulfide according to claim 3, characterized in that after mixing the dehydrating liquid, the dichloro aromatic compound, and the organic amide crystal slurry containing sodium chloride seed crystals, the total amount of the organic amide solvent is 2.0 to 2.5 mol per mol of the sulfur source in the system.

5. In step (2), the prepolymerization reactor is a two-stage reactor connected in series; The first stage reactor is selected from continuous stirred tank reactors, and the reaction temperature in the tank reactor is controlled at 180-220°C; The second stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled at 220-240°C; 2. The method for continuously producing polyarylene sulfide according to claim 1, characterized in that after the monomer conversion rate in the first-stage reactor reaches 60 to 80%, the reaction liquid is injected into the second-stage reactor.

6. In step (3), The polymerization reactor is selected from tubular reactors; 2. The method for continuously producing polyarylene sulfide according to claim 1, wherein the post-treatment includes drying, washing, filtering and re-drying.

7. A batch-type method for producing polyarylene sulfide, comprising the steps of: (a) charging an aqueous sulfur source solution, an aqueous alkali metal hydroxide solution, an optional auxiliary and an organic amide solvent into a reactor I, and continuously heating the reactor to 200-220°C to carry out a dehydration and saponification reaction, and the molar ratio of water to sulfur source after the completion of dehydration is 1.5-2.0, and then cooling the reactor to 150-180°C; (b) charging the organic amide crystal slurry containing sodium chloride seed crystals, the dichloro aromatic compound and the additional organic amide solvent into said reactor I, first heating to 205-220°C and keeping at the same temperature for 0.5-5h, then heating to 215-240°C and keeping at the same temperature for 0.5-5h, and then pouring the resulting reaction liquid into a filter to filter and remove sodium chloride particles; and step (c) of injecting the filtrate obtained through the filtration device into a preheated reaction vessel II, heating the filtrate to 240 to 280°C, and maintaining the temperature for 0.5 to 5 hours, and then subjecting the resulting reaction liquid to a flash evaporation treatment, and finally subjecting the resulting reaction liquid to a post-treatment to obtain polyarylene sulfide.

8. In step (a), The concentration of the sulfur source aqueous solution is 28 to 48 wt %, and the sulfur source is selected from sodium hydrosulfide and / or sodium sulfide; The alkali metal hydroxide aqueous solution is a sodium hydroxide aqueous solution having a concentration of 30 to 70 wt %, the dichloroaromatic compound is selected from one or more of paradichlorobenzene, dichloronaphthalene, dichlorofluorene, dichlorocarbazole; the molar ratio of the sulfur source to sodium oxide is 1:0.95-1.1; The auxiliary is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; The molar ratio of the sulfur source to the coagent is 1:0-0.5; the organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl-ε-caprolactam, N,N-dimethylformamide; The batch-type method for preparing polyarylene sulfide according to claim 7, characterized in that the molar ratio of the sulfur source and the organic amide solvent is 1:1.5-3.

9. In step (b), The concentration of the sodium chloride seed crystals in the organic amide crystal slurry is 0.1 to 1 mol / L, the particle size of the sodium chloride seed crystals is 1 to 20 μm, and the amount of the sodium chloride seed crystals added is 0.05 to 0.4 mol % of the sulfur source in the system at this time; the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0 to 1.1:1; After adding the organic amide solvent, the molar ratio of the total molar amount of the organic amide solvent to the sulfur source in the reaction vessel I is 2.5 to 3.5; First, the temperature is increased to 205-220°C at a rate of 0.2-0.5°C / min, and then increased to 215-240°C at a rate of 0.2-1°C / min, In step (c), The batch-type method for producing polyarylene sulfide according to claim 7, characterized in that the temperature is increased to 240-280°C at a rate of 0.2-1°C / min, and the post-treatment includes drying, washing, filtering and redrying.

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 in producing a crosslinked polyarylene sulfide, characterized in that the polyarylene sulfide is subjected to a thermal oxidation treatment.

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