CN109422408B

CN109422408B - Method for treating catalyst production wastewater

Info

Publication number: CN109422408B
Application number: CN201710752803.6A
Authority: CN
Inventors: 殷喜平; 李叶; 顾松园
Original assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Current assignee: China Petroleum and Chemical Corp; Sinopec Catalyst Co
Priority date: 2017-08-28
Filing date: 2017-08-28
Publication date: 2023-04-07
Anticipated expiration: 2037-08-28
Also published as: CN109422408A

Abstract

The invention relates to the field of sewage treatment, and discloses a method for treating catalyst production wastewater, wherein the catalyst production wastewater contains NH ₄ ⁺ 、SO ₄ ^2‑ 、Cl ^‑ And Na ⁺ The method is characterized by comprising the following steps of 1) carrying out first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution; 2) Cooling and crystallizing the first concentrated solution to obtain a crystallization solution containing sodium sulfate crystals; 3) Carrying out first solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals, and introducing a liquid phase obtained by the first solid-liquid separation into a second MVR evaporation device for second evaporation to obtain second ammonia-containing steam and a second concentrated solution containing the crystals; 4) Cooling the second concentrated solution containing the crystals to obtain a treatment solution containing sodium chloride crystals; 5) And carrying out second solid-liquid separation on the treatment liquid. The method can respectively recover ammonium, sodium sulfate and sodium chloride in the wastewater, and furthest recycle resources in the wastewater.

Description

Method for treating wastewater from catalyst production

Technical Field

The invention relates to the field of sewage treatment, in particular to a method for treating catalyst production wastewater, and especially relates to a catalyst containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method for treating wastewater from catalyst production.

Background

In the production process of the oil refining catalyst, a large amount of inorganic acid alkali salts such as sodium hydroxide, hydrochloric acid, sulfuric acid, ammonium salts, sulfates, hydrochlorides and the like are needed, and a large amount of mixed sewage containing ammonium, sodium sulfate, sodium chloride and aluminosilicate is generated. For such sewage, the common practice in the prior art is that the pH value is adjusted to be within the range of 6-9, most of suspended matters are removed, then the biochemical method, the blow-off method or the steam stripping method is adopted to remove ammonium ions, then the salt-containing sewage is subjected to pH value adjustment, most of suspended matters are removed, hardness, silicon and part of organic matters are removed, most of organic matters are removed through ozone biological activated carbon adsorption oxidation or other advanced oxidation methods, then the salt-containing sewage enters an ion exchange device for further hardness removal, enters an enrichment device (such as reverse osmosis or electrodialysis) for concentration, and then MVR evaporative crystallization or multiple-effect evaporative crystallization is adopted to obtain mixed miscellaneous salt of sodium sulfate and sodium chloride containing a small amount of ammonium salt; or is; firstly, adjusting the pH value to be within the range of 6.5-7.5, removing most suspended matters, then removing hardness, silicon and part of organic matters, removing most organic matters through ozone biological activated carbon adsorption oxidation or other advanced oxidation methods, then entering an ion exchange device for further removing hardness, entering a thickening device (such as reverse osmosis and/or electrodialysis) for concentration, and then adopting MVR (mechanical vapor recompression) evaporative crystallization or multi-effect evaporative crystallization to obtain the mixed salt of sodium sulfate and sodium chloride containing ammonium salt. However, these ammonium-containing mixed salts are currently difficult to treat or expensive to treat, and the process of removing ammonium ions at the early stage additionally increases the cost of wastewater treatment.

In addition, the biochemical deamination can only treat wastewater with low ammonium content, and can not directly carry out biochemical treatment due to insufficient COD content in the catalyst sewage, and organic matters such as glucose or starch and the like are additionally added in the biochemical treatment process, so that the ammonia nitrogen can be treated by the biochemical method. The most important problems are that the total nitrogen of the wastewater after the biochemical deamination treatment is not up to the standard (the contents of nitrate ions and nitrite ions exceed the standard), advanced treatment is needed, in addition, the salt content in the wastewater is not reduced (20-30 g/L), the wastewater cannot be directly discharged, and further desalination treatment is needed.

In order to remove ammoniacal nitrogen from wastewater by gas stripping deamination, a large amount of alkali is needed to adjust the pH value, the alkali consumption is high, the alkali in the wastewater after deamination cannot be recovered, the pH value of the treated wastewater is high, the treatment cost is high, the COD content in the catalyst wastewater after gas stripping does not change greatly, the salt content in the wastewater is not reduced (20-30 g/L), the wastewater cannot be directly discharged, further desalting treatment is needed, the wastewater treatment operation cost is high, a large amount of alkali remains in the treated wastewater, the pH value is high, waste is large, and the treatment cost is up to 50 yuan/ton.

Disclosure of Invention

The invention aims to overcome the defect of NH content in the prior art ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The wastewater treatment cost is high, and only mixed salt crystals can be obtained, and a low-cost and environment-friendly NH-containing catalyst is provided ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method for treating wastewater can respectively recover ammonium, sodium sulfate and sodium chloride in the wastewater generated in the catalyst production, and can furthest recycle resources in the wastewater.

In order to achieve the above objects, one aspect of the present invention provides a method for treating catalyst production wastewater containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method comprises the following steps of,

1) Performing first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution, wherein the wastewater to be treated contains the catalyst production wastewater;

2) Cooling and crystallizing the first concentrated solution to obtain a crystallization solution containing sodium sulfate crystals;

3) Carrying out first solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals, and introducing a liquid phase obtained by the first solid-liquid separation into a second MVR evaporation device for second evaporation to obtain second ammonia-containing steam and a second concentrated solution containing the crystals;

4) Cooling the second concentrated solution containing the crystals to obtain a treatment solution containing sodium chloride crystals;

5) Carrying out second solid-liquid separation on the treatment liquid;

whereinBefore the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9; SO in the first concentrated solution ₄ ^2- Has a concentration of 0.01mol/L or more, cl ^- The concentration of (B) is 5.2mol/L or less.

By the technical scheme, the method aims at the content of NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The pH value of the wastewater to be treated is adjusted to a specific range in advance, then, the wastewater is subjected to first evaporation to obtain concentrated ammonia water, then, sodium sulfate crystals are obtained by cooling crystallization separation, then, second concentrated solution containing sodium chloride crystals and sodium sulfate crystals and dilute ammonia water are obtained by second evaporation, then, the sodium sulfate crystals in the second concentrated solution are dissolved by cooling treatment, and sodium chloride is further crystallized and separated out to obtain sodium chloride crystals. The method can respectively obtain high-purity sodium sulfate and sodium chloride, avoids the difficulty in the processes of mixed salt treatment and recycling, simultaneously completes the process of separating ammonia and salt, simultaneously heats the wastewater and cools the ammonia-containing steam by adopting a heat exchange mode without a condenser, reasonably utilizes the heat in the evaporation process, saves energy, reduces the wastewater treatment cost, recovers the ammonium in the wastewater in the form of ammonia water, recovers the sodium chloride and the sodium sulfate in the form of crystals respectively, does not generate waste residues and waste liquid in the whole process, and achieves the purpose of changing waste into valuable.

Furthermore, the method obtains stronger ammonia water through first evaporation, is convenient for the utilization of the ammonia water, improves the ion concentration of the solution, improves the precipitation rate of cooling crystallization and improves the efficiency; through the cooperation of the second evaporation and the cooling treatment, the second evaporation process can be carried out at a higher temperature, the solid content and the second evaporation efficiency of the second evaporation are improved, the amount of circulating liquid in the treatment system is reduced, and meanwhile, the energy-saving effect can be achieved.

Drawings

FIG. 1 is a schematic flow diagram of a method for treating wastewater from catalyst production according to an embodiment of the present invention.

Description of the reference numerals

1. Second MVR evaporation plant 70, eleventh circulating pump

2. Cooling crystallization device 71 and first circulation pump

3. First MVR evaporation plant 72, second circulating pump

31. First heat exchange device 73 and third circulating pump

32. Second heat exchanger 74 and fourth circulating pump

33. Third heat exchange device 75 and fifth circulating pump

34. Fourth heat exchange device 76 and sixth circulating pump

35. Fifth heat exchange device 77 and seventh circulation pump

36. Sixth heat exchange device 78 and eighth circulating pump

38. Eighth heat exchange device 79 and ninth circulating pump

30. Tenth heat exchange device 80, tenth circulating pump

51. First ammonia water storage tank 81 and vacuum pump

52. Second ammonia storage tank 82 and circulating water tank

53. First mother liquor tank 83 and tail gas absorption tower

54. Second mother liquid tank 91 and first solid-liquid separation device

55. Low-temperature treatment tank 92 and second solid-liquid separation device

61. First pH value measuring device 101 and first compressor

62. Second pH value measuring device 102 and second compressor

60. Third pH value measuring device

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

The present invention will be described below with reference to fig. 1, but the present invention is not limited to fig. 1.

The invention provides a method for treating wastewater generated in catalyst production, which contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The method comprises the following steps of,

5) Carrying out second solid-liquid separation on the treatment liquid;

before the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9; SO in the first concentrated solution ₄ ^2- Has a concentration of 0.01mol/L or more, cl ^- The concentration of (B) is 5.2mol/L or less.

Preferably, the wastewater to be treated is the catalyst production wastewater; or the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by the second solid-liquid separation.

More preferably, the wastewater to be treated is a mixed solution of the catalyst production wastewater and at least part of a liquid phase obtained by the second solid-liquid separation.

The method provided by the invention can be used for the treatment of the ammonia-containing gas containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ Except that it contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ In addition, the catalyst production wastewater is not particularly limited.

The method provided by the invention can be used for the treatment of the compounds containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ Is treated except for containing NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ In addition, the catalyst production wastewater is not particularly limited. From the viewpoint of improving the efficiency of wastewater treatment, the first concentrated solution after concentration may be in the following range. SO contained in the first concentrated solution ₄ ^2- Is 0.01mol/L or more, more preferably 0.07mol/L or more, further preferably 0.1mol/L or more, even more preferably 0.2mol/L or more, particularly preferably 0.3mol/L or more, and may be, for example, 0.01mol/L, 0.03mol/L, 0.05mol/L, 0.1mol/L, 0.3mol/L, 0.5mol/L, 0.7mol/L, 0.9mol/L, 1mol/L, 1.2mol/L, 1.4mol/L, 1.5mol/L, 1.8mol/L, 2mol/L, 2.2mol/L, 2.5mol/L, 2.8mol/L, 3mol/L, 3.2mol/L, 3.5mol/L, 3.8mol/L, or 4mol/L, etc. And, cl in the first concentrated solution ^- The concentration of (B) is 5mol/L or less, preferably 4.5mol/L or less, and may be, for example, 0.5mol/L, 0.8mol/L, 1mol/L, 1.3mol/L, 1.5mol/L, 1.6mol/L, 1.7mol/L, 1.8mol/L, 2mol/L, 2.2mol/L, 2.4mol/L, 2.5mol/L, 2.8mol/L, 3mol/L, 3.2mol/L, 3.5mol/L, 3.8mol/L, 4mol/L, 4.2mol/L, 4.5mol/L, 4.8mol/L or 5 mol/L. By adding SO in the first concentrated solution ₄ ^2- 、Cl ^- The concentration of (3) is controlled in the above range, so that sodium sulfate can be separated out from the cooled crystal without separating out sodium chloride, thereby achieving the purpose of efficiently separating sodium sulfate.

In the present invention, it is understood that the first ammonia-containing steam and the second ammonia-containing steam are so-called secondary steam in the art. The pressures are all pressures in gauge pressure.

In the present invention, the sequence of the first heat exchange, the adjustment of the pH value of the wastewater to be treated, and the preparation of the wastewater to be treated (in the case where the wastewater to be treated contains a liquid phase obtained by the solid-liquid separation of the catalyst production wastewater and the second solid-liquid separation, the preparation of the wastewater to be treated is required) is not particularly limited, and may be appropriately selected as needed, and is completed before the first evaporation of the wastewater to be treated.

In the present invention, the first evaporation is to concentrate the wastewater to be treated, obtain relatively concentrated ammonia water, increase the concentration of ions, and increase the precipitation rate of cooling crystals, and the degree of the first evaporation is not particularly limited, and may be selected according to the requirements of the wastewater to be treated and the components of the wastewater to be treated, so as to meet the requirements of the cooling crystals on the first concentrated solution. For example, evaporation can be controlled to obtain only a small amount of ammonia-containing steam, so as to obtain ammonia water with higher concentration; and evaporation can be controlled to be fully performed, so that the wastewater to be treated is fully concentrated, and subsequent cooling crystallization is facilitated. Preferably, sodium chloride is not crystallized out in the first evaporation.

In the present invention, the first evaporation may be performed using an evaporation apparatus conventional in the art, such as an MVR evaporation apparatus, a single-effect evaporation apparatus, or a multi-effect evaporation apparatus. As the MVR evaporation means, for example, one or more selected from the group consisting of an MVR falling film evaporator, an MVR forced circulation evaporator, an MVR-FC continuous crystallization evaporator, and an MVR-OSLO continuous crystallization evaporator may be mentioned. Among them, preferred are an MVR forced circulation evaporator and an MVR-FC continuous crystallization evaporator, and more preferred is a falling film + forced circulation two-stage MVR evaporative crystallizer. The evaporator as the single-effect evaporator or the multiple-effect evaporator may be, for example, one or more selected from falling-film evaporators, rising-film evaporators, wiped-plate evaporators, central-circulation-tube-type multiple-effect evaporators, basket-suspended evaporators, external-heat evaporators, forced-circulation evaporators and lien evaporators. Among them, a forced circulation evaporator and an external heating evaporator are preferable. The respective evaporators of the multi-effect evaporation apparatus are composed of a heating chamber and an evaporation chamber, and may further include other evaporation auxiliary components as necessary, such as a demister for further separating liquid foam, a condenser for condensing all secondary steam, and a vacuum apparatus for pressure reduction operation. The number of evaporators included in the multi-effect evaporation apparatus is not particularly limited, and may be 2 or more, and more preferably 3 to 5. According to a preferred embodiment of the present invention, said first evaporation is carried out by means of a first MVR evaporation device 3.

According to the present invention, the conditions of the first evaporation are not particularly limited, and the purpose of concentrating the wastewater to be treated can be achieved. For example, the conditions of the first evaporation may include: the temperature is above 35 ℃ and the pressure is above-98 kPa. In order to improve the efficiency of evaporation, preferably, the first evaporation condition may include: the temperature is 75-130 ℃, and the pressure is-73 kPa-117 kPa; preferably, the conditions of the first evaporation include: the temperature is 85-130 ℃, and the pressure is-58 kPa-117 kPa; preferably, the conditions of the first evaporation include: the temperature is 95-110 ℃, and the pressure is-37 kPa-12 kPa; preferably, the conditions of the first evaporation include: the temperature is 100-110 ℃, and the pressure is-23 kPa-12 kPa.

In the present invention, the operation pressure of evaporation is preferably the saturated vapor pressure of the evaporated feed liquid. Further, the evaporation amount of the evaporation may be appropriately selected depending on the capacity of the apparatus to treat and the amount of the waste water to be treated, and may be, for example, 0.1m ³ More than h (e.g. 0.1 m) ³ /h～500m ³ /h)。

By appropriately controlling the conditions of the first evaporation, 80 mass% or more, preferably 90 mass% or more, of the ammonia contained in the wastewater to be treated can be obtained by evaporation, and for example, 80 mass%, 83 mass%, 85 mass%, 86 mass%, 87 mass%, 88 mass%, 89 mass%, 90 mass%, 91 mass%, 93 mass%, 95 mass%, or 98 mass% can be obtained, and the first aqueous ammonia can be directly recycled in the production process of the catalyst, or can be recycled after being neutralized with an acid to obtain an ammonium salt, or can be used by blending with water and a corresponding ammonium salt or aqueous ammonia.

In the present invention, the degree of progress of the first evaporation is performed by monitoring the concentration of the first evaporation-derived liquid, specifically, by controlling the concentration of the first evaporation-derived liquid within the above range, so that the first evaporation does not cause crystallization of sodium chloride in the wastewater to be treated. The concentration of the liquid resulting from the first evaporation is monitored by measuring the density, which may be carried out using a densitometer.

According to the present invention, the pH of the wastewater to be treated is adjusted to be greater than 9, preferably greater than 10.8 before the wastewater to be treated is subjected to the first evaporation, and the upper limit of the adjustment of the pH is not limited, and may be, for example, 14 or less, preferably 13.5 or less, more preferably 13 or less, further preferably 12 or less, and still further preferably 11.5 or less. By carrying out the first evaporation at the above pH, the evaporation of ammonia can be promoted to obtain aqueous ammonia of higher concentration, and it is convenient to obtain high purity sodium sulfate and sodium chloride crystals in the subsequent crystallization.

Specific examples of adjusting the pH of the wastewater to be treated before subjecting the wastewater to be treated to the first evaporation include: 9. 9.5, 9.6, 9.7, 9.8, 9.9, 10, 10.1, 10.2, 10.3, 10.4, 10.5, 10.6, 10.7, 10.8, 10.9, 11, 11.1, 11.2, 11.3, 11.4, 11.5, 11.6, 11.7, 11.8, 11.9, 12, 12.2, 12.4, 12.6, 12.8, 13, 13.5, or 14, etc.

In the present invention, the method of adjusting the pH is not particularly limited, and for example, the pH of the wastewater to be treated may be adjusted by adding an alkaline substance. The alkaline substance is not particularly limited, and the purpose of adjusting the pH value may be achieved. The alkaline substance is preferably NaOH in order not to introduce new impurities in the wastewater to be treated, increasing the purity of the crystals obtained. In addition, the second mother liquor (i.e., the liquid phase obtained by the second solid-liquid separation) contains NaOH at a relatively high concentration, and it is preferable to use the second mother liquor as the basic substance, and further NaOH may be added.

The manner of adding the basic substance may be any manner known in the art, but it is preferable to mix the basic substance with the wastewater to be treated in the form of an aqueous solution, and for example, an aqueous solution containing the basic substance may be introduced into a pipe through which the wastewater to be treated is introduced and mixed. The content of the alkaline substance in the aqueous solution is not particularly limited as long as the above-mentioned purpose of adjusting the pH value can be achieved. However, in order to reduce the amount of water used and further reduce the cost, it is preferable to use a saturated aqueous solution of an alkaline substance or a second mother liquor. In order to monitor the pH value of the wastewater to be treated, the pH value of the wastewater to be treated may be measured after the above-mentioned pH value adjustment.

According to a preferred embodiment of the present invention, the first evaporation is performed in the first MVR evaporation device 3, pH adjustment is performed by introducing and mixing the aqueous solution containing the alkaline substance in the pipe that feeds the wastewater to be treated to the first MVR evaporation device 3 before feeding the wastewater to be treated to the first MVR evaporation device 3, and the adjusted pH is measured by the first pH measurement device 61 and the third pH measurement device 60 after the adjustment.

According to the present invention, in order to fully utilize the heat of the first ammonia-containing steam, it is preferable that before the wastewater to be treated is subjected to the first evaporation, the wastewater to be treated is subjected to the first heat exchange with the first ammonia-containing steam to obtain the first ammonia water, and at the same time, the temperature of the wastewater to be treated is raised to facilitate the evaporation.

According to a preferred embodiment of the present invention, the first heat exchange between the wastewater to be treated and the first ammonia-containing steam is performed by the first heat exchange device 31 and the eighth heat exchange device 38, specifically, the ammonia-containing steam is sequentially passed through the eighth heat exchange device 38 and the first heat exchange device 31, meanwhile, the wastewater to be treated is heat exchanged with the first ammonia-containing steam condensate by the first heat exchange device 31, and then the wastewater to be treated is passed through the eighth heat exchange device 38 to exchange heat with the first ammonia-containing steam. Through the first heat exchange, the obtained first ammonia water is stored in the first ammonia water storage tank 51, and simultaneously, the temperature of the wastewater to be treated is raised to 74-129 ℃, preferably 94-109 ℃, so that the evaporation is convenient to carry out.

According to the present invention, in order to fully utilize the heat of the first concentrated solution, it is preferable that before the wastewater to be treated is subjected to the first evaporation, the wastewater to be treated is subjected to the first heat exchange with the first concentrated solution, so that the temperature of the first concentrated solution is lowered to facilitate cooling crystallization, and the temperature of the wastewater to be treated is raised to facilitate evaporation.

According to a preferred embodiment of the present invention, the first heat exchange between the wastewater to be treated and the first concentrated solution is performed by the tenth heat exchange device 30, and the wastewater to be treated is heat exchanged with the first concentrated solution by the tenth heat exchange device 30.

The first heat exchanger 31, the tenth heat exchanger 30 and the eighth heat exchanger 38 are not particularly limited, and various heat exchangers conventionally used in the art may be used to perform heat exchange. Specifically, it may be a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger of duplex stainless steel, titanium and titanium alloy, hastelloy can be selected as the material, and the heat exchanger containing plastic material can be selected when the temperature is lower.

In the present invention, in order to increase the solid content in the first MVR evaporation device 3 and reduce the ammonia content in the liquid, it is preferable that a part of the liquid after evaporation by the first MVR evaporation device 3 (i.e. the liquid located inside the first MVR evaporation device, hereinafter also referred to as the first circulation liquid) is heated and then returned to the first MVR evaporation device 3 for evaporation. The above-mentioned process of returning the first circulation liquid to the first MVR evaporation device 3 is preferably to return the first circulation liquid to the first MVR evaporation device 3 after mixing with the wastewater to be treated after the first pH adjustment and before the second pH adjustment, for example, the first circulation liquid may be returned to the wastewater delivery pipeline between the first heat exchange device 31 and the eighth heat exchange device 38 by the fifth circulation pump 75 to be mixed with the wastewater to be treated, and then after the second pH adjustment, heat exchange is performed by the eighth heat exchange device 38, and finally the wastewater is sent to the first MVR evaporation device 3. Here, the first reflux ratio means: the ratio of the amount of reflux to the total amount of liquid fed to the first MVR evaporator 3 minus the amount of reflux. As a ratio of returning a part of the liquid evaporated by the first MVR evaporation device 2 to the MVR evaporation device 2, there is no particular limitation, and for example, the first return ratio of the first evaporation may be 10 to 200, preferably 40 to 150.

According to the present invention, preferably, the method further comprises compressing the first ammonia-containing vapor before the first heat exchange. The compression of the first ammonia-containing vapor may be performed by a first compressor 101. Through right first ammonia vapor that contains compresses, for input energy among the MVR vaporization system, guarantee that waste water intensification-evaporation-cooling's process goes on in succession, need input start-up steam when MVR vaporization process starts, only need pass through first compressor 101 energy supply after reaching continuous running state, no longer need input other energy. The first compressor 101 may employ various compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, or a roots compressor. After being compressed by the first compressor 101, the temperature of the first ammonia-containing vapor is raised by 5 to 20 ℃.

In the present invention, the purpose of the cooling crystallization is to precipitate sodium sulfate, but sodium chloride is not precipitated, and sodium sulfate can be separated from wastewater favorably. The cooling crystallization only precipitates sodium sulfate and does not exclude sodium chloride entrained by or adsorbed on the surface of the sodium sulfate crystals. In the present invention, the content of sodium sulfate in the obtained sodium sulfate crystals is preferably 92% by mass or more, more preferably 96% by mass or more, and further preferably 98% by mass or more), it is understood that the amount of the obtained sodium sulfate crystals is based on a dry basis. When the content of sodium sulfate in the obtained sodium sulfate crystal is within the above range, it is considered that only sodium sulfate is precipitated.

Adjusting the pH value of the first concentrated solution to be more than 9 through the pH value before evaporation, wherein the NH ₄ ⁺ Most of the ammonia molecules are evaporated during the first evaporation, so that ammonium sulfate and/or ammonium chloride are not separated out in the cooling crystallization process, and the sodium sulfate separation rate can be improved due to the increase of the concentration of sodium chloride.

In the present invention, the conditions for the cooling crystallization are not particularly limited and may be appropriately selected as necessary, and the effect of crystallizing the sodium sulfate may be obtained. The cooling crystallization conditions may include: the temperature is-21.7-17.5 ℃, preferably-20-5 ℃, more preferably-10-5 ℃, further-10-0 ℃, and particularly preferably-4-0 ℃; the time (in terms of the residence time in the cooling crystallization apparatus 2) is 5min or more, preferably 60min to 180min, more preferably 90min to 150min, and still more preferably 120min to 150min. By controlling the cooling crystallization conditions within the above range, sodium sulfate can be sufficiently precipitated.

Specific examples of the temperature for cooling crystallization include: -21 ℃, -20 ℃, -19 ℃, -18 ℃, -17 ℃, -16 ℃, -15 ℃, -14 ℃, -13 ℃, -12 ℃, -11 ℃, -10 ℃, -9 ℃, -8 ℃, -7 ℃, -6 ℃, -5 ℃, -4 ℃, -3 ℃, -2 ℃, -1 ℃ or 0 ℃.

Specific examples of the time for cooling crystallization include: 5min, 6min, 7min, 8min, 10min, 15min, 20min, 25min, 30min, 35min, 40min, 45min, 50min, 52min, 54min, 56min, 58min, 60min, 65min, 70min, 75min, 80min, 85min, 90min, 95min, 100min, 105min, 110min, 115min, 120min, 130min, 140min, 150min, or 160min.

According to the present invention, the cooling crystallization is carried out in a continuous or batch manner, and the temperature of the first concentrated solution is lowered to precipitate sodium sulfate crystals, and the continuous cooling crystallization is preferably carried out. The cooling crystallization of sodium sulfate can be carried out by various cooling crystallization devices conventionally used in the art, for example, by using a continuous cooling crystallizer with an external cooling heat exchanger, or by using a crystallization tank having a cooling means, such as the cooling crystallization device 2. The cooling part can lead the first concentrated solution in the cooling crystallization device to be cooled to the condition required by cooling crystallization by introducing a cooling medium. The cooling crystallization device is preferably provided with a mixing part, such as a stirrer, and the like, so that the first concentrated solution is mixed to achieve the effect of uniform cooling, sodium sulfate in the first concentrated solution can be fully precipitated, and the size of crystal grains can be increased. The cooling crystallization device is preferably provided with a circulating pump, so as to avoid generating a large amount of fine crystal nuclei and prevent crystal grains in circulating crystal slurry from colliding with the impeller at a high speed to generate a large amount of secondary crystal nuclei, and the circulating pump is preferably a low-rotation-speed centrifugal pump, more preferably a high-flow and low-rotation-speed guide pump impeller or a high-flow, low-lift and low-rotation-speed axial pump.

By carrying out the cooling crystallization under the above conditions, sodium sulfate can be sufficiently precipitated in the cooling crystallization without precipitating sodium chloride, thereby achieving the purpose of separating and purifying sodium sulfate.

According to the present invention, before the first concentrated solution is cooled and crystallized, the concentration of sodium chloride in the first concentrated solution may be adjusted as necessary so that the concentration of sodium chloride in the crystal solution obtained by cooling and crystallization is X or less, where X is the concentration of sodium chloride at which both sodium sulfate and sodium chloride in the crystal solution are saturated under the condition of cooling and crystallization. By adjusting the concentration of sodium chloride in the first concentrated solution to the above range, it can be ensured that sodium chloride is not precipitated in the cooled crystals. Preferably, the concentration of sodium chloride in the crystallization liquid is made to be 0.95X-0.999X. The method of adjusting the sodium chloride concentration may be performed by mixing the second mother liquor, water, the catalyst production wastewater and/or other waste liquid generated during the treatment, etc. at an appropriate concentration, as long as the concentration of sodium chloride in the first concentrated solution is adjusted to be within the above range. To avoid introducing more liquid, it is preferred that the concentration of sodium chloride in the first concentrate is adjusted by mixing the second mother liquor, the catalyst production wastewater and/or the wash liquor after leaching the sodium sulfate crystals. By adjusting the concentration of the sodium chloride in the first concentrated solution to the range, the precipitation of the sodium chloride in the cooling crystallization process is avoided, the purity of the sodium sulfate precipitated in the cooling crystallization process is improved, and the cooling crystallization efficiency is improved.

In the present invention, in order to control the crystal size distribution in the cooling crystallization device 2 and reduce the content of fine crystal grains, it is preferable that a part of the liquid crystallized by the cooling crystallization device 2 (i.e., the liquid located inside the cooling crystallization device 2, hereinafter also referred to as cooling circulation liquid) is mixed with the catalyst production wastewater and returned to the cooling crystallization device 2 for cooling crystallization again. The process of returning the cooling circulation liquid to the cooling crystallization device 2 for crystallization can be implemented by, for example, returning the cooling circulation liquid to the sixth heat exchange device 36 through the second circulation pump 72, mixing the cooling circulation liquid with the first concentrated liquid, and then entering the cooling crystallization device 2 again for cooling crystallization. The return amount of the cooling circulation liquid can be defined by a circulation ratio of cooling, and the circulation ratio of the cooling crystallization is as follows: the ratio of the amount of the liquid circulated to the total amount of the liquid fed to the cooling crystallization device 2 minus the amount of the reflux. The circulation ratio may be appropriately set according to the supersaturation degree of sodium sulfate in the cooling crystallization apparatus 2 to ensure the particle size of sodium sulfate crystals. In order to control the particle size distribution of crystals obtained by cooling crystallization and to reduce the content of fine crystal grains, it is preferable to control the supersaturation degree to less than 1.5g/L, more preferably to less than 1g/L.

In the invention, the sodium sulfate crystal and the first mother liquor (i.e. the liquid phase obtained by the first solid-liquid separation) are obtained after the first solid-liquid separation is carried out on the crystallization liquid containing the sodium sulfate crystal. The method of the first solid-liquid separation is not particularly limited, and may be selected from, for example, one or more of centrifugation, filtration, and sedimentation.

According to the present invention, the first solid-liquid separation may be performed by using a first solid-liquid separation device (for example, a centrifuge, a belt filter, a plate filter, or the like) 91. After the first solid-liquid separation, the first mother liquor obtained by the first solid-liquid separation device 91 is temporarily stored in the first mother liquor tank 53, and can be sent to the second MVR evaporation device 1 by the sixth circulation pump 76 for second evaporation. In addition, it is difficult to avoid that the obtained sodium sulfate crystals adsorb impurities such as chlorine ions, free ammonia, and hydroxide ions to a certain extent, and in order to remove the adsorbed impurities, reduce the odor of solid salts, reduce corrosiveness, and improve the purity of the crystals, the sodium sulfate crystals are preferably subjected to a first washing with water or a sodium sulfate solution, and may be dried when anhydrous sodium sulfate is required to be obtained. The first washing method is preferably rinsing, and rinsing is preferably performed after solid-liquid separation.

The form of the first solid-liquid separation and the first washing is not particularly limited, and may be carried out, for example, by using a solid-liquid separation apparatus which is conventional in the art, or may be carried out in a staged solid-liquid separation apparatus. The washing is not particularly limited and may be carried out by a method conventional in the art. The number of washing is not particularly limited, and may be 1 or more, and is preferably 2 to 4 times in order to obtain sodium sulfate crystals with higher purity. The first washing is preferably carried out using an aqueous sodium sulfate solution (the concentration of which is preferably the concentration of sodium sulfate in the aqueous solution at which the sodium chloride and sodium sulfate reach saturation simultaneously at the temperature corresponding to the sodium sulfate crystals to be washed). As for the liquid resulting from the washing, it is preferable that the water or the washing solution of sodium sulfate aqueous solution is returned to the cooling crystallization device 2, for example, may be returned to the cooling crystallization device 2 by the eighth circulation pump 78.

According to a preferred embodiment of the present invention, after cooling and crystallizing the obtained crystal liquid containing sodium sulfate, solid-liquid separation is performed by a solid-liquid separation apparatus, and the crystal obtained by the solid-liquid separation is rinsed again with an aqueous sodium sulfate solution (the concentration of the aqueous sodium sulfate solution is the concentration of sodium sulfate in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at a temperature corresponding to the sodium sulfate crystal to be washed), and the rinsed liquid is returned to the cooling and crystallizing apparatus 2. By the above washing process, the purity of the obtained sodium sulfate crystals can be improved.

According to the present invention, in order to fully utilize the refrigeration capacity of the first mother liquor, it is preferable that the first mother liquor and the first concentrated solution are subjected to the second heat exchange before the first concentrated solution is subjected to the cooling crystallization.

According to a preferred embodiment of the present invention, the second heat exchange is performed by the second heat exchange device 32, and specifically, the first mother liquid and the first concentrated liquid are respectively passed through the second heat exchange device 32 to perform heat exchange, so that the temperature of the first concentrated liquid is lowered to facilitate the cooling crystallization, and the temperature of the first mother liquid is raised to facilitate the second evaporation. After the second heat exchange is performed by the second heat exchange device 32, the temperature of the first concentrated solution is-20.7 ℃ to 16.5 ℃, preferably-5 ℃ to 10 ℃, and is close to the temperature of cooling crystallization.

According to the invention, in order to facilitate the cooling crystallization, the first concentrated solution and the refrigerating fluid are subjected to second heat exchange. According to a preferred embodiment of the present invention, the second heat exchange between the first concentrated solution and the refrigerating fluid is performed by the sixth heat exchange device 36, and specifically, the refrigerating fluid and the first concentrated solution are respectively passed through the sixth heat exchange device 36 to perform heat exchange therebetween, so that the temperature of the first concentrated solution is lowered to facilitate the cooling crystallization. The refrigerating fluid can be the refrigerating fluid which is used for reducing the temperature conventionally in the field, as long as the temperature of the first concentrated solution can meet the requirement of cooling crystallization.

The second heat exchanger 32 and the sixth heat exchanger 36 are not particularly limited, and various heat exchangers conventionally used in the art may be used to perform heat exchange. Specifically, a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like may be mentioned, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger made of duplex stainless steel, titanium alloy and hastelloy can be selected, and the heat exchanger made of plastic can be selected when the temperature is lower. The second heat exchange device 32 is preferably a heat exchanger made of plastic.

In the present invention, the second evaporation is intended to separate the ammonia and salts in the wastewater by precipitating sodium chloride and/or sodium sulfate and evaporating further ammonia. According to the invention, by controlling the conditions of the second evaporation, with the progressive reduction of the solvent, sodium chloride is first precipitated, possibly sodium sulphate, obtaining a second concentrated solution containing crystals. In order to reduce the amount of circulating water in the treatment system and to increase the efficiency of the second evaporation and thus of the wastewater treatment, said second evaporation is preferably carried out to such an extent that sodium chloride and sodium sulfate are precipitated simultaneously, that is to say that the second evaporation preferably results in a second concentrated solution comprising sodium sulfate crystals and sodium chloride crystals.

In the present invention, the second MVR evaporation device 1 is not particularly limited, and may be various MVR evaporation devices conventionally used in the art. For example, it may be one or more selected from the group consisting of an MVR falling film evaporator, an MVR forced circulation evaporator, an MVR-FC continuous crystallization evaporator, and an MVR-OSLO continuous crystallization evaporator. Among them, preferred are an MVR forced circulation evaporator and an MVR-FC continuous crystallization evaporator, and more preferred is a falling film + forced circulation two-stage MVR evaporative crystallizer.

In the present invention, the evaporation conditions of the second evaporation are not particularly limited, and may be appropriately selected as necessary to achieve the purpose of precipitating crystals. The conditions of the second evaporation may include: the temperature is above 17.5 ℃ and the pressure is above-101 kPa; preferably, the conditions of the second evaporation include: the temperature is 35-110 ℃, and the pressure is-98 kPa-12 kPa; preferably, the conditions of the second evaporation include: the temperature is 45-110 ℃, and the pressure is-95 kPa-12 kPa; preferably, the conditions of the second evaporation include: the temperature is 105-110 ℃, and the pressure is-8 kPa-11.34 kPa.

In the present invention, the operating pressure of the second evaporation is preferably the saturated vapor pressure of the evaporated feed liquid. Further, the evaporation amount of the second evaporation may be appropriately selected depending on the capacity of the apparatus to treat and the amount of the waste water to be treated, and may be, for example, 0.1m ³ More than h (e.g. 0.1 m) ³ /h～500m ³ /h)。

In order to ensure that sodium chloride crystals are precipitated as much as possible in the evaporation process without precipitation or precipitation of a small amount of sodium sulfate and can be dissolved at the time of temperature reduction treatment, it is preferable that the amount of SO contained in the liquid phase obtained by the first solid-liquid separation is 1mol based on 1mol of SO contained in the liquid phase ₄ ^2- Cl contained in the liquid phase obtained by the first solid-liquid separation ^- Is 7.15mol or more, preferably 20mol or more, more preferably 44mol or more, more preferably 50mol or more, more preferably 74mol or more, preferably 460mol or less, more preferably 230mol or less. For example, the amount of the organic solvent may be 7.5mol, 7.8mol, 8mol, 8.2mol, 8.4mol, 8.6mol, 8.8mol, 9mol, 9.2mol, 9.4mol, 9.5mol, 10.5mol, 11mol, 11.5mol, 12mol, 12.5mol, 13mol, 13.5mol, 14mol, 14.5mol, 15mol, 15.5mol, 16mol, 16.5mol, 17mol, 17.5mol, 18mol, 18.5mol, 19mol, 19.5mol, 20mol, 21mol, 22mol, 23mol, 25mol, 27mol, 29mol, 31mol, 35mol, 40mol, 45mol, 50mol, 60mol, 65mol or the like. By reacting SO ₄ ^2- And Cl ^- The molar ratio of the sodium sulfate to the sodium chloride is controlled within the range, and relatively pure sodium chloride crystals can be obtained through evaporation and cooling treatment, so that the sodium sulfate and the sodium chloride are obtainedAnd the energy consumption in the cooling crystallization process is reduced.

According to the present invention, from the viewpoint of improving the efficiency of wastewater treatment, the higher the degree of progress of the second evaporation, the better; however, if the second evaporation exceeds a certain level, the treatment solution containing only sodium chloride crystals cannot be obtained by the temperature reduction treatment, and in this case, the crystals can be dissolved by adding water to the treatment solution, but the efficiency of wastewater treatment is impaired. Therefore, the second evaporation is preferably performed to such an extent that sodium chloride crystals and sodium sulfate crystals are simultaneously precipitated, that is, preferably, the second concentrated solution containing crystals obtained in step 3) is a concentrated solution containing sodium chloride crystals and sodium sulfate crystals, and the temperature reduction treatment dissolves the sodium sulfate crystals in the concentrated solution containing sodium chloride crystals and sodium sulfate crystals. In order to dissolve the sodium sulfate crystals in the concentrated solution containing the sodium chloride crystals and the sodium sulfate crystals in the temperature reduction treatment, for example, the evaporation degree of the second evaporation may be controlled so that the concentration of sodium sulfate in the treatment solution is Y or less (where Y is the concentration of sodium sulfate when both sodium sulfate and sodium chloride in the treatment solution are saturated under the temperature reduction treatment). In the subsequent temperature-lowering treatment step, the concentration of sodium sulfate in the treatment solution is preferably 0.9Y to 0.99Y, more preferably 0.95Y to 0.98Y, from the viewpoint of precipitating sodium chloride as much as possible and completely dissolving sodium sulfate. By controlling the degree of the second evaporation within the above range, sodium chloride can be precipitated as much as possible in the second evaporation process, and in the cooling treatment, the sodium sulfate crystals are completely dissolved, and finally pure sodium chloride crystals are separated. By crystallizing sodium chloride in the second evaporation as much as possible, the wastewater treatment efficiency can be improved, and energy can be saved.

In the present invention, the proceeding degree of the second evaporation is performed by monitoring the evaporation amount of the second evaporation to obtain the amount of the liquid, specifically, the concentration multiple is controlled by controlling the evaporation amount of the second evaporation, that is, the amount of the ammonia water, so that the sodium sulfate crystals precipitated in the second concentrated solution obtained by the second evaporation can be dissolved during the temperature reduction treatment. The degree of the second evaporation concentration is monitored by measuring the second evaporation amount, and specifically, a mass flow meter can be used for flow measurement, the amount of the secondary steam can be measured, and the amount of the condensate can also be measured.

In the present invention, in order to increase the liquid salt concentration in the second MVR evaporation device 1 and reduce the ammonia content in the liquid, it is preferable to return part of the liquid evaporated by the second MVR evaporation device 1 (i.e. the liquid located inside the second MVR evaporation device, hereinafter also referred to as circulating liquid) to the second MVR evaporation device 1 for evaporation, and it is preferable to return the liquid to the second MVR evaporation device 1 after heating for evaporation. The above-described process of returning the circulating liquid to the second MVR evaporating device 1 may be returned to the third heat exchanging process by, for example, the seventh circulating pump 77. The second reflux ratio of the second evaporation is: the ratio of the amount of reflux to the total amount of liquid fed to the second MVR evaporator 1 minus the amount of reflux. The second reflux ratio can be set properly according to the evaporation capacity, so that the first mother liquor after temperature rise can be ensured to evaporate water and ammonia with required amount at the given second evaporation temperature after entering the second MVR evaporation device. The second reflux ratio of the second evaporation may be, for example, 10 to 200, preferably 40 to 150.

According to the present invention, preferably, the method further comprises compressing the second ammonia-containing vapor before the third heat exchange. The compression of the second ammonia-containing vapor may be performed by a second compressor 102. The second ammonia-containing steam is compressed, energy is input into the MVR evaporation system, the continuous process of waste water heating, evaporation and cooling is guaranteed, starting steam needs to be input when the MVR evaporation process is started, energy is supplied only through the second compressor 102 after the continuous operation state is achieved, and other energy does not need to be input any more. The second compressor 102 may employ various second compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, a roots compressor, or the like. After compression by the second compressor 102, the temperature of the second ammonia-containing vapor is raised by 5 to 20 ℃.

According to the present invention, in order to fully utilize the heat in the second ammonia-containing vapor obtained by the second evaporation, it is preferable to subject the first mother liquor to a third heat exchange with the second ammonia-containing vapor before sending the first mother liquor to the second MVR evaporation device 1.

According to a preferred embodiment of the present invention, the third heat exchange of the first mother liquor and the second ammonia-containing vapor is carried out by a third heat exchange means 33 and a fourth heat exchange means 34, respectively. Specifically, the first mother liquor passes through the third heat exchange device 33 and the fourth heat exchange device 34 in sequence, and the second ammonia-containing steam passes through the fourth heat exchange device 34 and the third heat exchange device 33 in sequence, so that the temperature of the first mother liquor is raised to facilitate second evaporation, and the second ammonia-containing steam is condensed to obtain ammonia water. After heat exchange is carried out by the third heat exchange device 33, the temperature of the first mother liquor is raised to 34-109 ℃, preferably 44-109 ℃; after heat exchange by the fourth heat exchange device 34, the temperature of the first mother liquor is raised to 42 ℃ to 117 ℃, preferably 52 ℃ to 117 ℃.

According to the present invention, in order to sufficiently utilize the heat in the second concentrated liquid containing crystals obtained by the second evaporation, it is preferable to subject the second concentrated liquid containing crystals to the third heat exchange with the first mother liquor before the second evaporation.

According to a preferred embodiment of the present invention, the third heat exchange of the second concentrate containing crystals with the first mother liquor is carried out by means of a fifth heat exchange means 35. Specifically, the first mother liquor and the second concentrated solution containing crystals are respectively passed through the fifth heat exchange device 35, so that the first mother liquor is heated to facilitate the second evaporation, and the second concentrated solution containing crystals is cooled to facilitate the cooling treatment. After the heat exchange is carried out by the fifth heat exchange device 35, the temperature of the first mother liquor is raised to 34-109 ℃, preferably 44-109 DEG C

The third heat exchange device 33, the fourth heat exchange device 34 and the fifth heat exchange device 35 are not particularly limited, and various heat exchangers conventionally used in the field can be used to achieve the purpose of heat exchange between the second ammonia-containing steam and the first mother liquor. Specifically, a jacketed heat exchanger, a plate heat exchanger, a shell-and-tube heat exchanger, or the like may be mentioned, with the plate heat exchanger being preferred. The material of the heat exchanger can be specifically selected according to the needs, for example, in order to resist the corrosion of chloride ions, the heat exchanger of duplex stainless steel, titanium and titanium alloy, hastelloy can be selected as the material, and the heat exchanger containing plastic material can be selected when the temperature is lower. Preferably, a duplex stainless steel plate heat exchanger is used.

According to the present invention, it is preferred that only the pH of the first mother liquor is monitored, for example by the second pH measuring device 62, before it is passed into the second MVR evaporator 1.

In the present invention, the purpose of the temperature reduction treatment is to dissolve sodium sulfate crystals that may be contained in the second concentrated solution containing crystals, thereby further precipitating sodium chloride. The temperature reduction treatment for dissolving the sodium sulfate crystals in the second concentrated solution containing the crystals means that the degree of the second evaporation needs to be properly controlled in order to obtain pure sodium chloride crystals, that is, the sodium sulfate in the mixed system does not exceed the solubility under the condition of the corresponding temperature reduction treatment. In addition, sodium chloride crystals entrain or adsorb sodium sulfate crystals on the surface during the temperature reduction treatment. In the present invention, the content of sodium sulfate in the obtained sodium chloride crystals is preferably 8% by mass or less, and preferably 4% by mass or less, and in the present invention, it is considered that sodium sulfate is dissolved when the content of sodium sulfate crystals in the obtained sodium chloride crystals is 8% by mass or less.

The conditions for performing the temperature reduction treatment are not particularly limited, and the sodium sulfate crystals in the second concentrated solution containing the crystals may be completely dissolved in the temperature reduction treatment process, for example, the conditions for performing the temperature reduction treatment may include: the temperature is 13 to 100 ℃, preferably 15 to 45 ℃, more preferably 15 to 35 ℃, and still more preferably 17.9 to 35 ℃. In order to ensure the effect of the temperature reduction treatment, preferably, the conditions of the temperature reduction treatment include: the time is more than 5min, preferably 5min to 120min, and more preferably 30min to 90min; more preferably 50 to 60min.

Specific examples of the temperature lowering treatment include: 13 deg.C, 14 deg.C, 15 deg.C, 15.5 deg.C, 16 deg.C, 16.5 deg.C, 17 deg.C, 17.5 deg.C, 17.9 deg.C, 18 deg.C, 18.5 deg.C, 19 deg.C, 19.5 deg.C, 20 deg.C, 21 deg.C, 23 deg.C, 25 deg.C, 27 deg.C, 30 deg.C, 31.5 deg.C, 32 deg.C, 33 deg.C, 34 deg.C, 40 deg.C, 45 deg.C, 50 deg.C, 55 deg.C, etc.

Specific examples of the time for the temperature reduction treatment include: 5min, 6min, 7min, 8min, 10min, 15min, 20min, 25min, 30min, 35min, 40min, 45min, 50min, 52min, 54min, 56min, 58min, 60min, 70min, 100min, 120min.

According to the present invention, the temperature reduction treatment is performed in the low-temperature treatment tank 55, and the treatment solution containing sodium chloride crystals is obtained after the temperature reduction treatment of the second concentrated solution containing crystals is performed in the low-temperature treatment tank 55. The low-temperature treatment tank 55 is not particularly limited, and may be, for example, a thickener, a crystallization tank with agitation, a crystallization tank with external circulation, or the like, and among them, a crystallization tank with agitation is preferable. The low-temperature treatment tank 55 is preferably provided with a kneading means for bringing the second concentrated solution into a kneaded state in the temperature reduction treatment, and for example, a conventionally used mechanical stirring, electromagnetic stirring and/or an external circulation device may be used, and it is preferable that the solid-liquid distribution in the second concentrated solution is brought into a uniform state. All parts of the second concentrated solution are kept at uniform temperature and uniform concentration through uniform mixing, so that the problem that the dissolution of sodium sulfate crystals cannot be fully carried out is avoided, and the efficiency of cooling treatment is improved. The low-temperature treatment tank 55 is preferably provided with a cooling means for cooling the low-temperature treatment tank 55 to a temperature required for the temperature reduction treatment by introducing a cooling medium, for example.

According to the present invention, the second solid-liquid separation may be performed by a second solid-liquid separation device (e.g., a centrifuge, a belt filter, a plate filter, etc.) 92. After the second solid-liquid separation, the second mother liquor obtained by the second solid-liquid separation device 92 (i.e. the liquid phase obtained by the second solid-liquid separation) is returned to the first MVR device 3 for evaporation again, and specifically, the second mother liquor can be returned to the place before the completion of the pH adjustment before the first evaporation by the ninth circulation pump 79. In addition, it is difficult to avoid that the obtained sodium chloride crystals adsorb certain impurities such as sulfate ions, free ammonia, hydroxide ions, etc., and in order to remove the adsorbed impurities, reduce the odor of solid salts, reduce corrosiveness, and improve the purity of the crystals, the sodium chloride crystals are preferably subjected to secondary washing with water, the catalyst production wastewater, or a sodium chloride solution and dried. In order to avoid dissolution of the sodium chloride crystals during washing, the sodium chloride crystals are preferably washed with an aqueous sodium chloride solution. More preferably, the concentration of the aqueous sodium chloride solution is preferably the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulphate reach saturation simultaneously at the temperature corresponding to the sodium chloride crystals to be washed. The second washing method is preferably performed by performing the elutriation before the rinsing. The second washing solution obtained in the above washing process is preferably returned to the second MVR evaporation device 1 by the tenth circulation pump 80 to perform the second evaporation again.

The manner of the second solid-liquid separation and the second washing is not particularly limited, and may be carried out by using, for example, a combination of an elutriation device and a solid-liquid separation device which are conventional in the art, or may be carried out on a staged solid-liquid separation device such as a belt filter. The above-mentioned elutriation and rinsing are not particularly limited and may be carried out by a method conventional in the art. The number of elutriation and rinsing is not particularly limited, and may be 1 or more, and is preferably 2 to 4 times in order to obtain sodium chloride crystals of higher purity. In the elutriation process, the washing liquid recovered by the second washing can be used in a countercurrent manner when used as the elutriation liquid. Before the elutriation, it is preferable to perform preliminary solid-liquid separation by sedimentation to obtain a slurry containing sodium chloride crystals (the liquid content may be 35% by mass or less). In the elutriation process, 1 to 20 parts by weight of a liquid is used for elutriation with respect to 1 part by weight of a slurry containing sodium chloride crystals. The rinsing is preferably carried out using an aqueous sodium chloride solution, the concentration of which is preferably the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at the temperature corresponding to the sodium chloride crystals to be washed. In order to further enhance the elutriation effect and obtain sodium chloride crystals with higher purity, it is preferable to perform elutriation using the eluted solution. As for the liquid generated by washing, preferably, water or an aqueous sodium chloride solution washing liquid and an elutriation liquid are returned to the second MVR evaporation device 1.

According to a preferred embodiment of the present invention, the treatment liquid containing sodium chloride crystals obtained by the temperature reduction treatment is subjected to preliminary solid-liquid separation by settling, then elutriated in another elutriation tank using a liquid obtained when sodium chloride crystals are subsequently washed, the elutriated treatment liquid containing sodium chloride crystals is sent to a solid-liquid separation apparatus for solid-liquid separation, the crystals obtained by the solid-liquid separation are washed with an aqueous sodium chloride solution (the concentration of the aqueous sodium chloride solution is the concentration of sodium chloride in an aqueous solution in which sodium chloride and sodium sulfate are simultaneously saturated at a temperature corresponding to the sodium chloride crystals to be washed), and the washed liquid is returned to the elutriation as an elutriation liquid. Through the washing process combining elutriation and leaching, the purity of the obtained sodium chloride crystal is improved, washing liquid is not excessively introduced, and the efficiency of wastewater treatment is improved.

According to a preferred embodiment of the invention, the tail gas generated by cooling crystallization is discharged after ammonia removal; discharging the tail gas which is remained by the third heat exchange condensation after ammonia removal; and discharging the tail gas which is remained by the condensation of the first heat exchange after ammonia removal. The tail gas generated by the cooling crystallization is the tail gas discharged by the cooling crystallization device 2, and the third heat exchange condenses the residual tail gas which is the non-condensable gas discharged by the fourth heat exchange device 34; the first heat exchange condenses the remaining tail gas, i.e., the tail gas discharged from the eighth heat exchange device 38. The ammonia in the tail gas is removed, so that the pollutant content in the discharged tail gas can be further reduced, and the tail gas can be directly discharged.

As the method of removing ammonia, absorption may be performed by the off-gas absorption tower 83. The off-gas absorption column 83 is not particularly limited, and may be any of various absorption columns conventionally used in the art, such as a plate-type absorption column, a packed absorption column, a falling film absorption column, or an empty column. Circulating water is arranged in the tail gas absorption tower 83, the circulating water circulates in the tail gas absorption tower 83 under the action of the fourth circulating pump 74, water can be supplemented into the tail gas absorption tower 83 from the circulating water tank 82 through the third circulating pump 73, fresh water can be supplemented into the circulating water tank 82, and meanwhile the temperature of working water of the vacuum pump 81 and the ammonia content can be reduced. The flow of the off-gas and the circulating water in the off-gas absorption tower 83 may be in a counter-current or co-current flow, preferably in a counter-current flow. The circulating water can be supplemented by additional fresh water. In order to ensure the sufficient absorption of the tail gas, dilute sulfuric acid may be further added to the tail gas absorption tower 83 to absorb a small amount of ammonia and the like in the tail gas. The circulating water can be used as ammonia water or ammonium sulfate solution for production or direct sale after absorbing tail gas. The off gas may be introduced into the off gas absorption tower 83 by a vacuum pump 81.

In the present invention, the catalyst production wastewater is not particularly limited as long as it contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The wastewater is obtained. In addition, the method is particularly suitable for treating high-salt ammonium-containing wastewater. The wastewater from the catalyst production of the present invention may be specifically wastewater from the production of a molecular sieve, alumina or an oil refining catalyst, or wastewater from the production of a molecular sieve, alumina or an oil refining catalyst after the following impurity removal and concentration. It is preferable that the wastewater from the production of molecular sieves, alumina or refinery catalysts is subjected to the following impurity removal and concentration.

As NH in the catalyst production wastewater ₄ ⁺ May be 8mg/L or more, preferably 300mg/L or more.

As Na in the wastewater from the catalyst production ⁺ May be 510mg/L or more, preferably 1g/L or more, more preferably 2g/L or more, further preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more.

As SO in wastewater from the production of said catalyst ₄ ^2- May be 1g/L or more, preferably 2g/L or more, more preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more, further preferably 70g/L or more.

As Cl in the catalyst production wastewater ^- May be 970mg/L or more, more preferably 2g/L or more, further preferably 4g/L or more, further preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferablyMore preferably 50g/L or more, and still more preferably 60g/L or more.

NH contained in the catalyst production wastewater ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ The upper limit of (3) is not particularly limited. SO in the wastewater from catalyst production from the viewpoint of easy wastewater treatment ₄ ^2- 、Cl ^- And Na ⁺ The upper limit of (b) is 200g/L or less, preferably 150g/L or less, respectively; NH in catalyst production wastewater ₄ ⁺ Is 50g/L or less, preferably 40g/L or less.

From the viewpoint of reducing energy consumption in the treatment process, SO contained in the catalyst production wastewater ₄ ^2- Preferably in a concentration higher than Cl ^- Relative to 1mol of SO ₄ ^2- ，Cl ^- Preferably 10mol or less, more preferably 8mol or less, for example, 1.3 to 6.5mol, from the viewpoint of improving the purity of the sodium sulfate product, cl contained in the wastewater from the catalyst production ^- Is 5.5mol/L or less, preferably 5mol/L or less, more preferably 4.5mol/L or less, for example, 1.8 to 3.9mol/L. By adding SO contained in the wastewater generated in the catalyst production ₄ ^2- And Cl ^- The concentration of the sodium sulfate is limited in the range, pure sodium sulfate can be obtained in the cooling and crystallizing process, energy is saved, and the treatment process is more economical.

In the present invention, the inorganic salt ions contained in the catalyst production wastewater are other than NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ In addition, it may contain Mg ²⁺ 、Ca ²⁺ 、K ⁺ 、Fe ³⁺ Inorganic salt ions such as rare earth element ions, mg ²⁺ 、Ca ²⁺ 、K ⁺ 、Fe ³⁺ The content of each inorganic salt ion such as a rare earth element ion is preferably 100mg/L or less, more preferably 50mg/L or less, still more preferably 10mg/L or less, and particularly preferably no other inorganic salt ion is contained. By controlling the other inorganic salt ions within the above range, the purity of the sodium sulfate crystals and sodium chloride crystals finally obtained can be further improved. In order to reduce other inorganic matters in the catalyst production wastewaterThe content of salt ions is preferably subjected to the following impurity removal.

The TDS of the catalyst production wastewater may be 1.6g/L or more, preferably 4g/L or more, more preferably 8g/L or more, further preferably 16g/L or more, further preferably 32g/L or more, further preferably 40g/L or more, further preferably 50g/L or more, further preferably 60g/L or more, further preferably 100g/L or more, further preferably 150g/L or more, further preferably 200g/L or more.

In the present invention, the pH of the catalyst production wastewater is preferably 4 to 8, for example 6 to 7.

In addition, since the COD of the wastewater may block a membrane during concentration, affect the purity and color of a salt during evaporative crystallization, etc., the COD of the wastewater from the catalyst production is preferably as small as possible (preferably 20mg/L or less, more preferably 10mg/L or less), and is preferably removed by oxidation during pretreatment, specifically, by biological method, advanced oxidation method, etc., and is preferably oxidized by an oxidizing agent such as Fenton's reagent when the COD content is very high.

In the invention, in order to reduce the concentration of impurity ions in the wastewater, ensure the continuous and stable operation of the treatment process and reduce the equipment operation and maintenance cost, the catalyst production wastewater is preferably subjected to impurity removal before being treated by the treatment method. Preferably, the impurity removal is selected from one or more of solid-liquid separation, chemical precipitation, adsorption, ion exchange and oxidation.

As the solid-liquid separation, filtration, centrifugation, sedimentation, or the like may be mentioned; as the chemical precipitation, pH adjustment, carbonate precipitation, magnesium salt precipitation, and the like may be mentioned; the adsorption can be physical adsorption and/or chemical adsorption, and the specific adsorbent can be selected from activated carbon, silica gel, alumina, molecular sieve, natural clay and the like; as the ion exchange, any one of a strongly acidic cation resin and a weakly acidic cation resin can be used; as the oxidation, various oxidizing agents conventionally used in the art, such as ozone, hydrogen peroxide, potassium permanganate, and in order to avoid introduction of new impurities, ozone, hydrogen peroxide, and the like are preferably used.

The specific impurity removal mode can be specifically selected according to the types of impurities contained in the catalyst production wastewater. For suspended matters, a solid-liquid separation method can be selected for removing impurities; for inorganic matters and organic matters, chemical precipitation, ion exchange and adsorption methods can be selected for removing impurities, such as weak acid cation exchange, activated carbon adsorption and the like; for organic matters, impurities can be removed by adopting an adsorption and/or oxidation mode, wherein an ozone biological activated carbon adsorption oxidation method is preferred. According to a preferred embodiment of the invention, the catalyst production wastewater is subjected to impurity removal by filtration, a weak acid cation exchange method and an ozone biological activated carbon adsorption oxidation method in sequence. Through the impurity removal process, most suspended matters, hardness, silicon and organic matters can be removed, the scaling risk of the device is reduced, and the continuous and stable operation of the wastewater treatment process is ensured.

In the present invention, the wastewater having a low salt content may be concentrated to have a salt content within a range required for the wastewater of the present invention before the wastewater is treated by the treatment method of the present invention (preferably, after the above-mentioned impurity removal). Preferably, the concentration is selected from ED membrane concentration and/or reverse osmosis; more preferably, the concentration is performed by ED membrane concentration and reverse osmosis, and the order of performing the ED membrane concentration and the reverse osmosis is not particularly limited. The ED membrane concentration and reverse osmosis treatment apparatus and conditions may be performed in a manner conventional in the art, and may be specifically selected according to the condition of wastewater to be treated. Specifically, as the concentration of the ED membrane, a one-way electrodialysis system or a reversed electrodialysis system can be selected for carrying out; as the reverse osmosis, a roll membrane, a plate membrane, a disc-tube membrane, a vibrating membrane or a combination thereof can be selected for use. Through the concentration can improve the efficiency of waste water treatment, avoid the energy waste that a large amount of evaporations caused.

In a preferred embodiment of the invention, the catalyst production wastewater is wastewater generated by chemical precipitation, filtration, weak acid cation exchange and ozone biological activated carbon adsorption oxidation of wastewater generated in the molecular sieve production process, and is concentrated by an ED membrane and a reverse osmosis method.

The conditions for the above chemical precipitation are preferably: sodium carbonate is used as a treating agent, 1.2-1.4mol of sodium carbonate is added relative to 1mol of calcium ions in the wastewater, the pH value of the wastewater is adjusted to be more than 7, the reaction temperature is 20-35 ℃, and the reaction time is 0.5-4h.

The conditions for the filtration are preferably: the filtering unit adopts a double-layer filtering material multi-medium filter consisting of anthracite and quartz sand, the grain diameter of the anthracite is 0.7-1.7mm, the grain diameter of the quartz sand is 0.5-1.3mm, and the filtering speed is 10-30m/h. After the filter material is used, the regeneration method of 'gas back flushing-gas and water back flushing-water back flushing' is adopted to regenerate the filter material, and the regeneration period is 10-15h.

The conditions for the weak acid cation exchange method are preferably: the pH value range is 6.5-7.5; the temperature is less than or equal to 40 ℃, the height of the resin layer is 1.5-3.0m, the HCl concentration of the regeneration liquid is as follows: 4.5-5 mass%; the dosage of the regenerant (calculated by 100%) is 50-60kg/m ³ Wet resin; the flow rate of the regeneration liquid HCl is 4.5-5.5m/h, and the regeneration contact time is 35-45min; the forward washing flow rate is 18-22m/h, and the forward washing time is 2-30min; the running flow rate is 15-30m/h; as the acidic cation exchange resin, for example, there can be used a Gallery Senno chemical Co., ltd, SNT brand D113 acidic cation exchange resin.

The conditions for the above-mentioned ozone biological activated carbon adsorption oxidation method are preferably as follows: the retention time of the ozone is 50-70min, and the empty bed filtration rate is 0.5-0.7m/h.

The conditions for the concentration of the ED membrane are preferably: the current is 145-155A, and the voltage is 45-65V. The ED membrane may be, for example, an ED membrane manufactured by astone corporation of japan.

The conditions for the reverse osmosis are preferably: the operation pressure is 5.4-5.6MPa, the water inlet temperature is 25-35 ℃, and the pH value is 6.5-7.5. The reverse osmosis membrane is, for example, a seawater desalination membrane TM810C manufactured by Dongli corporation of Lanxingdong.

According to the invention, when the wastewater treatment is started, the catalyst production wastewater can be used for directly starting operation, and if the ion content of the catalyst production wastewater meets the conditions of the invention, the first evaporation and cooling crystallization can be carried out firstly, and then the second evaporation and cooling treatment can be carried out according to the conditions of the invention; if the ion content of the catalyst production wastewater does not satisfy the conditions of the present invention, the second distillation may be performed firstAnd performing cooling treatment to obtain a treatment solution, performing solid-liquid separation to obtain sodium sulfate crystals and a second mother solution, mixing the second mother solution with the catalyst production wastewater to adjust the ion content of the wastewater to be treated to be within the range required by the invention, and performing first evaporation, cooling and crystallization to obtain the sodium sulfate crystals. Of course, the ion content of the wastewater to be treated can be adjusted by using sodium sulfate or sodium chloride in the initial stage as long as the wastewater to be treated satisfies the SO content in the wastewater to be treated in the present invention ₄ ^2- 、Cl ^- The requirements are met.

The present invention will be described in detail below by way of examples.

In the following examples, the catalyst production wastewater is wastewater from a molecular sieve production process, which is subjected to chemical precipitation, filtration, weak acid cation exchange and ozone biological activated carbon adsorption oxidation in sequence to remove impurities, and is subjected to ED membrane concentration and reverse osmosis concentration in sequence.

Example 1

As shown in FIG. 1, the catalyst production wastewater (containing NaCl 45g/L and Na) ₂ SO ₄ 99.1g/L、NH ₄ Cl 20.1g/L、(NH ₄ ) ₂ SO ₄ 45g/L, pH 6.8) at a feed rate of 10m ³ The method comprises the steps of feeding the wastewater to be treated into a pipeline of a treatment system at a speed of/h, introducing 45.16 mass percent of sodium hydroxide aqueous solution into the pipeline for carrying out first pH value adjustment, monitoring the adjusted pH value through a first pH value measuring device 61 (a pH meter) (the measured value is 7.8), feeding one part of the catalyst production wastewater into a first heat exchange device 31 (a two-phase stainless steel plate heat exchanger) through a first circulating pump 71 for carrying out heat exchange with a first ammonia-containing steam condensate, mixing the rest part of the catalyst production wastewater with a second mother liquor returned by a ninth circulating pump 79, feeding the mixture into a tenth heat exchange device 30 (the two-phase stainless steel plate heat exchanger) for carrying out heat exchange with the first condensate, combining the two parts to obtain wastewater to be treated, feeding the wastewater to be treated into a pipeline of an eighth heat exchange device 38 for introducing 45.16 mass percent of sodium hydroxide aqueous solution for carrying out second pH value adjustment, feeding the wastewater to be treated into the eighth heat exchange device 38 (the two-phase stainless steel plate heat exchanger) for carrying out heat exchange with the first ammonia-containing steam so as to heat the wastewater to be heated to 11After 2 ℃, the adjusted pH was monitored by a third pH measuring device 60 (pH meter) (measurement value 11).

The first evaporation is carried out in a first MVR evaporation device 3 (a falling film and forced circulation two-stage MVR second evaporation crystallizer), the temperature of the first evaporation is 105 ℃, the pressure is-7.01 kPa, and the evaporation capacity is 2.72m ³ And/h, obtaining first ammonia-containing steam and first concentrated solution. After the first ammonia-containing steam is compressed by the first compressor 101 (the temperature is raised by 18 ℃), the first ammonia-containing steam exchanges heat with the wastewater to be treated and the catalyst production wastewater in the eighth heat exchange device 38 and the first heat exchange device 31 in sequence to obtain first ammonia water, and the first ammonia water is stored in the first ammonia water storage tank 51. In addition, in order to increase the solid content in the first MVR evaporation device 3, part of the liquid evaporated in the first MVR evaporation device 3 is circulated as a first circulation liquid to the eighth heat exchange device 38 by the fifth circulation pump 75, and then enters the first MVR evaporation device 3 again for first evaporation (the first reflux ratio is 24.4). The degree of the first evaporation is monitored by a densimeter arranged on the first MVR evaporation device 3, and the concentration of the first concentrated solution is controlled to be NaCl 107.3g/L and Na ₂ SO ₄ 191.7g/L(Cl ^- Has a concentration of 1.833mol/L, SO ₄ ^2- The concentration of (b) is 1.35 mol/L).

The first concentrated solution is sent to a tenth heat exchange device 30 to exchange heat with the catalyst production wastewater, then sent to a second heat exchange device 32 (a heat exchanger made of plastic) to exchange heat with the first mother solution to cool the first concentrated solution to 21 ℃, then mixed with the circulating solution of the cooling crystallization device 2 conveyed by a second circulating pump 72, further cooled through heat exchange with the refrigerating solution by a sixth heat exchange device 36, and sent to the cooling crystallization device 2 (a freezing crystallization tank) to be cooled and crystallized, so as to obtain the crystallization solution containing sodium sulfate crystals. Wherein the cooling crystallization temperature is-2 deg.C, the cooling crystallization time is 120min, and the circulation amount of the cooling crystallization is controlled to be 1251m ³ And h, controlling the supersaturation degree of the sodium sulfate in the cooling crystallization process to be not more than 1.1g/L.

The obtained crystal liquid containing sodium sulfate crystals was sent to a first solid-liquid separation apparatus 91 (centrifuge) to conduct solid-liquid separation, thereby obtaining 3.15m per hour ³ Contains 272g/L NaCl and Na ₂ SO ₄ 18.6g/L、NaOH 6.6g/L、NH ₃ 0.33g/L of the first mother liquor was temporarily stored in the first mother liquor tank 53, and 6257.79kg of a sodium sulfate decahydrate crystal cake having a purity of 98 mass% and a water content of 76 mass% was obtained per hour.

The second evaporation process is carried out in a second MVR evaporation plant 1 (falling film + forced circulation two-stage MVR second evaporation crystallizer). After the first mother liquor exchanges heat with the first concentrated solution in the second heat exchange device 32, one part of the first mother liquor is sent into a third heat exchange device 33 (a duplex stainless steel plate type heat exchanger) to exchange heat with a second ammonia-containing steam condensate, the other part of the first mother liquor is sent into a fifth heat exchange device 35 (a duplex stainless steel plate type heat exchanger) to exchange heat with a second concentrated solution containing crystals, then the two parts of the first mother liquor are combined and sent into a fourth heat exchange device 34 (a duplex stainless steel plate type heat exchanger) to exchange heat with the second ammonia-containing steam, 45.16 mass percent of sodium hydroxide aqueous solution is led into a pipeline sent into the second MVR evaporation device 1 to carry out third pH value adjustment, the adjusted pH value is monitored through a second pH value measuring device 62 (a pH meter) (the measured value is 11.4), and second evaporation is carried out in the second MVR evaporation device 1 to obtain a second concentrated solution containing sodium sulfate crystals and sodium chloride crystals and second ammonia-containing steam. The temperature of the second evaporation was 105 deg.C, the pressure was-7.02 kPa, and the evaporation capacity was 2.54m ³ H is used as the reference value. After the second ammonia-containing steam is compressed by the second compressor 102 (the temperature is raised by 17 ℃), the second ammonia-containing steam exchanges heat with the first mother liquor in the fourth heat exchange device 34 and the third heat exchange device 33 in sequence to obtain second ammonia water, and the second ammonia water is stored in the second ammonia water storage tank 52. In addition, in order to increase the solid content in the second MVR evaporation device 1, part of the liquid after the second evaporation in the second MVR evaporation device 1 is sent again to the second MVR evaporation device 1 as a circulating liquid through the seventh circulating pump 77 for the second evaporation (the second reflux ratio is 51.3). The degree of the second evaporation is monitored by a mass flow meter arranged on the second MVR evaporation device 1, and the second evaporation amount is controlled to be 2.54m ³ H (corresponding to the control of the sodium sulfate concentration in the treatment solution to 0.976Y (81.8 g/L)).

And (3) cooling the second concentrated solution containing the sodium sulfate crystals and the sodium chloride crystals obtained by the second evaporation in a low-temperature treatment tank 55 at the temperature of 17.9 ℃ for 60min to obtain a treatment solution containing the sodium chloride crystals. The low-temperature treatment tank 55 is internally provided with a stirring paddle, and the rotating speed of the stirring paddle is 60r/min.

The treatment liquid containing sodium chloride crystals obtained by the temperature reduction treatment is sent to a second solid-liquid separation device 92 (centrifugal machine) for second solid-liquid separation and leaching, and then 0.717m is obtained per hour ³ Contains 262.7g/L NaCl and Na ₂ SO ₄ 81.8g/L、NaOH 29g/L、NH ₃ 0.0004g/L of the second mother liquor was temporarily stored in the second mother liquor tank 54. The obtained sodium chloride solid (756.42 kg of sodium chloride crystal cake with a water content of 15 mass% obtained per hour, wherein the sodium sulfate content was 1.5 mass% or less) was washed with 262g/L of a sodium chloride solution having a mass equivalent to the dry mass of sodium chloride, and dried in a dryer to obtain 642.96kg of sodium chloride (purity of 99.5 mass%) per hour, and the second washing liquid obtained by the washing was circulated to the third heat exchange apparatus 33 by the tenth circulation pump 80.

In this example, 2.72m of ammonia water having a concentration of 5.6 mass% was obtained per hour in the first ammonia water tank 51 ³ (ii) a 2.54m of 0.04 mass% ammonia water is obtained per hour in the second ammonia water tank 52 ³ 。

In addition, the tail gas discharged by the eighth heat exchange device 38, the cooling crystallization device 2 and the fourth heat exchange device 34 is introduced into a tail gas absorption tower 83 through a vacuum pump 81 for absorption, circulating water is introduced into the tail gas absorption tower 83, the circulating water circulates in the tail gas absorption tower 83 under the action of the fourth circulating pump 74, water is supplemented into the tail gas absorption tower 83 from a circulating water tank 82 through a third circulating pump 73, and fresh water is supplemented into the circulating water tank 82, so that the temperature and the ammonia content of the working water of the vacuum pump 81 are reduced. Dilute sulfuric acid is introduced into the tail gas absorption tower 83 to absorb ammonia and the like in the tail gas. The starting phase of MVR evaporation was initiated by steam at a temperature of 143.3 ℃.

Example 2

The treatment of the catalyst production wastewater was carried out in the same manner as in example 1 except that: for NaCl-containing 70g/L, na ₂ SO ₄ 66g/L、NH ₄ Cl 30g/L、(NH ₄ ) ₂ SO ₄ 28.7g/L of catalyst production wastewater with pH of 6.9 is treated, and the wastewater to be treated is subjected to heat exchange through the eighth heat exchange device 38The temperature of (2) was 107 ℃.

The first evaporation temperature is 100 deg.C, the pressure is-22.82 Pa, and the evaporation capacity is 2.98m ³ H; cooling and crystallizing at-4 deg.C for 120min; the temperature of the second evaporation is 100 ℃, the pressure is-22.83 Pa, and the evaporation capacity is 4.02m ³ H; the temperature of the temperature reduction treatment is 20 ℃, and the time is 58min. The obtained first concentrated solution contained NaCl 161.6g/L and Na ₂ SO ₄ 131.3g/L(Cl ^- Has a concentration of 2.763mol/L, SO ₄ ^2- The concentration of (b) is 0.9249 mol/L).

The first solid-liquid separation device 91 obtains 4066.13kg of sodium sulfate decahydrate crystal cake containing 76 mass% of water with a purity of 98.3 mass%) per hour; yield 4.82m per hour ³ The concentration of NaCl 268.5g/L and Na ₂ SO ₄ 16.9g/L、NaOH 3.6g/L、NH ₃ 0.2g/L of the first mother liquor.

The second solid-liquid separation device 92 obtained 1194.92kg of sodium chloride crystal cake with a water content of 14.5 mass% per hour, and finally 1021.66kg of sodium chloride (purity 99 mass%) per hour; the second solid-liquid separation device 92 obtains 0.977m per hour ³ The concentration is NaCl 271.2g/L and Na ₂ SO ₄ 83.3g/L、NaOH 18.0g/L、NH ₃ 0.0002g/L of second mother liquor.

2.98m of ammonia water having a concentration of 4.8 mass% is obtained per hour in the first ammonia water tank 51 ³ (ii) a 4.02m of ammonia water having a concentration of 0.02 mass% is obtained in the second ammonia water tank 52 every hour ³ 。

Example 3

The treatment of the catalyst production wastewater was carried out in the same manner as in example 1 except that: for NaCl 100g/L, na ₂ SO ₄ 40g/L、NH ₄ Cl 40g/L、(NH ₄ ) ₂ SO ₄ 16.3g/L of catalyst production wastewater with the pH value of 6.3 is treated, and the temperature of the wastewater to be treated is 117 ℃ after heat exchange is carried out by the eighth heat exchange device 38.

The temperature of the first evaporation is 110 ℃, the pressure is 11.34Pa, and the evaporation capacity is 3.39m ³ H; cooling and crystallizing at 0 deg.C for 120min; the temperature of the second evaporation is 110 ℃, the pressure is 11.33Pa,the evaporation capacity is 4.98m ³ H; the temperature of the cooling treatment is 25 ℃, and the time is 55min. The obtained first concentrated solution contains NaCl 227g/L and Na ₂ SO ₄ 85.5g/L(Cl ^- Has a concentration of 3.881mol/L, SO ₄ ^2- The concentration of (B) is 0.6018 mol/L).

The first solid-liquid separation device 91 obtained 2404.50kg of sodium sulfate decahydrate crystal cake containing 76 mass% of water per hour, with a purity of 98.4 mass%); yield 6.11m per hour ³ The concentration of NaCl 297.2g/L and Na ₂ SO ₄ 17.8g/L、NaOH 3.5g/L、NH ₃ 0.16g/L of the first mother liquor.

The second solid-liquid separation device 92 yielded 1688.61kg of sodium chloride crystal cake with a water content of 15 mass% per hour, and finally 1435.32kg of sodium chloride (purity of 99.5 mass%) per hour; the second solid-liquid separation device 92 gave a particle size of 1.39m per hour ³ The concentration of NaCl is 273.2g/L and Na ₂ SO ₄ 78.4g/L、NaOH 15.6g/L、NH ₃ 0.0001g/L of second mother liquor.

3.39m of ammonia water having a concentration of 4.3% by mass per hour was obtained in the first ammonia water tank 51 ³ (ii) a 4.98m of ammonia water having a concentration of 0.02 mass% is obtained in the second ammonia water tank 52 every hour ³ 。

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. Method for treating wastewater generated in catalyst production, wherein the wastewater generated in catalyst production contains NH ₄ ⁺ 、SO ₄ ^2- 、Cl ^- And Na ⁺ Characterized in that the method comprises the following steps,

1) Performing first evaporation on wastewater to be treated to obtain first ammonia-containing steam and first concentrated solution;

5) Carrying out second solid-liquid separation on the treatment liquid;

before the wastewater to be treated is subjected to first evaporation, adjusting the pH value of the wastewater to be treated to be more than 9;

SO in the first concentrated solution ₄ ^2- The concentration of (b) is more than 0.01 mol/L;

the conditions of the second evaporation include: the temperature is 45-110 ℃, and the pressure is-95 kPa-12 kPa; the temperature of the temperature reduction treatment is 15-45 ℃;

the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by the second solid-liquid separation; NH in the catalyst production wastewater ₄ ⁺ Is more than 8mg/L, SO ₄ ^2- Is more than 1g/L, cl ^- Over 970mg/L of Na ⁺ Is more than 510 mg/L.

2. The method of claim 1, wherein the SO contained in the first concentrate is ₄ ^2- The concentration of (B) is 0.1mol/L or more.

3. The method of claim 2, wherein the first concentrate contains SO ₄ ^2- The concentration of (B) is 0.2mol/L or more.

4. The method of claim 1, wherein the first concentrate comprises Cl ^- The concentration of (B) is 5mol/L or less.

5. The method of claim 1, whereinRelative to 1mol of SO contained in a liquid phase obtained by the first solid-liquid separation ₄ ^2- Cl contained in the liquid phase obtained by the first solid-liquid separation ^- Is 7.15mol or more.

6. The method as claimed in claim 1, wherein the pH of the wastewater to be treated is adjusted to be greater than 10.8 before the wastewater to be treated is subjected to the first evaporation.

7. The method of claim 1, wherein adjusting the pH is performed with NaOH.

8. The method of claim 1, wherein SO is in said first concentrate prior to said cooling for crystallization ₄ ^2- Has a concentration of 0.01mol/L or more and Cl ^- The concentration of (b) is 5.2mol/L or less.

9. The method of claim 8, wherein the cooling crystallization does not crystallize sodium chloride.

10. The process of claim 9, wherein the concentration of sodium chloride in the first concentrate is adjusted prior to the cooling crystallization.

11. The process according to claim 10, wherein the adjustment of the concentration of sodium chloride in the first concentrated solution is performed by mixing a liquid phase obtained by the second solid-liquid separation, the catalyst production wastewater and/or a washing liquid for washing sodium sulfate crystals.

12. The method according to claim 1, wherein the second concentrated solution containing crystals obtained in step 3) is a concentrated solution containing sodium chloride crystals and sodium sulfate crystals, and the temperature reduction treatment dissolves the sodium sulfate crystals in the concentrated solution containing sodium chloride crystals and sodium sulfate crystals.

13. The method of any one of claims 1-12, wherein the conditions of the first evaporation comprise: the temperature is above 35 ℃ and the pressure is above-98 kPa.

14. The method of claim 13, wherein the conditions of the first evaporation comprise: the temperature is 75-130 ℃, and the pressure is-73-117 kPa.

15. The method of claim 14, wherein the conditions of the first evaporation comprise: the temperature is 85-130 ℃, and the pressure is-58-117 kPa.

16. The method of claim 15, wherein the conditions of the first evaporation comprise: the temperature is 95-110 ℃, and the pressure is-37 kPa-12 kPa.

17. The method of claim 16, wherein the first evaporation is performed using an MVR evaporation device.

18. The method according to any one of claims 1 to 12, wherein the temperature of the cooling crystallization is from-21.7 ℃ to 17.5 ℃.

19. The method according to claim 18, wherein the temperature of the cooling crystallization is from-20 ℃ to 5 ℃.

20. The method of claim 19, wherein the temperature of the cooling crystallization is from-10 ℃ to 5 ℃.

21. The method of claim 20, wherein the temperature of the cooling crystallization is from-10 ℃ to 0 ℃.

22. The method according to claim 18, wherein the cooling crystallization time is 5min or more.

23. The method of claim 22, wherein the cooling crystallization time is 60min to 180min.

24. The method of claim 23, wherein the cooling crystallization time is 90min to 150min.

25. The method according to any one of claims 1 to 12, wherein the temperature of the temperature reduction treatment is 15 ℃ to 35 ℃.

26. The method according to claim 25, wherein the temperature of the temperature reduction treatment is 17.9-35 ℃.

27. The method according to claim 25, wherein the time of the temperature reduction treatment is 5min or more.

28. The method of claim 27, wherein the time of the temperature reduction treatment is 5min to 120min.

29. The method of claim 28, wherein the time of the temperature reduction treatment is 45min to 90min.

30. The process according to any one of claims 1 to 12, wherein the wastewater to be treated is subjected to a first heat exchange with the first ammonia-containing vapor and a first aqueous ammonia is obtained before the wastewater to be treated is subjected to a first evaporation.

31. The process as claimed in claim 30, wherein the first concentrated solution is subjected to a second heat exchange with a liquid phase obtained by the first solid-liquid separation before the first concentrated solution is subjected to cooling crystallization.

32. The process of claim 31, wherein the second ammonia-containing vapor is subjected to a third heat exchange with the first solid-liquid separated liquid phase to obtain ammonia water before passing the first solid-liquid separated liquid phase to a second MVR evaporation plant.

33. The method according to any one of claims 1 to 12, further comprising subjecting the sodium sulfate crystal-containing crystal liquid to a first solid-liquid separation to obtain sodium sulfate crystals.

34. The method of claim 33, further comprising washing the resulting sodium sulfate crystals.

35. The method according to any one of claims 1 to 12, further comprising subjecting the treatment liquid containing sodium chloride crystals to a second solid-liquid separation to obtain sodium chloride crystals.

36. The method of claim 35, further comprising washing the obtained sodium chloride crystals.

37. The process of any of claims 1-12, wherein the catalyst production wastewater is wastewater from a molecular sieve, alumina, or refinery catalyst production process.

38. The method of claim 37, further comprising removing impurities and concentrating the catalyst process wastewater.