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 more particularly to continuous and batch processes for the production of polyarylene sulfide resins and their products and uses. [Background technology]

[0002] The typical production method for polyarylene sulfide (PAS) involves the polymerization of sodium sulfide / sodium hydrosulfide with a dihaloaromatic compound in the presence of an alkali metal hydroxide in an organic amide solvent. The solvent is typically N-methylpyrrolidone (NMP), and the alkali metal hydroxide is typically sodium hydroxide. Currently, industrial production systems for polyphenylene sulfide (PPS) using this method are batch processes. A portion of the reactants are first dehydrated in a reactor, cooled, and then the remaining reactants are added. The reactor is then sealed and heated for high-temperature polycondensation. After the reaction is complete, the crude product is obtained by cooling filtration or flash evaporation, and then purified to obtain the final product. Batch production requires repeated material addition, heating and cooling, removal, and reactor cleaning, which is time-consuming, significantly reduces production efficiency, and also results in poor process stability and reliability. Therefore, researchers have been striving to develop continuous PAS production methods to improve production efficiency.

[0003] Patents US4060520, US4056515, and US4066632 provide a continuous production process in which multiple reactors are connected in series, with each reactor having a different reaction temperature, the reactor pressure decreasing continuously from front to rear, and the pressure difference between adjacent reactors causing the raw materials to flow from front to rear. Chinese patent publication number CN108779253A provides a continuous PAS production apparatus and method, which includes a chamber containing multiple reaction vessels, and supplies an organic amide solvent, a sulfur source, and a dihaloaromatic compound to the chamber, where the sulfur source and the dihaloaromatic compound are polymerized in the organic amide solvent to form a reaction mixture. The multiple reaction vessels are connected in series and communicate with each other via the gas phase, and the reaction mixture flows sequentially through each reaction vessel. A Chinese patent document 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 by metering pumps to series-connected kettle or tubular reactors.

[0004] The sodium sulfide process for PPS production generates a large amount of sodium chloride as a by-product, producing 1.08 tons of sodium chloride for every ton of PPS produced. Because sodium chloride is formed and precipitates in large quantities early in the polymerization reaction, the reaction slurry contains a large amount of salt particles. Based on a typical NMP / sulfur source molar ratio of 3.5 in the reaction solution, the by-product salt particle content in the polymerization solution after the reaction is completed is approximately 20.5 wt%. The large amount of salt particles in the slurry tends to settle during long-distance transport or in reaction pipelines, especially in narrow tubular reactors, causing clogging and production stoppages. Increasing the concentration of the reactant, whether in a batch or continuous process, is beneficial for improving production efficiency. A high concentration of the reactant is also beneficial for the efficient production of high-molecular-weight polyphenylene sulfide. However, the by-product salt particle content also increases significantly, and the viscosity of the reaction solution also increases sharply, exacerbating the clogging and mixing problems during transport. The presence of a large amount of by-product salt particles in the PPS reaction slurry severely limits research and development of continuous PPS production methods, and also limits the efficient production of higher molecular weight polyphenylene sulfide.

[0005] Whether using a batch or continuous process, the by-product sodium chloride remains on the product after the reaction is complete. This requires multiple washes using large amounts of pure water, resulting in the generation of large amounts of high-salt wastewater. Currently, industrial production consumes more than 15 tons of water per ton of PPS. The by-product salt dissolved in the water is evaporated and recovered, further increasing energy consumption. Furthermore, the corrosion caused by saltwater promotes corrosion, placing high requirements on the materials used for polyphenylene sulfide production equipment. The high water and energy consumption and corrosion-promoting characteristics of the polyphenylene sulfide production process significantly increase the production and environmental costs of polyphenylene sulfide, significantly limiting the ecological and sustainability aspects of the polyphenylene sulfide production process and limiting the competitiveness and development potential of polyphenylene sulfide resin.

[0006] These problems can be solved by removing the precipitated salt particles before solid-liquid separation of the product, thereby reducing the content of salt particles in subsequent equipment and in the crude product.For example, Chinese patent document publication number CN104371103A discloses a method for desalination of polyphenylene sulfide, in which after the polymerization reaction is completed, the reactor is kept at a constant temperature and pressure, and the inorganic salts and inorganic auxiliary agents settle to the bottom of the reactor. After that, the pressure difference is used to extrude the upper layer of polyphenylene sulfide solution from the reactor through the inner extension pipe at the bottom, and the solution is further 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 processes, and the patent document does not further explain how to further treat the inorganic salts in the sediment layer in the reactor, the large amount of reaction solvent, and the product.In fact, it is not possible to restart the stirring of the polymerization reactor before the settled inorganic salt layer is completely treated.

[0007] Furthermore, for example, Chinese patent document publication number CN107964098A relates to a method and apparatus for desalination in the synthesis of polyphenylene sulfide, in which during polymerization, a material conveying pump is used to pass the raw material from the reactor through a filter with a screen, the salt remains in the filter, and the remaining polyphenylene sulfide raw material is returned to the reactor, circulating in this manner 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 stages of the reaction, filtration becomes more difficult as the thickness of the salt filter cake layer and the viscosity of the reaction liquid increase. Summary of the Invention

[0008] In response to the above-mentioned problems in the prior art, the present invention discloses a process for producing polyarylene sulfide resin, which can significantly reduce the salt content in the crude polyarylene sulfide resin produced, significantly reduce the generation of high-salt wastewater, reduce water consumption and energy consumption, and is useful for producing high-molecular-weight polyarylene sulfide resin.More importantly, this process can be applied not only to batch processes but also to continuous production, which is expected to significantly 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 optional auxiliary agent, 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 dehydration is completed, the molar ratio of water to sulfur source is 1.5-2.0. The reaction mixture is cooled to 150-180°C, and the resulting dehydrated liquid is kept warm in a storage vessel.

[0011] (2) The organic amide crystal slurry containing sodium chloride seed crystals, the dichloroaromatic compound, and the dehydrated liquid are poured 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 poured into a filtration device and filtered to remove sodium chloride particles.

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

[0013] The inventors conducted extensive research into the growth process and particle morphology of salt particles in PPS reaction solutions. After all raw materials are added and the temperature is raised to a predetermined level, the polymerization reaction begins, and sodium chloride is continuously produced. The presence of water in the polyphenylene sulfide reaction system allows a small amount of by-product sodium chloride to dissolve in the reaction solvent. Once saturated, the continuously produced sodium chloride nucleates and begins to precipitate. The reaction continues rapidly, and the large amount of sodium chloride produced rapidly nucleates and precipitates. The inventors' research findings include the following: (1) In conventional batch PPS production processes, the salt particles produced in the reaction solution are small, around 20 μm in diameter, making them difficult to filter, a point that has rarely been reported. (2) In the early prepolymerization stage, the reaction conversion rate exceeds 90%, and sodium chloride crystal growth is nearly complete. However, the product still has a low molecular weight and the reaction solution is not yet viscous. The molecular weight must be increased by entering the high-temperature polymerization stage. Therefore, filtering of the by-product salt is best performed when sodium chloride crystal growth is essentially complete in the prepolymerization stage. After the high-temperature polymerization stage, the reaction solution becomes more viscous and difficult to filter. (3) Controlling the water content in the system within an appropriate range can, on the one hand, slow down the rate of sodium chloride precipitation and nucleation and, on the other hand, accelerate the rate of sodium chloride crystal growth. (4) Introducing a small amount of sodium chloride seed crystals can effectively inhibit primary nucleation of sodium chloride.

[0014] Based on these research findings, the present invention proposes a technical solution. Testing results show that in step (1), in order to control the particle size of the sodium chloride particles to 100 μm or more and facilitate filtration, it is necessary to control the water content after dehydration to an appropriate ratio (the molar ratio of water to sulfur source must be controlled between 1.5 and 2.0). The sodium chloride can be removed by filtration, significantly reducing the generation of high-salt wastewater and reducing water and energy consumption. At the same time, the amount of solvent in the reaction solution can be further reduced, contributing 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, and the concentration is 30 to 70 wt %;

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

[0017] The auxiliary agent is selected from one or more of sodium acetate, sodium benzoate, and sodium salts of C5-C6 fatty acids; The molar ratio of the sulfur source to the auxiliary agent 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, and 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 saponification reaction may be carried out batchwise according to the method of step (1), or may be carried out using a continuous dehydration method reported in other patents. The key point is to control the water content after completion of dehydration within the above-mentioned 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 to 1 mol / L and the particle size of the sodium chloride seed crystals is 1 to 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 sodium chloride seed crystals, the amount of the sodium chloride seed crystals added is controlled to be 0.05 to 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 to 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 appropriately 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 test, it was found that the weight average molecular weight of the PAS produced can be significantly improved by controlling the molar ratio of the dichloroaromatic compound to the sulfur source and the NMP / S ratio within the above more preferred ranges.

[0027] The inventors further conducted intensive research into the polymerization kinetics of polyphenylene sulfide and discovered the following: (1) The main polymerization reaction of polyarylene sulfide is a highly exothermic reaction, and temperature runaway during prepolymerization has a significant impact on reaction safety and product quality. Furthermore, the polymerization reaction is a second-order kinetic reaction. Therefore, in a continuous production process in which multiple reactors are connected in series, in order to meet the requirements for reaction temperature control and reaction efficiency, the first reactor is preferably a continuous stirred tank reactor, and the conversion rate is controlled within an appropriate range. (2) Increasing the water content in the reaction system can slow the reaction rate, i.e., the rate of sodium chloride production. However, too high a water content can cause problems such as low reaction efficiency, high reaction pressure, and increased energy consumption during flash evaporation.

[0028] To summarize the above research results, large amounts of by-product salts are generated during the conventional batch production process of PAS and during the development of a continuous production process. (1) Removing the by-product salts by washing results in high water consumption and discharge of high-salt wastewater. (2) The problem of clogging is a major limitation on the development of a continuous PAS production process. (3) The increase in substrate concentration is limited, and the main polymerization reaction is a second-order kinetic, highly exothermic reaction, which places high demands on temperature control during the reaction process. In response to these various issues, the present invention provides an environmentally friendly, low-cost, and highly efficient continuous PAS production process.

[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 to 220°C, preferably 190 to 210°C.

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

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

[0032] Preferably, the heating medium used in the jacketed tubular reactor is heat transfer oil, the flow direction of the heat transfer oil is the same as the flow direction of the reaction liquid, the oil temperature at the inlet is controlled to 220 to 230° C., and the oil temperature at the outlet is 230 to 240° C. Preferably, the monomer conversion rate at the outlet reaches 90% or more, more preferably 93% or more, and at this time, the growth of sodium chloride crystals is almost complete, the particle size has grown 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, and is applied to continuous filtration of high-temperature slurries containing volatile components and crystals that are prone to precipitate.

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

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

[0036] Preferably, the flash evaporation treatment uses superheated steam at 260-300°C as auxiliary heat supply, and 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 before entering 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 subjected to one continuous high-temperature, high-pressure water wash. The mass ratio of the added 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 device may be a continuous tank type washing device or a continuous pipe type washing device, and is preferably a continuous pipe type washing device 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 continuous filters such as a duplex filter press, a decanter type centrifugal separator and a belt filter, and is preferably a belt filter for continuous vacuum suction filtration.

[0042] The dryer is selected from continuous feed and discharge dryers. The production process of the present invention allows for the pre-filtration and removal of large salt particles, significantly reducing the salt concentration in the subsequent crude product and significantly reducing the pressure required for the post-treatment washing process, thereby requiring only one continuous high-temperature, high-pressure water wash. Without pre-filtration of the salt particles, several cycles of atmospheric pressure and high-pressure water washes would be required to thoroughly wash out the sodium chloride in the product, significantly increasing the amount of washing water and the amount of salt-containing wastewater produced. Without the addition of seed crystals and the inability to properly control the water content of the reaction solution, the sodium chloride particles are too small to be effectively filtered using a filter with a large pore size. Using a filter with a small pore size results in excessive pressure loss in the filter, making it prone to clogging.

[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 further below.

[0044] Therefore, the present invention further discloses a batchwise method for producing polyarylene sulfide, comprising the steps of:

[0045] (a) An aqueous solution of a sulfur source, an aqueous solution of an alkali metal hydroxide, an optional auxiliary agent, and an organic amide solvent are charged into a reactor I, and the temperature is continuously raised to 200-220°C to carry out a dehydration and saponification reaction, and after the dehydration, the molar ratio of water to sulfur source is 1.5-2.0, and then the reaction mixture 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 reactor I, and the temperature is first raised to 205-220°C and maintained at that temperature for 0.5-3 hours, and then raised to 215-240°C and maintained 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 preheated reactor II, heated to 240 to 280°C, and kept at that temperature for 0.5 to 5 hours. The resulting reaction liquid is then subjected to flash evaporation, 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, and the concentration is 30 to 70 wt %; the dichloroaromatic compound is selected from one or more of paradichlorobenzene, dichloronaphthalene, dichlorofluorene, and dichlorocarbazole, and is preferably paradichlorobenzene; the molar ratio of the sulfur source to sodium hydroxide is 1:0.95 to 1.1; The auxiliary agent is selected from one or more of sodium acetate, sodium benzoate, and sodium salts of C5-C6 fatty acids; 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-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl-ε-caprolactam, and 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 dehydrated 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 reactor 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 further 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 a heat transfer oil jacket and a coil.

[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 for 10 minutes to 1 hour in the flash evaporator, 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 essence of which is to add sodium chloride seed crystals to the PAS reaction solution, while simultaneously controlling the water content in the system within an appropriate range, and controlling the reaction temperature in the prepolymerization stage to a relatively low temperature, thereby slowing the nucleation rate of sodium chloride, accelerating the crystal growth rate, suppressing primary nucleation, and promoting secondary nucleation, ultimately improving the size of the by-product sodium chloride crystals. Another key point in this process is the timing of filtering to remove salt particles, which is performed at the end of the prepolymerization reaction stage. The production process disclosed in this invention significantly reduces the salt content in the resulting polyarylene sulfide resin crude product, significantly reducing the generation of high-salt washing wastewater, reducing water and energy consumption, and contributing to the efficient production of high-molecular-weight polyarylene sulfide resin. More importantly, this process can be applied not only to batch processes but also to continuous production, which is expected to significantly reduce the production efficiency and production costs of the production system and be more environmentally friendly. When using a continuous production process, in order to balance temperature control and production efficiency in the prepolymerization stage, the prepolymerization stage is divided into two. For the first stage, a tank reactor suitable for stable heat transfer is selected, and by controlling the conversion rate to an appropriate value, most of the reaction heat is removed while preventing the reaction rate from becoming too low. For the second stage, a jacketed tubular reactor is selected, and the remaining heat is used to promote the reaction and continue to increase the temperature to maintain reaction efficiency. For the high-temperature polymerization stage, a coil reactor is selected after removing salt particles, which is more suitable for high conversion and high viscosity conditions (high molecular weight and high concentration). [Brief explanation of the drawings]

[0058] [Figure 1] FIG. 1 is a schematic diagram of the process flow for continuous production of polyarylene sulfide of the present invention, omitting post-treatment processes. DETAILED DESCRIPTION OF THE INVENTION

[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: A porcelain crucible is placed in a constant-temperature muffle furnace at 750°C and fired to a constant weight. It is then removed and placed in a desiccator to cool and then weighed. This is designated M0. A 3g sample is weighed and placed in the crucible. 10mL of nitric acid is added, and the mixture is placed over an alcohol burner and burned until no more smoke is produced. Finally, the crucible is placed in a constant-temperature muffle furnace at 750°C and fired for 1 hour. It is then removed and placed in a desiccator to cool and then weighed. This is designated M1. The formula for calculating resin ash content is as follows: (M1-M0) / 3*100%.

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

[0062] Example 1 Into a 100L reactor I, 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) were added, and the temperature was slowly raised to 200°C under a nitrogen atmosphere, constantly removing water during the process, and 10.00kg of distillate (water content 98.0%) was taken out. After dehydration was completed, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 147.0mol.

[0063] To 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 the sodium chloride seed crystals is 15 μm) were added, and 4.46 kg of NMP (45.0 mol) was further added. 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. After that, the temperature was increased to 230°C at 0.2°C / min and kept at that temperature for 1 hour. After that, the reaction solution 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 a rate of 0.5℃ / min, kept at this temperature for 1 hour, and then slowly released into an atmospheric flash evaporator for flash evaporation. Finally, the mixture was dried to further remove the solvent, yielding 10.4kg of crude product. The mixture was then beaten with 60kg of water, heated to 180℃, kept at this 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 after the reaction mixture was transferred and filtered. The D50 of the sodium chloride particles measured using a laser particle size analyzer was 130 μm. 60.0 kg of water was used in the washing process, and 50.0 kg of saline wastewater was produced, with a salt content of 0.6 wt%. The final product produced was 10.1 kg, with a weight-average molecular weight of 22,500 and an ash content of 0.53 wt%.

[0065] Comparative Example 1 Into a 100L reactor I, 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) were added, and the temperature was slowly raised to 200°C under a nitrogen atmosphere, constantly removing water during the process, and 10.00kg of distillate (water content 98.0%) was taken out. After dehydration was completed, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 147.0mol.

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

[0067] In this comparative example, 10.5 kg of sodium chloride was obtained after the reaction solution was transferred and filtered. The D50 of the sodium chloride particles measured with a laser particle size analyzer was 30 μm. 60.0 kg of water was used in the washing process, and 51.0 kg of saline wastewater was produced, with a salt content of 1.3 wt%. The final product produced was 10.1 kg, with a weight-average molecular weight of 22,500 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 significantly reduces production efficiency and significantly increases production costs.

[0069] Comparative Example 2 Into a 100L reactor I, 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) were added, and the temperature was slowly raised to 200°C under a nitrogen atmosphere, constantly removing water during the process, and 10.00kg of distillate (water content 98.0%) was taken out. After dehydration was completed, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 147.0mol.

[0070] Next, 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, the D50 of the sodium chloride seed crystals is 15 μm) was added to reactor I, and 4.46 kg of NMP (45.0 mol) was added. The reactor was sealed and heated to 215 ° C at 0.3 ° C / min. and kept at that temperature for 3 hours. Then, the temperature was raised to 230 ° C at 0.15 ° C / min. and kept at that temperature for 1 hour. Then, the temperature was raised to 260 ° C at 0.5 ° C / min. and kept at that temperature for 1 hour. After that, the temperature was slowly released into a normal pressure flash evaporator for flash evaporation, and further dried 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. After beating with an additional 60.0 kg of water, the mixture was 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, 180.0 kg of water was used in the washing process, 182.0 kg of saline wastewater was produced, and the salt content was 6.0 wt%. The final product produced was 10.0 kg, with a weight-average molecular weight of 21,500 and an ash content of 0.50 wt%.

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

[0073] Example 2 A 100L reactor was charged with 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (100.0mol of NaOH), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (100.0mol of NaHS). The temperature was slowly raised to 200°C under a nitrogen atmosphere, continuously removing water during the process. 10.00kg of distillate (water content 98.0%) was removed. After dehydration was complete, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol and the water content was 147.0mol. The dehydrated solution produced in the above batches was transferred to a storage vessel at 170°C for warm storage and prepared for subsequent continuous feed.

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

[0075] The production rate in 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 A batch of dehydrated liquid was produced according to the method of Example 2 and placed in a storage tank and kept at 170°C for subsequent continuous supply. The PAS continuous production apparatus shown in Figure 1 was used. Paradichlorobenzene 1 (PDCB), dehydrated liquid 2 (sodium sulfide content 25.48 wt%), and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of the sodium chloride seed crystals was 15 μm, and the concentration of the crystal slurry was 0.4 mol / L) were 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 pumped into a first polymerization reactor 4 (P / S=2.5), the temperature of which was 215°C, and the residence time was 3.5 hours. 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 residence for 1 hour, entered a sealed pressure filter 8 (screen pore size 80μm) to remove salt particles. The D50 of the filtered salt particles 9 was 130μm. The obtained second reaction mixture 10 passes through a second heat exchanger 11, where it is heated to 265°C, and enters a third tubular reactor 12 (the reactor temperature is 265°C), where it has a residence time of 1 hour.It then passes through a heat exchanger (not shown) and is heat-exchanged up to 280°C, and then enters an atmospheric pressure flash evaporator 14 to undergo 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 in a condenser 20, whereby waste liquid 21 to be recovered is produced. The undried crude product 15 first passes through a dryer supply chamber 16 and enters a continuous dryer 17. After further drying, the resulting crude product 18 continuously enters a beating tank, where water is added at a rate of 6.0 kg / h and beaten. After heat exchange to 180°C, the crude product enters a pipe washer equipped with a heat transfer oil jacket. The temperature of the pipe washer is controlled at 180°C, and the residence time is 10 minutes. The crude product is then cooled to 60°C via heat exchange, and is continuously filtered and dried (not shown) to obtain a final product.

[0077] The production rate in 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 Example 2 and Example 3, by further reducing the NMP / S molar ratio in the prepolymerization reaction stage, specifically in the first polymerization reactor, the amount of solvent used and the amount of superheated steam used in the flash evaporation stage can be reduced, which reduces energy consumption and waste liquid and significantly reduces production costs, improves reaction efficiency, shortens the residence time of the reaction liquid in each reactor stage, and appropriately improves the weight-average molecular weight of the final product.

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

[0080] Example 4 A 100L reactor was charged with 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide (100.0mol of sodium hydroxide), 11.93kg of 47wt% aqueous sodium hydrosulfide (100.0mol of sodium hydrosulfide), and 1.6kg of anhydrous sodium acetate. The temperature was then continuously increased under a nitrogen atmosphere to dehydrate the mixture. 10.00kg of aqueous solution (water content 98.0wt%) was removed. After dehydration was complete, the temperature was lowered to 170°C. At this time, the amount of sulfur in the dehydrated solution was 98.0mol, and the water content was 147.0mol. The dehydrated solution produced in the above batches was transferred to a storage vessel at 170°C for warm storage and prepared for subsequent continuous feed.

[0081] The PAS continuous production equipment shown in Figure 1 was used. Preheated paradichlorobenzene (PDCB) 1, dehydrated liquid 2 (sodium sulfide content 24.22 wt%), and NMP crystal slurry 3 containing sodium chloride seed crystals (particle size of the sodium chloride seed crystals was 15 μm, and the concentration of the crystal slurry was 0.4 mol / L) were 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 The temperature of the first polymerization reactor 4 was 215°C, and the residence time was 3.5 hours. 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). After residence for 1 hour, it entered a sealed pressure filter 8 (screen pore size 80μm) to remove salt particles. The D50 of the filtered salt particles 9 was 120μm. The obtained second reaction mixture 10 passes through a second heat exchanger 11, where it is heated to 265°C, and then enters a third tubular reactor 12 (the reactor temperature is 265°C), where it has a residence time of 1 hour.It then passes through a heat exchanger (not shown) where it is heat exchanged up to 280°C, and then enters an atmospheric pressure flash evaporator 14 where it undergoes 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 3.3 kg / h.The vaporized solvent and steam 19 are condensed and liquefied in a condenser 20, and the waste liquid to be recovered 21 is produced. The undried crude product 15 first passes through a dryer supply chamber 16 and enters a continuous dryer 17. After further drying, the resulting crude product 18 continuously enters a beating tank, where water is added at a rate of 6.0 kg / h and beaten. After heat exchange to 180°C, the crude product enters a pipe washer equipped with a heat transfer oil jacket. The temperature of the pipe washer is controlled at 180°C, and the residence time is 10 minutes. The crude product is then cooled to 60°C via heat exchange, and is continuously filtered and dried (not shown) to obtain a final product.

[0082] The production rate in 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 A 100L reactor was charged with 19.82kg (200.0mol) of N-methylpyrrolidone, 8.33kg of 48wt% aqueous sodium hydroxide solution (100.0mol of NaOH), and 11.93kg of 47wt% aqueous sodium hydrosulfide solution (100.0mol of NaHS). The temperature was slowly raised to 195°C under a nitrogen atmosphere, continuously removing water during the process. 9.07kg of distillate (water content 98.0%) was removed. After dehydration was complete, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol and the water content was 196.0mol. The dehydrated solution produced in the above batches was transferred to a storage vessel at 170°C for warm storage and prepared for subsequent continuous feed.

[0084] The PAS continuous production process flow chart shown in Figure 1 was used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content 24.66 wt%), and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of the sodium chloride seed crystals is 15 μm, and the concentration of the crystal slurry is 0.136 mol / L) were fed to a first polymerization reactor 4 (PDCB / S = 1.050, The resulting 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), where it remained for 1.5 hours before entering a sealed pressure filter 8 (screen pore size 80μm) to remove salt particles. The D50 of the filtered salt particles 9 was 150μm. The obtained second reaction mixture 10 passes through a second heat exchanger 11, where it is heated to 265°C, and then enters a third tubular reactor 12 (the reactor temperature is 265°C) for a residence time of 1.5 hours.Then, it passes through a heat exchanger (not shown) and is heat-exchanged up to 280°C, and then enters an 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 in a condenser 20, and a waste liquid 21 to be recovered is produced. The undried crude product 15 first passes through a dryer supply chamber 16 and enters a continuous dryer 17. After further drying, the resulting crude product 18 continuously enters a beating tank, where water is added at a rate of 6.0 kg / h and beaten. After heat exchange to 180°C, the crude product enters a pipe washer equipped with a heat transfer oil jacket. The temperature of the pipe washer is controlled at 180°C, and the residence time is 10 minutes. The crude product is then cooled to 60°C via heat exchange, and is continuously filtered and dried (not shown) to obtain the final product.

[0085] The production rate in 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 A 100L reactor was charged with 14.85kg (150.0mol) of N-methylpyrrolidone, 8.33kg of 48.0wt% aqueous sodium hydroxide solution (100.0mol of NaOH), and 11.93kg of 47.0wt% aqueous sodium hydrosulfide solution (100.0mol of NaHS). The temperature was slowly raised to 203°C under a nitrogen atmosphere, continuously removing water during the process. 9.47kg of distillate (water content 98.0%) was removed. After dehydration was complete, the temperature was lowered to 170°C. At this time, the amount of sulfur in the system was 98.0mol, and the water content was 176.4mol. The dehydrated solution produced in the above batches was transferred to a storage vessel at 170°C for warm storage and prepared for subsequent continuous feed.

[0087] The continuous PAS production process flow chart shown in Figure 1 was used. Paradichlorobenzene (PDCB) 1 preheated to 200°C, dehydration liquid 2 (sodium sulfide content 29.83 wt%), and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of the sodium chloride seed crystals is 15 μm, and the concentration of the crystal slurry is 0.2 mol / L) were fed to a 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 resulting first reaction mixture 5 was pumped into a first polymerization reactor 4 (MP / S=2.0), the temperature of which was 215°C, and the residence time was 3.5 hours. The resulting 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), where it remained for 1 hour before entering a sealed pressure filter 8 (screen pore size: 80 μm) to remove salt particles. The filtered salt particles 9 had a D50 of 180 μm. The obtained second reaction mixture 10 passes through a second heat exchanger 11, where it is heated to 265°C, and then enters a third tubular reactor 12 (the reactor temperature is 265°C) for a residence time of 1 hour.Then, it passes through a heat exchanger (not shown) and is heat-exchanged up to 280°C, and then enters an 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 in a condenser 20, and the waste liquid to be recovered 21 is produced. The undried crude product 15 first passes through the dryer feed chamber 16 and enters the continuous dryer 17. After further drying, the resulting crude product 18 continuously enters a beating tank, where water is added at a rate of 6.0 kg / h to be beaten. After heat exchange to 180°C, the crude product enters a pipe washer equipped with a heat transfer oil jacket. The temperature of the pipe washer is controlled at 180°C, and the residence time is 10 minutes. The crude product is then cooled to 60°C through heat exchange, and is continuously filtered and dried (not shown) to obtain the final product.

[0088] The production rate in 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 32,000 and an ash content of less than 0.52 wt%.

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

[0090] Comparing Example 6 with Example 2 as an example, the consumption of superheated steam is reduced by about half, the amount of waste liquid generated to be recovered is reduced by more than one-third, and the total reaction residence time is reduced by about 20%. These technical effects can significantly reduce the production costs of this continuous production.

[0091] Comparative Example 3 A 100L reactor was charged with 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. The temperature was then continuously increased under a nitrogen atmosphere to dehydrate the mixture. 10.51kg of aqueous solution (water content 98.0wt%) was removed. Once dehydration was complete, the temperature was lowered to 170°C. At this time, the amount of sulfur in the dehydrated solution system was 98.0mol, and the water content was 117.6mol. The dehydrated solution produced in the above batch process was transferred to a storage vessel at 170°C for warm storage and prepared for subsequent continuous feed.

[0092] The continuous PAS production process flow chart shown in Figure 1 was used. Paradichlorobenzene (PDCB) 1, preheated to 200°C, dehydration liquid 2 (sodium sulfide content 25.86 wt%), and NMP crystal slurry 3 containing sodium chloride seed crystals (D50 of the sodium chloride seed crystals is 15 μm, and the concentration of the crystal slurry is 0.136 mol / L) were 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 resulting 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), where it remained for 1 hour before entering a sealed pressure filter 8 (screen pore size 30μm) to remove salt particles. The filtered salt particles 9 had a D50 of 60μm. The obtained second reaction mixture 10 passes through a second heat exchanger 11, where it is heated to 265°C, and then enters a third tubular reactor 12 (the reactor temperature is 265°C) for a residence time of 1.5 hours.Then, it passes through a heat exchanger (not shown) and is heat-exchanged up to 280°C, and then enters an 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 in a condenser 20, and a waste liquid 21 to be recovered is produced. The undried crude product 15 first passes through a dryer supply chamber 16 and enters a continuous dryer 17. After further drying, the resulting crude product 18 continuously enters a beating tank, where water is added at a rate of 6.0 kg / h and beaten. After heat exchange to 180°C, the crude product enters a pipe washer equipped with a heat transfer oil jacket. The temperature of the pipe washer is controlled at 180°C, and the residence time is 10 minutes. The crude product is then cooled to 60°C via heat exchange, and is 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 saline 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 waste liquid to be recovered was 8.7 kg / h. The final product produced was measured and found to have a weight average molecular weight of approximately 23,000 and an ash content of less than 0.6 wt%.

[0094] In addition, in the process of 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, clogging occurred from time to time, and the stability of continuous operation was low. [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. Sealed 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 vapors 20 Condenser 21 Waste liquid to be collected

Claims

1. A continuous method for producing polyarylene sulfide, comprising: Step (1) of charging an aqueous sulfur source solution, an aqueous alkali metal hydroxide solution, an auxiliary agent, and an organic amide solvent into a reaction vessel, continuously raising the temperature to 200-220°C to carry out a dehydration and saponification reaction, so that the molar ratio of water to sulfur source after the completion of dehydration is 1.5-2.0, and cooling the resulting dehydrated solution to 150-180°C, and keeping the solution warm in a storage vessel; Step (2) of injecting the organic amide crystal slurry containing sodium chloride seed crystals, the dichloroaromatic compound, and the dehydrated 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 step (3) of injecting the filtrate obtained through the filtration device into a polymerization reactor, controlling the reaction temperature at 235 to 280°C, and carrying out the reaction until the monomer conversion rate reaches 95% or more, subjecting the resulting reaction liquid to flash evaporation treatment, and finally obtaining polyarylene sulfide through post-treatment, the auxiliary agent is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; In step (2), The method for continuous production of polyarylene sulfide is characterized in that the concentration of 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.

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, and the concentration is 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 hydroxide is 1:0.95-1.1; the molar ratio of the sulfur source to the coagent 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, and 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 to 2.

5.

3. In step (2), the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydrated liquid is 1.0 to 1.1:1; 2. The method for continuously producing polyarylene sulfide according to claim 1, wherein after mixing the dehydration liquid, the dichloroaromatic 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 the 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 dehydrated liquid is 1.0 to 1.025:1; 4. The method for continuously producing polyarylene sulfide according to claim 3, wherein after mixing the dehydration liquid, the dichloroaromatic 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 1 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 to 180-220°C; The second-stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled to 220-240°C; 2. The method for continuously producing polyarylene sulfide according to claim 1, wherein the reaction liquid is injected into the second-stage reactor after the monomer conversion rate in the first-stage reactor reaches 60 to 80%.

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 comprises drying, washing, filtering and redrying.

7. A batchwise method for producing polyarylene sulfide, comprising: Step (a) of charging a sulfur source aqueous solution, an alkali metal hydroxide aqueous solution, an auxiliary agent and an organic amide solvent into a reactor I, and continuously raising the temperature to 200-220°C to carry out a dehydration and saponification reaction, so that the molar ratio of water to sulfur source after the dehydration is 1.5-2.0, and then cooling to 150-180°C; Step (b) is to add the organic amide crystal slurry containing sodium chloride seed crystals, the dichloroaromatic compound and the additional organic amide solvent into the reactor I, and then heat the mixture to 205-220°C and keep the temperature for 0.5-5 hours, then heat the mixture to 215-240°C and keep the temperature for 0.5-5 hours, and then pour the resulting reaction liquid into a filter and filter it to remove sodium chloride particles; and step (c) of pouring the filtrate obtained through the filtration device into a preheated reactor II, heating it to 240-280°C, and maintaining the temperature for 0.5-5 hours, and then subjecting the resulting reaction liquid to flash evaporation treatment, and finally subjecting it to post-treatment to obtain polyarylene sulfide; the auxiliary agent is selected from one or more of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acid; In step (b), A batch-type method for producing polyarylene sulfide, characterized in that the concentration of 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.

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, and the concentration is 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 hydroxide is 1:0.95-1.1; the molar ratio of the sulfur source to the coagent 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, and N,N-dimethylformamide; 8. The batchwise method for producing polyarylene sulfide according to claim 7, wherein the molar ratio of the sulfur source to the organic amide solvent is 1:1.5-3.

9. In step (b), 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 dehydrated 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 reactor 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), 8. The batchwise method for producing polyarylene sulfide according to claim 7, characterized in that the temperature is increased to 240 to 280°C at a temperature increase rate of 0.2 to 1°C / min, and the post-treatment includes drying, washing, filtering and redrying.

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