WO2024021816A1 - 聚亚芳基硫醚树脂的制造工艺及其产品和应用 - Google Patents

聚亚芳基硫醚树脂的制造工艺及其产品和应用 Download PDF

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WO2024021816A1
WO2024021816A1 PCT/CN2023/096116 CN2023096116W WO2024021816A1 WO 2024021816 A1 WO2024021816 A1 WO 2024021816A1 CN 2023096116 W CN2023096116 W CN 2023096116W WO 2024021816 A1 WO2024021816 A1 WO 2024021816A1
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reaction
sulfur source
sodium chloride
reactor
temperature
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PCT/CN2023/096116
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English (en)
French (fr)
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陈志荣
尹红
贾艳宇
张雄伟
连明
陈兴
蒋杰
李沃源
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浙江大学
浙江新和成特种材料有限公司
浙江新和成股份有限公司
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Publication of WO2024021816A1 publication Critical patent/WO2024021816A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C08—ORGANIC MACROMOLECULAR COMPOUNDS; THEIR PREPARATION OR CHEMICAL WORKING-UP; COMPOSITIONS BASED THEREON
    • C08G—MACROMOLECULAR COMPOUNDS OBTAINED OTHERWISE THAN BY REACTIONS ONLY INVOLVING UNSATURATED CARBON-TO-CARBON BONDS
    • C08G75/00—Macromolecular compounds obtained by reactions forming a linkage containing sulfur with or without nitrogen, oxygen, or carbon in the main chain of the macromolecule
    • C08G75/02—Polythioethers
    • C08G75/0204—Polyarylenethioethers
    • C08G75/025—Preparatory processes
    • C08G75/0259—Preparatory processes metal hydrogensulfides

Definitions

  • the present invention relates to the field of synthesis of polyarylene sulfide resin, and in particular to a continuous and intermittent manufacturing process of polyarylene sulfide resin and its products and applications.
  • the representative manufacturing method of polyarylene sulfide is to polymerize sodium sulfide/sodium hydrosulfide and dihaloaromatic compounds in the presence of alkali metal hydroxides in an organic amide solvent.
  • the solvent is generally preferably
  • NMP N-methylpyrrolidone
  • the alkali metal hydroxide is generally preferably sodium hydroxide.
  • the industrial production equipment for synthesizing polyphenylene sulfide (PPS) using the above method is intermittent operation. Some of the reaction raw materials are first dehydrated in the reactor. After the temperature is lowered and the remaining reaction raw materials are added, the reactor is sealed and the temperature is raised to perform high-temperature polycondensation.
  • the crude product is obtained through cooling filtration or flash evaporation, and the crude product is purified to obtain the final product.
  • the intermittent production process requires repeated loading, heating and cooling, unloading and cleaning of the reactor, which takes a long time, greatly reduces production efficiency, and the process stability and reliability are also poor. To this end, researchers have been working hard to develop continuous manufacturing methods of PAS to improve manufacturing efficiency.
  • Patents US4060520, US4056515, and US4066632 propose a continuous production process with multiple kettles connected in series. Each kettle has a different reaction temperature, and the kettle pressure continuously decreases from front to back. The pressure difference between adjacent kettles promotes the flow of materials from front to back.
  • the Chinese patent document with publication number CN 108779253 A provides a continuous manufacturing device and a continuous manufacturing method of PAS.
  • the PAS continuous manufacturing device is equipped with an accommodation chamber that accommodates multiple reaction tanks, and organic amide solvents, sulfur are supplied to the accommodation chamber.
  • the source and the dihaloaromatic compound are polymerized in an organic amide solvent to form a reaction mixture; multiple reaction tanks are connected in sequence and connected to each other via the gas phase, and the reaction mixture is sequentially supplied to each reaction tank. move.
  • the public number is CN 113667122A
  • the national patent document provides a gradient temperature-controlled continuous condensation method of polyarylene sulfide resin.
  • the temperature of each reactor is step-wise controlled (stepwise temperature rise) to adapt to different progress of the polymerization reaction.
  • the reaction raw materials are continuously sent to the reactor through a metering pump. in series kettle or tubular reactors.
  • the above-mentioned sodium sulfide method for producing PPS will produce a large amount of by-product sodium chloride. For every ton of PPS produced, 1.08 tons of sodium chloride will be produced.
  • Sodium chloride begins to be generated and precipitated in large quantities in the early stage of the polymerization reaction, causing the reaction slurry to contain It contains a large amount of salt particles.
  • the content of by-product salt particles in the polymerization reaction solution after the reaction is about 20.5 wt%.
  • a large number of salt particles in the slurry are easy to settle and block in long-distance transportation or reaction pipelines, especially in tubular reactors with smaller sizes, causing production shutdown.
  • reaction substrate concentration is conducive to improving production efficiency.
  • High reaction substrate concentration is also conducive to the efficient production of high molecular weight polyphenylene sulfide, but the content of by-product salt particles will also increase significantly.
  • the viscosity of the reaction solution also increased sharply, worsening the above-mentioned transportation blockage and stirring problems.
  • the presence of a large number of by-product salt particles in the PPS reaction slurry severely limits the research and development of PPS continuous manufacturing methods and limits the efficient preparation of higher molecular weight polyphenylene sulfide.
  • the by-product sodium chloride will be wrapped in the product after the reaction is completed. It needs to be washed and removed with a large amount of pure water multiple times, resulting in a large amount of high-salt content washing wastewater.
  • every unit produced in the industry One ton of PPS requires more than 15 tons of water, and the recovery of by-product salts dissolved in water through evaporation will increase energy consumption.
  • the high corrosiveness of salt water also places high demands on the materials of polyphenylene sulfide manufacturing equipment.
  • the polyphenylene sulfide manufacturing process has the characteristics of high water consumption, high energy consumption and high corrosiveness, which greatly increases the manufacturing cost and environmental protection cost of polyphenylene sulfide, and seriously limits the greenness and sustainability of the polyphenylene sulfide production process. This limits the competitiveness and vitality of polyphenylene sulfide resin.
  • the Chinese patent document with publication number CN104371103A discloses a method for desalting polyphenylene sulfide. After the polymerization reaction is completed, the reaction kettle is kept at a temperature and pressure, and the inorganic salts and inorganic additives are allowed to settle into the kettle. After bottoming, the upper polyphenylene sulfide solution is pressed out of the reaction kettle through the bottom extension tube by using the pressure difference, and then the solution is lowered.
  • the present invention discloses a manufacturing process of polyarylene sulfide resin.
  • This process can greatly reduce the salt content in the prepared crude polyarylene sulfide resin, which greatly improves the production of polyarylene sulfide resin. It reduces the generation of high-salt wastewater, reduces water and energy consumption, and helps to prepare high molecular weight polyarylene sulfide resin; more importantly, this process is not only suitable for batch processes, but also for Continuous production is expected to significantly improve the production efficiency of the production system, reduce production costs, and be more environmentally friendly.
  • a continuous manufacturing method of polyarylene sulfide including:
  • the inventor conducted in-depth research on the growth process and particle morphology of salt particles in the PPS reaction solution. After all the raw materials are added and raised to a certain temperature, the polymerization reaction begins and sodium chloride is continuously generated. Due to the presence of water in the polyphenylene sulfide reaction system This allows the by-product sodium chloride to dissolve in a small amount in the reaction solvent. When the solution is saturated, the sodium chloride that continues to be generated begins to nucleate and precipitate. The reaction continues to proceed rapidly, and a large amount of sodium chloride that is generated quickly nucleates and precipitates. The study found that: (1) In the previous typical batch method PPS manufacturing process, the salt particles generated in the reaction solution were small in size, about 20 ⁇ m, which was not conducive to filtration.
  • a suitable water content can, on the one hand, reduce the precipitation and nucleation rate of sodium chloride, and on the other hand, increase the growth rate of sodium chloride crystals; (4) The introduction of a small amount of sodium chloride seed crystals can effectively inhibit the growth of sodium chloride. Primary nucleation.
  • step (1) in order to control the particle size of the sodium chloride particles above 100 ⁇ m to facilitate filtration, the water content after dehydration needs to be controlled at an appropriate ratio.
  • the molar ratio of water and sulfur source is controlled at 1.5-2.0; since sodium chloride can be filtered out, the generation of high-salt wastewater is greatly reduced, water consumption and energy consumption are reduced, and the amount of solvent in the reaction solution can be further reduced. , more conducive to the efficient preparation of high molecular weight PPS.
  • step (1)
  • the concentration of the sulfur source aqueous solution is 28-48wt%, 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 with a concentration of 30 to 70wt%;
  • the dichloroaromatic compound is selected from one or more of p-dichlorobenzene, dichloronaphthalene, dichlorofluorene and dichlorocarbazole, preferably p-dichlorobenzene;
  • the molar ratio of sulfur source to sodium hydroxide is 1:0.95 ⁇ 1.1;
  • the auxiliary agent is selected from one or more types of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acids;
  • the molar ratio of sulfur source to additive is 1:0 ⁇ 0.5;
  • the organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl- ⁇ -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.
  • the dehydration saponification reaction can be carried out in batches according to the method of step (1), or the continuous dehydration scheme reported in other patents can be used.
  • the key is to control the water content after the dehydration is completed to be within the appropriate range.
  • 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 ⁇ m;
  • the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0-1.1:1; preferably, it is 1.0-1.08:1.
  • the addition amount of the sodium chloride seed crystals is controlled based on 1 mole of sulfur source in the prepolymerization reactor system at this time. It is 0.05 ⁇ 0.4mol%. At this time, the total amount of organic amide solvent in the system is 2.0 ⁇ 4.0mol.
  • the molar ratio of the dichloroaromatic compound and the sulfur source is controlled to be 1.0 to 1.025:1.
  • the inventor also conducted in-depth research on the polymerization reaction kinetics of polyphenylene sulfide and found that: (1) the main polymerization reaction of polyphenylene sulfide is a strongly exothermic reaction, and temperature control during the prepolymerization process will seriously affect the reaction. Safety and product quality, and the polymerization reaction is a second-order kinetic reaction.
  • the first reactor is preferably a continuous stirred tank reactor and Control the conversion rate within an appropriate range; (2) Increasing the water content in the reaction system can reduce the reaction speed, that is, the production speed of sodium chloride, but too high water content will lead to low reaction efficiency, high reaction pressure and flash evaporation process Issues such as the need for higher energy consumption.
  • step (2) the prepolymerization reactor is a two-stage reactor in series;
  • the first-stage reactor is selected from a continuously stirred tank reactor, and the reaction temperature in the tank reactor is controlled to be 180 to 220°C, preferably 190 to 210°C.
  • the reaction liquid is injected into the second-stage reactor.
  • the second-stage reactor is selected from a jacketed tubular reactor, and the reaction temperature in the reactor is controlled to be 220 to 240°C.
  • the heating medium used in the jacketed tubular reactor is thermal oil
  • the flow direction of the thermal oil is consistent with the flow direction of the reaction liquid
  • the inlet oil temperature is controlled to be 220-230°C
  • the outlet oil temperature is controlled to be 230-240°C
  • the monomer conversion rate at the outlet reaches more than 90%, more preferably more than 93%.
  • the growth of sodium chloride crystals is basically completed, the particle size grows to more than 60 ⁇ m, preferably more than 100 ⁇ m, and the size is uniform.
  • the filtering device is a continuous filtering device, specifically selected from a closed pressurized filtering device, preferably It is a pressurized drum filter, suitable for continuous filtration of high-temperature slurries containing volatile components and crystals that are easy to precipitate.
  • the pore size of the filter material in the filter device is preferably 60 to 80 ⁇ m. At this size, sodium chloride particles can be sufficiently removed without the pressure drop over filtration being too high.
  • step (3)
  • the polymerization reactor is selected from a tubular reactor.
  • the tubular reactor is selected from a coil reactor immersed in a high-temperature oil bath.
  • the reaction temperature is controlled at 235-280°C to complete further molecular weight growth. , ensuring that the monomer conversion rate reaches more than 95%, preferably more than 98%.
  • the flash evaporation process uses superheated steam at 260 to 300°C for auxiliary heating.
  • the amount of superheated steam is 0.1 to 1kg/mol of sulfur source.
  • the steam and vaporized solvent are taken out from the upper part of the flash evaporator, and then condensed into the solvent.
  • the salty PAS crude product is continuously discharged from the lower part.
  • the post-processing includes drying, washing, filtering and re-drying, specifically:
  • the salt-containing PAS crude product is continuously dried in a dryer to further remove residual solvent.
  • the salt-containing PAS crude product After drying, the salt-containing PAS crude product is beaten with water and washed with water at a continuous high temperature and high pressure. The mass ratio of the amount of water added to the crude product is 3 to 10:1. The final product is obtained after continuous filtration and drying.
  • the equipment for continuous high-temperature and high-pressure water washing can be continuous kettle washing equipment or continuous tube washing equipment; it is preferably a continuous tube washing equipment with a thermal oil jacket.
  • the residence time of the water washing slurry is preferably 1 to 30 minutes, and the water washing temperature Preferably it is 150-220 degreeC.
  • the continuous filtration equipment is selected from continuous filters, such as plate and frame filters, decanter centrifuges, belt filters, etc., and is preferably a belt filter for continuous vacuum filtration.
  • the dryer is selected from the group consisting of dryers that continuously feed in and out materials.
  • the salt concentration in the subsequent crude product is greatly reduced, and the pressure of the washing process in the post-processing is significantly reduced. Therefore, only one continuous high-temperature and high-pressure water washing is required; If the salt particles are not pre-filtered, it will take several times of normal pressure water washing + high pressure water washing to fully remove the sodium chloride in the product, which greatly increases the amount of washing water and the amount of salty wastewater generated; if no seed crystals are added and the reaction solution is removed
  • the medium water content is controlled at an appropriate level.
  • the particle size of sodium chloride is small. It is difficult to effectively intercept it when filtering with large pore size filter material. Use small pore size. If the filter material has a large filtration pressure drop, it is easy to block the filter.
  • the present invention also discloses a batch production method of polyarylene sulfide, which includes:
  • reaction kettle I Put the organic amide slurry containing sodium chloride seed crystals, dichloroaromatic compounds and additional organic amide solvent into reaction kettle I, first heat it to 205 ⁇ 220°C, keep it warm for 0.5 ⁇ 3h, and then heat it to 215 ⁇ 240°C, keep for 0.5 ⁇ 5h, then inject the obtained reaction solution into the filter device, and filter to remove sodium chloride particles;
  • the concentration of the sulfur source aqueous solution is 28-48wt%, 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 with a concentration of 30 to 70wt%;
  • the dichloroaromatic compound is selected from one or more of p-dichlorobenzene, dichloronaphthalene, dichlorofluorene, and dichlorocarbazole, preferably p-dichlorobenzene; the molar ratio of the sulfur source to sodium hydroxide It is 1:0.95 ⁇ 1.1;
  • the auxiliary agent is selected from one or more types of sodium acetate, sodium benzoate, and sodium C5-C6 fatty acids;
  • the molar ratio of sulfur source to additive is 1:0 ⁇ 0.5;
  • the organic amide solvent is selected from one or more of N-methylpyrrolidone, hexamethylphosphoric triamide, N-methyl- ⁇ -caprolactam, and N,N-dimethylformamide;
  • the molar ratio of the sulfur source to the organic amide solvent is 1:1.5-3.
  • the concentration of sodium chloride seed crystals in the organic amide crystal slurry is 0.1-1mol/L, and the particle size of the sodium chloride seed crystals is 1-20 ⁇ m; the added amount of the sodium chloride seed crystals is the sulfur content in the system at this time. source 0.05 ⁇ 0.4mol%;
  • the molar ratio of the dichloroaromatic compound to the sulfur source in the dehydration liquid is 1.0 to 1.1:1;
  • the molar ratio of the total molar amount of the organic amide solvent to the sulfur source in the reactor I is 2.5 to 3.5;
  • the temperature is raised to 205-220°C at a heating rate of 0.2-0.5°C/min, and then to 215-240°C at a heating rate of 0.2-1°C/min.
  • the reaction kettle used in steps (a) and (b) is a stirred reactor equipped with a thermal oil jacket and a coil.
  • the temperature is raised to 240-280°C at a heating rate of 0.2-1°C/min.
  • drying is continued in the flash evaporator for 10 minutes to 1 hour, and then the temperature is cooled to room temperature before the material is discharged.
  • the post-processing includes drying, washing, filtering and drying again.
  • the invention also discloses polyarylene sulfide respectively prepared according to the above two processes.
  • the invention also discloses the use of the polyarylene sulfide in preparing cross-linked polyarylene sulfide, specifically by subjecting the polyarylene sulfide to thermal oxygen treatment.
  • the present invention has the following beneficial effects:
  • the invention discloses a manufacturing process of polyarylene sulfide resin.
  • the core is to add sodium chloride seed crystals to the PAS reaction liquid, while controlling the water content in the system to be within an appropriate range and controlling the reaction temperature of the prepolymerization section.
  • the nucleation rate of sodium chloride is reduced and the growth rate of crystals is increased to inhibit primary nucleation, promote secondary nucleation, and ultimately increase the size of by-product sodium chloride crystals; the process
  • Another key lies in the timing of filtration to remove salt particles.
  • the filtration selection of salt particles is carried out at the end of the prepolymerization stage; using the manufacturing process disclosed in the present invention, the salt content in the crude polyarylene sulfide resin can be reduced.
  • the prepolymerization section is divided into two sections.
  • the first section uses a kettle reactor suitable for stable heat transfer, and controls the conversion rate at an appropriate value.
  • a jacketed tube reactor is used in the second stage, and the remaining heat is used to promote the reaction to continue to heat up to maintain the reaction efficiency;
  • the high-temperature polymerization stage uses salt particles to remove more Coiled tube reactor suitable for high conversion rate and high viscosity state (high molecular weight and high concentration).
  • Figure 1 is a schematic process flow diagram of the continuous manufacturing of polyarylene sulfide of the present invention, in which the post-treatment process is omitted;
  • Ash content test Place the porcelain crucible in a muffle furnace with a constant temperature of 750°C and burn it to a constant weight, then take it out and place it in a desiccator to cool and then weigh it, recorded as M0. Weigh 3g of the sample into the crucible, add 10mL of nitric acid, and then place it on an alcohol blowtorch to burn until no more smoke comes out. Finally, place the crucible in a muffle furnace with a constant temperature of 750°C and burn it for 1 hour. Take it out and place it in a desiccator to cool and then weigh it, recorded as M1. The calculation formula of resin ash content is: (M1-M0)/3*100%.
  • the molecular weight of PPS is measured using gel permeation chromatography with polystyrene as the standard sample.
  • the mobile phase is 1-chloronaphthalene
  • the column temperature is 220°C
  • the flow rate is 1mL/mL
  • the detector is a refractive index detector.
  • reaction kettle I After adding 15.14kg (103.0mol) of p-dichlorobenzene and 10kg of NMP crystal slurry containing 11.7g of sodium chloride seed crystal (0.2mol, D50 of sodium chloride seed crystal is 15 ⁇ m) in reaction kettle I, and then add 4.46kg NMP (45.0mol), sealed reactor I, heated to 215°C at 0.3°C/min, kept for 3 hours, then heated to 230°C at 0.2°C/min, kept for 1 hour, and then the reaction solution was passed through a filter (filter pore size 80 ⁇ m) was filtered to remove sodium chloride particles and then pumped into a 100L reactor II that had been preheated to 230°C.
  • a filter filter pore size 80 ⁇ m
  • the reactor II was then heated to 260°C at 0.5°C/min, kept for 1 hour, and then slowly discharged to normal pressure. Perform flash evaporation treatment in the evaporator, and finally dry to further remove the solvent to obtain 10.4kg of crude product. Add 60kg of water to beat, raise the temperature to 180°C, keep it for 0.5h, then cool down, filter and dry to obtain the final product.
  • the filter pressure drop is high and the flux is small.
  • the filtration step takes a lot of time, greatly reduces the production efficiency, and significantly increases the production cost. cost.
  • the crude product is first washed and filtered under normal pressure twice, adding 60.0kg of water each time. After adding 60.0kg of water to beat, the temperature is raised to 180°C, kept for 0.5h, filtered after cooling, and the final product is obtained after filtration and drying.
  • Example 1 shows that removing salt particles by filtration can significantly reduce the amount of washing water and the amount of waste brine produced.
  • the obtained second reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then Enter the third tubular reactor 12 (reactor temperature is 265°C), the residence time is 1.5h, and then the heat is exchanged to 280°C through the heat exchanger (not shown) and then enters the normal pressure flash evaporator 14 for flash evaporation.
  • the flow rate of the superheated steam 13 is 5.0kg/h.
  • the vaporized solvent and steam 19 are condensed and liquefied by the condenser 20 to generate waste liquid 21 to be recycled.
  • the undried crude product 15 first enters the continuous dryer 17 through the dryer feed bin 16.
  • the crude product 18 obtained after further drying is continuously discharged into the beating kettle and added with water for beating.
  • the water addition amount is 6.0kg/h.
  • After heat exchange to 180°C enter the tubular scrubber with a heat transfer oil jacket, control the temperature of the tubular scrubber to 180°C, stay for 10 minutes, then cool to 60°C through heat exchange, continuously filter and dry (not shown), Obtain the final product.
  • the production rate of this embodiment is 1.01kg/h; 5.0kg/h of salty wastewater is produced with a salt content of 0.6wt%, 5.0kg/h of superheated steam is consumed, and 8.7kg/h of waste liquid to be recycled is produced.
  • the prepared final product has a tested weight average molecular weight of approximately 23,000 and an ash content of less than 0.55wt%.
  • the temperature of the polymerization reactor 4 is 215°C, and the residence time is 3.5h.
  • the obtained first reaction mixture 5 is heated to 230°C through the first heat exchanger 6 and enters the second tubular reactor 7 (the reactor temperature is 230°C). , after a residence time of 1 hour, it is passed into the closed pressurized filter 8, and the salt particles are removed through the closed pressurized filter 8 (filter cloth pore size is 80 ⁇ m).
  • the D50 of the filtered salt particles 9 is 130 ⁇ m; the resulting second reaction mixture
  • the liquid 10 is heated to 265°C through the second heat exchanger 11, enters the third tubular reactor 12 (reactor temperature is 265°C), stays for 1 hour, and then exchanges heat to 280°C through the heat exchanger (not shown) Then it enters the normal pressure flash evaporator 14 for flash evaporation.
  • 265°C superheated steam 13 is used to assist the flash evaporation.
  • the flow rate of the superheated steam 13 is 3.5kg/h, the vaporized solvent and steam 19 are condensed and liquefied by the condenser 20 to generate waste liquid 21 to be recycled.
  • the undried crude product 15 first enters the continuous dryer 17 through the dryer feed bin 16.
  • the crude product 18 obtained after further drying is continuously discharged into the beating kettle and added with water for beating.
  • the water addition amount is 6.0kg/h.
  • enter the tubular scrubber with a heat transfer oil jacket control the temperature of the tubular scrubber to 180°C, stay for 10 minutes, then cool to 60°C through heat exchange, continuously filter and dry (not shown), Get the final product.
  • the production rate of this embodiment is 1.01kg/h; 5.0kg/h of salty wastewater is produced with a salt content of 0.6wt%, 3.5kg/h of superheated steam is consumed, and 6.2kg/h of waste liquid to be recycled is produced.
  • the prepared final product has a tested weight average molecular weight of approximately 26,000 and an ash content of less than 0.55wt%.
  • Example 2 Comparing the records of Example 2 and Example 3, it can be seen that further reducing the prepolymerization stage, specifically the molar ratio of NMP/S in the first polymerization reactor, can reduce the amount of solvent and the amount of superheated steam in the flash evaporation stage, and reduce energy consumption. and waste liquid volume, significantly reducing production costs; it can improve reaction efficiency and reduce the residence time of reaction liquid in each section of the reactor; it can also appropriately increase the weight average molecular weight of the final product.
  • the amount of sulfur in the dehydration liquid system is 98.0 mol, containing The amount of water is 147.0mol.
  • the PAS continuous manufacturing apparatus shown in Figure 1 was used.
  • PDCB Preheated p-dichlorobenzene 1
  • dehydrated liquid 2 sodium sulfide content 24.22wt%
  • NMP slurry 3 containing sodium chloride seed crystals the particle size of the sodium chloride seed crystal
  • the obtained first reaction mixture 5 is heated to 230°C through the first heat exchanger 6 and enters the second tubular reactor 7 (Reactor temperature is 230°C). After a residence time of 1 hour, it is passed into the closed pressurized filter 8. The salt particles are removed through the closed pressurized filter 8 (filter pore size 80 ⁇ m). The D50 of the filtered salt particles 9 is 120 ⁇ m; 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 (reactor temperature is 265°C), with a residence time of 1 hour, and then passes through the heat exchanger (not shown). (out) after heat exchange to 280°C, it enters the normal pressure flash evaporator 14 for flash evaporation.
  • 265°C superheated steam 13 is used to assist the flash evaporation.
  • the flow rate of the superheated steam 13 is 3.3kg/h.
  • the vaporized solvent and steam 19 are condensed.
  • the device 20 is condensed and liquefied to produce waste liquid 21 to be recycled.
  • the undried crude product 15 first enters the continuous dryer 17 through the dryer feed bin 16, and the crude product 18 obtained after further drying is continuously discharged into the beating kettle and added with water for beating.
  • the water addition amount is 6.0kg/h.
  • Heat to 180°C enter the tubular scrubber with a thermal oil jacket, control the temperature of the tubular scrubber to 180°C, stay for 10 minutes, then cool to 60°C through heat exchange, continuously filter and dry (not shown) , to obtain the final product.
  • the production rate of this embodiment is 1.01kg/h; it produces 5.0kg/h salty wastewater with a salt content of 0.65wt%, consumes 3.3kg/h of superheated steam, and produces 6.0kg/h of waste liquid to be recycled.
  • the prepared final product has a tested weight average molecular weight of approximately 68,000 and an ash content of less than 0.5wt%.
  • the obtained first reaction mixture 5 is heated to 225°C through the first heat exchanger 6 and enters the second tubular reactor 7 (the reactor temperature is 225°C). , after a residence time of 1.5 hours, it is passed into the closed pressurized filter 8, and the salt particles are removed through the closed pressurized filter 8 (filter cloth pore size is 80 ⁇ m,), and the D50 of the filtered salt particles 9 is 150 ⁇ m; the obtained second The reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (reactor temperature is 265°C), with a residence time of 1.5 hours, and then is exchanged through the heat exchanger (not shown).
  • the production rate of this embodiment is 1.01kg/h; 5.0kg/h of salty wastewater is produced with a salt content of 0.6wt%, 5.0kg/h of superheated steam is consumed, and 8.78kg/h of waste liquid to be recycled is produced.
  • the prepared final product has a tested weight average molecular weight of approximately 21,000 and an ash content of less than 0.55wt%.
  • the obtained first reaction mixture 5 is heated to 225°C through the first heat exchanger 6 and enters the second tubular reactor 7 (the reactor temperature is 225°C ), after a residence time of 1 hour, it is passed into the closed pressurized filter 8, and the salt particles are removed through the closed pressurized filter 8 (filter cloth pore size is 80 ⁇ m), and the D50 of the filtered salt particles 9 is 180 ⁇ m; the obtained second The reaction mixture 10 is heated to 265°C through the second heat exchanger 11, and then enters the third tubular reactor 12 (reactor temperature is 265°C), with a residence time of 1 hour, and then exchanges heat through the heat exchanger (not shown) After reaching 280°C, it enters the normal pressure flash evaporator 14 for flash evaporation.
  • superheated steam 13 at 265°C is used to assist the flash evaporation.
  • the flow rate of the superheated steam 13 is 2.8kg/h.
  • the vaporized solvent and steam 19 are condensed and liquefied by the condenser 20 Waste liquid 21 to be recycled is generated.
  • the undried crude product 15 first enters the continuous dryer 17 through the dryer feed bin 16.
  • the crude product 18 obtained after further drying is continuously discharged into the beating kettle and added with water for beating.
  • the water addition amount is 6.0kg/h.
  • the production rate of this embodiment is 1.01kg/h; the salty wastewater produced is 5.0kg/h, with a salt content of 0.58wt%, consumes 2.8kg/h of superheated steam, and produces 5.1kg/h of waste liquid to be recycled.
  • the prepared final product has a tested weight average molecular weight of approximately 32,000 and an ash content of less than 0.52wt%.
  • the NMP/S molar ratio can be 2.5 (Example 4) and the NMP/S molar ratio can be 2.0. (Example 6) successfully prepared polyphenylene sulfide, which not only significantly reduced the energy consumption (superheated steam consumption) and the amount of three wastes (the amount of waste liquid to be recovered), but also reduced the residence time (reaction time) and significantly improved the reaction efficiency. .
  • Example 6 Comparing Example 6 and Example 2 as an example, the superheated steam consumption was reduced by nearly half, the amount of waste liquid to be recycled was reduced by more than 1/3, and the total reaction residence time was reduced by about 20%.
  • the above technical effects can significantly reduce the continuous Production costs of chemical production.
  • the obtained first reaction mixture 5 is heated to 225°C through the first heat exchanger 6 and enters the second tubular reactor 7 (the reactor temperature is 225°C). , after a residence time of 1 hour, it is passed into the closed pressurized filter 8, and the salt particles are removed through the closed pressurized filter 8 (filter cloth pore size is 30 ⁇ m,), and the D50 of the filtered salt particles 9 is 60 ⁇ m; the second reaction obtained
  • the mixed liquid 10 passes through the second heat exchanger 11 is heated to 265°C, then enters the third tubular reactor 12 (reactor temperature is 265°C), the residence time is 1.5h, and then is heated to 280°C through a heat exchanger (not shown) before entering the normal pressure flash evaporator 14 is flash evaporated.
  • 265°C superheated steam 13 is used to assist the flash evaporation.
  • the flow rate of the superheated steam 13 is 5.0kg/h.
  • the vaporized solvent and steam 19 are condensed and liquefied by the condenser 20 to generate waste liquid 21 to be recycled.
  • the undried crude product 15 first enters the continuous dryer 17 through the dryer feed bin 16.
  • the crude product 18 obtained after further drying is continuously discharged into the beating kettle and added with water for beating.
  • the water addition amount is 6.0kg/h.
  • the production rate of this comparative example is 1.01kg/h; it produces 5.0kg/h salty wastewater with a salt content of 0.65wt%, consumes 5.0kg/h superheated steam, and produces 8.7kg/h waste liquid to be recycled.
  • the prepared final product has a tested weight average molecular weight of approximately 23,000 and an ash content of less than 0.6wt%.

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Abstract

本发明公开了一种聚亚芳基硫醚的制造工艺及其产品和应用，该制造工艺以硫源、二氯代芳香化合物为原料，采用有机酰胺溶剂，经脱水、预聚、聚合、闪蒸及后处理等一系列工艺制备得到，本发明创造性地在预聚阶段向反应液中加入氯化钠晶种，又在预聚反应结束后对反应液增加了一步过滤处理。经该制造工艺可将制备的聚亚芳基硫醚树脂粗品中的含盐量大幅降低，极大地降低了高含盐废水的产生，减少了水耗和能耗，且有助于高效制备高分子量的聚亚芳基硫醚树脂；更为重要的是，该工艺不仅适用于间歇工艺，同样适用于连续化生产，有望大幅提高该生产体系的生产效率，并降低生产成本，且更加环保绿色。

Description

聚亚芳基硫醚树脂的制造工艺及其产品和应用 技术领域
本发明涉及聚亚芳基硫醚树脂的合成领域，尤其涉及一种聚亚芳基硫醚树脂的连续化与间歇式制造工艺及其产品和应用。
背景技术
聚亚芳基硫醚(PAS)的代表性制造方法是在有机酰胺溶剂中使硫化钠/硫氢化钠与二卤芳香族化合物在碱金属氢氧化物的存在下进行聚合反应，溶剂一般优选为N-甲基吡咯烷酮(以下，简称为NMP)，碱金属氢氧化物一般优选为氢氧化钠。目前，以前述方法合成聚苯硫醚(PPS)的工业化生产装置均为间歇操作，部分反应原料首先在反应釜中进行脱水反应，降温补加其余反应原料后，密闭反应釜、升温进行高温缩聚，反应结束后经降温过滤或闪蒸获得粗产品，对粗产品进行纯化处理获得最终产品。间歇式生产过程中需要反复装料、升降温、卸料以及清洗反应釜，耗时长，大大降低了生产效率，且工艺稳定性和可靠性也较差。为此，研究者们一直在努力开发PAS的连续化制造方法以提高制造效率。
专利US4060520、US4056515、US4066632提出一种多釜串联的连续化生产工艺，每个釜的反应温度不同，釜压从前至后不断降低，通过相邻釜间压差推动物料从前至后流动。公开号为CN 108779253 A的中国专利文献中提供了一种PAS的连续制造装置和连续制造方法，该PAS连续制造装置具备容纳多个反应槽的容纳室，向容纳室中供给有机酰胺溶剂、硫源以及二卤芳香族化合物，在有机酰胺溶剂中进行硫源与二卤芳香族化合物的聚合反应，并形成反应混合物；多个反应槽依次连接，经由气相相互连通，反应混合物依次向各反应槽移动。公开号为CN 113667122A的中 国专利文献中提供了一种聚亚芳硫醚树脂的梯度控温连续缩合方法，对各反应器温度进行梯级控制(逐步升温)以适配聚合反应不同进度，反应原料通过计量泵连续送到串联的釜式或管式反应器中。
以上述硫化钠法制造PPS会产生大量副产氯化钠，每生产一吨PPS，便会产生1.08吨的氯化钠，氯化钠在聚合反应前期便开始大量生成、析出，使得反应浆液中含有大量的盐颗粒，以反应液中NMP/硫源的摩尔比为3.5的典型值计，反应结束聚合反应液中副产盐颗粒的含量约为20.5wt％。浆液中大量的盐颗粒容易在长距离输送或反应管路中，特别是尺寸较细的管式反应器处沉降、堵塞，导致生产停车。无论是间歇还是连续工艺，提高反应底物浓度均有利于提高生产效率，高反应底物浓度也有利于高效制造高分子量的聚苯硫醚，但副产盐颗粒的含量也会随之显著增加，反应液粘度也急剧增加，使得上述输送堵塞以及搅拌问题更加恶化。PPS反应浆液中大量副产盐颗粒的存在严重限制了PPS连续制造方法的研究开发，并且限制了更高分子量的聚苯硫醚的高效制备。
无论是间歇还是连续制造工艺，副产氯化钠在反应结束后会被包裹在产品中，需要使用大量纯水多次洗涤除去，产生大量高盐含量的洗涤废水，目前，工业中每生产一吨PPS需要消耗15吨以上的水，而溶于水中的副产盐通过蒸发回收又会增加能耗。此外，盐水的高腐蚀性也对聚苯硫醚制造设备的材质提出了很高要求。聚苯硫醚制造工艺高水耗、高能耗和高腐蚀性的特点，大大提高了聚苯硫醚的制造成本和环保成本，严重限制了聚苯硫醚生产工艺的绿色化和可持续性，限制了聚苯硫醚树脂的竞争力和生命力。
前述这些问题的解决，可以通过在产品固液分离前除去析出的盐颗粒以减少后续设备以及粗产品中盐颗粒含量。如公开号为CN104371103A的中国专利文献中公开了一种聚苯硫醚除盐的方法，是在聚合反应结束后，在反应釜中保温保压静置，待无机盐和无机助剂沉降至釜底后，利用压力差，把上层聚苯硫醚溶液通过底部內伸管压出反应釜，再对该溶液进行降 温，析出的聚苯硫醚经纯化获得高纯度聚苯硫醚产品。但该专利所述方法只适用于间歇工艺，反应釜中沉降层中的无机盐、大量反应溶剂和产品如何进一步处理，在该专利文献中未做进一步描述，而且事实上沉降的无机盐层在未完全清理前会使聚合釜搅拌无法再次启动。
又如公开号为CN107964098A的中国专利文献中涉及一种聚苯硫醚合成过程中的除盐方法及装置，聚合过程中通过输料泵使反应釜物料通过装有滤网的过滤装置，盐留在过滤装置中，剩余的聚苯硫醚物料返回至反应釜，如此循环直至反应完成，出料洗涤干燥即得聚苯硫醚树脂。该方法同样只适用于间歇工艺，而且反应后期，随着盐饼层厚度以及反应液粘度的增加，过滤变得更加困难。
发明内容
针对现有技术存在的上述问题，本发明公开了一种聚亚芳基硫醚树脂的制造工艺，该工艺可将制备的聚亚芳基硫醚树脂粗品中的含盐量大幅降低，极大地降低了高含盐废水的产生，减少了水耗和能耗，且有助于制备高分子量的聚亚芳基硫醚树脂；更为重要的是，该工艺不仅适用于间歇工艺，同样适用于连续化生产，有望大幅提高该生产体系的生产效率，并降低生产成本，且更加环保绿色。
具体技术方案如下：
一种聚亚芳基硫醚的连续化制造方法，包括：
(1)将硫源水溶液、碱金属氢氧化物水溶液、可选择性加入的助剂与有机酰胺溶剂投入反应釜中，不断升温至200～220℃进行脱水皂化反应，脱水结束后水和硫源的摩尔比为1.5～2.0，冷却至150～180℃，将所得脱水液保温在储存釜中；
(2)将含有氯化钠晶种的有机酰胺晶浆、二氯代芳香化合物和所述脱水液注入预聚反应器中，反应温度控制在180～240℃，反应至单体转化率达到90％及以上后，将得到的反应液注入过滤装置，过滤除去氯化钠颗 粒；
(3)经过滤装置后得到的滤液注入聚合反应器中，反应温度控制在235～280℃，反应至单体转化率达到95％及以上，将得到的反应液进行闪蒸处理，最后经后处理得到聚亚芳基硫醚。
发明者对PPS反应液中盐颗粒的生长过程、颗粒形态进行了深入研究，原料全部加入升至一定温度后，聚合反应开始，氯化钠不断生成，由于聚苯硫醚反应体系中水的存在使得副产氯化钠能够少量溶于反应溶剂，当达到溶解饱和后，继续生成的氯化钠开始成核析出，反应继续快速进行，大量生成的氯化钠快速成核析出。研究发现：(1)之前的典型间歇法PPS制造工艺中，反应液中生成的盐颗粒粒径较小，在20μm左右，不利于过滤，这一点前人鲜有报道；(2)在前期预聚反应段，反应转化率已超过90％，此时氯化钠晶体的生长已经基本完成，但此时产品的分子量仍然较低，反应液粘度也不高，需要进入高温聚合段继续进行反应以提高分子量，因此副产盐的过滤去除选在预聚反应段氯化钠晶体基本完成生长时是比较合适，高温聚合段后反应液粘度较高，过滤会更加困难；(3)控制体系中的水含量在合适范围一方面可以降低氯化钠的析出成核速度，另一方面还可提高氯化钠晶体的生长速度；(4)少量氯化钠晶种的引入可以有效抑制氯化钠的初级成核。
基于以上研究成果提出了本发明的技术方案，经试验发现，步骤(1)中，为了将氯化钠颗粒的粒径控制在100μm以上以利于过滤，脱水结束后含水量需要控制在合适比例下(水和硫源的摩尔比控制在1.5～2.0)；由于氯化钠可以过滤除去，极大减少了高含盐废水的产生，减少水耗和能耗，同时反应液中溶剂用量可以进一步降低，更有助于高效制备高分子量的PPS。
步骤(1)中：
所述硫源水溶液的浓度为28～48wt％，硫源选自硫氢化钠和/或硫化钠；
所述碱金属氢氧化物水溶液为氢氧化钠水溶液，浓度为30～70wt％；
所述二氯代芳香化合物选对二氯苯、二氯代萘、二氯代芴、二氯咔唑中一种或多种，优选对二氯苯；
硫源与氢氧化钠的摩尔比为1：0.95～1.1；
所述助剂选自醋酸钠、苯甲酸钠、C5～C6的脂肪酸钠中的一种或多种；
硫源与助剂的摩尔比为1：0～0.5；
所述有机酰胺溶剂选自N-甲基吡咯烷酮、六甲基磷酰三胺、N-甲基-ε-己内酰胺、N,N-二甲基甲酰胺中的一种或多种，优选N-甲基吡咯烷酮；
硫源与有机酰胺溶剂的摩尔比为1：1.5～2.5。
脱水皂化反应可以是按照步骤(1)的方法间歇式分批进行，也可以采用其它专利报道的连续脱水方案，关键是控制脱水结束后水含量在所述合适范围内。
步骤(2)中：
所述有机酰胺晶浆由氯化钠晶种和所述有机酰胺溶剂配制得到，其中氯化钠晶种的浓度为0.1～1mol/L，氯化钠晶种的粒径为1～20μm；
所述二氯代芳香化合物与所述脱水液中硫源的摩尔比为1.0～1.1：1；优选为1.0～1.08：1。
脱水液、二氯代芳香化合物、含有氯化钠晶种的有机酰胺晶浆混合后，以此时预聚反应器体系内的1摩尔硫源计，控制所述氯化钠晶种的加入量为0.05～0.4mol％，此时体系内，有机酰胺溶剂的总量为2.0～4.0mol。
优选的，以此时预聚反应器体系内的1摩尔硫源计，有机酰胺溶剂的总量为2.0～2.5mol(即控制体系内NMP/S＝2.0～2.5)。
经试验发现，优选后不仅可以减少生产工艺产生的闪蒸能耗以及待回收的废液量，还可以适当提高制备的PAS的重均分子量。
进一步优选，控制二氯代芳香化合物和硫源的摩尔比为1.0～1.025：1。
经试验发现，当二氯代芳香化合物和硫源的摩尔比与NMP/S的比值均控制在上述进一步优选的范围内，可以显著提高制备的PAS的重均分子量。
发明者还对聚苯硫醚的聚合反应动力学进行了深入的研究，发现：(1)聚亚芳基硫醚的主聚合反应为强放热反应，预聚过程的温度失控会严重影响反应安全和产品质量，且聚合反应为二级动力学反应，因此为满足反应温度控制及反应效率的要求，对于多反应器串联的连续制造工艺，第一反应器优选为连续搅拌釜式反应器并控制转化率在合适范围；(2)提高反应体系中的水含量可以降低反应速度亦即氯化钠的生成速度，但含水量过高会带来反应效率过低，反应压力高以及闪蒸过程需要更高能耗等问题。
综合以上研究成果，针对PAS现有间歇制造工艺生产过程及连续制造工艺开发过程由于大量副产盐产生的各种问题：(1)洗涤除去副产盐，导致高水耗和高含盐废水排放；(2)易堵塞问题限制PAS连续制造工艺的开发；(3)限制底物浓度的提高；以及主聚合反应为二级动力学强放热反应对反应过程温度控制的高要求，本发明提供一种更加环保绿色、低成本和高效的PAS连续制造方法。
优选的：
步骤(2)中，所述预聚反应器为串联的两级反应器；
为兼顾反应温度控制和反应效率，第一级反应器选自连续搅拌釜式反应器，控制釜式反应器中反应温度为180～220℃，优选为190～210℃。
当第一级反应器中单体转化率达到60～80％后，更优选为70～80％后，将反应液注入第二级反应器。
第二级反应器选自夹套式管式反应器，控制反应器中反应温度为220～240℃。
优选的，所述夹套式管式反应器采用的加热介质为导热油，导热油流动方向与反应液流动方向一致，控制入口油温为220～230℃，出口油温为230～240℃；优选出口处单体转化率达达到90％以上，更优选为93％以上，此时，氯化钠晶体的生长基本完成，粒径成长至60μm以上，优选为100μm以上，尺寸均一。
所述过滤装置为连续过滤装置，具体选自密闭式加压过滤装置，优选 为加压转鼓过滤机，适用于含易挥发组分和易析出晶体的高温浆料的连续过滤。
优选的，所述过滤装置中滤材的孔径优选为60～80μm。该尺寸下氯化钠颗粒可以充分被除去，同时过滤的压力降不会太高。
步骤(3)中，
所述聚合反应器选自管式反应器，优选的，所述管式反应器选自浸在高温油浴中的盘管式反应器，控制反应温度在235～280℃，完成进一步的分子量增长，保证单体转化率达到95％以上，优选为98％以上。
优选的，所述闪蒸处理使用260～300℃的过热蒸汽进行辅助供热，过热蒸汽的用量为0.1～1kg/mol硫源，蒸汽和汽化的溶剂从闪蒸器上部采出，冷凝后进入溶剂回收系统，含盐PAS粗产品从下部连续放出。
所述后处理包括烘干、洗涤、过滤和再烘干，具体为：
含盐PAS粗产品经烘干机连续烘干进一步除去残留溶剂。
烘干后含盐PAS粗产品加水打浆后经一次连续高温高压水洗，加水量与粗产品的质量比为3～10：1，连续过滤、烘干后获得最终产品。
所述连续高温高压水洗的设备可以是连续釜式洗涤设备或连续管式洗涤设备；优选为带导热油夹套的连续管式洗涤设备，水洗浆料的停留时间优选为1～30min，水洗温度优选为150～220℃。
所述连续过滤的设备选自连续的过滤器，如板框过滤器、沉降式离心机和带式过滤器等，优选为连续真空抽滤的带式过滤器。
所述烘干机选自连续进出料的烘干机。
采用本发明的制备工艺，因为预先将大颗粒的盐过滤去除，大幅减少了后续粗品中的盐浓度，显著降低了后处理中洗涤过程的压力，因此，仅需一次连续高温高压水洗即可；若不预先过滤盐颗粒，则需要若干次常压水洗+高压水洗才能充分洗去产品中氯化钠，大幅增加了洗涤用水量和产生的含盐废水量；若不加入晶种以及将反应液中水含量控制在合适水平，氯化钠颗粒尺寸较小，采用大孔径滤材过滤则难以有效拦截，采用小孔径 滤材则过滤压降过大，容易堵塞过滤器。
上述工艺同样适用于聚亚芳基硫醚的间歇化生产，各工艺参数对间歇化生产制备的PAS性能的影响与连续化生产中类似，下面将不再赘述。
因此，本发明还公开了一种聚亚芳基硫醚的间歇式制造方法，包括：
(a)将硫源水溶液、碱金属氢氧化物水溶液、可选择性加入的助剂与有机酰胺溶剂投入反应釜Ⅰ中，不断升温至200～220℃进行脱水皂化反应，脱水结束后水和硫源的摩尔比为1.5～2.0，然后冷却至150～180℃；
(b)将含有氯化钠晶种的有机酰胺晶浆、二氯代芳香化合物和补加的有机酰胺溶剂投入反应釜Ⅰ中，先升温至205～220℃，保温0.5～3h，再升温至215～240℃，保温0.5～5h，然后将得到的反应液注入过滤装置，过滤除去氯化钠颗粒；
(c)经过滤装置后得到的滤液注入预热好的反应釜Ⅱ中，升温至240～280℃，保温0.5～5h，然后将所得反应液进行闪蒸处理，最后经后处理得到聚亚芳基硫醚。
步骤(a)中：
所述硫源水溶液的浓度为28～48wt％，硫源选自硫氢化钠和/或硫化钠；
所述碱金属氢氧化物水溶液为氢氧化钠水溶液，浓度为30～70wt％；
所述二氯代芳香化合物选自对二氯苯、二氯代萘、二氯代芴、二氯咔唑中一种或多种，优选对二氯苯；硫源与氢氧化钠的摩尔比为1：0.95～1.1；
所述助剂选自醋酸钠、苯甲酸钠、C5～C6的脂肪酸钠中的一种或多种；
硫源与助剂的摩尔比为1：0～0.5；
所述有机酰胺溶剂选自N-甲基吡咯烷酮、六甲基磷酰三胺、N-甲基-ε-己内酰胺、N,N-二甲基甲酰胺中的一种或多种；
硫源与有机酰胺溶剂的摩尔比为1：1.5～3。
步骤(b)中：
所述有机酰胺晶浆中氯化钠晶种的浓度为0.1～1mol/L，氯化钠晶种的粒径为1～20μm；所述氯化钠晶种的加入量为此时体系内硫源的 0.05～0.4mol％；
所述二氯代芳香化合物与所述脱水液中硫源的摩尔比为1.0～1.1：1；
补加有机酰胺溶剂后，反应釜Ⅰ中，有机酰胺溶剂的总摩尔量与硫源的摩尔比为2.5～3.5；
优选的：
先以0.2～0.5℃/min的升温速率升温至205～220℃，再以0.2～1℃/min的升温速率升温至215～240℃。
步骤(a)、(b)中采用的反应釜选自带导热油夹套和盘管的搅拌式反应釜。
步骤(c)中：
优选的，以0.2～1℃/min的升温速率升温至240～280℃。
优选的，闪蒸结束后，在闪蒸器中继续烘干10min～1h，然后降温至室温后放料。
所述后处理包括烘干、洗涤、过滤和再烘干。
本发明还公开了根据上述两种工艺分别制备的聚亚芳基硫醚。
本发明还公开了所述聚亚芳基硫醚在制备交联型聚亚芳基硫醚中的应用，具体为通过将所述聚亚芳基硫醚进行热氧处理。
与现有技术相比，本发明具有如下有益效果：
本发明公开了一种聚亚芳基硫醚树脂的制造工艺，核心是通过向PAS反应液中加入氯化钠晶种，同时控制体系中含水量在合适范围并将预聚段的反应温度控制在较低温度下，以降低氯化钠的成核速度、提高晶体的生长速度，达到抑制初级成核，促进次级成核的目的，最终提高副产氯化钠晶体的尺寸；该工艺的另一关键还在于过滤除去盐颗粒的时机，盐颗粒的过滤选择在预聚反应阶段结束时进行；采用本发明公开的制造工艺可将制备的聚亚芳基硫醚树脂粗品中的含盐量大幅降低，极大地降低了高含盐洗涤废水的产生，减少了水耗和能耗，且有助于高效制备高分子量的聚亚芳基硫醚树脂；更为重要的是，该工艺不仅适用于间歇工艺，同样适用于连 续化生产，有望大幅降低该生产体系的生产效率与生产成本，且更加环保绿色。当采用连续化生产工艺时，为兼顾预聚段的温度控制和生产效率，预聚段分为两段，第一段选用适于稳定移热的釜式反应器，并控制转化率在合适值以移除大部分反应热但同时反应速率又没有太低，第二段选用夹套管式反应器，并利用剩余放热推动反应继续升温以维持反应效率；高温聚合段选用盐颗粒去除后更加适合高转化率、高粘度状态(高分子量和高浓度)的盘管式反应器。
附图说明
图1为本发明的聚亚芳基硫醚的连续化制造的工艺流程示意图，其中省略了后处理工艺；
图中，1-对二氯苯，2-脱水液，3-含有氯化钠晶种的NMP晶浆，4-第一聚合反应釜，5-第一反应混合液，6-第一换热器，7-第二管式反应器，8-密闭式加压过滤机，9-盐颗粒，10-第二反应混合液，11-第二换热器，12-第三管式反应器，13-过热蒸汽，14-常压闪蒸器，15-未烘干粗品，16-烘干机进料仓，17-连续烘干机，18-烘干后粗产品，19-汽化的溶剂和蒸汽，20-冷凝器，21-待回收废液。
具体实施方式
下面结合实施例和对比例对本发明作进一步详细的描述，但本发明的实施方式不限于此。
灰分测试：将瓷坩埚放入750℃恒温的马弗炉中烧至恒重，然后取出置于干燥器中冷却后称重，记为M0。称取3g样品加入坩埚中，加入10mL硝酸，然后置于酒精喷灯上烧至再无烟气冒出。最后再将坩埚置于750℃恒温的马弗炉中灼烧1h，取出置于干燥器中冷却后称重，记为M1。树脂灰分计算公式为：(M1-M0)/3*100％。
分子量测试：PPS的分子量使用凝胶渗透色谱仪以聚苯乙烯为标样测 定，流动相为1-氯萘，柱温220℃，流动速度1mL/mL，检测器为折光指数检测器。
实施例1
在100L反应釜Ⅰ内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至200℃，过程中水份不断被脱除，出馏液10.00kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为147.0mol。
在反应釜Ⅰ中继续加入对二氯苯15.14kg(103.0mol)，含11.7g氯化钠晶种(0.2mol，氯化钠晶种的D50为15μm)的NMP晶浆共10kg，再补加4.46kg NMP(45.0mol)，密封反应釜Ⅰ，以0.3℃/min升温至215℃，保温3h，然后以0.2℃/min升温至230℃，保温1h，然后将反应液经过滤器(滤网孔径80μm)过滤除去氯化钠颗粒后泵入已经预热至230℃的100L反应釜Ⅱ中，然后将反应釜Ⅱ以0.5℃/min升温至260℃，保温1h，然后缓慢泄放至常压闪蒸器中进行闪蒸处理，最后经烘干进一步除去溶剂后获得粗产品10.4kg，加入60kg水打浆，升温至180℃，保温0.5h后降温，过滤烘干后得到最终产物。
本实施例中，反应液转移过滤时得到氯化钠10.9kg，经激光粒度分析仪测得氯化钠颗粒的D50为130μm；洗涤过程用水60.0kg，产生含盐废水50.0kg，含盐量为0.6wt％；制备的最终产物为10.1kg，重均分子量为22500，灰分为0.53wt％。
对比例1
在100L反应釜Ⅰ内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至 200℃，过程中水份不断被脱除，出馏液10.00kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为147.0mol。
在反应釜Ⅰ中加入对二氯苯15.14kg(103.0mol)，NMP 14.46kg，密封反应釜Ⅰ，以0.3℃/min升温至215℃，保温3h，然后以0.2℃/min升温至230℃，保温1h，然后将反应液经过滤器(滤网孔径20μm)过滤除去氯化钠颗粒后泵入已经预热至230℃的100L反应釜Ⅱ中，然后将反应釜Ⅱ以0.5℃/min升温至260℃，保温1h，然后缓慢泄放至常压闪蒸器中进行闪蒸处理，最后经烘干进一步除去溶剂后获得粗产品10.8kg，加入60kg水打浆，升温至180℃，保温0.5h后降温，过滤烘干后得到最终产物
本对比例中，反应液转移过滤时得到氯化钠10.5kg，经激光粒度分析仪测得氯化钠颗粒的D50为30μm；洗涤过程用水60.0kg，产生含盐废水51.0kg，含盐量为1.3wt％；制备的最终产物为10.1kg，重均分子量为22500，灰分为0.56wt％。
需要说明的是，由于盐粒径和与之相匹配的过滤器滤网孔径较小，过滤器压降高、通量小，过滤步骤花费了大量时间，大大降低了生产效率，显著增加了生产成本。
对比例2
在100L反应釜Ⅰ内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至200℃，过程中水份不断被脱除，出馏液10.00kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为147.0mol。
在反应釜Ⅰ中继续加入对二氯苯15.14kg(103.0mol)，含11.7g氯化钠晶种(0.2mol，氯化钠晶种的D50为15μm)的NMP晶浆共10kg，再补加4.46kg NMP(45.0mol)，密封反应釜，以0.3℃/min升温至215℃，保温3h，然后以0.15℃/min升温至230℃，保温1h，然后以0.5℃/min升 温至260℃，保温1h，然后缓慢泄放至常压闪蒸器中进行闪蒸处理，再经烘干进一步除去溶剂后获得粗产品21.2kg。粗产品先经两次常压水洗-过滤，每次加水量为60.0kg，再加水60.0kg打浆后升温至180℃，保温0.5h，降温后过滤，过滤烘干后得到最终产品。
本对比例中，洗涤过程用水180.0kg，产生含盐废水182.0kg，含盐量6.0wt％；制备的最终产物为10.0kg，重均分子量为21500，灰分为0.50wt％。
实施例1和对比例2比较表明，通过过滤除去盐颗粒可以显著减少洗涤用水量和产生的废盐水量。
实施例2
在100L反应釜内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至200℃，过程中水份不断被脱除，出馏液10.00kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为147.0mol。将按上述方法分批制备的脱水液放入170℃的储存釜中保温存放，供后续连续进料使用。
使用图1所示的PAS连续制造工艺流程图。将预热至200℃的对二氯苯1(PDCB)、脱水液2(硫化钠含量25.48wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的D50为15μm，晶浆的浓度为0.136mol/L)依次以25.21g/min、50g/min和24.5g/min的流量泵至第一聚合反应釜4(控制第一聚合反应釜4中PDCB/S＝1.050(摩尔比，下同)，NMP/S＝3.50(摩尔比，下同)，第一聚合反应釜4温度为215℃，停留时间为5h，所得第一反应混合液5经第一换热器6升温至225℃，进入第二管式反应器7(反应器温度为225℃)，停留时间1h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤网孔径为80μm)除去盐颗粒，过滤出的盐颗粒9的D50为140μm；所得第二反应混合液10经第二换热器11升温至265℃，再进 入第三管式反应器12(反应器温度为265℃)，停留时间1.5h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为5.0kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化，产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，获得最终产物。
本实施例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量0.6wt％，消耗过热蒸汽5.0kg/h，产生待回收废液8.7kg/h。制备的最终产物经测试重均分子量约为23000，灰分低于0.55wt％。
实施例3
按照实施例2中的方法分批制备脱水液，放入储罐中保温在170℃中供后续连续进料使用。使用图1所示的PAS连续制造装置。将预热至200℃的对二氯苯1(PDCB)、脱水液2(硫化钠含量25.48wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的D50为15μm，晶浆浓度为0.4mol/L)依次以25.21g/min、50g/min和8.1g/min的流量泵至第一聚合反应釜4(PDCB/S＝1.050，NMP/S＝2.5)，第一聚合反应釜4温度为215℃，停留时间为3.5h，所得第一反应混合液5经第一换热器6升温至230℃，进入第二管式反应器7(反应器温度为230℃)，停留时间1h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤布孔径为80μm)除去盐颗粒，过滤出的盐颗粒9的D50为130μm；所得第二反应混合液10经第二换热器11升温至265℃，进入第三管式反应器12(反应器温度为265℃)，停留时间1h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为 3.5kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化，产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，获得最终产品。
本实施例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量0.6wt％，消耗过热蒸汽3.5kg/h，产生待回收废液6.2kg/h。制备的最终产物经测试重均分子量约为26000，灰分低于0.55wt％。
对比实施例2和实施例3的记载可知，进一步降低预聚反应阶段，具体为第一聚合反应釜内NMP/S的摩尔比，可以降低溶剂用量以及闪蒸阶段过热蒸汽的用量，减少能耗和废液量，显著降低生产成本；可以提高反应效率，降低反应液在各段反应器中的停留时间；还可以适当提高最终产物的重均分子量。
但降低NMP/S的摩尔比意味着提高体系内PPS和氯化钠的浓度，随着体系中PPS和氯化钠的浓度，以及PPS分子量的增加，反应液的粘度快速升高，采用现有技术中的生产工艺，无论是间歇式还是连续式工艺，都很难在低NMP/S下顺利实施生产。这恰恰说明了本发明中公开的生产工艺适应性强，尤其适用于低NMP/S的摩尔比的生产工况，因此更利于高效率制备高分子量的PAS。
实施例4
在100L反应釜内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(氢氧化钠摩尔数为100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(硫氢化钠摩尔数为100.0mol)，无水醋酸钠1.6kg，然后在氮气保护下不断升温脱水，脱除10.00kg水溶液(含水量98.0wt％)，脱水完毕后，降温至170℃。此时，脱水液体系中硫的量为98.0mol，含 水量为147.0mol。将按上述方法分批制备的脱水液放入170℃的储存釜中保温存放，供后续连续进料使用。
使用图1所示的PAS连续制造装置。将预热后的对二氯苯1(PDCB)、脱水液2(硫化钠含量24.22wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的粒径为15μm，晶浆浓度为0.4mol/L)依次以24.61g/min、52.6g/min和8.1g/min的流量泵至第一聚合反应釜4(PDCB/S＝1.025，NMP/S＝2.5，CH3COONa/S＝0.2)，第一聚合反应釜温度4为215℃，停留时间为3.5h，所得第一反应混合液5经第一换热器6升温至230℃，进入第二管式反应器7(反应器温度为230℃)，停留时间1h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤网孔径80μm)除去盐颗粒，过滤出的盐颗粒9的D50为120μm；所得第二反应混合液10经第二换热器11升温至265℃后进入第三管式反应器12(反应器温度为265℃)，停留时间1h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为3.3kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化，产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜中加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，获得最终产品。
本实施例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量0.65wt％，消耗过热蒸汽3.3kg/h，产生待回收废液6.0kg/h。制备的最终产物经测试重均分子量约为68000，灰分低于0.5wt％。
实施例5
在100L反应釜内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47wt％硫氢化 钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至195℃，过程中水份不断被脱除，出溜液9.07kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为196.0mol。将按上述方法分批制备的脱水液放入170℃的储存釜中保温存放，供后续连续进料使用。
使用图1所示的PAS连续制造工艺流程图。将预热200℃的对二氯苯1(PDCB)、脱水液2(硫化钠含量24.66wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的D50为15μm，晶浆的浓度为0.136mol/L)依次以25.19g/min、51.66g/min和24.5g/min的流量泵至第一聚合反应釜4(PDCB/S＝1.050，NMP/S＝3.50)，第一聚合反应釜4温度为215℃，停留时间为6h，所得第一反应混合液5经第一换热器6升温至225℃，进入第二管式反应器7(反应器温度为225℃)，停留时间1.5h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤布孔径为80μm，)除去盐颗粒，过滤出的盐颗粒9的D50为150μm；所得第二反应混合液10经第二换热器11升温至265℃，再进入第三管式反应器12(反应器温度为265℃)，停留时间1.5h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为5.0kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化，产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，获得最终产物。
本实施例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量0.6wt％，消耗过热蒸汽5.0kg/h，产生待回收废液8.78kg/h。制备的最终产物经测试重均分子量约为21000，灰分低于0.55wt％。
实施例6
在100L反应釜内，加入N-甲基吡咯烷酮14.85kg(150.0mol)，48.0wt％氢氧化钠水溶液8.33kg(NaOH摩尔数为100.0mol)，加入47.0wt％硫氢化钠水溶液11.93kg(NaHS摩尔数为100.0mol)，在氮气保护下缓慢升温至203℃，过程中水份不断被脱除，出溜液9.47kg(含水量98.0％)，脱水完毕，降温至170℃。此时，体系中硫的量为98.0mol，含水量为176.4mol。将按上述方法分批制备的脱水液放入170℃的储存釜中保温存放，供后续连续进料使用。
使用图1所示的PAS连续制造工艺流程图。将预热200℃的对二氯苯1(PDCB)、脱水液2(硫化钠含量29.83wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的D50为15μm，晶浆的浓度为0.2mol/L)依次以24.60g/min、42.71g/min和8.1g/min的流量泵至第一聚合反应釜4(PDCB/S＝1.025，NMP/S＝2.0)，第一聚合反应釜4温度为215℃，停留时间为3.5h，所得第一反应混合液5经第一换热器6升温至225℃，进入第二管式反应器7(反应器温度为225℃)，停留时间1h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤布孔径为80μm，)除去盐颗粒，过滤出的盐颗粒9的D50为180μm；所得第二反应混合液10经第二换热器11升温至265℃，再进入第三管式反应器12(反应器温度为265℃)，停留时间1h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为2.8kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，然后烘干获得最终产物。
本实施例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量 0.58wt％，消耗过热蒸汽2.8kg/h，产生待回收废液5.1kg/h。制备的最终产物经测试重均分子量约为32000，灰分低于0.52wt％。
对比常规NMP/S摩尔比的实施例2(摩尔比为3.5)，采用本发明中的连续化制备工艺，可以在NMP/S摩尔比为2.5(实施例4)和NMP/S摩尔比为2.0(实施例6)下成功制备聚苯硫醚，不仅显著降低了能耗(过热蒸汽消耗量)和三废量(待回收废液量)，还会降低停留时间(反应时间)，显著提高反应效率。
对比实施例6和实施例2为例，过热蒸汽消耗量减少近半，产生待回收废液量减少了超过1/3，反应总停留时间减少了约20％，以上技术效果可大幅降低该连续化生产的生产成本。
对比例3
在100L反应釜内，加入N-甲基吡咯烷酮19.82kg(200.0mol)，48wt％氢氧化钠水溶液8.33kg(100.0mol)，加入47wt％硫氢化钠水溶液11.93kg(100.0mol)，然后在氮气保护下不断升温脱水，脱除10.51kg水溶液(含水量98.0wt％)，脱水完后毕，降温至170℃。此时，脱水液体系中硫的量为98.0mol，含水量为117.6mol。将按上述方法分批制备的脱水液放入170℃的储存釜中保温存放，供后续连续进料使用。
使用图1所示的PAS连续制造工艺流程图。将预热200℃的对二氯苯1(PDCB)、脱水液2(硫化钠含量25.86wt％)和含有氯化钠晶种的NMP晶浆3(氯化钠晶种的D50为15μm，晶浆的浓度为0.136mol/L)依次以25.21g/min、49.27g/min和24.5g/min的流量泵至第一聚合反应釜4(PDCB/S＝1.050，NMP/S＝3.50)，第一聚合反应釜4温度为215℃，停留时间为5h，所得第一反应混合液5经第一换热器6升温至225℃，进入第二管式反应器7(反应器温度为225℃)，停留时间1h后通入密闭式加压过滤机8，经密闭式加压过滤机8(滤布孔径为30μm，)除去盐颗粒，过滤出的盐颗粒9的D50为60μm；所得第二反应混合液10经第二换热器 11升温至265℃，再进入第三管式反应器12(反应器温度为265℃)，停留时间1.5h，然后经换热器(未画出)换热至280℃后进入常压闪蒸器14进行闪蒸，闪蒸时使用265℃过热蒸汽13辅助闪蒸，过热蒸汽13的流量为5.0kg/h，汽化的溶剂和蒸汽19经冷凝器20冷凝液化，产生待回收废液21。未烘干粗品15先经烘干机进料仓16进入连续烘干机17，经进一步烘干后得到的粗产品18连续排入打浆釜加水打浆，加水量为6.0kg/h，经换热至180℃，进入带导热油夹套的管式洗涤器，控制管式洗涤器温度为180℃，停留时间10min，然后经换热降温至60℃，连续过滤、烘干(未画出)，获得最终产物。
本对比例的生产速率为1.01kg/h；产生含盐废水5.0kg/h，含盐量0.65wt％，消耗过热蒸汽5.0kg/h，产生待回收废液8.7kg/h。制备的最终产物经测试重均分子量约为23000，灰分低于0.6wt％。
需要说明的是，本对比例实施过程中，由于盐粒径和与之相匹配的滤布孔径较小，过滤器的阻力较大，偶尔会发生堵塞，连续实施的稳定性较差。

Claims (11)

  1. 一种聚亚芳基硫醚的连续化制造方法，其特征在于，包括：
    (1)将硫源水溶液、碱金属氢氧化物水溶液、可选择性加入的助剂与有机酰胺溶剂投入反应釜中，不断升温至200～220℃进行脱水皂化反应，脱水结束后水和硫源的摩尔比为1.5～2.0，冷却至150～180℃，将所得脱水液保温在储存釜中；
    (2)将含有氯化钠晶种的有机酰胺晶浆、二氯代芳香化合物和所述脱水液注入预聚反应器中，反应温度控制在180～240℃，反应至单体转化率达到90％及以上后，将得到的反应液注入过滤装置，过滤除去氯化钠颗粒；
    (3)经过滤装置后得到的滤液注入聚合反应器中，反应温度控制在235～280℃，反应至单体转化率达到95％及以上，将得到的反应液进行闪蒸处理，最后经后处理得到聚亚芳基硫醚。
  2. 根据权利要求1所述的聚亚芳基硫醚的连续化制造方法，其特征在于，步骤(1)中：
    所述硫源水溶液的浓度为28～48wt％，硫源选自硫氢化钠和/或硫化钠；
    所述碱金属氢氧化物水溶液为氢氧化钠水溶液，浓度为30～70wt％；
    所述二氯代芳香化合物选自对二氯苯、二氯代萘、二氯代芴、二氯咔唑中一种或多种；
    硫源与氢氧化钠的摩尔比为1：0.95～1.1；
    所述助剂选自醋酸钠、苯甲酸钠、C5～C6的脂肪酸钠中的一种或多种；
    硫源与助剂的摩尔比为1：0～0.5；
    所述有机酰胺溶剂选自N-甲基吡咯烷酮、六甲基磷酰三胺、N-甲基-ε-己内酰胺、N,N-二甲基甲酰胺中的一种或多种；
    硫源与有机酰胺溶剂的摩尔比为1：1.5～2.5。
  3. 根据权利要求1所述的聚亚芳基硫醚的连续化制造方法，其特征 在于，步骤(2)中：
    所述有机酰胺晶浆中氯化钠晶种的浓度为0.1～1mol/L，氯化钠晶种的粒径为1～20μm；
    所述二氯代芳香化合物与所述脱水液中硫源的摩尔比为1.0～1.1：1；
    脱水液、二氯代芳香化合物、含有氯化钠晶种的有机酰胺晶浆混合后，以此时体系内的1摩尔硫源计，所述氯化钠晶种的加入量为0.05～0.4mol％，此时体系内，有机酰胺溶剂的总量为2.0～4.0mol。
  4. 根据权利要求3所述的聚亚芳基硫醚的连续化制造方法，其特征在于：
    所述二氯代芳香化合物与所述脱水液中硫源的摩尔比为1.0～1.025：1；
    脱水液、二氯代芳香化合物、含有氯化钠晶种的有机酰胺晶浆混合后，以此时体系内的硫源为1摩尔计，有机酰胺溶剂的总量为2.0～2.5mol。
  5. 根据权利要求1所述的聚亚芳基硫醚的连续化制造方法，其特征在于，步骤(2)中，所述预聚反应器为串联的两级反应器；
    第一级反应器选自连续搅拌釜式反应器，控制釜式反应器中反应温度为180～220℃；
    第二级反应器选自夹套式管式反应器，控制反应器中反应温度为220～240℃；
    当第一级反应器中单体转化率达到60～80％后，将反应液注入第二级反应器。
  6. 根据权利要求1所述的聚亚芳基硫醚的连续化制造方法，其特征在于，步骤(3)中，
    所述聚合反应器选自管式反应器；
    所述后处理包括烘干、洗涤、过滤和再烘干。
  7. 一种聚亚芳基硫醚的间歇式制造方法，其特征在于，包括：
    (a)将硫源水溶液、碱金属氢氧化物水溶液、可选择性加入的助剂与有机酰胺溶剂投入反应釜Ⅰ中，不断升温至200～220℃进行脱水皂化反应， 脱水结束后水和硫源的摩尔比为1.5～2.0，然后冷却至150～180℃；
    (b)将含有氯化钠晶种的有机酰胺晶浆、二氯代芳香化合物和补加的有机酰胺溶剂投入上述反应釜Ⅰ中，先升温至205～220℃，保温0.5～5h，再升温至215～240℃，保温0.5～5h，然后将得到的反应液注入过滤装置，过滤除去氯化钠颗粒；
    (c)经过滤装置后得到的滤液注入预热好的反应釜Ⅱ中，升温至240～280℃，保温0.5～5h，然后将所得反应液进行闪蒸处理，最后经后处理得到聚亚芳基硫醚。
  8. 根据权利要求7所述的聚亚芳基硫醚的间歇式制造方法，其特征在于，步骤(a)中：
    所述硫源水溶液的浓度为28～48wt％，硫源选自硫氢化钠和/或硫化钠；
    所述碱金属氢氧化物水溶液为氢氧化钠水溶液，浓度为30～70wt％；
    所述二氯代芳香化合物选自对二氯苯、二氯代萘、二氯代芴、二氯咔唑中一种或多种；
    硫源与氢氧化钠的摩尔比为1：0.95～1.1；
    所述助剂选自醋酸钠、苯甲酸钠、C5～C6的脂肪酸钠中的一种或多种；
    硫源与助剂的摩尔比为1：0～0.5；
    所述有机酰胺溶剂选自N-甲基吡咯烷酮、六甲基磷酰三胺、N-甲基-ε-己内酰胺、N,N-二甲基甲酰胺中的一种或多种；
    硫源与有机酰胺溶剂的摩尔比为1：1.5～3。
  9. 根据权利要求7所述的聚亚芳基硫醚的间歇式制造方法，其特征在于，步骤(b)中：
    所述有机酰胺晶浆中氯化钠晶种的浓度为0.1～1mol/L，氯化钠晶种的粒径为1～20μm；所述氯化钠晶种的加入量为此时体系内硫源的0.05～0.4mol％；
    所述二氯代芳香化合物与所述脱水液中硫源的摩尔比为1.0～1.1：1；
    补加有机酰胺溶剂后，反应釜Ⅰ中，有机酰胺溶剂的总摩尔量与硫源 的摩尔比为2.5～3.5；
    先以0.2～0.5℃/min的升温速率升温至205～220℃，再以0.2～1℃/min的升温速率升温至215～240℃；
    步骤(c)中：
    以0.2～1℃/min的升温速率升温至240～280℃；所述后处理包括烘干、洗涤、过滤和再烘干。
  10. 一种根据权利要求1～6任一项所述的连续化制造方法或7～9任一项所述的间歇式制造方法制备的聚亚芳基硫醚。
  11. 一种根据权利要求10所述的聚亚芳基硫醚在制备交联型聚亚芳基硫醚中的应用，其特征在于，通过将所述聚亚芳基硫醚进行热氧处理。
PCT/CN2023/096116 2022-07-27 2023-05-24 聚亚芳基硫醚树脂的制造工艺及其产品和应用 WO2024021816A1 (zh)

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