CN115260496B

CN115260496B - Process for producing polyarylene sulfide resin, and product and use thereof

Info

Publication number: CN115260496B
Application number: CN202210888486.1A
Authority: CN
Inventors: 尹红; 贾艳宇; 张雄伟; 陈志荣; 连明; 陈兴; 蒋杰; 李沃源
Original assignee: Zhejiang University ZJU; Zhejiang NHU Co Ltd; Zhejiang NHU Special Materials Co Ltd
Current assignee: Zhejiang University ZJU; Zhejiang NHU Co Ltd; Zhejiang NHU Special Materials Co Ltd
Priority date: 2022-07-27
Filing date: 2022-07-27
Publication date: 2024-03-29
Anticipated expiration: 2042-07-27
Also published as: CN115260496A; WO2024021816A1

Abstract

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

Description

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

Technical Field

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

Background

A typical process for producing polyarylene sulfide (PAS) comprises polymerizing sodium sulfide/sodium hydrosulfide and a dihalo-aromatic compound in an organic amide solvent in the presence of an alkali metal hydroxide, preferably N-methylpyrrolidone (hereinafter, abbreviated as NMP), and preferably sodium hydroxide. At present, industrial production devices for synthesizing polyphenylene sulfide (PPS) by the method are all intermittent operation, part of reaction raw materials are dehydrated in a reaction kettle, after cooling and supplementing other reaction raw materials, the reaction kettle is closed, the temperature is raised for high-temperature polycondensation, after the reaction is finished, a crude product is obtained through cooling and filtering or flash evaporation, and the crude product is purified to obtain a final product. The intermittent production process needs repeated loading, temperature rising and reducing, unloading and cleaning of the reaction kettle, and has long time consumption, greatly reduced production efficiency and poor process stability and reliability. For this reason, researchers have been struggling to develop continuous processes for producing PAS to improve the production efficiency.

Patent US4060520, US4056515 and US4066632 propose a continuous production process with multiple kettles connected in series, wherein the reaction temperature of each kettle is different, the kettle pressure is continuously reduced from front to back, and the material is pushed to flow from front to back by the pressure difference between adjacent kettles. In chinese patent publication No. CN 108779253A, there is provided a continuous production apparatus and continuous production method of PAS, wherein the continuous production apparatus comprises a housing chamber for housing a plurality of reaction tanks, an organic amide solvent, a sulfur source and a dihalo-aromatic compound are supplied into the housing chamber, and polymerization reaction of the sulfur source and the dihalo-aromatic compound is performed in the organic amide solvent to form a reaction mixture; the reaction tanks are connected in sequence, and are communicated with each other through gas phase, and the reaction mixture moves to each reaction tank in sequence. In chinese patent publication No. CN 113667122a, a gradient temperature-controlled continuous condensation method of polyarylene sulfide resin is provided, the temperature of each reactor is controlled stepwise (temperature is raised stepwise) to adapt to different progress of polymerization reaction, and the reaction raw materials are continuously fed into a series of kettle-type or pipe-type reactors by a metering pump.

The sodium sulfide method produces a large amount of sodium chloride as a byproduct, and each ton of PPS is produced, 1.08 tons of sodium chloride is produced, and the sodium chloride begins to be produced and precipitated in a large amount in the early stage of the polymerization reaction, so that the reaction slurry contains a large amount of salt particles, and the content of the byproduct salt particles in the reaction-terminated polymerization reaction solution is about 20.5wt% based on a typical molar ratio of NMP/sulfur source in the reaction solution of 3.5. A large amount of salt particles in the slurry are prone to settling and clogging in long distance transport or reaction lines, especially at the finer sized tubular reactors, resulting in production shutdowns. Whether the process is a batch process or a continuous process, the improvement of the concentration of the reaction substrate is beneficial to the improvement of the production efficiency, the high-efficiency production of the high-molecular-weight polyphenylene sulfide is also beneficial to the high-efficiency production of the high-molecular-weight polyphenylene sulfide, but the content of byproduct salt particles is also obviously increased along with the high-molecular-weight polyphenylene sulfide, and the viscosity of the reaction solution is also sharply increased, so that the problems of conveying blockage and stirring are further worsened. The presence of a large amount of byproduct salt particles in PPS reaction slurry severely limits the research and development of continuous PPS production processes, and limits the efficient production of higher molecular weight polyphenylene sulfide.

Whether the intermittent or continuous manufacturing process is adopted, the byproduct sodium chloride is wrapped in the product after the reaction is finished, a large amount of pure water is needed to be used for washing and removing for many times, a large amount of washing wastewater with high salt content is generated, at present, more than 15 tons of water is needed to be consumed for producing one ton of PPS in industry, and the energy consumption is increased due to the fact that the byproduct salt dissolved in the water is recovered through evaporation. In addition, the high corrosiveness of brine also places high demands on the materials of polyphenylene sulfide manufacturing equipment. The preparation process of the polyphenylene sulfide has the characteristics of high water consumption, high energy consumption and high corrosiveness, greatly improves the preparation cost and the environmental protection cost of the polyphenylene sulfide, severely limits the greenness and the sustainability of the production process of the polyphenylene sulfide, and limits the competitiveness and the vitality of the polyphenylene sulfide resin.

The foregoing problems can be solved by removing precipitated salt particles before solid-liquid separation of the product to reduce the salt particle content in subsequent equipment and the crude product. The method for desalting polyphenylene sulfide is disclosed in Chinese patent document with publication number of CN104371103A, wherein Wen Baoya is kept in a reaction kettle after the polymerization reaction is finished, after inorganic salt and inorganic auxiliary agent settle to the bottom of the kettle, the upper layer polyphenylene sulfide solution is extruded out of the reaction kettle through an inner bottom extension pipe by utilizing pressure difference, the solution is cooled, and the precipitated polyphenylene sulfide is purified to obtain a high-purity polyphenylene sulfide product. However, the method described in this patent is only suitable for batch processes, how the inorganic salts, the reaction solvent and the product in the settled layer in the reactor are further treated, is not described further in this patent document, and in fact the settled inorganic salt layer does not allow the stirring of the reactor to be started again until it is not completely cleaned.

In another example, chinese patent publication No. CN107964098A relates to a method and a device for desalting during the synthesis of polyphenylene sulfide, in the polymerization process, the material in the reaction kettle passes through a filtering device equipped with a filter screen by a feed pump, the salt is left in the filtering device, the rest polyphenylene sulfide material returns to the reaction kettle, and the polyphenylene sulfide resin is obtained after the reaction is completed, discharged, washed and dried. The method is also only suitable for batch process, and the later stage of the reaction, the filtration becomes more difficult along with the increase of the thickness of the salt cake layer and the viscosity of the reaction liquid.

Disclosure of Invention

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

The specific technical scheme is as follows:

a continuous process for producing polyarylene sulfide, comprising:

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

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

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

The inventor has conducted intensive studies on the growth process and particle morphology of salt particles in PPS reaction liquid, after all raw materials are added and heated to a certain temperature, polymerization reaction starts, sodium chloride is continuously generated, and due to the existence of water in a polyphenylene sulfide reaction system, a small amount of byproduct sodium chloride can be dissolved in a reaction solvent, and after the dissolution saturation is achieved, the continuously generated sodium chloride starts to nucleate and separate out, the reaction continues to be rapidly carried out, and a large amount of generated sodium chloride rapidly nucleates and separates out. The research finds that: (1) In the prior typical batch method PPS manufacturing process, the particle size of salt particles generated in the reaction liquid is small, about 20 mu m, which is unfavorable for filtration, which has been reported recently; (2) In the pre-polymerization reaction section in the earlier stage, the reaction conversion rate is over 90 percent, the growth of sodium chloride crystals is basically finished at the moment, but the molecular weight of the product is still lower at the moment, the viscosity of the reaction solution is not high, and the reaction needs to be carried out in a high-temperature polymerization section for further reaction to improve the molecular weight, so that the filtering and removal of byproduct salt are proper when the growth of sodium chloride crystals in the pre-polymerization reaction section is basically finished, the viscosity of the reaction solution is higher after the high-temperature polymerization section, and the filtering can be more difficult; (3) The water content in the system is controlled in a proper range, so that on one hand, the precipitation nucleation speed of sodium chloride can be reduced, and on the other hand, the growth speed of sodium chloride crystals can be improved; (4) The introduction of a small amount of sodium chloride seed crystal can effectively inhibit the primary nucleation of sodium chloride.

Based on the research results, experiments show that in the step (1), in order to control the particle size of sodium chloride particles to be more than 100 mu m to facilitate filtration, the water content needs to be controlled at a proper ratio (the molar ratio of water to sulfur source is controlled to be 1.5-2.0) after dehydration is finished; the sodium chloride can be filtered and removed, so that the generation of high-salt-content wastewater is greatly reduced, the water consumption and the energy consumption are reduced, and meanwhile, the solvent consumption in the reaction liquid can be further reduced, thereby being more beneficial to efficiently preparing the PPS with high molecular weight.

In step (1):

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

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

the dichloro aromatic compound is one or more of paradichlorobenzene, dichloro naphthalene, dichloro fluorene and dichloro carbazole, preferably paradichlorobenzene;

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

the auxiliary agent is one or more of sodium acetate, sodium benzoate and C5-C6 fatty acid sodium;

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

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

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

The dehydration saponification reaction can be carried out in batch mode according to the method of the step (1), or a continuous dehydration scheme reported in other patents can be adopted, and the key point is to control the water content within the proper range after dehydration is finished.

In the step (2):

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

the molar ratio of the dichloro aromatic compound to the sulfur source in the dehydration liquid is 1.0-1.1: 1, a step of; preferably 1.0 to 1.08:1.

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

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

Experiments show that the energy consumption of flash evaporation generated in the production process and the amount of waste liquid to be recovered can be reduced after optimization, and the weight average molecular weight of PAS prepared can be properly improved.

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

it was found through experiments that when the molar ratio of the dichloro aromatic compound to the sulfur source to the NMP/S ratio is controlled within the above further preferred range, the weight average molecular weight of the PAS prepared can be significantly increased.

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

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

Preferably:

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

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

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

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

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

The filtering device is a continuous filtering device, is particularly selected from closed type pressurized filtering devices, is preferably a pressurized rotary drum filter, and is suitable for continuous filtering of high-temperature slurry containing volatile components and easy to precipitate crystals.

Preferably, the pore diameter of the filter medium in the filter device is preferably 60 to 80 μm. At this size the sodium chloride particles can be removed sufficiently while the pressure drop of the filtration is not too high.

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

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

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

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

and continuously drying the PAS crude product containing salt by a dryer to further remove the residual solvent.

Adding water into the dried salt PAS crude product, pulping, and carrying out one-time continuous high-temperature high-pressure water washing, wherein the mass ratio of the water adding amount to the crude product is 3-10: and 1, continuously filtering and drying to obtain a final product.

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

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

The dryer is selected from continuous feeding and discharging dryers.

By adopting the preparation process, as the salt with large particles is filtered and removed in advance, the salt concentration in the subsequent crude product is greatly reduced, and the pressure in the washing process in the post-treatment is obviously reduced, so that the preparation process can be carried out only by one continuous high-temperature high-pressure water washing; if salt particles are not filtered in advance, sodium chloride in the product can be fully washed out by a plurality of times of normal pressure water washing and high pressure water washing, so that the water consumption for washing and the produced salt-containing wastewater are greatly increased; if seed crystal is not added and the water content in the reaction liquid is controlled at a proper level, the size of sodium chloride particles is smaller, the sodium chloride particles are difficult to effectively intercept by adopting a large-aperture filter material for filtration, and the pressure drop of the filter is overlarge by adopting a small-aperture filter material, so that the filter is easy to be blocked.

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

Accordingly, the present invention also discloses a batch production method of polyarylene sulfide, comprising:

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

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

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

In step (a):

the dichloro aromatic compound is one or more of paradichlorobenzene, dichloro naphthalene, dichloro fluorene and dichloro carbazole, preferably paradichlorobenzene; the molar ratio of the sulfur source to the sodium hydroxide is 1:0.95 to 1.1;

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

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

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

In step (b):

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

the molar ratio of the dichloro aromatic compound to the sulfur source in the dehydration liquid is 1.0-1.1: 1, a step of;

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

preferably:

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

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

In step (c):

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

Preferably, after the flash evaporation is finished, the material is continuously dried in a flash evaporator for 10min to 1h, and then cooled to room temperature and discharged.

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

The invention also discloses polyarylene sulfide prepared according to the two processes.

The invention also discloses application of the polyarylene sulfide in preparing crosslinked polyarylene sulfide, in particular to application of the polyarylene sulfide in performing thermal oxidation treatment.

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

the invention discloses a process for preparing polyarylene sulfide resin, which comprises the steps of adding sodium chloride seed crystal into PAS reaction liquid, controlling the water content in the system to be in a proper range and controlling the reaction temperature of a prepolymerization section to be at a lower temperature, so as to reduce the nucleation speed of sodium chloride, improve the growth speed of crystals, achieve the purposes of inhibiting primary nucleation and promoting secondary nucleation, and finally improve the size of byproduct sodium chloride crystals; another key to the process is the timing of the removal of salt particles by filtration, which is optionally done at the end of the prepolymerization stage; the preparation process disclosed by the invention can greatly reduce the salt content in the prepared polyarylene sulfide resin crude product, greatly reduce the generation of high-salt washing wastewater, reduce the water consumption and the energy consumption, and is beneficial to efficiently preparing the polyarylene sulfide resin with high molecular weight; more importantly, the process is not only suitable for batch process, but also suitable for continuous production, is expected to greatly reduce the production efficiency and the production cost of the production system, and is more environment-friendly and green. When the continuous production process is adopted, the temperature control and the production efficiency of the prepolymerization section are considered, the prepolymerization section is divided into two sections, the first section adopts a kettle type reactor suitable for stable heat transfer, the conversion rate is controlled at a proper value to remove most of reaction heat, but the reaction rate is not too low, and the second section adopts a jacket type reactor, and the residual heat release is utilized to push the reaction to continuously heat so as to maintain the reaction efficiency; the high-temperature polymerization section is a coiled tube reactor which is more suitable for high conversion rate and high viscosity state (high molecular weight and high concentration) after salt particles are removed.

Drawings

FIG. 1 is a schematic process flow diagram of the continuous production of polyarylene sulfide of the present invention, wherein the post-treatment process is omitted;

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

Detailed Description

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

Ash test: the porcelain crucible is put into a muffle furnace with constant temperature of 750 ℃ to be burned to constant weight, then is taken out and put into a dryer to be cooled, and is weighed and marked as M0. 3g of the sample is weighed and added into a crucible, 10mL of nitric acid is added, and then the mixture is placed on an alcohol burner to burn until no smoke is emitted. Finally, the crucible is placed in a muffle furnace with constant temperature of 750 ℃ to be burned for 1h, taken out and placed in a dryer to be cooled, and then weighed, and marked as M1. The resin ash calculation formula is: (M1-M0)/3 x 100%.

Molecular weight testing: the molecular weight of PPS was measured using a gel permeation chromatograph using polystyrene as a standard, the mobile phase was 1-chloronaphthalene, the column temperature was 220 ℃, the flow rate was 1mL/mL, and the detector was a refractive index detector.

Example 1

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

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

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

Comparative example 1

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

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

It should be noted that, because the salt particle size and the pore diameter of the filter screen matched with the salt particle size are smaller, the pressure drop of the filter is high, the flux is small, the filtering step takes a lot of time, the production efficiency is greatly reduced, and the production cost is remarkably increased.

Comparative example 2

15.14kg (103.0 mol) of paradichlorobenzene, 10kg of NMP crystal slurry containing 11.7g of sodium chloride crystal seeds (0.2 mol, and D50 of the sodium chloride crystal seeds being 15 mu m) are continuously added into a reaction kettle I, 4.46kg of NMP (45.0 mol) is added, the reaction kettle is sealed, the temperature is raised to 215 ℃ at 0.3 ℃/min, the temperature is kept for 3 hours, the temperature is raised to 230 ℃ at 0.15 ℃/min, the temperature is kept for 1 hour, the temperature is raised to 260 ℃ at 0.5 ℃/min, the temperature is kept for 1 hour, then slow discharge is carried out in an atmospheric flash evaporator, and the flash evaporation treatment is carried out, and the crude product 21.2kg is obtained after the solvent is further removed through drying. The crude product is washed and filtered twice under normal pressure, the water adding amount is 60.0kg each time, 60.0kg of water is added, the temperature is raised to 180 ℃ after pulping, the temperature is kept for 0.5h, the temperature is reduced, the filtration is carried out, and the final product is obtained after the filtration and the drying.

In the comparative example, 180.0kg of water is used for the washing process, 182.0kg of salt-containing wastewater is generated, and the salt content is 6.0wt%; the final product prepared was 10.0kg, weight average molecular weight 21500 and ash content 0.50wt%.

Comparison of example 1 with comparative example 2 shows that removal of salt particles by filtration can significantly reduce the amount of wash water and the amount of waste salt water produced.

Example 2

19.82kg (200.0 mol) of N-methylpyrrolidone, 8.33kg (NaOH mol number is 100.0 mol) of 48wt% sodium hydroxide aqueous solution, 11.93kg (NaHS mol number is 100.0 mol) of 47wt% sodium hydrosulfide aqueous solution, and under the protection of nitrogen, slowly heating to 200 ℃ and continuously removing water in the process, discharging 10.00kg (water content 98.0%) of distillate, dehydrating, and cooling to 170 ℃. At this time, the amount of sulfur in the system was 98.0mol and the water content was 147.0mol. The dehydrated liquid prepared in batches according to the method is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

A continuous process flow diagram of PAS shown in fig. 1 was used. P-dichlorobenzene 1 (PDCB), a dehydration liquid 2 (sodium sulfide content 25.48 wt%) and an NMP slurry 3 containing sodium chloride seeds (D50 of sodium chloride seeds is 15 μm, concentration of slurry is 0.136 mol/L) preheated to 200 ℃ are pumped to a first polymerization reactor 4 (PDCB/s=1.050 (molar ratio, same below) controlled in the first polymerization reactor 4, NMP/s=3.50 (molar ratio, same below), the temperature of the first polymerization reactor 4 is 215 ℃ and the residence time is 5 hours, the obtained first reaction mixture 5 is heated to 225 ℃ by a first heat exchanger 6, enters a second tubular reactor 7 (reactor temperature is 225 ℃) and is fed to a closed pressure filter 8 after the residence time is 1 hour, salt particles are removed by the closed pressure filter 8 (pore diameter of 80 μm), and D50 of the filtered salt particles 9 is 140 μm; the obtained second reaction mixed solution 10 is heated to 265 ℃ by a second heat exchanger 11, then enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃), the residence time is 1.5 hours, then enters an atmospheric flash evaporator 14 for flash evaporation after heat exchange to 280 ℃ by a heat exchanger (not shown), the flash evaporation is assisted by using 265 ℃ superheated steam 13, the flow of the superheated steam 13 is 5.0kg/h, vaporized solvent and steam 19 are condensed and liquefied by a condenser 20, waste liquid 21 to be recovered is generated, the undried crude product 15 firstly enters a continuous dryer 17 by a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle for pulping, the water adding amount is 60 kg/h, the heat exchange is carried out to 180 ℃, and (3) entering a tubular scrubber with a heat conducting oil jacket, controlling the temperature of the tubular scrubber to be 180 ℃ and the residence time to be 10min, then cooling to 60 ℃ through heat exchange, continuously filtering and drying (not shown), and obtaining the final product.

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

Example 3

The dehydration solution was prepared batchwise as in example 2, and placed in a storage tank and incubated at 170℃for subsequent continuous feeding. The PAS continuous production apparatus shown in FIG. 1 was used. P-dichlorobenzene 1 (PDCB), a dehydration liquid 2 (sodium sulfide content 25.48 wt%) and an NMP crystal slurry 3 containing sodium chloride seeds (D50 of sodium chloride seeds is 15 μm, crystal slurry concentration is 0.4 mol/L) preheated to 200 ℃ are pumped to a first polymerization reactor 4 (PDCB/s=1.050, NMP/s=2.5) sequentially at a flow rate of 25.21g/min, 50g/min and 8.1g/min, the temperature of the first polymerization reactor 4 is 215 ℃ and the residence time is 3.5 hours, the obtained first reaction mixture 5 is heated to 230 ℃ by a first heat exchanger 6, enters a second tubular reactor 7 (reactor temperature is 230 ℃) and is fed into a press filter 8 after the residence time of 1 hour, salt particles are removed by the press filter 8 (filter cloth aperture is 80 μm), and the D50 of the filtered salt particles 9 is 130 μm; the obtained second reaction mixed solution 10 is heated to 265 ℃ by a second heat exchanger 11, enters a third tubular reactor 12 (the reactor temperature is 265 ℃) and stays for 1h, then enters an atmospheric flash evaporator 14 for flash evaporation after heat exchange to 280 ℃ by a heat exchanger (not shown), and is subjected to auxiliary flash evaporation by using 265 ℃ superheated steam 13 in the flash evaporation, wherein the flow rate of the superheated steam 13 is 3.5kg/h, and vaporized solvent and steam 19 are condensed and liquefied by a condenser 20 to generate waste liquid 21 to be recovered. The undried crude product 15 firstly enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water adding amount is 6.0kg/h, the raw product enters a tubular scrubber with a heat conducting oil jacket after heat exchange to 180 ℃, the temperature of the tubular scrubber is controlled to be 180 ℃ and the residence time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and the finished product is obtained after continuous filtration and drying (not shown).

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

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

However, decreasing the molar ratio of NMP/S means increasing the concentration of PPS and sodium chloride in the system, and as the concentration of PPS and sodium chloride in the system and the molecular weight of PPS increase, the viscosity of the reaction solution rapidly increases, and it is difficult to smoothly perform the production under low NMP/S by the production process of the prior art, either the batch type or the continuous type. This precisely illustrates the strong adaptability of the production process disclosed in the present invention, which is particularly suitable for production conditions of low NMP/S molar ratio, thus being more advantageous for the efficient preparation of PAS of high molecular weight.

Example 4

In a 100L reactor, 19.82kg (200.0 mol) of N-methylpyrrolidone, 8.33kg (the molar number of sodium hydroxide is 100.0 mol) of 48wt% aqueous sodium hydroxide solution, 11.93kg (the molar number of sodium bisulfide is 100.0 mol) of 47wt% aqueous sodium bisulfide solution and 1.6kg of anhydrous sodium acetate were added, and then, the mixture was dehydrated by continuously heating under nitrogen protection to remove 10.00kg of aqueous solution (the water content is 98.0 wt%), and after the dehydration was completed, the mixture was cooled to 170 ℃. At this time, the amount of sulfur in the dehydrated liquid system was 98.0mol, and the water content was 147.0mol. The dehydrated liquid prepared in batches according to the method is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

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

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

Example 5

19.82kg (200.0 mol) of N-methylpyrrolidone, 8.33kg (NaOH mol number is 100.0 mol) of 48wt% sodium hydroxide aqueous solution, 11.93kg (NaHS mol number is 100.0 mol) of 47wt% sodium hydrosulfide aqueous solution, and the mixture is slowly heated to 195 ℃ under the protection of nitrogen, water is continuously removed in the process, 9.07kg (water content is 98.0%) of distillate is discharged, and the mixture is cooled to 170 ℃ after dehydration. At this time, the amount of sulfur in the system was 98.0mol and the water content was 196.0mol. The dehydrated liquid prepared in batches according to the method is put into a storage kettle at 170 ℃ for heat preservation and storage for subsequent continuous feeding.

A continuous process flow diagram of PAS shown in fig. 1 was used. P-dichlorobenzene 1 (PDCB) preheated to 200 ℃, dehydration liquid 2 (sodium sulfide content 24.66 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (D50 of sodium chloride crystal seeds is 15 mu m, concentration of crystal slurry is 0.136 mol/L) are pumped to a first polymerization reaction kettle 4 (PDCB/S=1.050, NMP/S=3.50) sequentially at a flow rate of 25.19g/min, 51.66g/min and 24.5g/min, the temperature of the first polymerization reaction kettle 4 is 215 ℃ and the retention time is 6h, the obtained first reaction mixture 5 is heated to 225 ℃ through a first heat exchanger 6, enters a second tubular reactor 7 (reactor temperature is 225 ℃) and is fed into a closed pressure filter 8 after the retention time is 1.5h, salt particles are removed through the closed pressure filter 8 (the aperture of 80 mu m), and the D50 of the filtered salt particles 9 is 150 mu m; the obtained second reaction mixture 10 is heated to 265 ℃ by a second heat exchanger 11, enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃) and has a residence time of 1.5 hours, then enters an atmospheric flash evaporator 14 for flash evaporation after heat exchange to 280 ℃ by a heat exchanger (not shown), the flash evaporation is assisted by using 265 ℃ superheated steam 13 during the flash evaporation, the flow of the superheated steam 13 is 5.0kg/h, and vaporized solvent and steam 19 are condensed and liquefied by a condenser 20 to generate waste liquid 21 to be recovered. The undried crude product 15 firstly enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water adding amount is 6.0kg/h, the raw product enters a tubular scrubber with a heat conducting oil jacket after heat exchange to 180 ℃, the temperature of the tubular scrubber is controlled to be 180 ℃ and the residence time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and the finished product is obtained after continuous filtration and drying (not shown).

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

Example 6

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

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

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

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

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

Comparative example 3

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

A continuous process flow diagram of PAS shown in fig. 1 was used. P-dichlorobenzene 1 (PDCB) preheated to 200 ℃, dehydration liquid 2 (sodium sulfide content 25.86 wt%) and NMP crystal slurry 3 containing sodium chloride crystal seeds (D50 of sodium chloride crystal seeds is 15 μm, concentration of crystal slurry is 0.136 mol/L) are pumped to a first polymerization reactor 4 (PDCB/S=1.050, NMP/S=3.50) sequentially at a flow rate of 25.21g/min, 49.27g/min and 24.5g/min, the temperature of the first polymerization reactor 4 is 215 ℃ and the retention time is 5h, the obtained first reaction mixture 5 is heated to 225 ℃ by a first heat exchanger 6, enters a second tubular reactor 7 (reactor temperature is 225 ℃) and is fed into a press filter 8 after the retention time of 1h, salt particles are removed by the press filter 8 (filter cloth aperture is 30 μm), and D50 of filtered salt particles 9 is 60 μm; the obtained second reaction mixture 10 is heated to 265 ℃ by a second heat exchanger 11, enters a third tubular reactor 12 (the temperature of the reactor is 265 ℃) and has a residence time of 1.5 hours, then enters an atmospheric flash evaporator 14 for flash evaporation after heat exchange to 280 ℃ by a heat exchanger (not shown), the flash evaporation is assisted by using 265 ℃ superheated steam 13 during the flash evaporation, the flow of the superheated steam 13 is 5.0kg/h, and vaporized solvent and steam 19 are condensed and liquefied by a condenser 20 to generate waste liquid 21 to be recovered. The undried crude product 15 firstly enters a continuous dryer 17 through a dryer feeding bin 16, the crude product 18 obtained after further drying is continuously discharged into a pulping kettle, water is added for pulping, the water adding amount is 6.0kg/h, the raw product enters a tubular scrubber with a heat conducting oil jacket after heat exchange to 180 ℃, the temperature of the tubular scrubber is controlled to be 180 ℃ and the residence time is 10min, then the temperature is reduced to 60 ℃ through heat exchange, and the finished product is obtained after continuous filtration and drying (not shown).

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

In the comparative example, the salt particle diameter and the pore diameter of the filter cloth matched with the salt particle diameter are smaller, so that the resistance of the filter is larger, the filter is blocked occasionally, and the stability of continuous operation is poor.

Claims

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

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

The prepolymerization reactor is a two-stage reactor connected in series;

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

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

after the monomer conversion rate in the first-stage reactor reaches 60-80%, injecting the reaction liquid into the second-stage reactor;

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

2. The continuous production method of polyarylene sulfide according to claim 1, wherein in 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, and the concentration is 30-70wt%;

the dichloro aromatic compound is selected from one or more of p-dichlorobenzene, dichloro naphthalene, dichloro fluorene and dichloro carbazole;

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

The auxiliary agent is selected from one or more of sodium acetate, sodium benzoate and sodium fatty acid of C5-C6;

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

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

3. The continuous production method of polyarylene sulfide according to claim 1, wherein in step (2):

after the dehydration liquid, the dichloro aromatic compound and the organic amide crystal slurry containing sodium chloride seed crystal are mixed, the adding amount of the sodium chloride seed crystal is 0.05-0.4 mol percent based on 1mol of sulfur source in the system, and the total amount of the organic amide solvent in the system is 2.0-4.0 mol percent.

4. The continuous production method of polyarylene sulfide according to claim 3, wherein:

the molar ratio of the dichloro aromatic compound to the sulfur source in the dehydration liquid is 1.0-1.025: 1, a step of;

After the dehydration solution, the dichloro aromatic compound and the organic amide crystal slurry containing sodium chloride crystal seeds are mixed, the total amount of the organic amide solvent is 2.0-2.5 mol based on 1 mol of sulfur source in the system.

5. The continuous production method of polyarylene sulfide according to claim 1, wherein in the step (3),

the polymerization reactor is selected from a tubular reactor;

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

6. A batch process for producing polyarylene sulfide, comprising:

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

7. The batch production method of polyarylene sulfide according to claim 6, wherein in step (a):

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

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

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

8. The batch production method of polyarylene sulfide according to claim 6, wherein in step (b):

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

the molar ratio of the sulfur source in the dehydration liquid obtained after dehydration saponification reaction to the dichloro aromatic compound is 1.0-1.1: 1, a step of;

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

in step (c):

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

9. A polyarylene sulfide prepared by the continuous production method according to any one of claims 1 to 5 or the batch production method according to any one of claims 6 to 8.

10. Use of the polyarylene sulfide according to claim 9 for preparing a crosslinked polyarylene sulfide, wherein the polyarylene sulfide is subjected to a thermal oxidation treatment.