Method for recovering waste acid and waste salt in chlorine drug intermediate production wastewater

Abstract

The invention provides a method for recovering waste acid and waste salt in chlorine drug intermediate production wastewater, and belongs to the technical field of industrial wastewater treatment. The method comprises the following steps: adsorbing homogenized production wastewater by macroporous resin in an adsorption column to remove residual organic phases in the production wastewater so as to obtain inorganic waste liquid; inorganic waste liquid enters a chromatographic column, and free acid is adsorbed by an adsorption filler to obtain high-salt waste liquid; eluting with the chromatographic column, and recovering concentrated acid; the high-salt waste liquid meeting the expected pH value enters a multi-effect evaporation system to be subjected to evaporative crystallization, and sodium salt crystals are recycled respectively; the multi-effect evaporation system at least adopts a triple-effect parallel-flow feeding evaporation system, and the temperature adopted for evaporation in the material liquid flowing direction is gradually reduced; the macroporous resin adopts chloromethylated resin; the adsorption filler is a microsphere filler of which the surface is crosslinked with a biquaternary ammonium salt group. The waste acid and the waste salt in the chlorine drug intermediate production wastewater are efficiently recovered by adopting a mode of adsorbing an organic phase, separating free acid by a chromatographic column and performing multi-effect parallel flow feeding evaporative crystallization.

Description

Technical Field

[0001] The invention belongs to the technical field of industrial wastewater treatment, and particularly relates to a method for recovering waste acid and waste salt in wastewater produced by the production of chlorine-based drug intermediates. Background Art

[0002] Currently, there are a large number of chlorine drugs in medicines or a large number of chlorine compound intermediates are generated during the preparation of drugs, such as 3-chloropropionic acid, 2,6-dichloroaniline, and 3-chloropropyne; a large number of chlorine drug intermediates such as 2,5-dichlorohexanoyl chloride are produced during the synthesis of 2,5-thiophenedicarboxylic acid (an important intermediate for bis(benzoxazole)-type fluorescent whitening agents, high-efficiency agricultural fungicides and anticancer drugs).

[0003] The synthesis of these drugs requires substitution, addition, and other reactions, which generate large amounts of waste acids and waste salts, such as hydrochloric acid, sulfuric acid, phosphoric acid, and their sodium salts. The current mainstream process uses limewater to react and generate a water-insoluble or slightly water-soluble calcium precipitate to treat the waste acids and waste salts. However, for chlorine-containing waste acids and waste salts, treatment is difficult because calcium chloride is soluble in water. Furthermore, since large amounts of acid are required during the synthesis of drugs or drug intermediates, the acid salts in the wastewater are not recovered, resulting in a significant waste of resources.

[0004] Due to the volatility of hydrochloric acid, current methods for recovering spent acid from waste liquids primarily rely on a distillation-and-condensation process, and evaporation and crystallization of spent salts. While distillation is effective for treating high-concentration spent acid and salts, it suffers from low recovery efficiency, high costs, and high energy consumption for hydrochloric acid and salts from low-concentration waste liquids. Furthermore, the waste liquid's complex composition, including sodium chloride, sodium sulfate, and sodium phosphate, makes it difficult to separate and recover the high-value salt and acid. Summary of the Invention

[0005] In view of this, the purpose of the present invention is to provide a method for recovering waste acid and waste salts in wastewater produced by the production of chlorine-based pharmaceutical intermediates, aiming to solve at least one technical problem among the background technologies.

[0006] The present invention is achieved in that:

[0007] A method for recovering waste acid and waste salt from wastewater produced by the production of chlorine-based pharmaceutical intermediates, the method comprising the following steps:

[0008] The homogenized production wastewater is adsorbed by the macroporous resin in the adsorption column to remove the residual organic phase and obtain inorganic waste liquid;

[0009] The inorganic waste liquid enters the chromatographic column, where the adsorption filler adsorbs the free acid to obtain a high-salt waste liquid that meets the expected pH value; the concentrated acid is recovered after elution from the chromatographic column;

[0010] The high-salt wastewater that meets the expected pH value enters the multi-effect evaporation system for evaporation and crystallization, and the sodium salt crystals are recovered separately;

[0011] The multi-effect evaporation system adopts at least a three-effect parallel-flow feeding evaporation system, wherein the feed liquid flows from the previous effect evaporation system to the next effect evaporation system through the pressure difference between adjacent two-effect evaporation systems or through a pump, heating steam is added to the first effect evaporation system, and the secondary steam generated from the previous effect evaporation system is used as the heating steam for the next effect evaporation system; the temperature used for evaporation in the multi-effect evaporation system gradually decreases along the flow direction of the feed liquid;

[0012] The macroporous resin is chloromethylated resin; the adsorption filler is a microsphere filler with surface cross-linked diquaternary ammonium salt groups.

[0013] Preferably, in the three-effect parallel evaporation system, superheated saturated steam is introduced into the first-effect evaporation system, and the temperature is lowered stepwise to room temperature for crystallization, mainly precipitating sodium sulfate crystals;

[0014] The vacuum degree of evaporation in the second-effect evaporation system is set to -70kPa to -90kPa, and the temperature is set to 60℃ to 70℃, mainly precipitating sodium chloride crystals;

[0015] The vacuum degree of evaporation in the third-effect evaporation system is set to -50kPa ~ -45kPa, and the temperature is set to 40℃ ~ 50℃, mainly precipitating sodium phosphate crystals.

[0016] Preferably, a crystallization accelerator is added during crystallization in the third-effect evaporation system, and the crystallization accelerator is sodium phosphate seed crystals.

[0017] Preferably, the pH of the high-salt waste liquid is adjusted to 8-10 before entering the multiple-effect evaporation system.

[0018] Preferably, the heating steam or secondary steam enters the first-effect evaporation system or the second-effect evaporation system tangentially after steam concentration.

[0019] Preferably, the macroporous resin is a chloromethylated resin with a pore size of 20 nm to 100 nm.

[0020] Preferably, the adsorption filler is a microsphere filler with surface cross-linked diquaternary ammonium salt groups.

[0021] Preferably, the preparation steps of the adsorption filler include:

[0022] Use ethanol or acetone to remove impurities from the microspheres;

[0023] The diquaternary ammonium salt solution was added dropwise to the microspheres, the pH was adjusted to 9-11, the solution was heated to 70-80°C, and the reaction was stirred for 2-4 hours.

[0024] After cooling, the solid is separated by filtration under reduced pressure, repeatedly washed, and vacuum dried to obtain the adsorption filler.

[0025] Preferably, the mass concentration of the diquaternary ammonium salt solution is 0.1% to 5%; the molar ratio of the diquaternary ammonium salt to the microspheres is 0.25 to 1:1.

[0026] Preferably, the general structural formula of the diquaternary ammonium salt is:

[0027]

[0028] In the formula, R is a C2-C6 alkenyl group; R1, R2, R3, and R4 are each independently selected from a C1-C6 alkyl group.

[0029] Preferably, the expected pH value refers to the pH value of the high-salt waste liquid being not less than 5; the high-salt waste liquid with a pH value less than 5 is circulated to the chromatographic column for secondary treatment.

[0030] Compared with the prior art, the present invention has the following beneficial effects:

[0031] 1. The present invention adopts the method of adsorption of organic phase + separation of free acid by chromatographic column + multi-effect parallel flow feeding evaporation crystallization to efficiently recover waste acid and waste salt in the production wastewater of chlorine-based drug intermediates.

[0032] 2. The present invention uses chloromethylated resin when adsorbing the organic phase, which has the same polarity as the organic phase impurities in the chlorine-based drug intermediate production wastewater, thereby enhancing the adsorption effect and improving the impurity removal efficiency.

[0033] 3. The filler in the chromatographic column for separating free acid of the present invention adopts microsphere filler with surface cross-linked diquaternary ammonium salt groups, which efficiently intercepts free acid in waste liquid and realizes acid recovery.

[0034] 4. The present invention adopts multi-effect parallel flow feeding evaporation crystallization. The feed liquid flows from the previous effect evaporation system to the next effect evaporation system through the pressure difference of the adjacent two-effect evaporation systems or through a pump. Heating steam is added to the first effect evaporation system, and the secondary steam generated from the previous effect evaporation system is used as the heating steam for the next effect evaporation system. Saturated steam is used as the heat source, which effectively enhances the uniformity and efficiency of heat radiation, thereby improving the crystallization efficiency. BRIEF DESCRIPTION OF THE DRAWINGS

[0035] Figure 1 Schematic diagram of the structure of the recovery system of the present invention;

[0036] Figure 2 Schematic diagram of the structure of the regulating tank in the present invention;

[0037] Description of the drawings: 1-homogenizing tank, 2-adsorption column, 3-chromatographic column, 4-liquid separator, 5-circulating tank, 6-regulating tank, 601-tank body, 602-feeding cylinder, 603-pH meter, 604-stirrer, 7-first evaporation tank, 8-first crystallizer, 9-second evaporation tank, 10-second crystallizer, 11-third evaporation tank, 12-third crystallizer, 13-condenser, 14-recovery tank, 15-pH detector, 16-steam concentrator. DETAILED DESCRIPTION

[0038] In order to make the purpose, technical solutions and advantages of the present invention more clearly understood, the present invention is further described in detail below in conjunction with the embodiments. It should be understood that the specific implementation cases described herein are only used to explain the present invention and are not intended to limit the present invention.

[0039] The present invention adopts Figure 1 and Figure 2 The system device shown is used to recover waste acid and waste salt from the wastewater produced by the production of chlorine-based pharmaceutical intermediates.

[0040] The invention discloses a method for recovering waste acid and waste salt from wastewater produced by the production of chlorine-based pharmaceutical intermediates, and the specific steps include S1-S4.

[0041] S1, homogenization;

[0042] The production wastewater from different time periods and different units in the drug preparation process is fed into the homogenization tank 1 for homogenization treatment to overcome the unevenness of wastewater discharge caused by different time periods and process units, and to evenly adjust the changes in the water quality, water volume, water temperature, pH, and pollutant content of the wastewater to make the water intake uniform, thereby reducing the impact of the inconsistency of the wastewater on subsequent treatment facilities.

[0043] S2, impurity removal treatment;

[0044] The supernatant discharged from the drainage trough of the homogenization tank 1 is transported to the adsorption column 2 by a pump, and is adsorbed by the macroporous resin in the adsorption column to remove the residual organic phase and obtain inorganic waste liquid.

[0045] In a specific embodiment, the macroporous resin is a chloromethylated resin with a pore size of 20 nm to 100 nm. In the present invention, the organic phase in the wastewater produced by the production of chlorine-based drug intermediates mainly contains drugs or their intermediates, such as 3-chloropropionic acid, 2,6-dichloroaniline, 3-chloropropyne, adipoyl chloride, 2,5-dichlorohexanoyl chloride, sodium 2,5-dichloroadipate, 2,5-thiophenedicarboxylic acid, etc. The chloromethylated resin has the same polarity as the above-mentioned organic phase impurities, and the adsorption effect is enhanced.

[0046] S3, deacidification and desalination treatment;

[0047] After the impurity removal treatment in S2, the inorganic waste liquid enters the chromatographic column 3 through a pump, and the adsorption filler adsorbs the free acid to obtain a high-salt waste liquid that meets the expected pH value; the concentrated acid is recovered after elution from the chromatographic column;

[0048] The adsorption filler is loaded into the chromatographic column 3 as the stationary phase; the adsorption filler is a microsphere filler with surface cross-linked diquaternary ammonium salt groups, and its preparation steps include:

[0049] (1) Using ethanol or acetone to remove impurities from the microspheres; the microspheres can be any microsphere filler that can adsorb free acid permitted in the art, such as silica gel microspheres or weakly alkaline resins;

[0050] (2) adding a diquaternary ammonium salt solution having a mass concentration of 0.1% to 5% to the microspheres, adjusting the pH to 9 to 11, heating to 70° C. to 80° C., and stirring the reaction for 2 to 4 hours; the molar ratio of the diquaternary ammonium salt to the microspheres is 0.25 to 1:1;

[0051] (3) After cooling, the microsphere filler is separated by filtration under reduced pressure, repeatedly washed, and vacuum dried to obtain the adsorption filler.

[0052] In a specific implementation, the general structural formula of the diquaternary ammonium salt is:

[0053]

[0054] In the formula, R is a C2-C6 alkenyl group; R1, R2, R3, and R4 are each independently selected from a C1-C6 alkyl group. The examples are given for illustration, but are not limited to the listed structures, and other structures not listed are also applicable.

[0055] The specific operations of deacidification and desalination treatment are as follows:

[0056] (1) Equilibration: Before adsorbing the free acid, chromatographic column 1 is equilibrated with a weak base solution (e.g., 1 M sodium hydroxide), and then deionized water is pumped into chromatographic column 3 for solution replacement;

[0057] (2) Sampling: The inorganic waste liquid passes through the chromatographic column 3 at a certain flow rate (generally 10 mL / min to 100 mL / min); the free acid and the inorganic salt are separated. In this process, the main component of the discharged liquid is the inorganic salt. A pH detector 15 is installed on the outlet pipe of the chromatographic column 3. If the pH value of the discharged liquid is greater than or equal to 5, it indicates that most of the free acid is adsorbed, that is, it is a high-salt waste liquid, which can enter the next unit for salt concentration treatment; if the pH value of the discharged liquid is less than 5, it indicates that it still contains a large amount of free acid, and the discharged liquid is discharged into the circulation pool 5, and can be subsequently mixed with the waste liquid in the homogenization pool 1 or directly passed through the chromatographic column 3 again for secondary treatment; in addition, the pH value of the discharged liquid is less than 5, which is generally in the early stage of sample loading or after the chromatographic column 3 is saturated with adsorption. When the adsorption is saturated, the sample loading is terminated;

[0058] (3) Eluent: After loading, elute column 3 with deionized water at a flow rate of 5 mL / min to 50 mL / min.

[0059] (4) Elution: Use a dilute acid solution (such as 1M hydrochloric acid) as the eluent to elute the adsorbed free acid at a flow rate of 10mL / min to 100mL / min. The eluted liquid is a high-concentration acid solution, which realizes the recovery of waste acid.

[0060] S4, salt concentration and crystallization

[0061] (1) The high-salt waste liquid that meets the expected pH value (≥5) first enters the tank body 601 of the regulating tank 6 for mixing, and the pH value is uniformly detected by the pH meter 603 installed therein, and alkali is added thereto by the alkali adding cylinder 602 installed outside the tank body 601, and the agitator 604 stirs the mixed liquid to adjust the pH to 8-10;

[0062] (2) The high-salt waste liquid after treatment in the regulating tank 6 enters the multi-effect evaporation system for evaporation and crystallization to recover the sodium salt crystals;

[0063] The multi-effect evaporation system employs at least a three-effect parallel-flow feed evaporation system. The feed liquid flows from the preceding evaporation system to the succeeding evaporation system through the pressure difference between adjacent two-effect evaporation systems or by a pump. Heating steam (preferably superheated saturated steam) is added to the first-effect evaporation system, and the secondary steam generated from the preceding evaporation system serves as the heating steam for the succeeding evaporation system. The evaporation temperature of the multi-effect evaporation system gradually decreases along the flow direction of the feed liquid. A three-effect or four-effect system is generally preferred. The following embodiments illustrate a three-effect parallel-flow feed evaporation system.

[0064] The workflow of this step is:

[0065] First-effect evaporation system: The high-salt waste liquid is first pumped into the first evaporation tank 7. At the same time, the superheated saturated steam after steam concentration enters the first evaporation tank 7 from the tangential direction, heating the incoming high-salt waste liquid and evaporating the water. After heating, the saturated liquid forms and falls into the first crystallizer 8 below. The temperature is slowly cooled to room temperature and crystallized. In this process, sodium sulfate crystals are mainly precipitated. The cooling crystallization of the first-effect evaporation system adopts step-by-step cooling crystallization because the solubility of sodium sulfate is greatly affected by temperature. The temperature is first reduced to 80℃~90℃ and kept warm until the crystallization stops, then the temperature is reduced to 50℃~60℃ and kept warm until the crystallization stops, and finally the temperature is reduced to room temperature.

[0066] Second-effect evaporation system: After solid-liquid separation in the first crystallizer 8, the mother liquor enters the second evaporator 9 through a pressure difference. Steam is discharged from the top of the first evaporator 7, condensed by the steam concentrator 16, and then enters the second evaporator 9 from a tangential direction. The incoming mother liquor is heated to evaporate water. The vacuum degree of the second evaporator 9 is set to -70kPa to -90kPa, and the temperature is set to 60°C to 70°C. After heating, the saturated liquid falls into the second crystallizer 10 below and is slowly cooled to room temperature. As the water evaporates, crystals precipitate. In this process, sodium chloride crystals are mainly precipitated.

[0067] Third-effect evaporation system: After solid-liquid separation in the second crystallizer 10, the mother liquor is pumped into the third evaporator 11. Steam is discharged from the top of the second evaporator 9, condensed by a steam concentrator 16, and then tangentially enters the third evaporator 11. The incoming mother liquor is heated to evaporate water. The vacuum level in the third evaporator 11 is set to -50kPa to -45kPa, and the temperature is set to 40°C to 50°C. After heating, the saturated liquid falls into the third crystallizer 12 below. As the water evaporates, crystals precipitate. In this process, sodium phosphate crystals primarily precipitate. In actual applications, the sodium phosphate content in wastewater from the production of chlorine-based pharmaceutical intermediates is relatively low, making precipitation difficult. Therefore, a small amount of crystallization promoter, such as sodium phosphate seed crystals, can be added to the third crystallizer 12 to promote the nucleation and crystallization of sodium phosphate in the wastewater. In practice, the crystallization promoter can be added in advance or during the evaporation and crystallization process, with a typical dosage of 0.5g / L to 10g / L.

[0068] Steam is discharged from the top of the third evaporation tank 11 and flows into the condenser 13. The liquid from the condenser 13 and the third crystallizer 12 flows into the recovery tank 14. The pollutant content in the liquid is low.

[0069] In the above-mentioned evaporation system, saturated steam / vapor can enter the corresponding evaporation system simultaneously with the liquid or after the liquid. A small amount of saturated steam / vapor is introduced first to ensure that the corresponding evaporation system reaches the desired temperature. If the temperature of the liquid fluctuates after entering the evaporation system, saturated steam / vapor is introduced simultaneously to maintain the temperature within the preset temperature range. Simultaneously, the steam enters tangentially, creating a mutual impact with the liquid to form a vapor-liquid mixed flow, making the temperature distribution more uniform while preventing crystallization blockage and scaling on the inner wall.

[0070] The present invention uses superheated saturated steam and secondary steam as heat sources, which saves energy on the one hand, and on the other hand forms vapor-liquid impact mixing compared to conventional heating evaporation, does not require stirring devices, and heat dissipation is more uniform.

[0071] Example 1

[0072] The method for recovering waste acid and waste salt from wastewater produced by chlorine-based drug intermediates comprises the following specific steps:

[0073] S1. The production wastewater is fed into the homogenization tank 1 for homogenization treatment for about 4 hours, and the solid sediments appearing in the process are scraped off, and the supernatant enters the drainage tank;

[0074] S2. The supernatant discharged from the homogenization tank 1 is transported to the adsorption column 2 by a pump. The adsorption column 2 is filled with chloromethylated resin (whose pore size is about 30nm~60nm). The supernatant flows through the adsorption column 2 at a flow rate of about 40ml / min. The organic phase impurities therein are adsorbed and retained by the chloromethylated resin, and the inorganic waste liquid is discharged.

[0075] S3, chromatographic column 3 is filled with weakly alkaline resin microspheres with surface cross-linked diquaternary ammonium salt groups as the stationary phase, and the inorganic waste liquid is pumped into the chromatographic column 3 after system balance and replacement treatment. If the pH value of the effluent from the chromatographic column 3 is greater than or equal to 5, the effluent is a high-salt waste liquid and enters the next unit for salt concentration treatment; if the pH value of the effluent is less than 5, it is discharged into the circulation pool 5, and can be subsequently mixed with the waste liquid in the homogenization pool 1 or directly passed through the chromatographic column 3 again for secondary treatment; when the chromatographic column 3 is saturated with adsorption, the sample is loaded; deionized water is used to elute the chromatographic column 3 at a flow rate of 20mL / min; finally, 1M hydrochloric acid is used to elute the adsorbed free acid in the chromatographic column 3 at a flow rate of 30mL / min, and the liquid after elution is a high-concentration acid solution, realizing waste acid recovery. The structural formula of the diquaternary ammonium salt is The molar ratio of the silica gel microspheres is 0.3:1.

[0076] The high-salt waste liquid discharged from S4 and S3 is adjusted to a pH of 8-10 and pumped into the first evaporation tank 7. The superheated saturated steam after steam concentration enters the first evaporation tank 7 from a tangential direction, heats the high-salt waste liquid entering, and evaporates water. The liquid after evaporation and heating falls into the first crystallizer 8 below. The temperature is first reduced to 80°C-90°C and kept warm until the crystals no longer increase, then cooled to 50°C-60°C and kept warm until the crystals no longer increase, and finally cooled to room temperature. The main component of the precipitated crystals is sodium sulfate crystals.

[0077] The vacuum degree in the second evaporation tank 9 is set to -70kPa to -80kPa. The steam discharged from the top of the first evaporation tank 7 is first condensed by the steam concentrator 16 and then enters the second evaporation tank 9 from a tangential direction, so that the internal temperature of the second evaporation tank 9 is preheated to 60°C to 65°C. The mother liquor after solid-liquid separation in the first crystallizer 8 is then passed into the second evaporation tank 9 through the pressure difference. At the same time, the steam discharged from the top of the first evaporation tank 7 is further compressed and then enters the second evaporation tank 9 from a tangential direction to heat and evaporate the mother liquor. The saturated liquid after the evaporation and heating treatment falls into the second crystallizer 10 below and is cooled to room temperature. The precipitated crystals are mainly sodium chloride crystals.

[0078] The vacuum level in the third evaporator 11 is set to -50 kPa to -45 kPa. Steam from the top of the second evaporator 9 is first condensed through a steam concentrator 16 and then tangentially fed into the third evaporator 11, preheating the internal temperature of the second evaporator 9 to 45°C to 50°C. The mother liquor from the second crystallizer 10 is then pumped into the third evaporator 11. Simultaneously, steam from the top of the second evaporator 9 is continuously condensed through a steam concentrator 16 and then tangentially fed into the third evaporator 11, where it is heated and evaporated. The saturated liquid after evaporation and heating falls into the third crystallizer 12 below (sodium phosphate seed crystals are added at a concentration of 0.5 g / L), where it is cooled to room temperature and sodium phosphate crystals are precipitated. Steam from the top of the third evaporator 11 flows into a condenser 13. Liquids from the condenser 13 and the third crystallizer 12 are then fed into a recovery tank 14.

[0079] Example 2

[0080] The method for recovering waste acid and waste salt from the production wastewater of chlorine-based drug intermediates comprises the following specific steps:

[0081] S1. The production wastewater is fed into the homogenization tank 1 for homogenization treatment for about 4 hours, and the solid sediments appearing in the process are scraped off, and the supernatant enters the drainage tank;

[0082] S2. The supernatant discharged from the homogenization tank 1 is transported to the adsorption column 2 by a pump. The adsorption column 2 is filled with chloromethylated resin (whose pore size is about 20nm~40nm). The supernatant flows through the adsorption column 2 at a flow rate of about 10ml / min. The organic phase impurities therein are adsorbed and retained by the chloromethylated resin, and the inorganic waste liquid is discharged.

[0083] S3, the silica gel microspheres with surface cross-linked diquaternary ammonium salt groups filled in the chromatographic column 3 are used as the stationary phase, and the inorganic waste liquid is pumped into the chromatographic column 3 after system balance and replacement treatment. If the pH value of the effluent from the chromatographic column 3 is greater than or equal to 5, the effluent is a high-salt waste liquid and enters the next unit for salt concentration treatment; if the pH value of the effluent is less than 5, it is discharged into the circulation pool 5, and can be subsequently mixed with the waste liquid in the homogenization pool 1 or directly passed through the chromatographic column 3 again for secondary treatment; when the chromatographic column 3 is saturated with adsorption, the sample is loaded; deionized water is used to elute the chromatographic column 3 at a flow rate of 20mL / min; finally, 1M hydrochloric acid is used to elute the adsorbed free acid in the chromatographic column 3 at a flow rate of 30mL / min, and the liquid after elution is a high-concentration acid solution, realizing waste acid recovery. The structural formula of the diquaternary ammonium salt is The molar ratio of the silica gel microspheres is 0.5:1.

[0084] The high-salt waste liquid discharged from S4 and S3 is adjusted to a pH of 8-10 and pumped into the first evaporation tank 7. The superheated saturated steam after steam concentration enters the first evaporation tank 7 from a tangential direction, heats the high-salt waste liquid entering, and evaporates water. The liquid after evaporation and heating falls into the first crystallizer 8 below. The temperature is first reduced to 80°C-90°C and kept warm until the crystals no longer increase, then cooled to 50°C-60°C and kept warm until the crystals no longer increase, and finally cooled to room temperature. The main component of the precipitated crystals is sodium sulfate crystals.

[0085] The vacuum degree in the second evaporation tank 9 is set to -80kPa to -90kPa. The steam discharged from the top of the first evaporation tank 7 is first condensed by the steam concentrator 16 and then enters the second evaporation tank 9 from a tangential direction, so that the internal temperature of the second evaporation tank 9 is preheated to 65°C to 70°C. The mother liquor after solid-liquid separation in the first crystallizer 8 is then passed into the second evaporation tank 9 through the pressure difference. At the same time, the steam discharged from the top of the first evaporation tank 7 is further compressed and then enters the second evaporation tank 9 from a tangential direction to heat and evaporate the mother liquor. The saturated liquid after the evaporated and heated treatment falls into the second crystallizer 10 below and is cooled to room temperature. The precipitated crystals are mainly sodium chloride crystals.

[0086] The vacuum level in the third evaporator 11 is set to -50 kPa to -45 kPa. Steam from the top of the second evaporator 9 is first condensed through a steam concentrator 16 and then tangentially fed into the third evaporator 11, preheating the internal temperature of the second evaporator 9 to 45°C to 50°C. The mother liquor from the second crystallizer 10 is then pumped into the third evaporator 11. Simultaneously, steam from the top of the second evaporator 9 is continuously condensed through a steam concentrator 16 and then tangentially fed into the third evaporator 11, heating and evaporating the incoming mother liquor. The vacuum level in the third evaporator 11 is set to above -90 kPa. The saturated liquid after evaporation and heating falls into the third crystallizer 12 below, where it is cooled to room temperature and sodium phosphate crystals are precipitated. Steam from the top of the third evaporator 11 flows into the condenser 13. The liquids from the condenser 13 and the third crystallizer 12 are then fed into the recovery tank 14.

[0087] Example 3

[0088] The method for recovering waste acid and waste salt from the production wastewater of chlorine-based drug intermediates comprises the following specific steps:

[0089] S1. The production wastewater is fed into the homogenization tank 1 for homogenization treatment for about 4 hours, and the solid sediments appearing in the process are scraped off, and the supernatant enters the drainage tank;

[0090] S2. The supernatant discharged from the homogenization tank 1 is transported to the adsorption column 2 by a pump. The adsorption column 2 is filled with chloromethylated resin (whose pore size is about 50nm~100nm). The supernatant flows through the adsorption column 2 at a flow rate of about 50ml / min. The organic phase impurities therein are adsorbed and retained by the chloromethylated resin, and the inorganic waste liquid is discharged.

[0091] S3, chromatographic column 3 is filled with weakly alkaline resin microspheres with surface cross-linked diquaternary ammonium salt groups as the stationary phase, and the inorganic waste liquid is pumped into the chromatographic column 3 after system balance and replacement treatment. If the pH value of the effluent from the chromatographic column 3 is greater than or equal to 5, the effluent is a high-salt waste liquid and enters the next unit for salt concentration treatment; if the pH value of the effluent is less than 5, it is discharged into the circulation pool 5, and can be subsequently mixed with the waste liquid in the homogenization pool 1 or directly passed through the chromatographic column 3 again for secondary treatment; when the chromatographic column 3 is saturated with adsorption, the sample is loaded; deionized water is used to elute the chromatographic column 3 at a flow rate of 20mL / min; finally, 1M hydrochloric acid is used to elute the adsorbed free acid in the chromatographic column 3 at a flow rate of 30mL / min, and the liquid after elution is a high-concentration acid solution, realizing waste acid recovery. The structural formula of the diquaternary ammonium salt is The molar ratio of the silica gel microspheres is 0.8:1.

[0092] The high-salt waste liquid discharged from S4 and S3 is adjusted to a pH of 8-10 and pumped into the first evaporation tank 7. The superheated saturated steam after steam concentration enters the first evaporation tank 7 from a tangential direction, heats the high-salt waste liquid entering, and evaporates water. The liquid after evaporation and heating falls into the first crystallizer 8 below. The temperature is first reduced to 80°C-90°C and kept warm until the crystals no longer increase, then cooled to 50°C-60°C and kept warm until the crystals no longer increase, and finally cooled to room temperature. The main component of the precipitated crystals is sodium sulfate crystals.

[0093] The vacuum degree in the second evaporation tank 9 is set to -70kPa to -80kPa. The mother liquor after solid-liquid separation in the first crystallizer 8 is first introduced into the second evaporation tank 9 through the pressure difference. The steam discharged from the top of the first evaporation tank 7 is then condensed through the steam concentrator 16 and then enters the second evaporation tank 9 from a tangential direction. The internal temperature of the controller is 60°C to 70°C. The mother liquor is heated and evaporated. The saturated liquid after the evaporation and heating treatment falls into the second crystallizer 10 below and is cooled to room temperature. The precipitated crystals are mainly composed of sodium chloride crystals.

[0094] The vacuum level in the third evaporator 11 is set to -50 kPa to -45 kPa. The mother liquor from the second crystallizer 10 is pumped into the third evaporator 11. The steam discharged from the top of the second evaporator 9 is then condensed through a steam concentrator 16 and then tangentially enters the third evaporator 11. The internal temperature of the third evaporator 11 is controlled at 45°C to 50°C. The incoming mother liquor is heated and evaporated. The saturated liquid after evaporation and heating falls into the third crystallizer 12 below (sodium phosphate seed crystals are added at a concentration of 0.5 g / L). The temperature is then cooled to room temperature, where sodium phosphate crystals are precipitated. The steam discharged from the top of the third evaporator 11 flows into the condenser 13. The liquids from the condenser 13 and the third crystallizer 12 are then fed into the recovery tank 14.

[0095] Comparative Example 1

[0096] In this comparative example 1, the impurity removal of the organic phase by the adsorption column in step S2 is deleted on the basis of example 1, and other reaction conditions and steps are consistent with example 1.

[0097] Comparative Example 2

[0098] In this comparative example 1, the adsorption of free acid by the chromatographic column in step S3 is deleted on the basis of Example 1, and other reaction conditions and steps are consistent with Example 1.

[0099] Comparative Example 3

[0100] In this comparative example 1, based on Example 1, the filler of the chromatographic column in step S3 was adjusted to conventional silica gel microspheres, i.e., silica gel microspheres with uncrosslinked diquaternary ammonium salt groups on the surface. Other reaction conditions and steps were consistent with Example 1.

[0101] Comparative Example 4

[0102] In this comparative example 1, based on Example 1, the uncrosslinked diquaternary ammonium salt groups on the surface of the silica gel microspheres used as the filler of the chromatographic column in step S3 are adjusted to conventional quaternary ammonium salts (illustrated with trimethylamine), and the other reaction conditions and steps are consistent with those of Example 1.

[0103] Comparative Example 5

[0104] In this comparative example 1, based on Example 1, the steam heating in step S4 was replaced with conventional heating, that is, a heater was set in the corresponding evaporation tank, and the steam discharged from the top of the evaporation tank was directly processed in the condenser. The other reaction conditions and steps were the same as those in Example 1.

[0105] The products collected in Examples 1 to 3 and Comparative Examples 1 to 5 were tested for purity, and the yields of waste acid and waste salt were calculated. The results are shown in Table 1.

[0106] in, Where C1 refers to the molar concentration of the free acid after recovery in step S3 (in terms of H + Calculation), V1 refers to the volume of the acid solution after recovery in step S3; C 废 Refers to the molar concentration of acid in the original production wastewater (in H + Calculation); V 废 Refers to the volume of original production wastewater.

[0107] Where n1, n2, and n3 refer to the molar mass of sodium (number of sodium atoms*crystallization mass / crystallization corresponding molecular weight) calculated from the crystals collected in the first crystallizer 10, the second crystallizer 10, and the third crystallizer 12 in step S4, respectively; p1, p2, and p3 refer to the purity of the corresponding crystals, respectively; n 废 Refers to the molar mass of sodium in the original wastewater (in Na + n4 refers to the total molar mass of sodium used to adjust pH during the recovery process and in the crystallization promoter.

[0108] The purity of sodium sulfate crystals, sodium chloride crystals, and sodium phosphate crystals refers to the purity test of the solids obtained by washing and vacuum low-temperature drying the crystals collected in the first crystallizer 10, the second crystallizer 10, and the third crystallizer 12 in step S4 using ion chromatography.

[0109] Table 1

[0110]

[0111]

[0112] As shown in Table 1, the present invention effectively recovers waste acid and waste salt from production wastewater and effectively separates the waste salt. Furthermore, the purity of sodium sulfate, sodium chloride, and sodium phosphate is greater than 90%, suitable for industrial production.

[0113] As can be seen from Example 2 compared with Example 1, along with the increase of the cross-linking degree of the diquaternary ammonium salt group, the spent acid recovery rate increases, and the reason is that its adsorptive power to the free acid increases; in the last effective evaporation system, sodium phosphate seed is not added, and the recovery rate of the sodium salt decreases, and the reason is that there is no crystallization promoter, and the crystallization generation amount of sodium phosphate decreases.

[0114] Comparing Example 3 with Example 1, it can be seen that in the evaporation system, after the steam intake time is adjusted, the purity and yield of sodium sulfate, sodium chloride, and sodium phosphate decrease slightly. The reason is that the steam and the feed liquid enter the evaporation system simultaneously, the two are mixed more evenly, and the heat transfer efficiency is higher.

[0115] Comparison of Comparative Example 1 with Example 1 shows that, before the recovery of waste acid and waste salt, if the organic phase is not removed from the impurities, the recovery efficiency of the acid solution, sodium salt, etc. is greatly reduced. The reason is that the organic phase interferes with the adsorption and crystallization effects of the free acid, resulting in irregular negative effects.

[0116] Comparison of Comparative Example 2 with Example 1 shows that the purity and yield of sodium sulfate, sodium chloride, and sodium phosphate decrease slightly by deleting the free acid adsorption step. The reason is that the free acid adsorption step can not only adsorb the free acid, but also other impurities such as small molecule organic solvents, thereby avoiding their adverse effects on subsequent crystallization.

[0117] Comparison of Comparative Examples 3 and 4 with Example 1 shows that the cross-linked diquaternary ammonium salt groups on the surface of the chromatographic column stationary phase filler in the free acid adsorption step have a much higher adsorption effect on free acid than conventional fillers or cross-linked ordinary quaternary ammonium salt groups.

[0118] Comparison of Comparative Example 5 with Example 1 shows that the crystallization effect of using superheated compressed steam for heating during the evaporation crystallization process is higher than that of conventional heating means.

[0119] The above-described embodiments merely illustrate several implementations of the present invention, and while their descriptions are relatively specific and detailed, they should not be construed as limiting the scope of the present invention. It should be noted that a person skilled in the art would be able to make numerous variations and improvements without departing from the spirit of the present invention, all of which fall within the scope of protection of the present invention. Therefore, the scope of protection of the present invention shall be determined by the appended claims.

Claims

1. A method for recovering waste acid and waste salt from waste water produced by the production of chlorine-based pharmaceutical intermediates, characterized in that: The method comprises the following steps: The homogenized production wastewater is adsorbed by the macroporous resin in the adsorption column to remove the residual organic phase and obtain inorganic waste liquid; The inorganic waste liquid enters the chromatographic column, where the adsorption filler adsorbs the free acid to obtain a high-salt waste liquid that meets the expected pH value; the concentrated acid is recovered after elution from the chromatographic column; The high-salt wastewater that meets the expected pH value enters the multi-effect evaporation system for evaporation and crystallization, and the sodium salt crystals are recovered separately; The multi-effect evaporation system adopts at least a three-effect parallel-flow feeding evaporation system, wherein the feed liquid flows from the previous effect evaporation system to the next effect evaporation system through the pressure difference between adjacent two-effect evaporation systems or through a pump, heating steam is added to the first effect evaporation system, and the secondary steam generated from the previous effect evaporation system is used as the heating steam for the next effect evaporation system; the temperature used for evaporation in the multi-effect evaporation system gradually decreases along the flow direction of the feed liquid; The macroporous resin is chloromethylated resin; the adsorption filler is a microsphere filler with surface cross-linked diquaternary ammonium salt groups.

2. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 1, characterized in that: In the three-effect parallel evaporation system, superheated saturated steam is introduced into the first-effect evaporation system, and the temperature is lowered step by step to room temperature for crystallization, mainly precipitating sodium sulfate crystals; The vacuum degree of evaporation in the second-effect evaporation system is set to -70kPa to -90kPa, and the temperature is set to 60℃ to 70℃, mainly precipitating sodium chloride crystals; The vacuum degree of evaporation in the third-effect evaporation system is set to -50kPa ~ -45kPa, and the temperature is set to 40℃ ~ 50℃, mainly precipitating sodium phosphate crystals.

3. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 2, characterized in that: A crystallization accelerator is added during crystallization in the third-effect evaporation system; the crystallization accelerator is sodium phosphate seed crystal.

4. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 2, wherein: The pH of the high-salt waste liquid is adjusted to 8-10 before entering the multiple-effect evaporation system.

5. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 2, characterized in that: The heating steam or secondary steam enters the first-effect evaporation system or the second-effect evaporation system tangentially after being condensed.

6. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 1, characterized in that: The pore size of the macroporous resin is 20nm to 100nm.

7. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 1, characterized in that: The preparation steps of the adsorption filler include: Use ethanol or acetone to remove impurities from the microspheres; The diquaternary ammonium salt solution was added dropwise to the microspheres, the pH was adjusted to 9-11, the solution was heated to 70-80°C, and the reaction was stirred for 2-4 hours. After cooling, the solid is separated by filtration under reduced pressure, repeatedly washed, and vacuum dried to obtain the adsorption filler.

8. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 7, characterized in that: The mass concentration of the diquaternary ammonium salt solution is 0.1% to 5%; the molar ratio of the diquaternary ammonium salt to the microspheres is 0.25 to 1:

1.

9. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 7, characterized in that: The general structural formula of the diquaternary ammonium salt is: In the formula, R is a C2-C6 alkenyl group; R1, R2, R3, and R4 are each independently selected from a C1-C6 alkyl group.

10. The method for recovering waste acid and waste salt from chlorine-based pharmaceutical intermediate production wastewater according to claim 1, characterized in that: The expected pH value means that the pH value of the high-salt waste liquid is not less than 5; the high-salt waste liquid with a pH value less than 5 is circulated to the chromatographic column for secondary treatment.

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