A method for recovering organic solvents and alkaline earth metal chlorides from aramid production waste liquid

Abstract

The present invention relates to a method for recovering organic solvents and alkaline earth metal chloride salts in aramid production waste liquid, belongs to the technical field of waste liquid treatment / recovery process, and solves at least one of the problems that the existing treatment / recovery process will cause secondary pollution, high energy consumption, and inability to simultaneously recover alkaline earth metal chloride salts. The present invention provides a method for recovering organic solvents and alkaline earth metal chloride salts in aramid production waste liquid: the recovery method includes three process sections, which are flocculation precipitation process section, azeotropic distillation process section, and "distillation + evaporation crystallization" process section, and finally obtains pure high-concentration organic solvent and alkaline earth metal chloride salt crystals. The recovery method provided by the present invention can realize the efficient recovery of organic solvents and alkaline earth metal chloride salts in aramid production waste liquid. Taking the DMAC-CaCl2 waste liquid system as an example, the recovery rate of DMAC can reach 98.3%, the recovery rate of CaCl2 can reach 94.5%, and the purity of DMAC is ≧99.99%.

Description

Technical Field

[0001] The invention relates to the technical field of waste liquid treatment / recovery process, and in particular to a method for recovering organic solvents and alkaline earth metal chloride salts in aramid production waste liquid. Background Art

[0002] A large amount of spinning waste liquid is generated during the aramid production process, and the waste liquid contains a large amount of organic solvents and alkaline earth metal chloride salts.

[0003] Among them, N,N-dimethylacetamide (DMAC) and N-methylpyrrolidone (NMP) are the most common organic solvents in aramid production wastewater. The above solvents are used in large quantities in the aramid spinning industry and have high market prices. They are also toxic to a certain extent. If they are discharged into water bodies during production, use and disposal, they will have a great impact on the water quality environment, endanger people's health, and cause waste. Calcium chloride and lithium chloride are the most common alkaline earth metal chloride salts in aramid production wastewater. Direct discharge will pollute the environment and poison animals and plants. Recycling and utilization is beneficial to environmental protection and saving production costs.

[0004] At present, the main methods for recycling aramid production wastewater are extraction distillation and evaporation distillation. Solution extraction is a method of transferring the solute from a low-solubility solvent to another high-solubility solvent based on the principle of different solubility of the solute in immiscible solvents, and then distilling and separating. Tsinghua University and the National Key Laboratory once simulated chloroform as an extractant to treat DMAC wastewater. The disadvantage of this method is that the extractant is introduced, the extractant is highly toxic, and the extractant is prone to cause secondary pollution to the environment; at the same time, in the extraction recovery, no recovery method for alkaline earth metal chloride salts is provided. Evaporation distillation is a method of separating and recovering organic solvents and alkaline earth metal chloride salts by evaporating all the waste liquids and then distilling them. This treatment process consumes huge energy, has low separation efficiency, and is only suitable for the treatment of high-concentration wastewater, not for the treatment of common aramid production wastewater.

[0005] In summary, it is necessary to develop a method for recovering organic solvents and alkaline earth metal chlorides in aramid production wastewater. Summary of the invention

[0006] In view of the above analysis, an embodiment of the present invention aims to provide a method for recovering organic solvents and alkaline earth metal chlorides in aramid production waste liquid, so as to solve at least one of the problems of existing treatment / recovery processes causing secondary pollution, high energy consumption, and inability to simultaneously recover organic solvents and alkaline earth metal chlorides.

[0007] The invention discloses a method for recovering organic solvents and alkaline earth metal chloride salts in aramid production waste liquid. The recovery method comprises three process sections, namely, a flocculation precipitation process section, an azeotropic distillation process section, and a "distillation + evaporation crystallization" process section, to obtain pure high-concentration organic solvents and alkaline earth metal chloride salt crystals.

[0008] Specifically, the specific process of the flocculation and sedimentation process section is: through flocculation and filtration processes, oligomers and mechanical impurities in the waste liquid are separated to obtain filter residue and filtrate.

[0009] Specifically, the specific process of the azeotropic distillation process section is: adding an azeotropic agent to the filtrate, conveying it to a primary azeotropic distillation tower, removing water from the waste liquid through an azeotropic distillation process, and obtaining a tower top distillate D1 containing the azeotropic agent and water and a tower bottom mixed liquid F1 containing the azeotropic agent, an organic solvent, and an alkaline earth metal chloride salt.

[0010] Specifically, the specific process of the "distillation + evaporation crystallization" process section is: the entrainer, organic solvent and alkaline earth metal chloride in the bottom mixed liquid F1 are separated by the "distillation + evaporation crystallization" process to obtain a gas phase discharge near the bottom of the tower, and a high-purity organic solvent is obtained after condensation, and an alkaline earth metal chloride solution is concentrated to obtain an alkaline earth metal chloride product after crystallization.

[0011] Specifically, the specific composition of the aramid production waste liquid is: 20% to 40% by mass of organic solvent, 0.2% to 3% of alkaline earth metal chloride, 0.1% to 2% of oligomer, and the rest is water.

[0012] Specifically, the organic solvent is N,N-dimethylacetamide (DMAC) or N-methylpyrrolidone (NMP); and the alkaline earth metal chloride is lithium chloride or calcium chloride.

[0013] Furthermore, the azeotropic agent is one or more of isobutyl alcohol, n-butyl alcohol, isopentanol, n-pentanol, xylene and ethylene dichloride.

[0014] Furthermore, the mass ratio of the azeotropic agent to the filtrate is 0.1 to 2:1.

[0015] Preferably, a "condensation + oil-water separation" process section is provided after the azeotropic distillation process section to separate the entrainer and the water phase in the overhead distillate D1, and the entrainer is recycled; after the "condensation + oil-water separation" process section, a secondary or multi-stage azeotropic distillation tower is provided for further separation and recovery of the entrainer.

[0016] Furthermore, the "rectification + evaporation crystallization" process section (referred to as process section S3) includes the following specific steps:

[0017] S3-1: The bottom mixed liquid F1 is subjected to primary distillation to obtain a top distillate D3 containing an azeotropic agent and an organic solvent and a bottom mixed liquid F3 containing an organic solvent and an alkaline earth metal chloride salt:

[0018] S3-2: transporting the mixed liquid F3 from the bottom of the tower to an evaporator to obtain organic solvent vapor and concentrated alkaline earth metal chloride solution;

[0019] S3-3: performing secondary distillation on the tower top distillate D3 and the organic solvent vapor to obtain the tower top distillate D4, i.e., the high-concentration entrainer; obtaining the gas phase discharge near the tower bottom, and condensing it to obtain a high-purity organic solvent;

[0020] S3-4: The concentrated alkaline earth metal chloride solution is transported to a crystallization kettle, cooled and crystallized, and the alkaline earth metal chloride product is obtained after filtration.

[0021] Compared with the prior art, the present invention can achieve at least one of the following beneficial effects:

[0022] 1. The recovery method provided by the present invention has high recovery efficiency and good product purity; and can realize the simultaneous recovery of organic solvent and alkaline earth metal chloride.

[0023] The present invention first uses a flocculant to act on oligomers in the waste liquid, so that the oligomers flocculate and precipitate, and then filters and separates the oligomers and mechanical impurities, preventing the oligomers and mechanical impurities from precipitating, accumulating and clogging the equipment when the waste liquid is further distilled, which can greatly improve the distillation effect; it also prevents the oligomers from entering the concentration and crystallization processing unit, affecting the purity of the crystallized salt, and also prevents the oligomers from entering the final distillation tower kettle, resulting in a large amount of tower kettle residues that need to be processed in the end, and the cost is greatly increased. The system introduces an azeotropic agent to reduce the boiling point of the material. At the same time, only water and the azeotropic agent in the material to be processed need to be distilled out of the tower according to the azeotropic ratio, and all the materials do not need to reach the boiling point distillation, which directly reduces energy consumption and improves the processing efficiency. Finally, in the distillation stage, it is mainly the separation of alcohol and organic solvent. The boiling points of the two materials are very different, the separation is easy, the separation effect is good, and the product purity is high; a salt solution concentration and crystallization processing unit is added to the system, and the solvent and salt are recovered at the same time, and the recovery of the salt uses a crystallization operation method to ensure the high purity of the final product.

[0024] 2. The recovery method provided by the present invention does not introduce an extractant, and the entrainer used has low toxicity and can be recycled, with a low risk of secondary pollution and is environmentally friendly.

[0025] Among the extraction and distillation routes that are currently more commonly used, chloroform with good extraction effect is often used. However, chloroform is harmful to both the human body and the environment. First, it is highly toxic and has certain teratogenic and mutagenic effects on humans and animals. It is a substance on the list of Class 2B carcinogens and a suspected human carcinogen. Secondly, it is a substance on the list of toxic and harmful water pollutants. It has an accumulation effect in groundwater, which will pollute drinking water, food and vegetables, and damage the respiratory system of sensitive aquatic organisms. It is also easy to make poisonous chemicals. In the residual phase of the extraction operation, there will be 100 to 5000 ppm of chloroform residues, which is harmful to humans or causes secondary pollution to the environment. The entrainer used in the present invention, such as isobutanol, has very low toxicity and can even be used as an edible flavoring. At the same time, it is relatively easy to be biodegraded.

[0026] In the evaporation-distillation process in the prior art, the wastewater needs to be completely evaporated to separate the high-boiling-point substances (oligomers, salts and other impurities) and low-boiling-point substances (mainly water and solvents) dissolved in the wastewater. Because the boiling point of the solvent is high, and the latent heat of evaporation of water is large and the content is also high, the evaporation time is long, the energy consumption is large, and the material will also decompose into small molecular acetic acid and dimethylamine after a long period of heating, which reduces the quality and yield of the product.

[0027] The entrainer used in the present invention is mostly alcohols containing 3 to 6 carbon atoms, and a few benzenes and ethylene dichloride. The boiling point is between 80 and 140 degrees, and the stability is good under normal conditions. The azeotropic point of the entrainer with water is lower than the boiling point of water, about 70 to 95 degrees, and the azeotropic ratio is high (the water content is 15 to 50%). It is not necessary to use evaporation during operation, and only the water or alcohol in the mixture needs to be taken out by azeotropy. At the same time, after azeotropy with water, the entrainer itself has low intermiscibility with water, which is conducive to subsequent separation and reuse (such as simple separation by an oil-water separator), and the entrainer is continuously recycled, thereby reducing energy consumption and improving the recovery rate.

[0028] Taking the above aspects into consideration, the azeotropic agent used in the present invention has a high azeotropic ratio, a low azeotropic point, and relatively low energy consumption. In addition, the azeotropic agent has low solubility in water at 40 to 60 degrees, can be quickly separated by stratification, and has a good separation effect. The toxicity is low, and several solvents such as isobutanol and n-butanol with low environmental impact are preferred.

[0029] 3. The recovery method provided by the present invention has reasonable process settings and significantly reduces energy consumption.

[0030] In the process of the present invention, a flocculant is first used to flocculate and filter the oligomers in the waste liquid to separate the oligomers, thereby preventing the oligomers from clogging in the recovery equipment, such as the filling material of the distillation tower, and improving the azeotropic separation effect and separation efficiency. In addition, the cost of processing the final distillation residue is reduced. If the oligomers are left in the final distillation tower as the bottom residue, the amount of the bottom residue will be increased many times. These aspects can significantly reduce the operating energy consumption and cost.

[0031] After the oligomers are removed, the present invention takes the simple materials into recycling, and the first stage azeotropic separation solvent and salt are extracted from the bottom of the tower; the tower top material is separated by an oil-water separator, and the entrainer is recovered, and the material containing a small amount of entrainer is further separated from water and the entrainer by a simple azeotropic step, and then the entrainer is recovered; and the solvent and salt are simply evaporated, and the tower bottom concentrate enters the crystallization to recover the salt; and the evaporated material is rectified to obtain a high-purity solvent product, and the entrainer is recovered at the same time. The whole process is efficient, reasonable and energy-saving.

[0032] 4. The relevant reagents and devices / equipment of the recovery method provided by the present invention are easily available and can be combined or appropriately modified on the existing basis to meet the implementation requirements, which is conducive to large-scale production / processing and wide promotion.

[0033] In the present invention, the above-mentioned technical solutions can also be combined with each other to achieve more preferred combination solutions. Other features and advantages of the present invention will be described in the subsequent description, and some advantages can become obvious from the description, or can be understood by practicing the present invention. The purpose and other advantages of the present invention can be realized and obtained through the contents particularly pointed out in the description and the drawings. BRIEF DESCRIPTION OF THE DRAWINGS

[0034] The drawings are only for the purpose of illustrating specific embodiments and are not to be considered limiting of the present invention. Like reference symbols denote like components throughout the drawings.

[0035] Figure 1 The present invention is a process flow chart of the method for recovering organic solvents and alkaline earth metal chloride salts from aramid production wastewater.

[0036] Reference numerals:

[0037] 1-pH regulator; 2-flocculating agent; 3-waste liquid; 4-stirring tank; 5-first filter; 6-mixer; 7-primary azeotropic distillation tower; 8-reboiler; 9-first condenser; 10-first oil-water separator; 11-first reboiler; 12-secondary azeotropic distillation tower; 13-second condenser; 14-second oil-water separator; 15-second reboiler; 16-primary distillation tower; 17-third condenser; 18-fourth condenser 1-5th condenser; 20-evaporator; 21-secondary distillation tower; 22-6th condenser; 23-crystallizer; 24-third reboiler; 25-organic solvent storage tank; 26-first delivery pump; 27-second filter; 28-organic solvent / alkaline earth metal chloride mixed storage tank; 29-second delivery pump; 30-entrainer storage tank; 31-third delivery pump; 32-oligomer filter residue; 33-bottom waste liquid; 34-filter residue. DETAILED DESCRIPTION

[0038] The preferred embodiments of the present invention are described in detail below in conjunction with the accompanying drawings, wherein the accompanying drawings constitute a part of this application and are used together with the embodiments of the present invention to illustrate the principles of the present invention, but are not used to limit the scope of the present invention.

[0039] The invention discloses a method for recovering organic solvents and alkaline earth metal chloride salts in aramid production waste liquid. The recovery method comprises three process sections, namely, a flocculation precipitation process section, an azeotropic distillation process section, and a "distillation + evaporation crystallization" process section, and finally obtains pure high-concentration organic solvents and alkaline earth metal chloride salt crystals.

[0040] Specifically, the specific process of the flocculation and sedimentation process section is: through flocculation and filtration processes, oligomers and mechanical impurities in the waste liquid are separated to obtain filter residue and filtrate.

[0041] Specifically, the specific process of the azeotropic distillation process section is: adding an azeotropic agent to the filtrate, conveying it to a primary azeotropic distillation tower, removing water from the waste liquid through an azeotropic distillation process, and obtaining a tower top distillate D1 containing the azeotropic agent and water and a tower bottom mixed liquid F1 containing the azeotropic agent, an organic solvent, and an alkaline earth metal chloride salt.

[0042] Specifically, the specific process of the "distillation + evaporation crystallization" process section is: the entrainer, organic solvent and alkaline earth metal chloride in the bottom mixed liquid F1 are separated by the "distillation + evaporation crystallization" process to obtain a gas phase discharge near the bottom of the tower, and a high-purity organic solvent is obtained after condensation, and an alkaline earth metal chloride solution is concentrated to obtain an alkaline earth metal chloride product after crystallization.

[0043] Specifically, the flocculation and sedimentation process section is referred to as process section S1, the azeotropic distillation process section is referred to as process section S2, and the "distillation + evaporation crystallization" process section is referred to as process section S3.

[0044] Specifically, the specific composition of the aramid production waste liquid is: 20% to 40% by mass of organic solvent, 0.2% to 3% of alkaline earth metal chloride, 0.1% to 2% of oligomer, and the rest is water.

[0045] Specifically, the organic solvent is N,N-dimethylacetamide (DMAC) or N-methylpyrrolidone (NMP).

[0046] Specifically, the alkaline earth metal chloride salt is lithium chloride or calcium chloride.

[0047] Specifically, a pH adjuster needs to be added before flocculation to adjust the pH value to neutral to weakly alkaline. During the synthesis of aramid polymers, an acidic byproduct, hydrogen chloride, is produced. In the subsequent treatment, some are neutralized with alkali, while others are not neutralized or not completely neutralized. Therefore, the waste liquid generated in the production will be neutral or acidic. In order to reduce the impact (corrosion) on equipment or processes, these waste liquids need to be treated with alkaline substances before treatment to make the waste liquid neutral or weakly alkaline.

[0048] Preferably, the pH regulator is an alkaline compound corresponding to an alkaline earth metal chloride salt, such as CaCl2 corresponding to Ca(OH)2, and LiCl corresponding to LiOH, to avoid introducing new ions / impurities into the waste liquid.

[0049] Specifically, the oligomers and mechanical impurities in the waste liquid may affect the stability of the azeotropic structure, and may block the gaps between the tower plates in the subsequent azeotropic distillation / rectification process, so they need to be removed in advance. Process section S1 adds a flocculant to flocculate / precipitate the oligomers, and removes the waste solid oligomers / filter residues (mechanical impurities) through a filtration process.

[0050] Specifically, the flocculant used in the flocculation process in process section S1 is one or more of polyaluminum sulfate, aluminum sulfate, polyaluminum chloride, aluminum chloride, polyferric sulfate, and polyacrylamide, preferably polyaluminum sulfate and polyaluminum chloride. The amount of flocculant added is 0.01% to 0.1% of the total mass of the waste liquid, preferably 0.01 to 0.02%.

[0051] The above flocculant has a good flocculation effect on this waste liquid system. After the waste liquid is neutralized and the pH is adjusted, more oligomers are precipitated from the solution to form very small suspended particles. The added flocculant quickly disperses in the solution and adsorbs or interacts with these suspended particles, so that these suspended particles can quickly aggregate into large particles or floccules. These floccules also have good filterability and can stay on the filter cake to achieve a good separation effect. When the amount of flocculant added is less than 0.01%, the flocculation time is long, the flocculent particles obtained are not large, and suspended matter is still found in the filtrate after filtration; if the amount of flocculant added is greater than 0.1%, the flocculation effect is somewhat reduced.

[0052] Exemplarily, the flocculation process can be carried out in a stirring tank, preferably a corrosion-resistant stainless steel stirring tank or a PP plastic stirring tank, with a stirring speed of 50 to 150 rpm.

[0053] Specifically, the filtrate contains an organic solvent, an alkaline earth metal chloride salt and water.

[0054] Specifically, in process section S2, an azeotropic agent is added to mix with the waste liquid, and the heating temperature rises to the azeotropic point. The steam formed by the azeotropic agent and water in the azeotropic ratio comes out from the top of the tower, and then is cooled into a liquid in the condenser. The liquid enters the stratifier. Since the azeotropic agent and water are not miscible (the azeotropic structure automatically collapses under low temperature conditions), the stratification is performed in the stratifier; the upper layer is an azeotropic agent containing a small amount of water, and the lower layer is water containing a small amount of azeotropic agent. The upper layer can be refluxed into the tower to enter azeotropy with the water in the system again and then take water out, and the lower layer of water can be discharged after further simple azeotropy, and the azeotropic agent of simple azeotropy in the secondary azeotropic distillation tower can also be recycled and reused. Compared with the extraction-distillation process, the present invention does not use an extractant that is highly toxic and harmful to humans and the environment. Compared with the evaporation-distillation process, the present invention does not require all the wastewater to be evaporated, and only requires the azeotropic agent to take the water out of the tower at a lower azeotropic point, and the material operation time is short, thereby reducing energy consumption and improving product quality.

[0055] Specifically, the azeotropic agent is one or more of isobutanol, n-butanol, isopentanol, n-pentanol, xylene, and ethylene dichloride; preferably isobutanol.

[0056] The selected entrainer has a moderate boiling point, stable physical properties within the operating temperature range, and less decomposition over a long period of time; the material has low toxicity and is less harmful to people and the environment; and at the same time, the azeotropic point is low, the water content in the azeotropic composition is high, the mutual solubility of the entrainer and water is not high, and it is easy to separate. Among them, isobutanol has low toxicity, a stable azeotropic system and a low boiling point (89.9℃), and is the best choice considering the separation effect and economic benefits.

[0057] In actual implementation, if multiple entrainers are mixed and used, it may cause the azeotropic system to be unstable or the boiling point of the azeotrope to fluctuate, which is not conducive to stable operation and separation and recovery effects. It is preferred to use one entrainer for azeotropic distillation.

[0058] Specifically, the weight ratio of the entrainer to the filtrate is 0.1 to 2: 1, preferably 0.5 to 1: 1. Too low an entrainer to waste liquid ratio cannot fully azeotropize with the water in the waste liquid and take the water out of the tower top for separation; too high a ratio will increase the burden on the tower and increase energy consumption.

[0059] Specifically, the reflux ratio of the primary azeotropic distillation tower is 0.5-5; the operating conditions are: normal pressure, tower top temperature is 89.9-95°C, tower bottom temperature is 100-110°C, and the number of tower plates is 12-18, preferably 12-15.

[0060] The selection of the number of plates and the reflux ratio is mainly based on whether the tower can meet the material separation effect. Secondly, after meeting the material separation effect, the energy consumption should be reduced as much as possible. At the same time, the investment in equipment should not be too high. Experiments show that the number of plates of the first-stage azeotropic tower can meet the requirements when it is more than 12, but the reflux required when there are 12 plates is relatively high, reaching more than 5; when the number of plates is 15 or more, or even 18 plates, the reflux ratio is lower.

[0061] Preferably, a "condensation + oil-water separation" process section is further provided after the process section S2 to separate the azeotropic agent and the water phase in the top distillate D1, and the azeotropic agent is recycled.

[0062] Preferably, a "secondary azeotropic distillation + condensation + oil-water separation" process section is further provided after the "condensation + oil-water separation" process section to separate the residual entrainer from the water phase, recycle the entrainer, and discharge the water phase as waste liquid. In actual operation, a multi-stage azeotropic distillation tower may also be provided, but the economic efficiency is poor and the effect of entrainer recovery is limited.

[0063] The mixed gas (tower top distillate D1) azeotropically distilled from the top of the primary azeotropic distillation tower is liquefied into a liquid state after passing through a condenser and enters an oil-water separator. Since the entrainer and water are immiscible (the azeotropic system automatically disintegrates under low temperature conditions), the mixture is phase-separated into two liquid layers in the oil-water separator. The aqueous solution layer with a trace of entrainer and the entrainer layer containing a small amount of water, the entrainer layer is refluxed to the primary azeotropic distillation tower for recycling. The aqueous solution layer with a trace of entrainer is extracted and enters the secondary azeotropic distillation tower, and is further separated by azeotropic distillation. The secondary azeotropic distillation tower top distillate D2 is condensed and liquefied and enters the water layer of the oil-water separator, and then discharged as waste liquid (tower bottom discharge liquid F2) in the tower bottom. The entrainer can be extracted to the entrainer storage tank for recycling. The recycling of the entrainer reduces the production / processing cost and reduces the discharge of waste liquid.

[0064] Specifically, the operating conditions of the secondary azeotropic distillation tower are: normal pressure, tower top temperature 89-92°C, tower bottom temperature 99-105°C, theoretical plate number 5-15, preferably 7-10; reflux ratio 0.05-2, preferably 0.1-0.5; condenser cooling water temperature 25-35°C.

[0065] The secondary azeotropic distillation tower is mainly used to separate a small amount of entrainer in the water layer and recycle it into the entrainer storage tank. Because the content of the entrainer required for azeotropy is not high, the tower is not heavily burdened; more than 5 plates can meet the requirements, and when the number of plates is 7 or more, or even 10 plates, the reflux ratio is lower.

[0066] Specifically, process section S3 includes the following specific steps:

[0067] S3-1: The bottom mixed liquid F1 is subjected to primary distillation to obtain a top distillate D3 containing an azeotropic agent and an organic solvent and a bottom mixed liquid F3 containing an organic solvent and an alkaline earth metal chloride salt:

[0068] S3-2: transporting the mixed liquid F3 from the bottom of the tower to an evaporator to obtain organic solvent vapor and concentrated alkaline earth metal chloride solution;

[0069] S3-3: subjecting the tower top distillate D3 and the organic solvent vapor to secondary distillation to obtain the tower top effluent D4, i.e., the high-concentration entrainer; obtaining the gas phase discharge at the tower bottom (near the tower bottom), and condensing it to obtain a high-purity organic solvent;

[0070] S3-4: The concentrated alkaline earth metal chloride salt solution is transported to a crystallization kettle, cooled and crystallized, and filtered to obtain a filter residue / alkaline earth metal chloride salt product.

[0071] Specifically, the primary distillation described in step S3-1 is carried out in a primary distillation tower, and the pressure is: vacuum degree 0.06-0.09 MPa, preferably 0.075-0.085 MPa; the tower top temperature is 60-90°C, and the tower bottom temperature is 100-135°C; when the organic solvent is DMAC, the tower bottom temperature is preferably 100-135°C; when the organic solvent is NMP, the tower bottom temperature is preferably 100-130°C.

[0072] Selecting negative pressure operating conditions can reduce the boiling point of materials and reduce operating costs. At the same time, different operating temperatures can be selected according to different material types, salt content, etc. to reduce the decomposition of the recovery solvent and improve the yield and quality of material recovery.

[0073] Specifically, the specific parameters of the evaporation process in step S3-2 are: the evaporator temperature is 100-135°C; when the organic solvent is DMAC, the evaporator temperature is preferably 100-135°C; when the organic solvent is NMP, the evaporator temperature is preferably 100-130°C; the evaporation pressure is -0.1-0Mpa (absolute pressure), the cooling water temperature in the condenser is 25-35°C, and the temperature of the organic solvent after cooling is 40-60°C.

[0074] Selecting negative pressure operating conditions can reduce the boiling point of materials and reduce operating costs. At the same time, different operating temperatures are selected according to different material types and salt content to reduce the decomposition of recovery solvents and improve the yield and quality of material recovery. The material is mainly condensed from steam into liquid by the main cooling water, and then enters the secondary distillation tower; the liquid temperature is kept at a high temperature or even directly steam fed; the main control is to stabilize the temperature and flow of the material.

[0075] Specifically, the secondary distillation described in step S3-3 is carried out in a secondary distillation tower to separate the organic solvent and the entrainer, and the top effluent D4 is an entrainer with a mass concentration of more than 96%; the gas phase discharge at the bottom of the tower (near the kettle) is condensed to obtain a high-purity organic solvent with a mass concentration of ≥99.9%.

[0076] Specifically, the process parameters in the secondary distillation tower are as follows: the tower top temperature is 60-80°C, the tower bottom temperature is 110-150°C, when the organic solvent is DMAC, the tower bottom temperature is preferably 110-130°C; when the organic solvent is NMP, the tower bottom temperature is preferably 120-150°C; the number of plates is 12-30, preferably 15-25; the reflux ratio is 1-5, preferably 1-3; the vacuum degree is 0.06-0.09Mpa, preferably 0.075-0.085Mpa; the cooling water temperature in the condenser is 25-30°C, and the temperature of the mixed liquid after cooling is 30°C-40°C.

[0077] The secondary distillation tower is the main equipment for separating the azeotropic agent from the solvent, and has high requirements. The number of plates must be more than 12 to achieve the separation requirements. When the number of plates is more than 15 or 20, the reflux rate is relatively small. Negative pressure operation is used to reduce energy consumption.

[0078] Specifically, the small amount of filtrate obtained in step S3-4 is an organic solvent solution containing alkaline earth metal chloride and some impurities, which can be returned to the evaporator to re-evaporate and recover the organic solution, or can be taken out for harmless treatment.

[0079] Example 1

[0080] The waste liquid composition is calculated by mass percentage: DMAC: 25% to 40%, lithium chloride: 0.2% to 2%, oligomer: 0.5% to 1.5%, and the rest is water.

[0081] Step 1: Add the aramid production waste liquid into the stirring tank at a stirring speed of 100 rpm. First, add a small amount of LiOH to the aramid production waste liquid in the stirring tank until the pH meter test is neutral, then add 0.015% aluminum sulfate while stirring, and the flocculant concentrates the oligomers in the waste liquid to flocculate, and filter to obtain the oligomer residue, which needs to be harmlessly treated. The filtrate is a mixed solution of DMAC-LiCl-water.

[0082] Step 2: After DMAC-LiCl-water mixed solution passes through a mixer at a flow rate of 100kg / h and isobutyl alcohol at a flow rate of 80kg / h, a primary azeotropic distillation tower 7 is added to the middle of the primary azeotropic distillation tower 7; the primary azeotropic distillation tower 7 has 15 theoretical plates; the tower bottom temperature is maintained at 100 degrees, and the tower top temperature is maintained at 90 degrees. The isobutyl alcohol and water in the tower form an azeotrope, which is distilled from the top of the tower, condensed by the condenser 9 and enters the oil-water separator 10, where the water and isobutyl alcohol are separated, and the upper layer is isobutyl alcohol containing a small amount of water, which is refluxed to the primary azeotropic distillation tower 7 at a flow rate of 208kg / h; the lower layer is transported to the secondary azeotropic distillation tower 12 for further separation; isobutyl alcohol-DMAC-LiCl containing a small amount of water is obtained at the bottom of the tower, the pressure is normal pressure, and the reflux ratio of the primary azeotropic distillation tower is 2.8.

[0083] Step 3: In step 2, the temperature of the top steam of the primary azeotropic distillation tower 7 is 60°C after condensation and liquefaction, and enters the oil-water separator 10, and the upper layer refluxes to the primary azeotropic distillation tower 7; the lower layer is water containing 5% isobutanol, which is transported to the secondary azeotropic distillation tower 12 with a flow rate of 70.9kg / h; after entering the secondary azeotropic distillation tower 12, isobutanol and water are azeotropically distilled from the top of the tower again, and after condensation, they enter the oil-water separator 14 to further separate isobutanol and water. The upper layer liquid is isobutanol containing 10% water, which can be extracted and placed in the isobutanol storage tank with a flow rate of 4.6kg / h; the lower layer is water containing 200ppm of alcohol, which refluxes to the secondary azeotropic distillation tower 12 with a flow rate of 0.86kg / h. The bottom of the secondary azeotropic distillation tower 12 is water containing 10 to 100ppm of alcohol and DMAC, which can be directly discharged to the wastewater treatment station. The secondary azeotropic distillation tower 12 has 9 plates, a top temperature of 90°C, a bottom temperature of 100°C, a reflux ratio of 0.19, and a tower pressure of normal pressure.

[0084] Step 4: The mixed liquid in the bottom of the first azeotropic distillation tower 7 in step 2 is extracted and transported to the primary distillation tower 16. The composition of the mixed liquid is as follows by mass percentage: isobutanol: 70.4%, DMAC: 23.8%, water: 3.1%, LiCl: 2.6%; the flow rate is 113.5 kg / h;

[0085] After separation in the primary distillation tower 16, the top distillate is water-isobutanol-DMAC, with a composition of 83.6% isobutanol, 12.6% DMAC, 3.75% water, and 0% LiCl, with a flow rate of 94.9 kg / h, and is transported to the secondary distillation tower 21;

[0086] The bottom of the primary distillation tower 16 is a trace isobutanol-DMAC-LiCl solution, with isobutanol of 3.2%, DMAC of 80.6%, water of 0.0004%, LiCl of 16.1%, and a flow rate of 18.61 kg / h. The top temperature of the primary distillation tower 16 is 60-75°C, the bottom temperature is 112-121°C, the number of plates is 8, the reflux ratio is 0.5, and the vacuum degree is 0.08Mpa.

[0087] Step 5: The top distillate D3 of the primary distillation tower 16 in step 4 is isobutanol-water-DMAC, which is liquefied through a condenser and enters the secondary distillation tower 21; the trace isobutanol-DMAC-LiCl solution (mixed liquid F3 in the bottom of the primary distillation tower 16) is extracted and input into an evaporator for evaporation, and the resulting steam can directly enter the secondary distillation tower 21 or enter the secondary distillation tower 21 after condensation;

[0088] The composition of liquefied steam is: 5.9% isobutanol, 94.1% DMAC, 0.007% water, 0% LiCl, and the flow rate is 10.11 kg / h. The residual material from the evaporator enters the crystallization kettle.

[0089] Step 6: In step 4 and step 5, the top distillate D3 of the primary distillation tower 16 and the evaporated steam of the evaporator 29 enter the secondary distillation tower 21 for separation. The top discharge of the secondary distillation tower 21 (top distillate D4) is high-concentration isobutanol, with isobutanol of 94.1%, DMAC of 1.68%, water of 4.18%, LICL of 0%, and a flow rate of 85.0 kg / h. The isobutanol storage tank is extracted and stored, and then transported to the inlet of the primary azeotropic distillation tower 7 by pump 31 to form the recycling of the isobutanol entrainer. The vapor phase discharge near the bottom of the tower is DMAC. After condensation, a high-purity DMAC product is obtained, which should be ≧99.99%, with water and alcohol contents of <200ppm, chroma ≤3°, acetic acid content ≤500ppm, dimethylamine ≤500ppm, and a flow rate of 18.82kg / h. It is stored in a DMAC product storage tank and can be returned to the workshop for direct use; the tower bottom discharge is tower bottom waste, which needs to be harmlessly treated.

[0090] The top temperature of the secondary distillation tower 21 is 70-80°C, the bottom temperature is 110-130°C, the number of plates is 25, the reflux ratio is 3.0, and the vacuum degree is 0.08Mpa.

[0091] Step 7: The concentrated solution in the evaporator in step 4 is a concentrated DMAC-LiCl solution, which can be transported to a glass-lined crystallization kettle with a stirrer and a screw discharger, and the jacket cooling water is introduced while stirring to cool to room temperature for crystallization. The crystals (i.e., high-purity lithium chloride) are sealed and squeezed from the crystallization kettle using a screw discharger to a filter to obtain a filtrate, with a flow rate of 6.48 kg / h, of which DMAC is 91.5%, lithium chloride is 7.98%, and the rest is impurities. The filtrate can be re-entered into the evaporator to evaporate DMAC.

[0092] The crystals in the sealed filter are used to recover lithium chloride, with a yield of 2.52 kg / h. These crystals can also be dissolved by adding qualified DMAC solvent, filtered and stored in DMAC-LiCl storage tanks, and then transported to the workshop for use.

[0093] The total recovery rate of DMAC is 97%, and the total recovery rate of lithium chloride is 95%. The detection purity of DMAC is ≥99.99%, the content of water and alcohol is respectively <200ppm, the chromaticity is ≤3°, the acetic acid content is ≤500ppm, and the dimethylamine is ≤500ppm.

[0094] Example 2

[0095] The waste liquid composition is calculated by mass percentage: DMAC: 20% to 40%, calcium chloride: 0.2 to 3%, oligomer: 1% to 2%, and the rest is water.

[0096] Step 1: Add the aramid production waste liquid into the stirring tank at a stirring speed of 100 rpm. First, add a small amount of Ca(OH)2 to the aramid production waste liquid in the stirring tank until the pH meter test is neutral, then add 0.13% aluminum chloride while stirring, and the flocculant concentrates the oligomers in the waste liquid to flocculate, and filter to obtain the oligomer filter residue, which needs to be harmlessly treated. The filtrate is a mixed solution of DMAC-CaCl2-water.

[0097] Step 2: After passing through a mixer at a flow rate of 100kg / h and isobutanol at a flow rate of 80kg / h in a DMAC-CaCl2-water mixed solution, a primary azeotropic distillation tower 7 is added to the middle of the primary azeotropic distillation tower 7; the primary azeotropic distillation tower 7 has 15 theoretical plates; the tower bottom temperature is maintained at 100 degrees, and the tower top temperature is maintained at 90 degrees. The isobutanol and water in the tower form an azeotrope, which is distilled from the top of the tower, condensed by a condenser 9 and enters an oil-water separator 10, where water and isobutanol are separated, and the upper layer is isobutanol containing a small amount of water, which is refluxed to the primary azeotropic distillation tower 7 at a flow rate of 208kg / h; the lower layer is transported to a secondary azeotropic distillation tower 12 for further separation; isobutanol-DMAC-CaCl2 containing a small amount of water is obtained at the bottom of the tower, the pressure is normal pressure, and the reflux ratio of the primary azeotropic distillation tower is 2.8.

[0098] Step 3: In step 2, the temperature of the top steam of the primary azeotropic distillation tower 7 is 60°C after condensation and liquefaction, and enters the oil-water separator 10, and the upper layer refluxes to the primary azeotropic distillation tower 7; the lower layer is water containing 5% isobutanol, which is transported to the secondary azeotropic distillation tower 12 with a flow rate of 70.9kg / h; after entering the secondary azeotropic distillation tower 12, isobutanol and water are azeotropically distilled from the top of the tower again, and after condensation, they enter the oil-water separator 14 to further separate isobutanol and water. The upper layer liquid is isobutanol containing 10% water, which can be extracted and placed in the isobutanol storage tank with a flow rate of 4.6kg / h; the lower layer is water containing 200ppm of isobutanol, which refluxes to the secondary azeotropic distillation tower 12 with a flow rate of 0.86kg / h. The bottom of the secondary azeotropic distillation tower 12 is water containing alcohol and DMAC at 10 to 100ppm, which can be directly discharged to the wastewater treatment station. The secondary azeotropic distillation tower 12 has 9 plates, a top temperature of 90°C, a bottom temperature of 100°C, a reflux ratio of 0.19, and a tower pressure of normal pressure.

[0099] Step 4: The mixed liquid F1 in the bottom of the first-stage azeotropic distillation tower 7 in step 2 is extracted and transported to the primary distillation tower 16. The composition of the mixed liquid is as follows by mass percentage: isobutanol: 70.4%, DMAC: 23.8%, water: 3.1%, CaCl2: 2.6%; the flow rate is 113.5 kg / h;

[0100] After separation in the primary distillation tower 16, the top distillate is isobutanol-DMAC, with a composition of 83.6% isobutanol, 12.6% DMAC, 3.75% water, 0% CaCl2, and a flow rate of 94.9 kg / h, which is transported to the secondary distillation tower 21;

[0101] The bottom of the primary distillation tower 16 is a trace isobutanol-DMAC-CaCl2 solution, with 3.2% isobutanol, 80.6% DMAC, 0.0004% water, 16.1% CaCl2, and a flow rate of 18.61 kg / h. The top temperature of the primary distillation tower 16 is 60-75°C, the bottom temperature is 112-121°C, the number of plates is 8, the reflux ratio is 0.5, and the vacuum degree is 0.08Mpa.

[0102] Step 5: The top distillate D3 of the primary distillation tower 16 in step 4 is isobutanol-water-DMAC, which is liquefied through a condenser and enters the secondary distillation tower 21; the trace alcohol-DMAC-CaCl2 solution (mixed liquid F3 in the bottom of the primary distillation tower 16) is extracted and input into an evaporator for evaporation, and the resulting steam can directly enter the secondary distillation tower 21 or enter the secondary distillation tower 21 after condensation;

[0103] The composition of liquefied steam is: 5.9% isobutanol, 94.1% DMAC, 0.007% water, 0% CaCl2, and the flow rate is 10.11 kg / h. The residual material from the evaporator enters the crystallization kettle.

[0104] Step 6: In step 4 and step 5, the top distillate D3 of the primary distillation tower 16 and the evaporated steam of the evaporator 29 enter the secondary distillation tower 21 for separation. The top discharge of the secondary distillation tower 21 (top distillate D4) is high-concentration isobutanol, with isobutanol of 94.1%, DMAC of 1.68%, water of 4.18%, CaCl2 of 0%, and a flow rate of 85.0 kg / h. The isobutanol storage tank is extracted and stored, and then transported to the inlet of the primary azeotropic distillation tower 7 by pump 31 to form the recycling of the isobutanol entrainer. The vapor phase discharge near the bottom of the tower is DMAC. After condensation, a high-purity DMAC product is obtained, which should be ≧99.99%, with water and alcohol contents of <200ppm, chroma ≤3°, acetic acid content ≤500ppm, dimethylamine ≤500ppm, and a flow rate of 18.82kg / h. It is stored in a DMAC product storage tank and can be returned to the workshop for direct use; the tower bottom discharge is tower bottom waste, which needs to be harmlessly treated.

[0105] The top temperature of the secondary distillation tower 21 is 70-80°C, the bottom temperature is 110-130°C, the number of plates is 25, the reflux ratio is 3.0, and the vacuum degree is 0.08Mpa.

[0106] Step 7: The concentrated solution in the evaporator in step 4 is a DMAC-CaCl2 concentrated solution, which can be transported to a glass-lined crystallization kettle with a stirrer and a screw discharger, and the jacket cooling water is introduced while stirring to cool to room temperature for crystallization. The crystals (i.e., high-purity calcium chloride) are sealed and squeezed from the crystallization kettle using a screw discharger to a filter to obtain a filtrate, with a flow rate of 6.48 kg / h, of which DMAC is 91.5%, lithium chloride is 7.98%, and the rest is impurities. The filtrate can re-enter the evaporator to evaporate DMAC.

[0107] The crystals in the sealed filter are recovered calcium chloride with a yield of 2.52 kg / h. These crystals can also be dissolved by adding qualified DMAC solvent, filtered and stored in DMAC-CaCl2 storage tanks, and then transported to the workshop for use.

[0108] The total recovery rate of DMAC is 98.3%, and the total recovery rate of calcium chloride is 94.5%. The detection purity of DMAC is ≥99.99%, the content of water and alcohol is respectively <200ppm, the chromaticity is ≤3°, the acetic acid content is ≤500ppm, and the dimethylamine is ≤500ppm.

[0109] Example 3

[0110] The waste liquid composition is calculated by mass percentage: NMP: 25% to 40%, calcium chloride: 0.2% to 2%, oligomer: 0.1% to 1%, and the rest is water.

[0111] Step 1: Add the aramid production waste liquid into the stirring tank at a stirring speed of 100 rpm. First, add a small amount of Ca(OH)2 to the aramid production waste liquid in the stirring tank until the pH meter test is neutral, then add 0.02% aluminum chloride while stirring, and the flocculant will concentrate the oligomers in the waste liquid to flocculate, and filter to obtain the oligomer filter residue, which needs to be harmlessly treated. The filtrate is a mixed solution of NMP-CaCl2-water.

[0112] Step 2: After passing through a mixer at a flow rate of 100kg / h and isobutanol at a flow rate of 80kg / h in a NMP-CaCl2-water mixed solution, a primary azeotropic distillation tower 7 is added to the middle of the primary azeotropic distillation tower 7; the primary azeotropic distillation tower 7 has 15 theoretical plates; the tower bottom temperature is maintained at 105 degrees, and the tower top temperature is maintained at 90 degrees. Isobutanol and water in the tower form an azeotrope, which is distilled from the top of the tower, condensed by a condenser 9 and enters an oil-water separator 10, in which water and isobutanol are separated, and the upper layer is isobutanol containing a small amount of water, which is refluxed to the primary azeotropic distillation tower 7, with a flow rate of 209kg / h; the lower layer is transported to a secondary azeotropic distillation tower 12 for further separation; isobutanol-NMP-CaCl2 containing a small amount of water is obtained at the bottom of the tower, the pressure is normal pressure, and the reflux ratio of the primary azeotropic distillation tower is 3.2.

[0113] Step 3: In step 2, the temperature of the top steam of the first-level azeotropic distillation tower 7 is 60°C after condensation and liquefaction, and enters the oil-water separator 10, and the upper layer refluxes to the first-level azeotropic distillation tower 7; the lower layer is water containing 5.57% of isobutanol, which is transported to the second-level azeotropic distillation tower 12 with a flow rate of 71.2kg / h; after entering the second-level azeotropic distillation tower 12, isobutanol and water are azeotropically distilled from the top of the tower again, and after condensation, they enter the oil-water separator 14 to further separate isobutanol and water. The upper layer liquid is isobutanol containing 8% water, which can be extracted and placed in the isobutanol storage tank with a flow rate of 4.5kg / h; the lower layer is water containing 200ppm of alcohol, which refluxes to the second-level azeotropic distillation tower 12 with a flow rate of 0.82kg / h. The bottom of the second-level azeotropic distillation tower 12 is water containing alcohol and NMP at 10 to 100ppm, which can be directly discharged to the wastewater treatment station. The secondary azeotropic distillation tower 12 has 9 plates, a top temperature of 89.9°C, a bottom temperature of 100°C, a reflux ratio of 0.21, and a tower pressure of normal pressure.

[0114] Step 4: The mixed liquid F1 in the bottom of the first-stage azeotropic distillation tower 7 in step 2 is extracted and transported to the primary distillation tower 16. The composition of the mixed liquid is as follows by mass percentage: isobutanol: 70.4%, NMP: 23.8%, water: 3.1%, CaCl2: 2.6%; the flow rate is 113.6 kg / h;

[0115] After separation in the primary distillation tower 16, the top distillate is isobutanol-NMP, with the composition of 83.6% isobutanol, 12.6% DMAC, 3.75% water, 0% CaCl2, and a flow rate of 94.9 kg / h, which is transported to the secondary distillation tower 21;

[0116] The bottom of the primary distillation tower 16 is a trace isobutanol-NMP-CaCl2 solution, with 3.2% isobutanol, 80.6% NMP, 0.0004% water, 16.1% CaCl2, and a flow rate of 18.6 kg / h. The top temperature of the primary distillation tower 16 is 65-80°C, the bottom temperature is 120-130°C, the number of plates is 9, the reflux ratio is 0.45, and the vacuum degree is 0.08Mpa.

[0117] Step 5: The top distillate D3 of the primary distillation tower 16 in step 4 is isobutanol-water-NMP, which is liquefied through a condenser and enters the secondary distillation tower 21; the trace alcohol-NMP-CaCl2 solution (mixed liquid F3 in the bottom of the primary distillation tower 16) is extracted and input into an evaporator for evaporation, and the resulting steam can directly enter the secondary distillation tower 21 or enter the secondary distillation tower 21 after condensation;

[0118] The composition of liquefied steam is: 5.9% isobutanol, 94.1% DMAC, 0.007% water, 0% CaCl2, and the flow rate is 10.11 kg / h. The residual material from the evaporator enters the crystallization kettle.

[0119] Step 6: In step 4 and step 5, the top distillate D3 of the primary distillation tower 16 and the evaporated steam of the evaporator 29 enter the secondary distillation tower 21 for separation. The top discharge of the secondary distillation tower 21 (top distillate D4) is high-concentration isobutanol, with 94.1% isobutanol, 1.68% NMP, 4.18% water, 0% CaCl2, and a flow rate of 85.0 kg / h. The isobutanol storage tank is extracted and stored, and then transported to the inlet of the primary azeotropic distillation tower 7 through a pump 31 to form a recycling of the isobutanol entrainer. The vapor phase discharge near the bottom of the tower is DMAC. After condensation, a high-purity DMAC product is obtained, which should be ≧99.99%, with water and alcohol contents of <200ppm, chroma ≤3°, acetic acid content ≤500ppm, dimethylamine ≤500ppm, and a flow rate of 18.82kg / h. It is stored in a DMAC product storage tank and can be returned to the workshop for direct use; the tower bottom discharge is tower bottom waste, which needs to be harmlessly treated.

[0120] The top temperature of the secondary distillation tower 21 is 70-80°C, the bottom temperature is 125-130°C, the number of plates is 25, the reflux ratio is 3.3, and the vacuum degree is 0.08Mpa.

[0121] Step 7: The concentrated solution in the evaporator in step 4 is a concentrated NMP-CaCl2 solution, which can be transported to a glass-lined crystallization kettle with a stirrer and a screw discharger, and the jacket cooling water is introduced while stirring to cool it to room temperature for crystallization. The crystals (i.e., high-purity calcium chloride) are sealed and squeezed from the crystallization kettle using a screw discharger to a filter to obtain a filtrate with a flow rate of 6.50 kg / h, of which NMP is 91.5%, calcium chloride is 8.2%, and the rest are impurities. The filtrate can re-enter the evaporator to evaporate NMP.

[0122] The crystals in the sealed filter are recovered calcium chloride with a yield of 2.6 kg / h. These crystals can also be dissolved by adding qualified NMP solvent, filtered and stored in NMP-CaCl2 storage tanks, and then transported to the workshop for use.

[0123] The total recovery rate of NMP is 97.3%, and the total recovery rate of calcium chloride is 94.0%. The purity of NMP detected is ≥ 99.99%, and the contents of water and alcohol are respectively less than 200 ppm.

[0124] In the existing DMAC recovery process, the evaporation-distillation process generally has a recovery rate of 90-92%, and can reach 95% in some periods, with a purity of 98% and no recovered salt. The extraction-distillation process has a recovery rate of 93-98%, and the solvent purity can reach 99%, without recovered salt.

[0125] The recovery rate of existing NMP recovery process is 90-99%, the purity can reach 99.5%, and no salt is recovered.

[0126] The DMAC recovery rate of the invention can reach more than 97%, the solvent purity can reach more than 99.9%, and the alkaline earth metal salt is recovered at the same time, thereby reducing the impact on the environment.

[0127] The NMP recovery rate of the invention can reach more than 97%, the solvent purity can reach more than 99.9%, and the alkaline earth metal salt and the organic solvent are recovered at the same time, thereby reducing the impact on the environment.

[0128] The above description is only a preferred specific implementation manner of the present invention, but the protection scope of the present invention is not limited thereto. Any changes or substitutions that can be easily conceived by any technician familiar with the technical field within the technical scope disclosed by the present invention should be covered within the protection scope of the present invention.

Claims

1. A method for recovering organic solvents and alkaline earth metal chlorides from aramid production wastewater, characterized in that: The recovery method includes three process sections, namely, a flocculation precipitation process section, an azeotropic distillation process section, and a "distillation + evaporation crystallization" process section, to obtain pure organic solvent and alkaline earth metal chloride crystals; The specific process of the flocculation and precipitation process section is: separating oligomers and mechanical impurities in the waste liquid through flocculation and filtration processes to obtain filter residue and filtrate; The specific process of the azeotropic distillation process section is as follows: adding an azeotropic agent to the filtrate, conveying it to a primary azeotropic distillation tower, removing water from the waste liquid through an azeotropic distillation process, and obtaining a tower top distillate D1 containing the azeotropic agent and water and a tower bottom mixed liquid F1 containing the azeotropic agent, an organic solvent, and an alkaline earth metal chloride salt; The specific process of the "rectification + evaporation crystallization" process section is: the entrainer, organic solvent and alkaline earth metal chloride salt in the tower bottom mixed liquid F1 are separated by the "rectification + evaporation crystallization" process to obtain a gas phase discharge material close to the tower bottom, and a high-purity organic solvent is obtained after condensation, and an alkaline earth metal chloride salt solution is concentrated, and an alkaline earth metal chloride salt product is obtained after crystallization; The organic solvent is N,N-dimethylacetamide or N-methylpyrrolidone; The azeotropic agent is one or more of isobutyl alcohol, n-butyl alcohol, isopentanol, n-pentanol, xylene and ethylene dichloride.

2. The recycling method according to claim 1, characterized in that: The specific composition of the aramid production waste liquid is: 20% to 40% by mass of organic solvent, 0.2% to 3% of alkaline earth metal chloride, 0.1% to 2% of oligomer, and the rest is water.

3. The recycling method according to claim 1, characterized in that: The alkaline earth metal chloride salt is lithium chloride or calcium chloride.

4. The recycling method according to claim 1, characterized in that: The mass ratio of the azeotropic agent to the filtrate is 0.1-2:

1.

5. The recycling method according to claim 1, characterized in that: A "condensation + oil-water separation" process section is also provided after the azeotropic distillation process section to separate the entrainer and the water phase in the overhead distillate D1, and the entrainer is recycled; after the "condensation + oil-water separation" process section, a secondary or multi-stage azeotropic distillation tower is also provided for further separation and recovery of the entrainer.

6. The recycling method according to claim 1, characterized in that: The "rectification + evaporation crystallization" process section includes the following specific steps: S3-1: The bottom mixed liquid F1 is subjected to primary distillation to obtain a top distillate D3 containing an azeotropic agent and an organic solvent and a bottom mixed liquid F3 containing an organic solvent and an alkaline earth metal chloride salt: S3-2: transporting the mixed liquid F3 from the bottom of the tower to an evaporator to obtain organic solvent vapor and concentrated alkaline earth metal chloride solution; S3-3: performing secondary distillation on the tower top distillate D3 and the organic solvent vapor to obtain the tower top distillate D4, i.e., the high-concentration entrainer; obtaining the gas phase discharge near the tower bottom, and condensing it to obtain a high-purity organic solvent; S3-4: The concentrated alkaline earth metal chloride solution is transported to a crystallization kettle, cooled and crystallized, and the alkaline earth metal chloride product is obtained after filtration.

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