High-salinity wastewater desalination crystallization process and system

Abstract

The present application relates to the field of industrial wastewater treatment, and particularly relates to a high-salinity wastewater salt separation and crystallization process and system, the process comprising hard and silicon removal treatment, turbidity removal treatment, softening treatment, membrane treatment, ozone oxidation treatment, evaporation and concentration treatment, and freezing crystallization treatment on the wastewater to obtain mirabilite and frozen mother liquor; wherein the mirabilite is subjected to melting crystallization treatment to obtain sodium sulfate product and melting mother liquor; the melting mother liquor is repeatedly subjected to melting crystallization treatment to obtain sodium sulfate product; the frozen mother liquor is subjected to cyclic sodium chloride crystallization treatment to obtain sodium chloride product and sodium chloride mother liquor; and the sodium chloride mother liquor is subjected to cyclic miscellaneous salt crystallization treatment and drum drying to obtain miscellaneous salt product. The process solves the problem of coal chemical industry wastewater crystallization and salt separation, realizes the recycling of coal chemical industry wastewater, and realizes the reduction and recycling of miscellaneous salt.

Description

Technical Field

[0001] This invention relates to the field of industrial wastewater treatment technology, and in particular to a high-salt wastewater desalination crystallization process and system. Background Technology

[0002] Wastewater from industries such as petrochemicals and coal chemicals has high salt content and complex composition, posing a significant threat to wastewater treatment systems. Strict treatment of this wastewater is essential to achieve near-zero discharge. However, the mixed salts generated during treatment are considered hazardous solid waste, resulting in high treatment costs and placing a burden on enterprises.

[0003] To address this issue, early-stage chemical plants aiming for zero emissions needed to replace their existing mixed salt crystallization technology with salt separation crystallization technology to achieve salt resource recovery. Among existing methods, nanofiltration membrane technology and traditional thermal crystallization separation technology are commonly used. However, with continuous technological advancements, a novel low-temperature critical freezing vacuum thermal salt separation crystallization process has emerged.

[0004] However, in practical applications, this new process has several problems: First, the operating temperature of the freeze crystallizer cannot reach the design standard (-4.5℃ to -5.0℃), resulting in a low extraction rate of sodium sulfate. Second, the freeze crystallizer experiences freezing blockage and "crystal bursting," leading to short and unstable operating cycles and frequent thawing. Third, due to the low sodium sulfate extraction rate (only about 95%), a large amount of sodium sulfate enters the sodium chloride crystallization system, resulting in substandard quality of the sodium chloride product salt, making it difficult to meet the requirements of first-grade refined products. These problems not only affect the company's production efficiency but also negatively impact product quality. Summary of the Invention

[0005] The purpose of this invention is to solve the problems of low sodium sulfate extraction rate, unqualified sodium chloride product salt quality, and short operating cycle of the cryo-crystallizer in the low-temperature critical freezing vacuum thermal salt separation crystallization process in the background technology, and to provide a high-salt wastewater salt separation crystallization process and system.

[0006] To achieve the above-mentioned objective, the first aspect of the present invention provides a process for desalination and crystallization of high-salt wastewater, the process comprising the following steps:

[0007] S1: High-salt wastewater is treated to remove hardness and silicon to obtain hardened and siliconized effluent;

[0008] S2: The hardening and silicating effluent is subjected to turbidity removal treatment to obtain turbidity-removed effluent;

[0009] S3: The turbidity-removed effluent is subjected to ion exchange softening treatment to obtain softened effluent;

[0010] S4: The softened effluent is subjected to membrane treatment to obtain membrane-treated effluent;

[0011] S5: The membrane-treated effluent is subjected to ozone oxidation treatment to obtain ozone-oxidized effluent;

[0012] S6: The ozone-oxidized effluent is subjected to evaporation and concentration treatment to obtain concentrated effluent;

[0013] S7: The evaporated and concentrated water is subjected to freeze crystallization treatment to obtain Glauber's salt and freeze mother liquor. The freeze mother liquor is divided into two parts, of which 85% to 90% is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to freeze crystallization treatment.

[0014] S8; The Glauber's salt is subjected to melt crystallization treatment to obtain sodium sulfate product and melt mother liquor, and the melt mother liquor is repeatedly subjected to melt crystallization treatment;

[0015] S9: The frozen mother liquor obtained in step S7 is subjected to sodium chloride crystallization treatment to obtain sodium chloride product and sodium chloride mother liquor. The sodium chloride mother liquor is divided into two parts, of which 5% to 9% is sent to step S10 for mixed salt crystallization treatment, and the remaining 91% to 95% is repeatedly subjected to sodium chloride crystallization treatment.

[0016] S10: The sodium chloride mother liquor obtained in step S9 is subjected to mixed salt crystallization treatment to obtain mixed salt product.

[0017] A second aspect of the present invention provides a system for the above-mentioned high-salt wastewater desalination crystallization process, the system comprising the following units according to the material flow:

[0018] The hardness and silica removal unit is used to remove most of the hardness and silica from high-salt wastewater, resulting in hardness and silica removal effluent.

[0019] The turbidity removal unit is used to remove suspended solids from the hardening and desiliconizing effluent to obtain turbidity-removed effluent;

[0020] The softening unit is used to remove the residual hardness from the turbid effluent to obtain softened effluent;

[0021] The membrane treatment unit is used to remove precipitated colloids, suspended solids, and reduce volume in softened effluent to obtain membrane-treated effluent;

[0022] The ozone oxidation unit is used to remove organic matter from the effluent of membrane treatment, resulting in ozone-oxidized effluent.

[0023] The evaporation and concentration unit is used to concentrate the ozone oxidation effluent to obtain concentrated evaporation effluent;

[0024] The freeze crystallization unit is used to freeze crystallize the concentrated evaporation water to obtain sodium sulfate and freeze mother liquor.

[0025] The melting crystallization unit is used to melt crystallize Glauber's salt to obtain sodium sulfate product and molten mother liquor;

[0026] The sodium chloride crystallization unit is used to crystallize sodium chloride from the frozen mother liquor to obtain sodium chloride product and sodium chloride mother liquor.

[0027] The mixed salt crystallization unit is used to process part of the mother liquor discharged from the sodium chloride crystallization unit to obtain mixed salt products.

[0028] The present invention has the following beneficial effects:

[0029] I. This invention solves the problem of salt separation and crystallization in high-salt coal chemical wastewater, realizing the recycling of coal chemical wastewater, reducing the amount of miscellaneous salts, and utilizing resources in a circular economy, producing sodium sulfate products that meet national first-class standards, and refined industrial salt products of first-class sodium chloride products.

[0030] Second, in this invention, calcium agent, magnesium agent, coagulant, and cation exchanger are used sequentially to treat high-salt wastewater so that the hardness of the resulting effluent is ≤5mg / L.

[0031] Third, in the membrane treatment of the present invention, ultrafiltration removes precipitated colloids and suspended solids from wastewater, reducing turbidity and sludge density, and then reverse osmosis is used to reduce the volume of wastewater, thereby reducing the processing capacity of each crystallization unit during subsequent salt separation and crystallization.

[0032] Fourth, in the evaporation and concentration process of the present invention, the ozone oxidation effluent is evaporated and concentrated by TVR and MVR devices, increasing the dissolved solids content in the brine to 200-230 g / L, thereby further reducing the scale of subsequent sodium sulfate and sodium chloride crystallization.

[0033] V. In the freeze crystallization of the present invention, when the actual salt-to-nitrate ratio is significantly lower than the designed salt-to-nitrate ratio of the freeze crystallizer, the chloride ion content in the pre-cooled concentrate is increased by increasing the degree of evaporation and concentration, replacing the step of adjusting the pH value with sulfuric acid with hydrochloric acid, and refluxing the sodium chloride mother liquor to the freeze crystallization unit. This increases the actual feed salt-to-nitrate ratio, bringing it closer to the designed ratio. This lowers the operating temperature of the freeze crystallizer to -5.5℃ to -4.5℃, while avoiding freeze blockage of the freeze crystallizer when the temperature of the pre-cooled concentrate drops to 0℃, ultimately improving the extraction rate of sodium sulfate (or sodium sulfate).

[0034] VI. In the freeze crystallization of the present invention, the heat exchange temperature difference between the precooled concentrated water and the heat exchange medium (i.e., ethylene glycol solution) is controlled to be ≤3℃ and the temperature difference of the heat exchange medium before and after heat exchange is controlled to be ≤1.5℃, so that the low-temperature crystallization process can be carried out continuously and stably, which ultimately greatly improves the extraction rate of sodium sulfate and thus improves the quality of sodium chloride products.

[0035] VII. In the melt crystallization process of the present invention, sodium sulfate produced by freeze crystallization is melt crystallized to obtain sodium sulfate product conforming to GB / T 6009-2014.

[0036] 8. In the sodium chloride crystallization of the present invention, the freezing mother liquor is crystallized by removing volatile substances, circulating flash crystallization and recrystallization, which can produce sodium chloride products that meet the requirements of first grade refined industrial salt.

[0037] 9. In the mixed salt crystallization of this invention, the sodium chloride mother liquor is crystallized using a circulating flash crystallization and rotary drum drying method, which can achieve zero wastewater discharge. The generated mixed salt product can be dissolved and returned to the sodium sulfate crystallization unit, thereby further improving the recovery rate of sodium chloride and sodium sulfate crystals and reducing the amount of mixed salt generated in the unit. Attached Figure Description

[0038] Figure 1 This is a schematic diagram of the high-salt wastewater desalination and crystallization process according to an embodiment of the present invention. Detailed Implementation

[0039] The present invention will now be described in detail with reference to specific embodiments, but this description is not intended to limit the scope of protection of the present invention. Those skilled in the art should understand that certain process steps and equipment used in the specific embodiments of the present invention can be appropriately adjusted or replaced without affecting the implementation of the present invention.

[0040] The first aspect of this invention provides a process for desalination and crystallization of high-salt wastewater, the process comprising the following steps:

[0041] S1: High-salt wastewater is treated to remove hardness and silicon to obtain hardened and siliconized effluent;

[0042] S2: The hardening and silicating effluent is subjected to turbidity removal treatment to obtain turbidity-removed effluent;

[0043] S3: The turbidity-removed effluent is subjected to ion exchange softening treatment to obtain softened effluent;

[0044] S4: The softened effluent is subjected to membrane treatment to obtain membrane-treated effluent;

[0045] S5: The membrane-treated effluent is subjected to ozone oxidation treatment to obtain ozone-oxidized effluent;

[0046] S6: The ozone-oxidized effluent is subjected to evaporation and concentration treatment to obtain concentrated effluent;

[0047] S7: The evaporated and concentrated water is subjected to freeze crystallization treatment to obtain Glauber's salt and freeze mother liquor. The freeze mother liquor is divided into two parts, of which 85% to 90% is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to freeze crystallization treatment.

[0048] S8; The Glauber's salt is subjected to melt crystallization treatment to obtain sodium sulfate product and melt mother liquor, and the melt mother liquor is repeatedly subjected to melt crystallization treatment;

[0049] S9: The frozen mother liquor obtained in step S7 is subjected to sodium chloride crystallization treatment to obtain sodium chloride product and sodium chloride mother liquor. The sodium chloride mother liquor is divided into two parts, of which 5% to 9% is sent to step S10 for mixed salt crystallization treatment, and the remaining 91% to 95% is repeatedly subjected to sodium chloride crystallization treatment.

[0050] S10: The sodium chloride mother liquor obtained in step S9 is subjected to mixed salt crystallization treatment to obtain mixed salt product.

[0051] In this scheme, the high-salinity wastewater refers to wastewater with a total salt content of 1 wt% (equivalent to 10 g / L). Its main components are sulfates and sodium salts among inorganic salts. This wastewater mainly originates from petrochemical, coal chemical plants, and natural gas extraction and processing. This high-salinity wastewater is usually weakly alkaline and is characterized by high salt content, hardness, silicon content, and organic matter content.

[0052] Furthermore, the high-salinity wastewater typically contains Na. + K + Ca 2+ Mg 2+ Fe 3+ Mn 2+ Al 3+ NH4 + Cu 2+ Ba 2 + 、Sr 2+ As 3+ Isocations and Cl - SO4 2- HCO3 - CO3 2- OH - NO3 - NO2 - PO4 3- F - S 2- SiO3 2- Anions.

[0053] In step S1, the hardening and silicon removal process includes:

[0054] (S1a) Add calcium agent to the high-salt wastewater to obtain the first clarified effluent;

[0055] (S1b) Add magnesium agent to the first clarified effluent to obtain the second clarified effluent;

[0056] (S1c) Add coagulant, sodium carbonate and coagulant aid to the second clarified effluent to obtain hardness- and silica-removed effluent.

[0057] Preferably, the high-salt wastewater can undergo homogenization operations well known in the art before hardening and desiliconization treatment. Its main functions are twofold: first, to regulate the water volume and store the wastewater to ensure stable water intake after the device is in operation; and second, to regulate the water quality by adjusting various different water qualities to stabilize the influent water quality.

[0058] As is well known to those skilled in the art, the hardness in the removal of hardness and silicon refers to the sum of the concentrations of various metal ions (such as calcium and magnesium ions) in water, and the silicon in the removal of hardness and silicon refers to silicon existing in the form of silicates, colloids, or particles.

[0059] In some embodiments, in step (S1a), the calcium agent is calcium chloride and / or calcium hydroxide, the purpose of which is to remove F from high-salinity wastewater. - CO3 2- PO4 3- Calcium is an insoluble anion.

[0060] Furthermore, sodium hydroxide may be added to the calcium agent to adjust the pH of the reaction between the calcium agent and the high-salinity wastewater to 11-11.5, thereby removing most of the hardness, such as Fe, from the high-salinity wastewater. 3+ Mn 2+ Al 3+ Cu 2+ 、Sr 2+ Plasma.

[0061] Further, the amount of calcium agent added is 0.8-1.5 g / L; further, the reaction time of the calcium agent with high-salt wastewater is 5-60 min, preferably 15-30 min.

[0062] In some embodiments, the hardness of the first clarified effluent is ≤65mg / L.

[0063] In some embodiments, in step (S1b), the magnesium agent is at least one of magnesium oxide, magnesium chloride, and magnesium sulfate, with the aim of removing silicates from high-salt wastewater.

[0064] Furthermore, the amount of magnesium agent added is 0.8–2.0 g / L.

[0065] In some embodiments, sulfuric acid or hydrochloric acid may be added to the magnesium agent to make the pH of the magnesium agent react with the first clarified effluent at 7-8;

[0066] Furthermore, considering that the freeze crystallizer used in the following freeze crystallization treatment needs to meet the feed salt-to-nitrate ratio (the ratio of sodium chloride to sodium sulfate) of 3:7 to 4:6 in its design, otherwise the freeze crystallizer is prone to freezing and clogging, the choice between sulfuric acid and hydrochloric acid when adjusting the pH can be based on the concentrations of Cl- and SO42- in the wastewater. 2- The relative content is used to determine this.

[0067] In some embodiments, in step (S1b), the reaction time between the first clarified effluent and the magnesium agent is 5 to 60 minutes, preferably 15 to 30 minutes.

[0068] Those skilled in the art will understand that adding magnesium to the first clarified effluent can generate new magnesium hydroxide precipitate, which can create a high suspended solids environment, maintain a high adsorption specific surface area, adsorb silicon in precipitate and colloidal form in wastewater, and convert magnesium hydroxide into magnesium silicate precipitate with lower solubility, thereby achieving the purpose of silicon removal.

[0069] In some embodiments, the silicon content of the second clarified effluent is ≤40 mg / L.

[0070] In some embodiments, in step (S1c), the coagulant is selected from 30% polyferric sulfate solution, 30% polyferrous sulfate solution, or 30% polyferric chloride solution, and the amount added is 0.1-0.25 g / L.

[0071] The coagulant is PAM (polyacrylamide), and the dosage is 0.2-0.5 mg / L;

[0072] The amount of sodium carbonate added is 0.1 to 0.5 g / L.

[0073] Those skilled in the art will understand that the coagulant can destabilize the colloids in the second clarified effluent; the coagulant aid can enhance the adsorption effect of the flocs and remove suspended solids from the second clarified effluent; and the sodium carbonate is used to remove calcium hardness from the second clarified effluent.

[0074] Furthermore, the reaction time of the second clarified effluent with the coagulant, sodium carbonate, and coagulant aid is 5 to 60 minutes, preferably 15 to 30 minutes.

[0075] In some embodiments, steps (S1a) to (S1c) can be carried out in a reaction sedimentation tank and a V-type filter tank. After steps (S1a) to (S1c), calcium hardness, magnesium hardness, silicates, fluorides and their complexes in the high-salt wastewater are removed, and the hardness of the hardness and silicate removal effluent is ≤50mg / L.

[0076] In step S2, the purpose of the turbidity removal treatment is to remove the insoluble suspended solids in the dehardening and desiliconizing effluent to obtain turbidity-removed effluent.

[0077] Furthermore, the turbidity removal treatment method involves sedimentation in a high-density sedimentation tank, under the following conditions: the upward flow velocity of water in the pipe area is 0.1–0.2 m / s, and the sludge solids loading is 11–12 kg / (m³). 2 The linear velocity of the outer edge of the scraper is 0.03–0.04 m / s (h).

[0078] In some embodiments, the high-density sedimentation tank is equipped with inclined plates, which can increase the hydraulic retention time and ultimately improve the turbidity removal efficiency.

[0079] Furthermore, the turbidity of the effluent after turbidity removal is ≤5 NTU.

[0080] In step S3, the softening treatment involves using a hydrogen-type weak acid cation exchanger to perform deep hardening removal on the turbid water to remove cations other than Na+ and K+, resulting in ion-exchanged water.

[0081] Furthermore, the hardness of the softened water is ≤5mg / L.

[0082] In some embodiments, the hydrogen-type weak acid cation exchanger includes a primary weak acid cation exchanger and a secondary weak acid cation exchanger.

[0083] Furthermore, the regeneration cycle of the primary weak acid cation exchanger is 24 hours, and the regeneration cycle of the secondary weak acid cation exchanger is 48 hours.

[0084] In some embodiments, the ion-exchange effluent obtained in step S3 can be adjusted to an acidic pH using sulfuric acid or hydrochloric acid, and then subjected to decarbonization treatment to reduce the residual CO2 in the ion-exchange effluent to below 5 mg / L, before proceeding to step S4.

[0085] Furthermore, the step of adjusting the pH to acidic can prevent the brine from causing alkaline corrosion to the evaporation equipment of the evaporation and concentration unit.

[0086] In some embodiments, the decarbonization treatment uses a decarbonator in the art, which removes free CO2 from the water by blowing air for degassing. Water is introduced from the top of the decarbonator, flows over the surface of the packing layer via a spray device, and air enters from the bottom vent and passes through the packing layer in the opposite direction. The free CO2 in the water is rapidly desorbed and enters the air, and is discharged from the top of the decarbonator.

[0087] Furthermore, the carbon remover is installed after the hydrogen-type weak acid cation exchanger. After the ion exchange effluent is degassed by the carbon remover, it flows into the product water tank of the softening unit. At this time, the residual CO2 in the water is reduced to ≤5mg / L.

[0088] In some embodiments, the softened effluent after the decarbonization treatment has the following characteristics: pH = 6-9, hardness ≤ 5 mg / L, SiO2 content ≤ 40 mg / L, and alkalinity ≤ 80 mg / L.

[0089] In step S4, the membrane treatment includes:

[0090] (S4a) The softened effluent is subjected to ultrafiltration treatment to obtain ultrafiltration effluent;

[0091] (S4b) The ultrafiltration effluent is subjected to reverse osmosis treatment to obtain membrane-treated effluent.

[0092] In some embodiments, in step (S4a), the ultrafiltration process may use an ultrafiltration device or system well known in the art, such as a spiral wound membrane module made of PVDF material.

[0093] Furthermore, the conditions for the ultrafiltration treatment are: inlet water pressure ≤ 0.3 MPa, transmembrane pressure difference ≤ 0.16 MPa, and recovery rate ≥ 95%.

[0094] In some embodiments, the ultrafiltration process can remove precipitated colloids and suspended solids from the softened effluent, reducing turbidity and sludge density, thereby improving the efficiency of subsequent reverse osmosis treatment.

[0095] Furthermore, the dissolved solids content (i.e., salt content) of the ultrafiltration effluent is 10-14 g / L, the turbidity of the ultrafiltration effluent is ≤0.2 NTU, and the pollution index (SDI) of the ultrafiltration effluent is ≤3.

[0096] In some implementations, in step (S4b), the reverse osmosis treatment can use reverse osmosis apparatus or systems well known in the art, such as reverse osmosis spiral wound membrane modules from DuPont, Toray, or Hydranautics.

[0097] Furthermore, the reverse osmosis treatment conditions are: inlet water pressure ≤ 3.0 MPa, inter-stage pressure ≤ 0.3 MPa, and recovery rate ≥ 75%.

[0098] In some embodiments, the reverse osmosis treatment can concentrate and reduce the volume of the ultrafiltration effluent, which, after being processed by the reverse osmosis membrane module, is separated into fresh water (which has passed through the reverse osmosis membrane) and concentrated water (which has not passed through the reverse osmosis membrane, i.e., reverse osmosis effluent).

[0099] Furthermore, the dissolved solids content (i.e. salt content) of the fresh water is 0.2 g / L, accounting for 75% of the total feed water of the reverse osmosis unit or system. It can be reused in the production unit as high-quality reclaimed water (e.g., for dissolving Glauber's salt, recrystallizing sodium chloride, or dissolving and reusing miscellaneous salts).

[0100] Furthermore, the concentrated water (i.e., reverse osmosis effluent) has a dissolved solids content (i.e., salt content) of 41-45 g / L, accounting for 25% of the total feed water entering the reverse osmosis device or system, and it enters step S5 as the final membrane treatment effluent.

[0101] Furthermore, the conductivity of the effluent from the membrane treatment is ≤200 μS / cm. 2 .

[0102] In some implementations, to reduce the size of subsequent crystallization units, the reverse osmosis recovery rate (freshwater production / total feedwater volume) of the reverse osmosis treatment needs to be controlled at ≥75%, and the system desalination rate (salt content of freshwater / salt content of total feedwater) needs to be controlled at ≥98%.

[0103] In some implementations, the reverse osmosis process may involve adding different agents (such as reducing agents, scale inhibitors, non-oxidizing bactericides, hydrochloric acid, sodium hydroxide, etc.) to the ultrafiltration effluent to reduce the cleaning frequency of the reverse osmosis unit and extend the life of the reverse osmosis membrane, depending on changes in the operating status.

[0104] In step S5, the ozone oxidation treatment involves using ozone catalytic oxidation technology to remove organic matter from the membrane treatment effluent, while simultaneously decomposing macrocyclic or long-chain molecules into easily decomposable small molecules, reducing (or even removing) the color of the softened effluent, thus obtaining ozone-oxidized effluent.

[0105] Furthermore, the catalyst used in the ozone catalytic oxidation technology can be an aluminum-based ozone catalyst, a silicon-aluminum ozone catalyst, or a silicon-aluminum-titanium ozone catalyst.

[0106] In some embodiments, the conditions for performing the ozone oxidation treatment are: an influent flow rate of 90–114 m³ / h. 3 The total ozone production is 60 kg / h, the ozone concentration is 148 mg / L (10 wt%), and the reaction time is 20–50 min.

[0107] Furthermore, the removal rate of total organic matter (COD) in the ozone oxidation treatment is above 40%, and the COD of the ozone oxidation effluent is ≤200mg / L.

[0108] In step S6, the evaporation and concentration process includes the following steps:

[0109] (S6a) Adjust the pH of the ozone-oxidized effluent to 5.5–6.0;

[0110] (S6b) Heat the solution obtained in (S6a) to a temperature close to the boiling point;

[0111] (S6c) The solution obtained in (S6b) is subjected to deoxygenation treatment to obtain deoxygenated effluent;

[0112] (S6d) The deoxygenated effluent is evaporated and concentrated to obtain evaporated and concentrated effluent.

[0113] In some implementations, the purpose of pH adjustment in step (S6a) is to prevent scale buildup in the evaporator during the evaporation and concentration process.

[0114] Furthermore, the pH adjustment can be achieved using sulfuric acid or hydrochloric acid, depending on the pH level in the wastewater after ozone oxidation. - and SO4 2- The relative content is used to determine the final feed salt-nitrate ratio during the following freeze-crystallization process, so that it is close to its design value.

[0115] In some embodiments, step (S6b) is a preheating step before the evaporation and concentration, which can reduce the solubility of oxygen and other non-condensable gases, and reduce the steam load of the deaerator during the deoxygenation process, thus ensuring the deaerator's operating performance.

[0116] In some embodiments, the purpose of deoxygenation in step (S6c) is to remove non-condensable gases from the water and to reduce the corrosion of the sodium chloride crystallization unit by Cl- during subsequent sodium chloride crystallization.

[0117] Furthermore, the deoxygenation treatment is carried out in a deaerator, specifically by spraying the solution obtained in (S6b) through a nozzle and allowing it to flow downwards in stages within the deaerator, while simultaneously contacting the upward-flowing steam to obtain the deoxygenated effluent.

[0118] Furthermore, the deoxygenation treatment is carried out in a deaerator. The solution obtained in (S6b) is sprayed through a nozzle and flows downward in stages in the deaerator, while contacting the steam flowing upward in the opposite direction. The removed gas is discharged to the atmosphere along with a small amount of vented steam, and finally deoxygenated effluent is obtained.

[0119] In some embodiments, during the evaporation and concentration in step (S6d), the deoxygenated effluent is divided into two streams and evaporated and concentrated in an MVR evaporation device and a TVR evaporation device, respectively, in this field. Finally, the concentrated liquids are combined into one stream, which is the evaporation and concentration effluent of step (S6d).

[0120] Furthermore, the TVR evaporator operates at a temperature of 98℃~99.5℃ and a pressure of ≤50kPa; the MVR evaporator operates at a temperature of 98℃~99.5℃ and a pressure of ≤50kPa.

[0121] Furthermore, the evaporation and concentration method can concentrate the ozone oxidation effluent by more than 5 times, meaning that the amount of evaporated and concentrated effluent is only 20% of the corresponding influent volume. At the same time, it can generate reusable steam condensate, reducing the scale of the subsequent sodium sulfate crystallization unit and sodium chloride crystallization unit.

[0122] Furthermore, the dissolved solids content in the evaporated and concentrated water is 200–230 g / L.

[0123] In step S7, the freeze crystallization includes:

[0124] (S7a) The evaporated and concentrated water is mixed with the frozen mother liquor that has been repeatedly subjected to freezing and crystallization treatment, and the mixture is pre-cooled to obtain pre-cooled concentrated water at a temperature of 28-40°C.

[0125] (S7b) The pre-cooled concentrated water is subjected to freeze crystallization to obtain a frozen crystal slurry with a temperature of -5.5 to -4.5℃;

[0126] (S7c) The frozen crystal slurry is subjected to solid-liquid separation to obtain sodium sulfate and frozen mother liquor. The frozen mother liquor is divided into two parts, of which 85% to 90% is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10% to 15% is repeatedly subjected to frozen crystallization treatment.

[0127] In some embodiments, before performing the pre-cooling treatment in step (S7a), the mixture obtained after mixing can be concentrated to improve the efficiency of freeze crystallization.

[0128] In some embodiments, during the pre-cooling process in step (S7a), the temperature of the mixture obtained after mixing is slowly reduced to 28-40°C to avoid rapid crystallization caused by rapid cooling, which could clog the heat exchange equipment.

[0129] Furthermore, if the temperature of the precooling concentrate is too low, it will cause blockage of the freeze crystallizer; if the temperature is too high, the freeze crystallization effect will not be ideal. The optimal temperature of the precooling concentrate is 28±0.2℃.

[0130] In some embodiments, in step (S7b), the freeze crystallization is carried out in a freeze crystallizer of the art, such as a crystal slurry internal circulation crystallizer, which has an internal guide tube to form a circulation channel, which is beneficial for continuous crystallization.

[0131] Furthermore, the cryo-crystallizer is preferably a DTB-type crystallizer, which is equipped with a guide tube. The circulating slurry enters the cryo-crystallizer from the top of the guide tube, and the circulating slurry comes into contact with the crystals in the cryo-crystallizer, causing the crystals to gradually grow. Large particles of crystalline salt settle in the settling zone around the annular baffle of the guide tube, while small particles of salt enter the circulation pipe with the circulating slurry to continue crystallizing. Preferably, the dissolved solids in the cryo-crystallizer are controlled to be below 15%.

[0132] Furthermore, the heat exchange medium of the cryogenic crystallizer is an ethylene glycol solution (also known in process as chilled water).

[0133] In some implementations, during step (S7b), if the heat exchange temperature difference is too large (exceeding 3°C) when the precooled concentrate is subjected to freeze crystallization, rapid crystal nucleation will occur, i.e., "crystal bursting". This results in a short operating cycle of the freeze crystallizer and the need for frequent thawing. Therefore, it is necessary to control the temperature difference between the precooled concentrate and the ethylene glycol solution during heat exchange to be ≤3°C.

[0134] Furthermore, an excessively large temperature difference between the precooling concentrate and the ethylene glycol solution will cause the precooling concentrate to cool rapidly, leading to the rapid precipitation of sodium sulfate crystals in the precooling concentrate, resulting in a "crystallization" phenomenon that ultimately blocks the cooler of the freeze crystallizer. On the other hand, if the temperature difference between the precooling concentrate and the ethylene glycol solution is too small, a larger amount of ethylene glycol solution is required for crystallization, which will increase the operating load of the refrigeration unit.

[0135] The inventors discovered that by adding an internal circulation pump to the ethylene glycol solution side, the heat-exchanged ethylene glycol solution is returned to the inlet of the freeze crystallizer, mixed with the ethylene glycol solution from the refrigeration unit, and then exchanged with the pre-cooled concentrate. This can avoid an excessive temperature difference between the pre-cooled concentrate and the ethylene glycol solution.

[0136] Furthermore, the circulation volume of the internal circulation pump is 850–950 m³. 3 / h.

[0137] The above measures, on the one hand, appropriately increase the flow rate of the ethylene glycol solution entering the cryogenic crystallizer, reducing the temperature difference between the ethylene glycol solution and the precooled concentrate during heat exchange; on the other hand, the ethylene glycol solution is circulated, ensuring the heat exchange effect with the precooled concentrate. These two aspects ensure that the crystallization process of sodium sulfate takes place in the metastable region, providing good conditions for its crystal growth, realizing a continuous and stable low-temperature crystallization process, and making it less likely for "crystallization explosion" to occur.

[0138] In some embodiments, the temperature difference before and after heat exchange of the precooled concentrate is ≤1.5℃, and preferably the temperature difference before and after heat exchange of the precooled concentrate is ≤0.5℃.

[0139] In some embodiments, in step (S7b), the present invention utilizes the solubility changes of sodium sulfate and sodium chloride in the solution at different temperatures. By cooling through heat exchange between brine and ethylene glycol solution, sodium sulfate precipitates from the solution in the form of sodium sulfate decahydrate crystals. When the temperature is reduced to -5°C, the precipitation rate of sodium sulfate is above 98%.

[0140] In some embodiments, the required salt-to-nitrate ratio (the ratio of sodium chloride to sodium sulfate) for the feed in the designed cryo-crystallizer is 3:7 to 4:6. The inventors have found that when the actual feed salt-to-nitrate ratio is too low (especially below 2.8:7.2), freezing blockage easily occurs when the temperature of the pre-cooled concentrate during cryo-crystallization drops below -2°C. However, at this time, the temperature for sodium sulfate crystal precipitation (-5°C) has not yet been reached, and effective separation of Glauber's salt and cryo-mother liquor cannot be achieved.

[0141] In other words, when the actual feed salt-to-nitrate ratio is too low, the operating temperature of the freeze crystallizer cannot be reduced to the design temperature (-5.0℃~-4.5℃), which ultimately leads to a decrease in the sodium sulfate extraction rate and the inability to produce qualified sodium chloride products.

[0142] In some embodiments, with the solution of the present invention, the salt-to-nitrate ratio of the freeze crystallizer designed for freeze crystallization at -5.0℃ to -4.5℃ can be reduced from the original (3:7) to (4:6) to a minimum of 2.8:7.2.

[0143] The inventors have taken the following measures: First, the degree of evaporation and concentration in step S6 is increased, and the TDS salt content in the concentrated water is controlled at 200-230 g / L (significantly higher than the 140-150 g / L in the prior art); Second, the water quality of the feed to the freeze crystallizer is sampled and analyzed. Based on the degree to which the water quality deviates from the designed salt-nitrate ratio (2.8:7.2), the step of adjusting the pH value with sulfuric acid can be replaced with the use of hydrochloric acid to increase the content of Cl- ions in the corresponding brine. Alternatively, in step S9, part of the sodium chloride mother liquor is refluxed to the freeze crystallization unit to further increase the chloride ion content in the pre-cooled concentrate.

[0144] Therefore, the above measures ensure that the actual salt-to-nitrate ratio of the feed to the cryo-crystallizer is within the designed ratio, reduce the operating temperature of the cryo-crystallizer to -5.5℃ to -4.5℃, avoid freezing blockage of the cooler in the cryo-crystallizer after the brine temperature drops to 0℃, and improve the extraction rate of Glauber's salt.

[0145] In some embodiments, in step (S7c), the solid-liquid separation is preferably performed by centrifugation, and the SO4 in the frozen mother liquor... 2- Ion content ≤5500mg / L.

[0146] In some embodiments, in step (S7c), the frozen crystal slurry may be thickened using a hydrocyclone separator before centrifugation to increase the sodium sulfate concentration and stabilize centrifuge operation.

[0147] Furthermore, the supernatant obtained after thickening is returned to the cryogenic crystallizer, and the resulting bottom thickened liquid is fed into a centrifuge to achieve solid-liquid separation.

[0148] In some embodiments, if the frozen mother liquor obtained by centrifugation contains a small amount of suspended solids (i.e., concentrated Glauber's salt crystals), a sedimentation device can be set up to separate the suspended solids from the frozen mother liquor, and then the suspended solids are sent to step S7 for repeated freeze crystallization treatment.

[0149] The above two methods can significantly reduce the content of sodium sulfate in the freezing mother liquor, thereby ensuring the recovery rate of the sodium chloride crystallization system and the quality of the sodium chloride product.

[0150] Furthermore, the frozen mother liquor is divided into two parts, of which 85%-90% is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10%-15% is repeatedly subjected to frozen crystallization treatment.

[0151] In some implementations, in step (S7c), the mother liquor for repeated freeze-crystallization is used both as a reflux to prevent crystal bursting and to balance the pressure during the freeze-crystallization process.

[0152] In existing nanofiltration membrane-based salt separation and crystallization techniques, the selective permeability of the nanofiltration membrane allows Cl- to pass through while SO42- passes through. 2- Remaining on the feed side, however SO4 2- Ultimately, it's impossible for all salts to be retained by a nanofiltration membrane; some will inevitably permeate through the membrane and enter the sodium chloride side, resulting in a short service life for the nanofiltration membrane. Furthermore, because water contains organic matter, the nanofiltration membrane becomes clogged during separation, further shortening its lifespan. However, the low-temperature critical freezing crystallization solubility method of this invention, which separates sodium sulfate and sodium chloride from brine, achieves a high sodium sulfate recovery rate, minimizes the generation of impurities, and exhibits strong resistance to water quality fluctuations, overcoming the degradation and instability of nanofiltration membrane salt separation efficiency.

[0153] In step S8, the melt crystallization process includes:

[0154] (S8a) Water vapor is passed into the Glauber's salt to dissolve the Glauber's salt into a supersaturated sodium sulfate solution;

[0155] (S8b) The supersaturated sodium sulfate solution is mixed with the molten mother liquor that has been repeatedly melt-crystallized, and then flash-evaporated to obtain a concentrated sodium sulfate slurry;

[0156] (S8c) The sodium sulfate concentrated salt slurry is subjected to solid-liquid separation to obtain anhydrous sodium sulfate crystals and molten mother liquor;

[0157] (S8d) The anhydrous sodium sulfate crystals are dried to obtain sodium sulfate product;

[0158] In some embodiments, in step (S8a), the supersaturated sodium sulfate solution contains 15 wt.% undissolved sodium sulfate crystals.

[0159] In some embodiments, in step (S8b), before performing the flash evaporation, it is preferable to heat the resulting mixture to near the boiling point.

[0160] Furthermore, the flash evaporation temperature is 96–110°C, and the atmospheric flash evaporation pressure is 1.00–5.00 kPa.

[0161] In some embodiments, in step (S8c), the solid-liquid separation is preferably performed by centrifugation;

[0162] Furthermore, the molten mother liquor is returned to step (S8b) and mixed with a supersaturated sodium sulfate solution for repeated melt crystallization treatment;

[0163] Furthermore, the moisture content of the anhydrous sodium sulfate crystals is ≤5%.

[0164] In some embodiments, in step (S8d), the sodium sulfate product has a sodium sulfate content ≥99.0g / 100g, a water-insoluble matter content ≤0.05g / 100g, a chloride content ≤0.35g / 100g, a TOC content ≤50mg / kg, a calcium and magnesium ion content ≤0.3g / 100g, and a moisture content ≤0.2%. That is, the various indicators of the sodium sulfate product at least meet the requirements of Class II Grade 1 in "GBT 6009-2014 Industrial Anhydrous Sodium Sulfate".

[0165] In step S9, the sodium chloride crystallization process includes:

[0166] (S9a) The frozen mother liquor obtained in step S7 is mixed with the sodium chloride mother liquor that has been repeatedly subjected to sodium chloride crystallization treatment, and then flash evaporated to obtain a concentrated sodium chloride slurry.

[0167] (S9b) The concentrated sodium chloride slurry is subjected to solid-liquid separation to obtain crude wet sodium chloride salt and sodium chloride mother liquor;

[0168] (S9c) The crude wet sodium chloride salt is recrystallized and dried to obtain sodium chloride product.

[0169] In some embodiments, the temperature of the frozen mother liquor obtained from step S7 is raised to 75℃±0.5℃ before sodium chloride crystallization. This has the advantage of reducing the heat exchange load on the subsequent sodium chloride crystallizer during sodium chloride crystallization and ensuring that the operating temperature of the sodium chloride crystallizer is stable at 110℃±1℃.

[0170] In some embodiments, prior to the mixing in step (S9a), a volatile matter removal tower can be used to remove volatile small-molecule organic matter from the feed, while also removing carbon dioxide, oxygen, and other non-condensable gases. This reduces the impact of organic matter on the quality of sodium chloride crystals and mitigates bubbling caused by organic matter accumulation in the sodium chloride crystallizer. Simultaneously, deoxygenation reduces the corrosion of equipment and pipelines by the brine.

[0171] Furthermore, the volatile substances are small-molecule organic compounds obtained through ozone oxidation treatment. The removal method can be to heat the freezing mother liquor with low-pressure steam and then use flash evaporation to separate the volatile organic compounds.

[0172] In some embodiments, in step (S9a), before the flash evaporation, the mixture obtained by mixing can be heated to near the boiling point.

[0173] In some embodiments, in step (S9a), the frozen mother liquor obtained in step S7 and the sodium chloride mother liquor repeatedly subjected to sodium chloride crystallization are concentrated by flash evaporation. Because of the supersaturation, the continuously precipitating crystallized salt gradually falls from the cone of the sodium chloride crystallizer into the salt leg. The washing brine in the salt leg enters from the bottom and slowly rises along the leg wall to countercurrently wash the salt crystals, carrying fine salt particles into the crystallizer to continue participating in circulation and crystallization. At the same time, soluble impurities are dissolved and removed, which is beneficial to improving the quality of the crystallized salt.

[0174] Furthermore, sight glasses can be installed at different heights of the salt leg to facilitate observation of the sedimentation of crystalline salt particles.

[0175] In some embodiments, in step (S9a), the flash temperature is 95°C to 100°C, and the flash pressure based on atmospheric pressure is -30 to 5 kPa.

[0176] In some embodiments, in step (S9b), the sodium chloride slurry is preferably thickened using a hydrocyclone separator before solid-liquid separation until the concentration of suspended solids in the sodium chloride slurry increases to more than 40 wt.%.

[0177] Furthermore, the solid-liquid separation method is preferably centrifugation, the moisture content of the crude wet sodium chloride salt is 5 wt.%, and the chloride ion content in the sodium chloride mother liquor is ≤120 g / L.

[0178] In some embodiments, in step (S9b), the sodium chloride mother liquor is sent to step S10 for mixed salt crystallization, and / or sent to step (S9a) for mixing with the freezing mother liquor, and / or sent to step S7 for freezing crystallization.

[0179] Preferably, 5% to 9% of the sodium chloride mother liquor is sent to step S10 for mixed salt crystallization treatment, and the remaining 91% to 95% is sent to step (S9a) to be mixed with the freezing mother liquor and then subjected to repeated sodium chloride crystallization treatment.

[0180] In some embodiments, in step (S9c), the recrystallization mother liquor obtained by recrystallization can also be sent to step (S7a) for mixing.

[0181] In step (S9b), the sodium chloride mother liquor in step S7 can increase the chloride ion content of the precooled concentrate, so that the salt-to-nitrate ratio of the feed to the freeze crystallizer is between 3:7 and 4:6, thereby reducing the freezing blockage of the freeze crystallizer and increasing the extraction rate of sodium sulfate during the freeze crystallization process.

[0182] In some embodiments, in step (S9c), the sodium chloride product has a sodium chloride content ≥98.5g / 100g, a water-insoluble matter content ≤0.1g / 100g, a sulfate ion content ≤0.5g / 100g, a moisture content ≤0.5%, and a calcium and magnesium ion content ≤0.4g / 100g, meaning that the obtained sodium chloride product meets the requirements of Grade 1 refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard.

[0183] In step S10, the mixed salt crystallization process includes:

[0184] (S10a) Flash evaporate the sodium chloride mother liquor obtained in step S9 to obtain a concentrated salt slurry of mixed salts.

[0185] (S10b) The mixed salt concentrated slurry is subjected to solid-liquid separation to obtain mixed salt I and mixed salt mother liquor;

[0186] (S10c) The mixed salt mother liquor is dried in a rotary drum to obtain mixed salt II.

[0187] (S10d) Mix the mixed salt I and mixed salt II and then dry them to obtain the mixed salt product.

[0188] In some embodiments, in step (S10a), the sodium chloride mother liquor obtained from step S9 can be heated to near the boiling point before flash evaporation, so as to reduce the heat exchange load of the subsequent mixed salt crystallization unit during mixed salt crystallization treatment, and also to ensure the stability of the operating temperature of the sodium chloride crystallizer.

[0189] Furthermore, the flash evaporation temperature is 90℃~105℃, and the flash evaporation pressure based on the atmosphere is -5~5kPa.

[0190] In some embodiments, in step (S10b), the solid-liquid separation is preferably performed by centrifugation, and the water content of the mixed salt I is ≤80%.

[0191] In some embodiments, in step (S10c), the drum drying is preferably a double-drum drying, and the surface temperature of the drum is 115-120°C;

[0192] Furthermore, the drum dryer can crystallize highly soluble salts such as nitrates, significantly reducing system energy consumption, footprint, and wastewater discharge.

[0193] In some embodiments, in step (S10c), it is preferable to dissolve 48% to 52% of the mixture obtained by mixing mixed salt I and mixed salt II, and then return to step S7 for freeze crystallization. The remaining 52% to 48% becomes the mixed salt product. This can further improve the recovery rate of sodium sulfate and sodium chloride crystal salts and reduce the production of mixed salts.

[0194] Furthermore, the production volume of the aforementioned miscellaneous salt products accounts for ≤12% of the total production volume of crystalline salts (including sodium sulfate products and sodium chloride products), i.e., the miscellaneous salt rate is ≤12%.

[0195] A second aspect of the present invention provides a system for the above-mentioned high-salt wastewater desalination crystallization process, the system comprising the following units according to the material flow:

[0196] The hardness and silica removal unit is used to remove most of the hardness and silica from high-salt wastewater, resulting in hardness and silica removal effluent.

[0197] The turbidity removal unit is used to remove suspended solids from the hardening and desiliconizing effluent to obtain turbidity-removed effluent;

[0198] The softening unit is used to remove the residual hardness from the turbid effluent to obtain softened effluent;

[0199] The membrane treatment unit is used to remove precipitated colloids, suspended solids, and reduce volume in softened effluent to obtain membrane-treated effluent;

[0200] The ozone oxidation unit is used to remove organic matter from the effluent of membrane treatment, resulting in ozone-oxidized effluent.

[0201] The evaporation and concentration unit is used to concentrate the ozone oxidation effluent to obtain concentrated evaporation effluent;

[0202] The freeze crystallization unit is used to freeze crystallize the concentrated evaporation water to obtain sodium sulfate and freeze mother liquor.

[0203] The melting crystallization unit is used to melt crystallize Glauber's salt to obtain sodium sulfate product and molten mother liquor;

[0204] The sodium chloride crystallization unit is used to crystallize sodium chloride from the frozen mother liquor to obtain sodium chloride product and sodium chloride mother liquor.

[0205] The mixed salt crystallization unit is used to process part of the mother liquor discharged from the sodium chloride crystallization unit to obtain mixed salt products; among which...

[0206] The hardening and silicon removal unit uses a reaction sedimentation tank and a V-type filter tank to treat the high-salt wastewater;

[0207] The turbidity removal unit uses a high-density sedimentation tank to treat the hardness and silica removal effluent;

[0208] The softening unit uses a hydrogen-type weak acid cation exchanger to treat the turbidity-removed effluent.

[0209] The membrane treatment unit uses an ultrafiltration device and a reverse osmosis device to treat the softened effluent;

[0210] The ozone oxidation unit uses an ozone generator to treat the membrane treatment effluent;

[0211] The evaporation and concentration unit uses a TVR evaporator and an MVR evaporator to treat the ozone oxidation effluent;

[0212] The freeze crystallization unit uses an internal crystallizer to treat the evaporated and concentrated effluent;

[0213] The melting and crystallization unit uses a forced circulation FC crystallizer to process the Glauber's salt;

[0214] The sodium chloride crystallization unit uses a negative pressure vacuum forced circulation FC crystallizer to treat the frozen mother liquor;

[0215] The mixed salt crystallization unit uses a forced circulation FC crystallizer and a rotary drum dryer to treat the sodium chloride mother liquor.

[0216] To make the objectives, technical solutions, and effects of this invention clearer and more explicit, the invention will be described in detail below through specific embodiments and comparative examples.

[0217] In the following embodiments and comparative examples, the process flow diagrams are as follows: Figure 1 As shown, where,

[0218] The water quality characteristics of the high-salinity wastewater from the coal chemical industry are as follows: the concentration of total dissolved solids (TDS) is 10625 mg / L, and the concentration of Cl... - The concentration was 1910 mg / L, SO4 2- The concentration was 3859 mg / L, Mg 2+ The concentration was 27 mg / L, Ca 2+ The concentration was 185.7 mg / L, the dissolved silica concentration was 154 mg / L, and F - The concentration was 54 mg / L, HCO3 - The concentration of [unspecified substance] was 537 mg / L, and the concentration of COD was 170 mg / L.

[0219] The following terms, sodium sulfate extraction rate and sodium chloride extraction rate, are commonly used in this field and can be calculated through sampling and analysis by quality control and analysis departments; among them,

[0220] Sodium sulfate extraction rate: SO4 in water 2- The amount of salt completely converted into sodium sulfate product;

[0221] Sodium chloride extraction rate: Cl in water - Salt content of the product converted into sodium chloride;

[0222] Mixed salt ratio: Mixed salt production / (Sodium chloride production + Sodium sulfate production + Mixed salt production);

[0223] Example 1

[0224] Step S1: High-salinity wastewater from coal chemical industry enters the hardening and silicon removal unit for treatment.

[0225] (S1a) High-salt wastewater from coal chemical industry (flow rate of 420 m³ / h) 3 The solution is fed into the primary reaction tank at a flow rate of 1.0 g / L, and calcium hydroxide is added. After reacting for 25 minutes, the first clarified effluent is overflowed into the secondary reaction tank.

[0226] Under these conditions, the hardness of the first clarified effluent was 52 mg / L.

[0227] (S1b) Magnesium oxide (flow rate 0.9 g / L) was added to the secondary reaction tank, and the pH of the tank was adjusted to 7.5 using hydrochloric acid (concentration 32%). After reacting for 30 minutes, the resulting second clarified effluent was overflowed into the tertiary reaction tank;

[0228] The silicon content of the second clarified effluent was 38 mg / L.

[0229] (S1c) Polyferric sulfate solution (concentration 30%, flow rate 0.15 g / L) was added to the three-stage reaction tank as a coagulant, PAM (flow rate 0.0003 g / L) as a coagulant aid, and sodium carbonate (flow rate 0.2 g / L). After reacting for 25 minutes, hardening and silica-removing effluent was obtained; among which,

[0230] Under these conditions, the hardness of the effluent after hardening and desilting was 45 mg / L.

[0231] Step S2: The water effluent from the hardness and silica removal process enters the turbidity removal unit for further treatment.

[0232] The de-hardening and desiliconizing effluent was passed into a high-density sedimentation tank, and after reacting for 145 minutes, turbidity-removed effluent was obtained; wherein...

[0233] The operating conditions of the high-density sedimentation tank are as follows: the upward flow velocity of water in the inclined tube zone is 0.167 m / s, the hydraulic retention time is 60 min, and the sludge solids loading is 12 kg / (m³). 2 ·h), the linear velocity of the outer edge of the sludge scraper is 0.038m / s.

[0234] Under these conditions, the turbidity of the effluent obtained after turbidity removal is 3.5 NTU.

[0235] Step S3: The turbid water effluent enters the softening unit for further treatment.

[0236] (S3a) The turbidity-removed effluent sequentially enters a primary weak cation exchanger (using hydrogen-form ion exchange resin) and a secondary weak cation exchanger (using hydrogen-form ion exchange resin) to obtain ion-exchanged effluent; wherein,

[0237] The theoretical exchange capacity of both the primary and secondary weak cation exchangers is 1.2 eq / l.

[0238] (S3b) Hydrochloric acid is added to the ion-exchange effluent to adjust its pH to 5.8, and then the effluent is obtained by stripping through a decarbonator; wherein,

[0239] The concentration of the hydrochloric acid is 32%.

[0240] Under these conditions, the hardness of the softened effluent is 0 mg / L, and the alkalinity (OH-) is... - The concentration of ions was 65 mg / L.

[0241] Step S4: The softened effluent enters the membrane treatment unit for further processing.

[0242] (S4a) The softened effluent enters the ultrafiltration unit (a spiral wound membrane module made of PVDF) for cross-flow filtration to obtain ultrafiltration effluent; wherein,

[0243] The operating conditions of the ultrafiltration device are: inlet water pressure of 0.2 MPa, transmembrane pressure difference of 0.11 MPa, and recovery rate of 97%.

[0244] The dissolved solids content (i.e., salt content) of the obtained ultrafiltration effluent was 10.5 g / L, the turbidity was 0.1 NTU, and the solidification index (SDI) was 2.

[0245] (S4b) The ultrafiltration effluent is fed into a high-pressure reverse osmosis unit for reverse osmosis treatment to obtain membrane-treated effluent; wherein...

[0246] The operating conditions of the high-pressure reverse osmosis unit are: inlet water pressure of 2.2 MPa, inter-stage pressure of 0.24 MPa, and recovery rate of 75.5%.

[0247] Under these conditions, the conductivity of the membrane-treated effluent was 115 μS / cm. 2 .

[0248] Step S5: The effluent from the membrane treatment enters the ozone oxidation unit for further treatment.

[0249] Low-pressure oxygen enters the ozone generator, where it is converted into ozone through high-voltage discharge. The ozone then enters the ozone dosing tank, where an aluminum-based catalyst is added. Finally, the membrane-treated effluent enters the ozone dosing tank to obtain ozone-treated water.

[0250] The operating conditions of the ozone dosing tank are: influent flow rate of 95m³. 3 The total ozone production was 28 kg / h, the ozone concentration was 148 mg / L (10 wt%), and the ozone reaction time was 45 min.

[0251] Under these conditions, the COD removal rate of the ozone oxidation unit is 43%, and the COD of the effluent from the ozone oxidation process is 187 mg / L.

[0252] Step S6: The ozone-oxidized effluent enters the evaporation and concentration unit for further treatment.

[0253] (S6a) Use hydrochloric acid (concentration of 32%) to adjust the pH of the ozone oxidation effluent to 5.8.

[0254] (S6b) The solution obtained in (S6a) is heated to 90°C using a plate heat exchanger.

[0255] (S6c) The solution obtained in (S6b) is sent to a deaerator to remove non-condensable gases from the solution. Inside the deaerator, the solution is sprayed through nozzles and flows downwards in stages, and comes into contact with the steam flowing upwards in the opposite direction. The removed gas is discharged to the atmosphere along with a small amount of vented steam, and finally deoxygenated effluent is obtained.

[0256] (S6d) The deoxygenated effluent is divided into two equal streams. One stream is sent to an MVR evaporator for evaporation and concentration, and the other stream is sent to a TVR evaporator for evaporation and concentration. Finally, the resulting concentrates are combined into one stream, yielding evaporated and concentrated effluent that is only 1 / 5 the weight of the original effluent, along with water vapor condensate.

[0257] The operating conditions of the TVR evaporator are: temperature 98.7℃ and pressure 38kPa.

[0258] The operating conditions of the MVR evaporator are: temperature 98.7℃ and pressure 38kPa.

[0259] Under these conditions, the dissolved solids content in the concentrated evaporated water was 225 g / L, with SO4 content at 100 g / L. 2- The ion content is 129 g / L, Cl - The ion content is 35.08 g / L.

[0260] Step S7: The concentrated brine is evaporated and then fed into the freeze crystallization unit for further processing.

[0261] (S7a) The water from the evaporation and concentration process is heat-exchanged in a precooler to obtain precooled concentrated water at 28°C.

[0262] (S7b) Pre-cooled concentrated water is transported to a DTB-type crystallizer to exchange heat with the circulating ethylene glycol solution (i.e., the heat exchange medium) to obtain a sodium sulfate slurry at a temperature of -5.0℃; wherein,

[0263] During the heat exchange process, the internal circulation flow rate on the ethylene glycol solution side is 850 m³ / s. 3 The heat exchange rate of the pre-cooled concentrate is 0.27℃ / cycle.

[0264] (S7c) Sodium sulfate slurry was thickened by a hydrocyclone and then centrifuged to obtain sodium sulfate and a freezing mother liquor; among which...

[0265] The frozen mother liquor still contains some suspended solids. These solids are separated by setting up a settling tank. The concentrated liquid at the bottom of the settling tank is used as circulating sodium sulfate slurry and sent to (S5b) for heat exchange again (i.e., repeating the freeze crystallization operation). The supernatant of the settling tank can then enter the following step S9 for sodium chloride crystallization.

[0266] Under these conditions, the resulting cryogenic mother liquor contains SO4 2- The ion content is 4580 mg / L.

[0267] Step S8: Glauber's salt enters the melting and crystallization unit for processing.

[0268] (S8a) The water vapor condensate (flow rate of 4.8 t / h) obtained in step (S6d) is passed into a sodium sulfate hot melt tank containing the sodium sulfate to obtain a supersaturated sodium sulfate solution; wherein,

[0269] The supersaturated sodium sulfate solution contains 15 wt.% anhydrous sodium sulfate.

[0270] (S8b) A supersaturated sodium sulfate solution is heated to 110°C and flash-evaporated in a sodium sulfate crystallizer to obtain a concentrated sodium sulfate slurry; wherein,

[0271] The flash evaporation temperature is 98.5°C, and the flash evaporation pressure based on the atmosphere is 3.8 kPa.

[0272] (S8c) The concentrated sodium sulfate slurry was concentrated using a hydrocyclone thickener and then centrifuged to obtain sodium sulfate crystals and sodium sulfate mother liquor; among which...

[0273] The sodium sulfate mother liquor is returned to step (S8b) for repeated melting and crystallization.

[0274] (S8d) Anhydrous sodium sulfate crystals were dried in a fluidized bed to obtain sodium sulfate product (flow rate of 2.83 t / h).

[0275] Under these conditions, the parameters of the obtained sodium sulfate product are as follows: sodium sulfate content 99.1g / 100g, water-insoluble matter 0.02g / 100g, chloride 0.25g / 100g, TOC content 25mg / kg, moisture content 0.04%, and calcium and magnesium content 0, which meets the requirements of Class II Grade 1 product in "GBT 6009-2014 Industrial Anhydrous Sodium Sulfate".

[0276] The extraction rate of the sodium sulfate product was 93.19%.

[0277] Step S9: The frozen mother liquor enters the sodium chloride crystallization unit for processing.

[0278] (S9a) The frozen mother liquor is heated to 75°C and then passed through a volatile matter removal tower for flash evaporation; wherein,

[0279] The flash evaporation temperature is 95°C, and the flash evaporation pressure based on the atmosphere is 10.5 kPa.

[0280] (S9b) The solution obtained in (S9a) is heated to 97.5°C and flash-evaporated in a sodium chloride crystallizer to obtain a concentrated sodium chloride slurry; wherein,

[0281] The flash evaporation temperature is 95°C, and the flash evaporation pressure based on the atmosphere is -28 kPa.

[0282] (S9c) The concentrated sodium chloride slurry is thickened by a hydrocyclone separator to increase the suspended solids concentration of the slurry to 42%. After centrifugation, crude wet sodium chloride and sodium chloride mother liquor are obtained; among them,

[0283] 5% of the sodium chloride mother liquor is sent to step S10 for mixed salt crystallization treatment, and 95% of it enters step (S9b) for repeated sodium chloride crystallization treatment.

[0284] The moisture content of the crude sodium chloride wet salt is 5%.

[0285] (S9d) The crude wet sodium chloride salt is dissolved in the water vapor condensate obtained in step (S6d), and then recrystallized in a sodium chloride recrystallizer to obtain refined wet sodium chloride salt.

[0286] (S9e) The wet salt of refined sodium chloride is dried in a fluidized bed to obtain sodium chloride product (flow rate of 1t / h).

[0287] Under these conditions, the parameters of the sodium chloride product are as follows: sodium chloride content 99.8g / 100g, water-insoluble matter 0.04g / 100g, sulfate ion 0.16g / 100g, moisture content 0.02%, and calcium and magnesium ion content 0g / 100g, which meets the requirements of Grade I refined industrial salt in the "GB / T5462-2015 Industrial Salt" standard.

[0288] The extraction rate of the sodium chloride product was 81.59%.

[0289] Step S10: The sodium chloride mother liquor is fed into the mixed salt crystallization unit for processing.

[0290] (S10a) The sodium chloride mother liquor is heated to 95°C, and then flash evaporated in a mixed salt crystallizer to obtain a concentrated mixed salt slurry;

[0291] (S10b) After centrifuging the concentrated salt slurry, mixed salt I and mixed salt mother liquor are obtained; wherein...

[0292] The water content of the mixed salt I is 65%;

[0293] (S10c) The mixed salt mother liquor is dried in a double-drum dryer to obtain mixed salt II; wherein,

[0294] The surface temperature of the drum is 118°C;

[0295] (S10d) Mix mixed salt I and mixed salt II, and dissolve 50% of them in the water vapor condensate obtained in step (S6d) and return to step S7, whereby the remaining 50% becomes mixed salt brine.

[0296] (S11d) The mixed salt brine is concentrated and dehydrated to obtain a mixed salt product with a flow rate of 0.28 t / h.

[0297] In summary, the extraction rates of sodium sulfate and sodium chloride from the high-salt wastewater from coal chemical plants were 93.19% and 81.59%, respectively. The resulting product flow rates were 2.83 t / h for sodium sulfate, 1 t / h for sodium chloride, and 0.28 t / h for miscellaneous salts, with a miscellaneous salt rate of 6.81%.

[0298] Comparative Example 1 (When adjusting the pH, hydrochloric acid was replaced with sulfuric acid, and the circulation rate of the ethylene glycol solution was reduced)

[0299] The only difference from Example 1 is that:

[0300] In steps (S1b), (S3b), and (S6a), the 32% hydrochloric acid is replaced with 98% sulfuric acid.

[0301] In step (S6), the SO4 in the evaporated and concentrated water... 2- The ion content increased from 129 g / L to 142 g / L, Cl - The ion content decreased from 35.08 g / L to 28.46 g / L.

[0302] In step (S7b), the temperature of the sodium sulfate slurry is increased from -5°C to -2.0°C, and the internal circulation flow rate is increased from 850 m³ / h. 3 / h reduced to 250m 3 / h, the heat exchange rate is increased from 0.27℃ / cycle to 0.5℃ / cycle;

[0303] Under these conditions

[0304] In the obtained cryogenic mother liquor, SO4 2- The ion content increased from 4580 mg / L to 22500 mg / L;

[0305] The crystallizer cannot be further reduced to the design temperature (-5.0℃ to -4.5℃) after the operating temperature reaches -2.0℃. Furthermore, when the temperature of the precooled concentrate drops to 0.5℃, a "crystal bursting" phenomenon occurs in the cooler of the freeze crystallizer, causing the cooler to become blocked.

[0306] In step (S8d), the flow rate of the sodium sulfate product is reduced from 2.83 t / h to 1.91 t / h;

[0307] The parameters of the sodium sulfate product are as follows: sodium sulfate content 98.5g / 100g, water-insoluble matter 0.35g / 100g, chloride 0.45g / 100g, moisture content 0.6%, calcium and magnesium content 0g / 100g, which meets the requirements for qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate";

[0308] The extraction rate of the sodium sulfate product decreased from 93.19% to 70.56%;

[0309] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.99 t / h;

[0310] The parameters of the sodium chloride product are as follows: sodium chloride content is 72g / 100g, water-insoluble matter content is 3.5g / 100g, SO42- 2- The ion content is 18.2g / 100g, the moisture content is 4.3%, and the calcium and magnesium ion content is 2.0g / 100g, which does not meet the requirements of Grade I refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard, and the sodium chloride product is also unqualified.

[0311] The extraction rate of the sodium chloride product decreased from 81.59% to 80.12%.

[0312] In step (S11d), the flow rate of the mixed salt product is increased from 0.28 t / h to 0.55 t / h.

[0313] In summary, the extraction rates of sodium sulfate and sodium chloride from the high-salt wastewater of the coal chemical industry were 70.56% and 80.12%, respectively. The resulting product flow rates were 1.91 t / h for sodium sulfate, 0.99 t / h for sodium chloride, and 0.55 t / h for miscellaneous salts, with a miscellaneous salt rate of 15.94%.

[0314] In the production process of Comparative Example 1, when sulfuric acid was used to adjust the pH and a lower ethylene glycol solution circulation rate was used, the freezer experienced crystal bursting and the extraction rates of sodium sulfate and sodium chloride were low. Therefore, Comparative Examples 2 and 3 were conducted to test these two variables respectively.

[0315] Comparative Example 2 (reducing the circulation rate of ethylene glycol solution)

[0316] The only difference from Example 1 is that:

[0317] In step (S7b), the temperature of the sodium sulfate slurry is increased from -5°C to -2.0°C, and the internal circulation flow rate is increased from 850 m³ / h. 3 / h reduced to 250m 3 / h, the heat exchange rate is increased from 0.27℃ / cycle to 0.5℃ / cycle;

[0318] Under these conditions

[0319] In the freezing mother liquor, SO4 2- The ion content increased from 4580 mg / L to 24500 mg / L.

[0320] The crystallizer cannot be further reduced to the design temperature (-5.0℃ to -4.5℃) after the operating temperature reaches -2.0℃. Furthermore, when the temperature of the precooled concentrate drops to 0.5℃, a "crystal bursting" phenomenon occurs in the cooler of the freeze crystallizer, causing the cooler to become blocked.

[0321] In step (S8d), the flow rate of the sodium sulfate product is reduced from 2.83 t / h to 1.86 t / h.

[0322] The parameters of the sodium sulfate product are as follows: sodium sulfate content is 98.6g / 100g, water-insoluble matter content is 0.38g / 100g, chloride content is 0.47g / 100g, moisture content is 0.5%, and calcium and magnesium ion content is 0g / 100g, which meets the requirements for qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate".

[0323] The extraction rate of the sodium sulfate product decreased from 93.19% to 70.56%, with a large amount of sodium sulfate being lost to the subsequent sodium chloride system.

[0324] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.99 t / h.

[0325] The parameters of the sodium chloride product are as follows: sodium chloride content is 71g / 100g, water-insoluble matter content is 3.4g / 100g, SO4 content is... 2- The ion content is 20.1g / 100g, the moisture content is 3.5%, and the calcium and magnesium ion content is 2.0g / 100g, which does not meet the requirements of Grade I refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard.

[0326] The extraction rate of the sodium chloride product decreased from 81.59% to 80.12%.

[0327] In step (S11d), the flow rate of the mixed salt product is increased from 0.28 t / h to 1.25 t / h.

[0328] In summary, the extraction rates of sodium sulfate and sodium chloride from the high-salt coal chemical wastewater were 70.56% and 80.12%, respectively. The resulting product flow rates were 1.86 t / h for sodium sulfate, 0.99 t / h for sodium chloride, and 1.25 t / h for miscellaneous salts, with a miscellaneous salt rate of 30.49%.

[0329] Comparative Example 3 (When adjusting the pH, hydrochloric acid was replaced with sulfuric acid).

[0330] The only difference from Example 1 is that:

[0331] In steps (S1b), (S3b), and (S6a), the 32% hydrochloric acid is replaced with 98% sulfuric acid, while everything else remains unchanged.

[0332] Under these conditions

[0333] In step S6, SO4 in the water is evaporated and concentrated. 2- The ion content increased from 129 g / L to 142 g / L, Cl - The ion content decreased from 35.08 g / L to 27.46 g / L;

[0334] In step S7, the obtained cryogenic mother liquor contains SO4 2- The ion content increased from 4580 mg / L to 24500 mg / L;

[0335] The extraction rate of the sodium sulfate product decreased from 93.19% to 76.12%.

[0336] In step (S8d), the flow rate of sodium sulfate product changes from 2.83 t / h to 1.99 t / h.

[0337] The parameters of the obtained sodium sulfate product are as follows: sodium sulfate content 98.0g / 100g, water-insoluble matter 0.30g / 100g, chloride 0.46g / 100g, moisture content 0.3%, calcium and magnesium content 0g / 100g, which meets the requirements for qualified products in "GB / T 6009-2014 Industrial Anhydrous Sodium Sulfate";

[0338] In step (S9e), the flow rate of the sodium chloride product is reduced from 1 t / h to 0.95 t / h.

[0339] The parameters of the sodium chloride product obtained in step S9 are as follows: sodium chloride content is 72.5 g / 100 g, water-insoluble matter content is 2.2 g / 100 g, SO42- 2- The ion content is 19.2g / 100g, the moisture content is 5.1%, and the calcium and magnesium ion content is 1.0g / 100g, which does not meet the requirements of Grade I refined industrial salt in the "GB / T 5462-2015 Industrial Salt" standard.

[0340] The extraction rate of the sodium chloride product decreased from 81.59% to 76.12%.

[0341] In step (S11d), the flow rate of the mixed salt product is increased from 0.28 t / h to 1.17 t / h.

[0342] In summary, the extraction rates of sodium sulfate and sodium chloride from the high-salt coal chemical wastewater were 76.12% and 76.12%, respectively. The resulting product flow rates were 1.99 t / h for sodium sulfate, 0.95 t / h for sodium chloride, and 1.17 t / h for miscellaneous salts, with a miscellaneous salt rate of 28.47%.

[0343] Sodium sulfate extraction rate 93.19％ 70.56％ 70.56％ 76.12％ Sodium chloride extraction rate 81.59％ 80.12％ 80.12％ 76.12％ Sodium chloride product grade Grade A Non-Grade 1 Product Non-Grade 1 Product Non-Grade 1 Product Mixed salt (product) rate 6.81％ 15.94％ 30.49％ 28.47％ Freeze crystallizer operation cycle More than 300 days 30-40 days 10 to 15 days or less 12-15 days

[0344] The preferred embodiments of the present invention have been described in detail above. However, these embodiments are for illustrative purposes only and are not intended to limit the scope of protection of the present invention. Those skilled in the art can make various other simple modifications and combinations within the scope of the present invention. These simple modifications and combinations should also be regarded as the contents disclosed in the present invention and are all within the scope of protection of the present invention.

Claims

1. A process for desalination and crystallization of high-salt wastewater, characterized in that, The process includes the following steps: S1: High-salt wastewater is treated to remove hardness and silicon to obtain hardened and siliconized effluent; S2: The hardening and silicating effluent is subjected to turbidity removal treatment to obtain turbidity-removed effluent; S3: The turbidity-removed effluent is subjected to ion exchange softening treatment to obtain softened effluent; S4: The softened effluent is subjected to membrane treatment to obtain membrane-treated effluent; S5: The membrane-treated effluent is subjected to ozone oxidation treatment to obtain ozone-oxidized effluent; S6: Adjust the pH of the ozone-oxidized effluent to 5.5-6.0 using sulfuric acid or hydrochloric acid, and then perform evaporation and concentration treatment to obtain evaporated and concentrated effluent with a dissolved solids content of 200-230 g / L and a salt-nitrate ratio of 3:7 to 4:6; S7: The evaporated and concentrated water is pre-cooled to obtain pre-cooled concentrated water at a temperature of 28℃~40℃. Then, the pre-cooled concentrated water is subjected to freeze crystallization to obtain a freeze crystal slurry at a temperature of -5.5℃~-4.5℃. The freeze crystal slurry is subjected to solid-liquid separation to obtain sodium sulfate and freeze mother liquor. The freeze mother liquor is divided into two parts. 85%~90% of the mother liquor is sent to step S9 for sodium chloride crystallization, and the remaining 10%~15% is repeatedly subjected to freeze crystallization. In the aforementioned freeze-crystallization process, ethylene glycol solution is used as the heat exchange medium, the heat exchange temperature difference between the pre-cooled concentrated water and the ethylene glycol solution is ≤3℃, and the temperature difference of the ethylene glycol solution before and after heat exchange is ≤1.5℃. S8: The Glauber's salt is subjected to melt crystallization treatment to obtain sodium sulfate product and melt mother liquor, and the melt mother liquor is repeatedly subjected to melt crystallization treatment; S9: The frozen mother liquor obtained in step S7 is subjected to sodium chloride crystallization treatment to obtain sodium chloride product and sodium chloride mother liquor. The sodium chloride mother liquor is divided into two parts, of which 5%~9% is sent to step S10 for mixed salt crystallization treatment, and the remaining 91%~95% is repeatedly subjected to sodium chloride crystallization treatment. S10: The sodium chloride mother liquor obtained in step S9 is subjected to mixed salt crystallization treatment to obtain mixed salt product.

2. The process according to claim 1, characterized in that, In step S1, the hardening and silicon removal process includes: (S1a) Add calcium agent to the high-salt wastewater to obtain the first clarified effluent; (S1b) Add magnesium agent to the first clarified effluent to obtain the second clarified effluent; (S1c) Add coagulant, sodium carbonate and coagulant aid to the second clarified effluent to obtain hardness- and silica-removed effluent.

3. The process according to claim 2, characterized in that, In step (S1a), the calcium agent is calcium chloride and / or calcium hydroxide, the amount of calcium agent added is 0.8~1.5 g / L, the reaction time of the calcium agent with the high-salt wastewater is 5~60 min, the pH of the calcium agent reacting with the high-salt wastewater is 11~11.5, and the hardness of the first clarified effluent is ≤65 mg / L.

4. The process according to claim 2, characterized in that, In step (S1b), the magnesium agent is at least one of magnesium oxide, magnesium chloride and magnesium sulfate, the amount of magnesium agent added is 0.8~2.0 g / L, the reaction time of the magnesium agent with the first clarified effluent is 5~60 min, the pH of the magnesium agent reacting with the first clarified effluent is 7~8, and the silicon content of the second clarified effluent is ≤40 mg / L.

5. The process according to claim 2, characterized in that, In step (S1c), the coagulant is selected from 30% polyferric sulfate solution, 30% polyferrous sulfate solution, or 30% polyferric chloride solution, and the amount of coagulant added is 0.1~0.25 g / L. The coagulant aid is polyacrylamide, and the amount of coagulant added is 0.2~0.5 mg / L. The amount of sodium carbonate added is 1.0~1.8 g / L. The reaction time of the second clarified effluent with the coagulant, sodium carbonate, and coagulant aid is 5~60 min. The hardness of the hardness and silica removal effluent is ≤50 mg / L.

6. The process according to claim 1, characterized in that, In step S2, the turbidity removal treatment is carried out by settling in a high-density sedimentation tank under the following conditions: the upward flow velocity of the water in the pipe area is 0.1~0.2 m / s, and the sludge solids load is 11~12 kg / (m³). 2 The linear velocity of the outer edge of the sludge scraper is 0.03~0.04 m / s (h).

7. The process according to claim 6, characterized in that, The turbidity of the effluent after turbidity removal is ≤5 NTU.

8. The process according to claim 1, characterized in that, In step S3, a hydrogen-type weak acid cation exchanger is used to soften the turbidity-removed effluent, wherein the softened effluent has a pH of 6-9, a hardness of ≤5 mg / L, a SiO2 content of ≤40 mg / L, and an alkalinity of ≤80 mg / L.

9. The process according to claim 8, characterized in that, The hydrogen-type weak acid cation exchanger includes a primary weak acid cation exchanger with a regeneration cycle of 24 hours and a secondary weak acid cation exchanger with a regeneration cycle of 48 hours.

10. The process according to claim 1, characterized in that, In step S4, the membrane treatment includes: (S4a) The softened effluent is subjected to ultrafiltration treatment to obtain ultrafiltration effluent; (S4b) The ultrafiltration effluent is subjected to reverse osmosis treatment to obtain membrane-treated effluent; In step (S4a), the conditions for ultrafiltration are: influent pressure ≤ 0.3 MPa, transmembrane pressure difference ≤ 0.16 MPa, and recovery rate ≥ 95%; In step (S4a), the dissolved solids content of the ultrafiltration effluent is 10~14 g / L, the turbidity is ≤0.2 NTU, and the pollution index is ≤3; In step (S4b), the reverse osmosis treatment conditions are: inlet water pressure ≤ 3.0 MPa, inter-stage pressure ≤ 0.3 MPa, and recovery rate ≥ 75%; In step (S4b), the dissolved solids content of the membrane-treated effluent is 41~43 g / L, and the conductivity of the membrane-treated effluent is ≤200 μs / cm.

11. The process according to claim 1, characterized in that, In step S5, The ozone oxidation treatment uses an aluminum-based ozone catalyst, a silicon-aluminum ozone catalyst, or a silicon-aluminum-titanium ozone catalyst. The conditions for the ozone oxidation treatment are: influent flow rate of 90~114 m³ / h. 3 / h, the total ozone production is 60 kg / h, the ozone concentration is 148 mg / L, and the reaction time is 20~50 min; The COD of the ozone-oxidized effluent is ≤200 mg / L.

12. The process according to claim 1, characterized in that, In step S6, the evaporation and concentration process includes: (S6a) Adjust the pH of the ozone-oxidized water to 5.5-6.0 using sulfuric acid or hydrochloric acid; (S6b) Heat the solution obtained in (S6a) to a temperature close to the boiling point; (S6c) The solution obtained in (S6b) is subjected to deoxygenation treatment to obtain deoxygenated effluent; (S6d) The deoxygenated effluent is subjected to evaporation and concentration treatment to obtain evaporated and concentrated effluent; In step (S6c), the deoxygenation treatment is carried out in a deaerator. Specifically, the solution obtained in (S6b) is sprayed through a nozzle and flows downward in stages in the deaerator, while contacting the steam flowing upward in the opposite direction to obtain the deoxygenated effluent. In step (S6d), during the evaporation and concentration, the deoxygenated effluent is divided into two streams, which are then evaporated and concentrated in an MVR evaporator and a TVR evaporator, respectively. Finally, the concentrated liquids are combined into one stream to obtain the evaporated and concentrated effluent.

13. The process according to claim 12, characterized in that, The operating temperature of the TVR evaporator is 98℃~99.5℃ and the pressure is ≤50 kPa; the operating temperature of the MVR evaporator is 98℃~99.5℃ and the pressure is ≤50 kPa.

14. The process according to claim 1, characterized in that, In step S7, the freeze-crystallization process includes: (S7a) The evaporated and concentrated water is mixed with the frozen mother liquor that has been repeatedly subjected to freeze crystallization treatment, and the mixture is pre-cooled to obtain pre-cooled concentrated water with a temperature of 28°C to 40°C. (S7b) The precooled concentrated water is subjected to freeze crystallization to obtain a frozen crystal slurry with a temperature of -5.5℃ to -4.5℃; (S7c) The frozen crystal slurry is subjected to solid-liquid separation to obtain sodium sulfate and frozen mother liquor. The frozen mother liquor is divided into two parts, of which 85%~90% is sent to step S9 for sodium chloride crystallization treatment, and the remaining 10%~15% is repeatedly subjected to frozen crystallization treatment. In step (S7c), the SO4 in the cryogenic mother liquor 2- Ion content ≤5500 mg / L.

15. The process according to claim 14, characterized in that, In step (S7c), the solid-liquid separation is centrifugation.

16. The process according to claim 1, characterized in that, In step S8, the melt crystallization process includes: (S8a) Pass water vapor into the Glauber's salt to dissolve the Glauber's salt into a supersaturated sodium sulfate solution; (S8b) The supersaturated sodium sulfate solution is mixed with the molten mother liquor that has undergone repeated melt crystallization treatment, and then flash evaporated to obtain a concentrated sodium sulfate slurry; (S8c) The sodium sulfate slurry is subjected to solid-liquid separation to obtain anhydrous sodium sulfate crystals and molten mother liquor; (S8d) The anhydrous sodium sulfate crystals are dried to obtain sodium sulfate product; In step (S8a), the supersaturated sodium sulfate solution contains 15 wt% undissolved sodium sulfate crystals; In step (S8b), before the flash evaporation, the mixture obtained after mixing is heated to near the boiling point; In step (S8c), the solid-liquid separation is performed by centrifugation, and the moisture content of the anhydrous sodium sulfate crystals is ≤5%. In step (S8c), the molten mother liquor is returned to step (S8b) and mixed with a supersaturated sodium sulfate solution.

17. The process according to claim 1, characterized in that, In step S9, the sodium chloride crystallization process includes: (S9a) The frozen mother liquor obtained in step S7 is mixed with the sodium chloride mother liquor that has been repeatedly subjected to sodium chloride crystallization treatment, and the mixture is flash evaporated to obtain a concentrated sodium chloride slurry. (S9b) The concentrated sodium chloride slurry is subjected to solid-liquid separation to obtain crude wet sodium chloride salt and sodium chloride mother liquor; (S9c) The crude wet sodium chloride salt is recrystallized and dried to obtain sodium chloride product; Before performing the step (S9a), the freezing mother liquor is heated to 75℃±0.5℃; In step (S9a), the flash temperature is 95°C to 100°C, and the flash pressure based on the atmosphere is -30 to 5 kPa; In step (S9b), before solid-liquid separation, the concentrated sodium chloride slurry is thickened using a hydrocyclone separator until the concentration of suspended solids is 40 wt%. In step (S9b), the moisture content of the crude wet sodium chloride salt is 5 wt%, and the chloride ion content in the sodium chloride mother liquor is ≤120 g / L.

18. The process according to claim 17, characterized in that, In step (S9b), the solid-liquid separation is performed by centrifugation.

19. The process according to claim 1, characterized in that, In step S10, the mixed salt crystallization process includes: (S10a) Flash evaporate the sodium chloride mother liquor obtained in step S9 to obtain a concentrated salt slurry of mixed salts. (S10b) The mixed salt concentrated slurry is subjected to solid-liquid separation to obtain mixed salt I and mixed salt mother liquor; (S10c) The mixed salt mother liquor is dried in a rotary drum to obtain mixed salt II; (S10d) Mix the mixed salt I and mixed salt II and then dry them to obtain the mixed salt product; In step (S10a), before the flash evaporation, the sodium chloride mother liquor is heated to near the boiling point; In step (S10b), the solid-liquid separation is performed by centrifugation; In step (S10c), during the drum drying process, the surface temperature of the drum is 115℃~120℃; In step (S10c), 50% of the mixture obtained by mixing mixed salt I and mixed salt II is dissolved, and then returned to step S7 for freeze crystallization, and the remaining 50% becomes the mixed salt product.

20. The process according to claim 19, characterized in that, In step (S10a), the flash temperature is 90~105℃ and the flash pressure based on the atmosphere is -5~5 kPa.

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