阅读说明：本技术 一种生产聚苯醚的工艺方法 (Technological method for producing polyphenyl ether ) 是由 李同军 张军利 冯海涛 李湛 于 2020-05-06 设计创作，主要内容包括：本发明公开了一种生产聚苯醚的工艺方法,涉及化工技术领域。生产聚苯醚的工艺方法包括聚合氧化反应阶段和嵌入反应阶段,聚合氧化反应阶段是在含氧气氛下,将单体、溶剂和催化剂混合反应,嵌入反应阶段是将聚合氧化反应后的产物与螯合剂混合反应,终止聚合氧化反应,并将聚合氧化反应过程生成的副产物重新嵌入到目标产物上,显著减少了低聚物的产生,并将反应后的产品进行三相分离后得到油相通过蒸发浓缩、加水调浓、晶体析出和粒径调整,明显减少了小粒径产品的含量,改善了产品的质量。(The invention discloses a process method for producing polyphenyl ether, and relates to the technical field of chemical industry. The technological process for producing polyphenyl ether includes polymerization oxidation reaction stage and embedding reaction stage, the polymerization oxidation reaction stage is that under the oxygen-containing atmosphere, monomer, solvent and catalyst are mixed and reacted, the embedding reaction stage is that the product after polymerization oxidation reaction is mixed and reacted with chelating agent, the polymerization oxidation reaction is stopped, the by-product produced in the polymerization oxidation reaction process is embedded into the target product again, the production of oligomer is reduced obviously, and the three-phase separation of the reacted product is carried out to obtain oil phase, and the oil phase is evaporated, concentrated by adding water, precipitated by crystal and adjusted in grain size, so that the content of small-grain-size product is reduced obviously, and the quality of the product is improved.)

1. A process for producing polyphenylene ether, characterized by comprising a polymerization oxidation reaction stage wherein a monomer, a solvent and a catalyst are mixed and reacted in an oxygen-containing atmosphere, and an intercalation reaction stage wherein a product of the polymerization oxidation reaction is mixed and reacted with a chelating agent.

2. The process for producing a polyphenylene ether according to claim 1, wherein said chelating agent is selected from any one of nitrilotriacetic acid and salts thereof, tartaric acid and salts thereof, nitrilotriacetic acid and salts thereof, glycine and salts thereof, aminocarboxylic acid and salts thereof, ethylenediaminetetraacetic acid and salts thereof, diethylenetriaminepentaacetic acid and salts thereof, and citric acid and salts thereof; preferably at least one of trisodium nitrilotriacetate and citric acid; more preferably trisodium nitrilotriacetate;

preferably, the reaction time of the intercalation reaction stage is 80-160min, and the reaction temperature is 40-80 ℃;

preferably, a pipeline mixer is arranged on a feeding pipeline of the reactor embedded in the reaction stage to mix the products after the polymerization oxidation reaction, the chelating agent and the external circulation materials, the feeding pipeline extends to the bottom of the reactor, and a turbine-propelled stirrer is used for mixing.

3. The process for producing a polyphenylene ether according to claim 1, wherein the reaction temperature in the step of the polymerization oxidation reaction is 10 to 80 ℃, preferably 20 to 70 ℃, more preferably 30 to 60 ℃;

preferably, the monomer is 2, 6-dimethylphenol;

preferably, the reaction time of the polymerization oxidation reaction stage is 100-140 min;

preferably, the solvent is selected from at least one of benzene, toluene, xylene, ethylbenzene, pyridine and chloroform, more preferably toluene;

preferably, nitrogen is introduced into the reactor in the polymerization oxidation reaction stage for pressure sealing, and the gauge pressure in the reactor is controlled to be 35-65 KPa;

preferably, a cooler and a heater are arranged on the circulating pipeline corresponding to the reactor of the polymerization oxidation reaction stage to control the temperature of the circulating material.

4. The process for producing a polyphenylene ether according to claim 3, wherein said catalyst comprises a first catalyst, a second catalyst and a third catalyst, and said first catalyst is an amine complex;

the second catalyst is selected from at least one of N, N-dimethylbutylamine, 1, 3-dimethylbutylamine, dimethylethylenediamine, di-tert-butylethylenediamine, dimethylpropylenediamine, ethylbutylenediamine, dimethylbutylenediamine, trimethylpentyldiamine, morpholino compounds, pyridine, gamma-methylpyridine and ethylmethylamine, and is preferably N, N-dimethylbutylamine;

the third catalyst is selected from at least one of tetrabutylammonium bromide, methyltrialkyl ammonium chloride, methyltrialkyl ammonium bromide, dodecyl trimethyl ammonium chloride, benzyl triethyl ammonium chloride and tetradecyl trimethyl ammonium chloride, and is preferably methyltrialkyl ammonium chloride;

preferably, the first catalyst is obtained by reacting a first component, a second component and a third component, wherein the first component is at least one selected from cuprous oxide, cupric oxide, manganese dioxide, manganese oxide, cobalt oxide and cobalt oxide; the second component is selected from an aqueous solution prepared from at least one of hydrogen iodide, sodium iodide, potassium iodide, hydrogen bromide, sodium bromide, potassium bromide, tetramethylammonium bromide, hydrogen chloride, hydrogen fluoride, sulfuric acid, acetic acid, lauric acid and benzoic acid; the third component is selected from at least one of dimethyl ethylenediamine, di-tert-butyl ethylenediamine, dimethyl propylenediamine, ethylbutylenediamine, dimethyl butylenediamine and trimethyl pentylenediamine;

more preferably, the first component is selected from at least one of cuprous oxide and cupric oxide; more preferably, the second component is selected from an aqueous solution made of at least one of hydrogen chloride, hydrogen bromide and hydrogen iodide; more preferably, the third component is at least one of di-tert-butylethylenediamine and dimethylethylenediamine;

more preferably, the first catalyst is prepared prior to introduction into the reactor of the polymeric oxidation reaction stage; more preferably, the reaction temperature for preparing the first catalyst is 15-80 ℃, and the reaction time is 5-60 min; further preferably, the reaction temperature for preparing the first catalyst is 20-60 ℃, and the reaction time is 10-50 min;

preferably, the mass ratio of the first component to the second component to the third component is 1:2-10:1-6, more preferably 1:3-8:2-5, and even more preferably 1:5-7: 3-4;

preferably, the molar ratio of metal ion to chelating agent in the first component is 1:2-7, more preferably 1:3-6, and even more preferably 1: 4-5.

5. The process for producing a polyphenylene ether according to claim 4, wherein the mass ratio of the monomer to the first catalyst is 1000:5 to 25, preferably 1000:10 to 20, more preferably 1000:14 to 16;

preferably, the mass ratio of the monomer to the second catalyst is 1000:9-72, more preferably 1000: 18-57;

preferably, the mass ratio of the monomer to the third catalyst is 1000:0.4-3.7, more preferably 1000: 0.6-2.6;

preferably, the mass ratio of the monomer to the solvent is 1:1 to 7, more preferably 1:3 to 5;

preferably, the molar ratio of the monomer to the oxygen is 16: 1-4.

6. The process for producing a polyphenylene ether according to any one of claims 1 to 5, further comprising separating the product obtained in the step of the intercalation reaction, and then subjecting the separated polyphenylene ether phase to crystallization;

preferably, the method further comprises the steps of filtering, purifying, drying and cooling the product obtained in the crystallization process.

7. The process for producing a polyphenylene ether according to claim 6, wherein said separation of the product obtained in said stage of the intercalation reaction is carried out by separating the product into a solid phase, an aqueous phase and an oil phase;

preferably, the aqueous phase is subjected to extraction and standing to form an extraction aqueous phase and an organic phase, and after separation, the extraction aqueous phase is subjected to electrolysis to recover the metal simple substance from the catalyst;

preferably, the extractant used in the extraction process is selected from at least one of benzene, toluene, xylene, ethylbenzene, pyridine and chloroform, preferably toluene;

preferably, the mass ratio of the extractant to the aqueous phase in the extraction process is 1:0.3-0.8, more preferably 1: 0.4-0.7.

8. The process for producing a polyphenylene ether according to claim 7, wherein said crystallization is carried out by subjecting said oil phase to concentration treatment, concentration adjustment treatment and precipitation treatment in this order;

preferably, the concentration treatment is to concentrate the oil phase to 30-35%; more preferably, the absolute pressure of a gas phase is kept to be 15-65 KPa in the concentration treatment process, and the temperature of a liquid phase is kept to be 68-102 ℃;

preferably, the concentration adjusting treatment is to mix the concentrated material with water, and the mass ratio of the water to the concentrated material is 1:30-80, more preferably 1:39-70, and further preferably 1: 48-60;

preferably, the precipitation treatment is to mix the material after the concentration treatment with an anti-solvent, and the mass ratio of the material after the concentration treatment to the anti-solvent is 1:1-2, more preferably 1:1.2-1.8, and further preferably 1: 1.4-1.6;

preferably, the anti-solvent is selected from at least one of ketones, alcohols and water, more preferably an alcohol of 1-6 carbon atoms, and further preferably at least one of methanol and ethanol;

preferably, the crystallization process is further performed with a grain conditioning treatment after the precipitation treatment; more preferably, the stirring speed in the grain-conditioning treatment process is 50-220 r/min, and the stirring time is 8-20 min.

9. The process for producing a polyphenylene ether according to claim 8, wherein the resultant material is subjected to filtration purification, drying and cooling after the precipitation treatment;

preferably, the drying temperature is 130-160 ℃, and the drying time is 20-40 min;

preferably, high-temperature moisture-carrying nitrogen is introduced as a heat medium in the drying process, and stirring is carried out in the drying process;

more preferably, nitrogen output from the dryer is introduced into an ejector with negative pressure and then enters a cooler for cooling; further preferably, the material cooled by the cooler is subjected to gas-liquid separation, a gas phase obtained by the gas-liquid separation is introduced into the dryer, and a liquid phase obtained by the gas-liquid separation is used for an ejector to generate negative pressure or is introduced into a filtering stage for recycling;

further preferably, the absolute pressure of the injector is 76 to 86 KPa.

10. The process for producing a polyphenylene ether according to claim 6, further comprising separating and recovering a solvent mixture produced in said crystallization process and said filtration and purification process;

preferably, the solvent mixture is separated into a heavy phase solvent and a light phase solvent by utilizing the density difference of the solvent mixture, the light phase solvent is rectified and purified, and the heavy phase solvent is sequentially rectified under high pressure and low pressure;

preferably, the separation of the solvent mixture is carried out at a temperature of from 10 to 50 ℃, more preferably from 20 to 40 ℃;

preferably, the residence time for separating the solvent mixture is from 90 to 180min, more preferably from 120-150 min;

preferably, the light phase solvent is rectified and purified and then is recycled as a solvent in a polymerization oxidation reaction stage, and the light phase solvent is rectified and purified and then is extruded to form an oligomer obtained at the bottom of the tower;

more preferably, in the process of rectifying the light phase solvent, the temperature at the top of the tower is 95-112 ℃, the temperature at the bottom of the tower is 140-160 ℃, the operation gauge pressure is 0.1-0.12MPa, and the reflux ratio is 1.5-2.1;

preferably, the tower top outlet material obtained in the high-pressure rectification process is used as a heat source in the low-pressure rectification process, and is mixed with the tower top outlet material in the low-pressure rectification process and introduced into a filtration and purification stage for recycling;

more preferably, in the high-pressure rectification process, the tower top temperature is 110-;

more preferably, in the low-pressure rectification process, the tower top temperature is 66-75 ℃, the tower bottom temperature is 103-113 ℃, the operation gauge pressure is 0.01-0.02MPa, and the reflux ratio is 2.3-3.6.

Technical Field

The invention relates to the technical field of chemical industry, and in particular relates to a process method for producing polyphenyl ether.

Background

The polyphenylene oxide is PPO or PPE for short, is one of five major engineering plastics, has already been industrially produced, and the main process method is formed in the eight and ninety years of the twentieth century, and the adopted production equipment is relatively old. With the development of science and technology and the progress of society, new application of polyphenyl ether is continuously discovered, the new application puts higher requirements on the quality of polyphenyl ether, the society is increasingly strict on safety and environmental protection, and production enterprises pay more attention to energy conservation, consumption reduction and cost reduction.

The prior production method of polyphenylene ether mainly has the problems of low yield of polyphenylene ether and large generation amount of oligomer. In addition, the production method of polyphenylene oxide has the problems of low operation efficiency, small product particle size, high wastewater treatment cost, potential safety hazard of deflagration of reaction waste gas and the like.

Disclosure of Invention

The invention aims to provide a process method for producing polyphenylene oxide, which aims to improve the yield of the polyphenylene oxide, reduce the generation amount of oligomer and obviously reduce the content of small-particle-size products.

The technical problem to be solved by the invention is realized by adopting the following technical scheme.

The invention provides a process method for producing polyphenyl ether, which comprises a polymerization oxidation reaction stage and an embedding reaction stage, wherein in the polymerization oxidation reaction stage, a monomer, a solvent and a catalyst are mixed and reacted under an oxygen-containing atmosphere, and in the embedding reaction stage, a product obtained after the polymerization oxidation reaction is mixed and reacted with a chelating agent.

The embodiment of the invention provides a process method for producing polyphenyl ether, which has the beneficial effects that: the method comprises the steps of carrying out polymerization oxidation reaction on a monomer, a solvent and a catalyst in the presence of oxygen, mixing a product obtained after the reaction in the polymerization oxidation reaction stage with a chelating agent for reaction, terminating the polymerization oxidation reaction, and re-embedding a byproduct generated in the polymerization oxidation reaction process into a target product, so that the generation of oligomers is remarkably reduced, and the quality of the product is improved.

Drawings

In order to more clearly illustrate the technical solutions of the embodiments of the present invention, the drawings needed to be used in the embodiments will be briefly described below, it should be understood that the following drawings only illustrate some embodiments of the present invention and therefore should not be considered as limiting the scope, and for those skilled in the art, other related drawings can be obtained according to the drawings without inventive efforts.

FIG. 1 is a schematic view of an apparatus used in a process for producing a polyphenylene ether according to an embodiment of the present invention.

Icon: 1-catalyst batching tank; 2-a polymeric oxidation reactor; 3-a reaction heater; 4-a reaction cooler; 5-insertion into a reactor; 6-a heater; 7-centrifuge feed tank; 8-a three-phase centrifuge; 9-extraction tank; 10-an electrolytic cell; 11-a concentration tank; 12-a heater; 13-a thickener; 14-a separating tank; 15-a granulation tank; 16-a filter; 17-a screw conveyor; 18-a dryer; 19-an ejector; 20-a tail gas cooler; 21-a gas-liquid separation tank; 22-a screw cooling conveyor; 23-product storage; 24-a solvent knockout drum; 241-a filler buffer section; 242-a material separating partition plate; 25-an anti-solvent rectification column; 251-an anti-solvent high-pressure rectifying tower; 252-an anti-solvent low pressure rectification column; 26-a solvent rectification column; 27-oligomer extruder.

Detailed Description

In order to make the objects, technical solutions and advantages of the embodiments of the present invention clearer, the technical solutions in the embodiments of the present invention will be clearly and completely described below. The examples, in which specific conditions are not specified, were conducted under conventional conditions or conditions recommended by the manufacturer. The reagents or instruments used are not indicated by the manufacturer, and are all conventional products available commercially.

The following will specifically explain a process for producing polyphenylene ether provided in the examples of the present invention.

The embodiment of the invention provides a process method for producing polyphenylene ether, and with reference to fig. 1, the process method specifically comprises the steps of preparing polyphenylene ether through a polymerization oxidation reaction stage and a chelation reaction stage, then sequentially carrying out product separation, crystallization, filtration and purification, drying and cooling, and in some embodiments, recycling a solvent and an oligomer in a synthesis process. The method specifically comprises the following steps:

(1) stage of polymerization oxidation reaction

The polymerization oxidation reaction stage is to mix and react monomers, a solvent and a catalyst in an oxygen-containing atmosphere, wherein the reaction temperature is 10-80 ℃, preferably 20-70 ℃, and more preferably 30-60 ℃; the reaction time is 100-140 min. The reaction temperature is low, the speed is slow, and the efficiency is low; the reaction temperature is high, the speed is high, the control is not easy, and the over-reaction is generated. The reaction is carried out in a temperature range, two stages of temperature rise and constant temperature are provided, the temperature is different, the reaction time is correspondingly different, the condition for judging the termination of the reaction is the viscosity of the reaction liquid, and the reaction time and the temperature are approximate parameter references.

Specifically, 2, 6-dimethylphenol is generally used as a monomer; the solvent is at least one selected from the group consisting of benzene, toluene, xylene, ethylbenzene, pyridine and chloroform, and toluene is more preferred.

Specifically, the catalyst comprises a first catalyst, a second catalyst and a third catalyst, wherein the first catalyst is an amine complex (such as a copper amine complex); the second catalyst is at least one selected from N, N-dimethylbutylamine, 1, 3-dimethylbutylamine, dimethylethylenediamine, di-tert-butylethylenediamine, dimethylpropylenediamine, ethylbutylenediamine, dimethylbutylenediamine, trimethylpentyldiamine, morpholino compounds, pyridine, gamma-picoline and ethylmethylamine, and is preferably N, N-dimethylbutylamine. The third catalyst is at least one selected from tetrabutylammonium bromide, methyltrialkylammonium chloride, methyltrialkylammonium bromide, dodecyltrimethylammonium chloride, benzyltriethylammonium chloride and tetradecyltrimethylammonium chloride, and is preferably methyltrialkylammonium chloride.

Preferably, the first catalyst is obtained by reacting a first component, a second component and a third component, wherein the first component is selected from at least one of cuprous oxide, cupric oxide, manganese dioxide, manganese oxide, cobalt oxide and cobalt oxide; the second component is selected from aqueous solution prepared from at least one of hydrogen iodide, sodium iodide, potassium iodide, hydrogen bromide, sodium bromide, potassium bromide, tetramethylammonium bromide, hydrogen chloride, hydrogen fluoride, sulfuric acid, acetic acid, lauric acid and benzoic acid; the third component is at least one selected from dimethyl ethylenediamine, di-tert-butyl ethylenediamine, dimethyl propylenediamine, ethyl butylenediamine, dimethyl butylenediamine and trimethyl pentylene diamine; more preferably, the first component is selected from at least one of cuprous oxide and cupric oxide; the second component is selected from an aqueous solution prepared from at least one of hydrogen chloride, hydrogen bromide and hydrogen iodide; the third component is at least one of di-tert-butyl ethylenediamine and dimethylethylenediamine.

In order to further improve the utilization rate of the raw materials, the use amounts of the raw materials such as monomers, catalysts, solvents, oxygen and the like are further optimized. The mass ratio of the monomer to the first catalyst is 1000:5-25, preferably 1000:10-20, more preferably 1000: 14-16; the mass ratio of the monomer to the second catalyst is 1000:9-72, preferably 1000: 18-57; the mass ratio of the monomer to the third catalyst is 1000:0.4-3.7, preferably 1000: 0.6-2.6; the mass ratio of the monomer to the solvent is 1:1-7, more preferably 1: 3-5; the molar ratio of the monomer to the oxygen is 16: 1-4. The solvent proportion is high, the concentration of the product after the reaction is low, and the disadvantages (long time and high energy consumption) are brought to the concentration; the solvent proportion is low, the concentration of the product after reaction is high, the fluidity is poor, and pipelines and equipment are seriously blocked.

Preferably, the mass ratio of the first component to the second component to the third component is 1:2-10: 1-6; more preferably 1:3 to 8:2 to 5, still more preferably 1:5 to 7:3 to 4. During the preparation process of the first catalyst, the dosage of the first component, the second component and the third component is controlled to further improve the catalytic activity of the first catalyst on the polymerization oxidation reaction.

In a preferred embodiment, the first catalyst is prepared prior to introduction into the reactor of the polymerization oxidation stage; the reaction temperature for preparing the first catalyst is 15-80 ℃, and the reaction time is 5-60 min; preferably, the reaction temperature for preparing the first catalyst is 20-60 ℃ and the reaction time is 10-50 min. Compared with the mode of adding all the raw materials into the polymerization oxidation reaction stage at one time, the mode of firstly preparing the first catalyst and then adding the first catalyst into the polymerization oxidation reaction stage can further shorten the reaction time, and the inventor makes further tests to find that: the reaction time can be shortened by about 6min, and the efficiency can be improved by more than 3.3%.

In some preferred embodiments, nitrogen is introduced into the reactor in the polymerization oxidation reaction stage to seal pressure, and the gauge pressure in the reactor is controlled to be 35-65 KPa. The polymerization oxidation reaction is carried out under the nitrogen-sealed micro-positive pressure, the pressure of the gas phase at the upper part of the reactor is controlled, the overflow of oxygen from the reaction liquid can be effectively inhibited, the oxygen concentration in the reaction liquid is increased, the reaction efficiency is improved, meanwhile, the oxygen concentration in the exhaust gas can be reduced, and the operation safety is improved. The inventor makes further tests, and compared with the mode of not controlling the gas phase pressure at the upper part of the reactor, the operation mode of sealing with nitrogen and controlling the gas phase pressure at the upper part of the reactor can further shorten the reaction time by 6min, improve the reaction efficiency by more than 3.3 percent and reduce the oxygen content in the exhaust gas by about 2 percent.

Referring to fig. 1, a first catalyst is prepared in a catalyst dosing tank 1 and then enters a polymerization oxidation reactor 2, and in some preferred embodiments, a reaction heater 3 (connected to a high-temperature water pipeline) and a reaction cooler 4 (connected to a low-temperature water pipeline) are disposed on a circulation pipeline corresponding to the polymerization oxidation reactor 2 to control the temperature of the circulation material. Compare in traditional single heat exchanger, can prevent to make low temperature water system outwards drainage always because of steam condensate sneaks into in the low temperature water system.

(2) Stage of intercalation reaction

The embedding reaction stage is to mix and react the product after the polyreaction with a chelating agent, the reaction time of the embedding reaction stage is 80-160min, and the reaction temperature is 40-80 ℃. The improved reaction unit of the invention changes one-step reaction into two-step reaction, can accurately control the operation parameters of each stage, carries out the polymerization of the first-step monomer through the polymerization oxidation reactor, terminates the polymerization oxidation reaction through the intercalation reaction, and re-intercalates the by-product generated in the polymerization oxidation reaction process into the target product, thereby obviously reducing the generation of oligomer and improving the purity and quality of the product.

Specifically, the chelating agent is any one selected from nitrilotriacetic acid and salts thereof, tartaric acid and salts thereof, nitrilotriacetic acid and salts thereof, glycine and salts thereof, aminocarboxylic acid and salts thereof, ethylenediaminetetraacetic acid and salts thereof, diethylenetriaminepentaacetic acid and salts thereof, and citric acid and salts thereof; preferably at least one of trisodium nitrilotriacetate and citric acid; more preferably trisodium nitrilotriacetate. Preferably, the molar ratio of metal ion to chelating agent in the first component is from 1:2 to 7, more preferably from 1:3 to 6, even more preferably from 1:4 to 5.

Preferably, a pipeline mixer is arranged on a feeding pipeline of the reactor embedded in the reaction stage to mix the products after the polyreaction, the chelating agent and the external circulation materials, the feeding pipeline extends to the bottom of the reactor, and a turbine propelling type stirrer is used for mixing. The high viscosity reaction polymer and the chelating agent can be mixed uniformly by modifying the feed line inserted into the reactor 5. Specifically, the discharge port of the polymerization oxidation reactor 2 communicates with the feed port of the embedded reactor 5, and the embedded reactor 5 is temperature-controlled by the heater 6.

(3) Catalyst separation and recovery stage

The product obtained in the embedding reaction stage is separated into a solid phase, a water phase and an oil phase, specifically, the product obtained in the embedding reaction stage enters a three-phase centrifuge 8 for separation after passing through a centrifuge feeding tank 7, the three-phase centrifuge 8 rotates at a high speed, the solid phase, the water phase and the oil phase are discharged from different outlets according to different densities, the target product polyphenylene oxide is in the oil phase, and the waste catalyst is in the water phase. The three-phase centrifuge 8 has three outlets, namely an organic phase, a water phase and a solid phase, and the discharged solid phase is merged into the water phase and enters the extraction tank 9 together because the solid phase amount is very small.

Preferably, the extractant used in the extraction process is selected from at least one of benzene, toluene, xylene, ethylbenzene, pyridine and chloroform, preferably toluene; the mass ratio of the extractant to the water phase in the extraction process is 1:0.3-0.8, more preferably 1: 0.4-0.7.

Preferably, the aqueous phase is extracted and stood to form an extraction aqueous phase and an organic phase, and after separation, the extraction aqueous phase is passed into an electrolytic bath 10 for electrolysis to recover elemental metals from the catalyst, such as blister copper. The distance between the positive electrode and the negative electrode of the electrolytic cell 10 is 100-300 mm, and the copper-containing water phase treated by the electrolytic method has the following advantages: (a) compared with the precipitation by adding a precipitator, the method has the advantages that: the by-product copper has high purity (the electrolysis is more than 99 percent, the precipitation is less than 40 percent) and can be directly sold as a commodity. (b) The advantages over triple effect evaporation: the purity of the by-product copper is high (triple effect is less than 60%), the by-product copper can be directly sold as a commodity, and the energy consumption electrolysis is about 10% of triple effect evaporation.

(4) Crystal precipitation stage

The crystallization process is to carry out concentration treatment, concentration adjustment treatment and precipitation treatment on the oil phase in sequence; that is, the mixture sequentially passes through a concentration tank 11 (temperature control by a heater 12), a thickener 13, and a precipitation tank 14.

Specifically, the concentration treatment is to concentrate the concentration of the polyphenylene ether in the oil phase to 30-35%; the gas phase pressure (absolute pressure) is kept at 15-65 KPa in the concentration process, and the liquid phase temperature is 68-102 ℃. And controlling the liquid phase temperature in the concentration process to further improve the quality of the polyphenyl ether product.

Specifically, the concentration adjusting treatment is to mix the concentrated material with water, and the mass ratio of the water to the concentrated material is 1:30-80, more preferably 1:39-70, and further preferably 1: 48-60. In the embodiment, the crystal nucleus formation speed of the polyphenyl ether is controlled by adding water into the concentrated material, but the water and the concentrated material are not compatible, and the method is carried out in a specially designed thickener 13 with high-speed stirring. The concentrated material and pure water are added into a thickener 13 in proportion, and the material and the pure water are fully mixed through high-speed stirring and then sent into an eduction tank 14. The proportion of water is low, and the thickening effect cannot be achieved; the proportion is high, precipitation is generated, the precipitation of the polyphenyl ether in a precipitation tank is influenced, and crystals are not easy to grow.

Specifically, the precipitation treatment is to mix the material after the concentration treatment with an anti-solvent, and the mass ratio of the material after the concentration treatment to the anti-solvent is 1:1-2, preferably 1:1.2-1.8, and more preferably 1: 1.4-1.6. The anti-solvent is at least one selected from ketones, alcohols and water, preferably an alcohol having 1 to 6 carbon atoms, and more preferably at least one selected from methanol and ethanol.

In some preferred embodiments, the crystallization process is further followed by a conditioning treatment, which is carried out after the precipitation treatment, and then enters the conditioning tank 15 after the precipitation tank 14; the stirring speed in the process of the granulation treatment is 50-220 r/min, and the stirring time is 8-20 min. By controlling the rotation speed of the stirrer on the granulation tank 15 and the residence time of the material in the granulation tank 15, a small amount of polyphenyl ether still dissolved in the solvent is precipitated on the existing crystals, and the small-particle crystals are combined into large-particle crystals, so that the polyphenyl ether crystals are enlarged, and the residence time of the material in the granulation tank 15 is realized by switching outlets arranged at different heights. In the embodiment, the crystallization process is improved, so that the particle size of the polyphenylene ether can be effectively increased, and the proportion of the particle size smaller than 200 meshes is reduced from 33.6% to below 25%, and is reduced by more than 8%.

(5) Crystal filtration purification stage

The filtering and purifying are carried out by a filter 16, the feed inlet of the filter 16 is communicated with the discharge outlet of the grain mixing tank 15, and the liquid outlet of the filter 16 is communicated with the anti-solvent feed inlet on the precipitation tank 14.

In some embodiments, filter 16 is a rotary drum filter, the polyphenylene ether crystal suspension is fed under pressure to the rotary drum filter, filtration of the polyphenylene ether solid is effected in the rotary drum filter, the filtrate is sent to a solvent and oligomer recovery unit for treatment, the cake is washed with an anti-solvent a plurality of times, the solvent and a small amount of the catalyst remaining in the polyphenylene ether are displaced and washed with the anti-solvent to purify the polyphenylene ether, the mixed solvent containing the solvent and the anti-solvent (collectively referred to as solvent) is then sent to an extraction tank 14, and the cake is then sent to a dryer by a screw conveyor 17.

(6) Product drying and cooling stage

The material from the filter 16 is dried by a dryer 18 to produce a product which is transported by a screw cooled conveyor 22 to a product storage bin 23. Preferably, the drying temperature is 130-160 ℃, and the drying time is 20-40 min. The temperature is low, and the drying effect is poor; high temperature affects the product quality.

In some preferred embodiments, high temperature moisture-carrying nitrogen is introduced as a heat medium during the drying process, and stirring is performed during the drying process. The drying can be carried out by adopting a dividing wall type dryer with adjustable speed stirring, so as to solve the problem that the polyphenyl ether crystal is small and is not easy to dry.

In some preferred embodiments, the nitrogen output from the dryer 18 is passed to an ejector 19 with negative pressure and then cooled in a tail gas cooler 20; introducing the material cooled by the cooler into a gas-liquid separation tank 21 for gas-liquid separation, introducing a gas phase obtained by gas-liquid separation into a dryer 18, and using a liquid phase obtained by gas-liquid separation for an ejector 19 to generate negative pressure or introducing the liquid phase into a filtering stage for recycling; the pressure (absolute pressure) of the injector is 76-86 KPa. The use of the negative pressure (venturi principle) generated by the ejector 19 allows hot moisture to be drawn into the ejector 19, reducing the pressure in the dryer 18, and also facilitating the evaporation of the solvent in the polyphenylene ether. The ejector 19 has large vacuum degree and high dust content in the moisture-carrying gas, and cannot form micro negative pressure in the drying gas if the vacuum degree is too small.

It is emphasized that the ejector can form micro negative pressure in the drying gas, the micro negative pressure in the dryer is compared with the micro positive pressure, the volatile matter in the product can be reduced by 8 percent (from 0.5 percent to below 0.46 percent), and the product quality is improved.

(7) Solvent and oligomer recovery stage

In some preferred embodiments, the solvent mixture generated in the crystallization process and the filtration process is separated and recycled to further reduce the process cost. Specifically, the solvent mixture is separated into a heavy phase solvent and a light phase solvent by utilizing the density difference of the solvent mixture, the light phase solvent is rectified and purified, and the heavy phase solvent is sequentially rectified under high pressure and low pressure.

Referring to fig. 1, after the packing buffer 241 of the solvent separation tank 24 is utilized, the solvent mixture is separated into a heavy-phase solvent (located at a side close to the packing buffer 241) and a light-phase solvent (located at a side far from the packing buffer 241) by a separating partition 242. Then, the heavy phase solvent sequentially passes through an anti-solvent high-pressure rectifying tower 251 and an anti-solvent low-pressure rectifying tower 252 (collectively called an anti-solvent rectifying tower 25), the top gas of the anti-solvent high-pressure rectifying tower 251 is a heat source of the anti-solvent low-pressure rectifying tower 252, and the purified anti-solvent is used for filtering and purifying crystals; after the light phase solvent passes through the solvent rectifying tower 26, the tower bottom material is extruded and molded by an oligomer extruder 27, and the light phase solvent is rectified and purified to be recycled as the solvent in the polymerization oxidation reaction stage.

Preferably, the separation of the solvent mixture is carried out at a temperature of from 10 to 50 ℃ and more preferably from 20 to 40 ℃; the residence time for separating the solvent mixture is 90-180min, more preferably 120-150min, and the solvent and the anti-solvent are better separated by controlling the operating temperature and residence time.

In some embodiments, during the rectification of the light phase solvent, the temperature at the top of the tower is 95-112 ℃, the temperature at the bottom of the tower is 140-160 ℃, the operating pressure (gauge pressure) is 0.1-0.12MPa, and the reflux ratio is 1.5-2.1; the tower top outlet material obtained in the high-pressure rectification process is used as a heat source in the low-pressure rectification process, and is mixed with the tower top outlet material in the low-pressure rectification process and introduced into the filtration stage for recycling.

In some embodiments, during the high-pressure rectification, the temperature at the top of the tower is 110-; in the low-pressure rectification process, the tower top temperature is 66-75 ℃, the tower bottom temperature is 103-113 ℃, the operating pressure (gauge pressure) is 0.01-0.02MPa, and the reflux ratio is 2.3-3.6.

The features and properties of the present invention are described in further detail below with reference to examples.