CN109422394B

CN109422394B - Method for treating catalyst production wastewater

Info

Publication number: CN109422394B
Application number: CN201710750916.2A
Authority: CN
Inventors: 殷喜平; 李叶; 顾松园; 王涛; 周岩; 苑志伟; 张志民; 刘夫足; 高晋爱; 安涛
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: CN109422394A

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 comprises the following steps of 1) carrying out 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) Carrying out second evaporation on the first concentrated solution to obtain second concentrated solution containing second ammonia vapor and sodium chloride crystals; 3) Carrying out first solid-liquid separation on the second concentrated solution containing the sodium chloride crystals, and cooling and crystallizing a liquid phase obtained by the first solid-liquid separation to obtain a crystallization liquid containing the sodium sulfate crystals; 4) And carrying out second solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals. The method can respectively obtain high-purity sodium sulfate and sodium chloride, and avoids the difficulty in the processes of mixed salt treatment and recycling.

Description

Method for treating catalyst production wastewater

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 ammoniacal 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 the NH-containing catalyst with low cost and environmental protection 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, and furthest recycle resources in the wastewater.

In order to achieve the above objects, an 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) Carrying out second evaporation on the first concentrated solution to obtain second concentrated solution containing second ammonia vapor and sodium chloride crystals;

3) Carrying out first solid-liquid separation on the second concentrated solution containing the sodium chloride crystals, and cooling and crystallizing a liquid phase obtained by the first solid-liquid separation to obtain a crystallization liquid containing the sodium sulfate crystals;

4) Carrying out second solid-liquid separation on the crystallization liquid containing the sodium sulfate crystals;

wherein, before the wastewater to be treated is subjected to first evaporation, the pH value of the wastewater to be treated is adjusted to be more than 9; the second evaporation did not crystallize sodium sulfate out.

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 stronger ammonia water is obtained through first evaporation, a second concentrated solution of sodium chloride and thinner ammonia water are obtained through second evaporation, and then sodium sulfate crystals are obtained through cooling crystallization separation. The method can respectively obtain high-purity sodium sulfate and sodium chloride, avoids the difficulties 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, and reduces the waste water treatmentThe cost is low, the ammonium in the wastewater is recovered in the form of ammonia water, the sodium chloride and the sodium sulfate are respectively recovered in the form of crystals, no waste residue and waste liquid are generated in the whole process, and the aim of changing waste into valuables is fulfilled.

Furthermore, according to needs, the method greatly reduces the sodium sulfate content in the cooling mother liquor through cooling crystallization, can improve the efficiency of preparing sodium chloride through second evaporation, and simultaneously, before the liquid phase obtained by the first solid-liquid separation is cooled and crystallized, cl in the liquid phase is adjusted by using the first concentrated solution, the sodium sulfate washing solution or the catalyst production wastewater ^- The concentration of the sodium sulfate avoids the precipitation of sodium chloride in the cooling process, and improves the purity of the sodium sulfate precipitated in the cooling process.

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. Cooling crystallization device 31 and first heat exchange device

2. Second MVR evaporation plant 32, second heat transfer device

3. First MVR evaporation plant 33, third heat transfer device

36. Sixth heat exchanger 76, sixth circulating pump

37. Seventh heat exchanger 77, seventh circulating pump

38. Eighth heat exchange device 78, eighth circulating pump

51. Second ammonia storage tank 79 and ninth circulating pump

52. First ammonia water storage tank 80, tenth circulating pump

53. First mother liquid tank 70 and eleventh circulating pump

54. Second mother liquid tank 81, vacuum pump

61. First pH value measuring device 82 and circulating water tank

62. Second pH value measuring device 83 and tail gas absorption tower

60. Fourth pH value measuring device 91 and first solid-liquid separating device

71. First circulating pump 92 and second solid-liquid separation device

72. Second circulation pump 101 and second compressor

73. Third circulating pump 102, first compressor

74. Fourth circulating pump

75. Fifth circulating pump

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 numerical ranges, each range between its endpoints and individual point values, and each individual point value can be combined with each other to give one or more new numerical ranges, and such numerical ranges should be construed 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,

2) Carrying out second evaporation on the first concentrated solution to obtain second ammonia-containing steam and second concentrated solution containing sodium chloride crystals;

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. From the viewpoint of improving the treatment efficiency of wastewater, it is preferable that relative to 1mol of SO contained in the wastewater to be treated ₄ ^2- Cl contained in the wastewater to be treated ^- 9.5mol or more, preferably 10 to 20mol. By mixing SO in the wastewater to be treated ₄ ^2- 、Cl ^- The concentration ratio of the sodium chloride is controlled in the range, so that more sodium chloride can be separated out in the second evaporation process, and the aim of efficiently separating the sodium chloride is fulfilled. In addition, as described above and below, it is also possible in the present invention to circulate the cooled mother liquor obtained in the cooling crystallization process to the first evaporation and thereby treat SO in the wastewater to be treated ₄ ^2- And Cl ^- Can be adjusted and the balance of sodium hydroxide can be maintained.

In the present invention, the order of the first heat exchange, the adjustment of the pH 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) which will be described later is not particularly limited, and may be appropriately selected as needed, and may be completed before the first evaporation of the wastewater to be treated.

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 first evaporation is performed to obtain a concentrated ammonia solution simultaneously with the concentration of the wastewater to be treated, and the degree of the first evaporation is not particularly limited, and may be selected according to the requirements and the components of the wastewater to be treated, so as to satisfy the requirements for the concentration of the first ammonia solution and to adjust the concentration of the chloride ions in the first mother liquor to satisfy the conditions for cooling crystallization. For example, the evaporation can be controlled to obtain only a small amount of ammonia-containing steam, thereby obtaining ammonia water with higher concentration.

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 evaporation 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-57.66 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 100-107 ℃ and the pressure is-23 kPa-0 kPa. When the multi-effect evaporation device is used for carrying out first evaporation and concurrent or countercurrent feeding is adopted, the first evaporation condition refers to the evaporation condition of the last evaporator of the multi-effect evaporation device; when advection feeding is employed, the conditions of the first evaporation include evaporation conditions of each effect evaporator of the multi-effect evaporation apparatus.

In the present invention, the operating pressure for 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 be treated and the amount of the wastewater 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 liquid obtained by the first evaporation, and specifically, by controlling the concentration of the liquid obtained by the first evaporation within the above range, the first evaporation does not cause crystallization of sodium sulfate 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 invention, the pH of the wastewater to be treated is adjusted to a value greater than 9, preferably greater than 10.8, before the wastewater to be treated is subjected to the first evaporation. The upper limit of the adjustment of the pH of the wastewater to be treated is not limited, and may be, for example, 14 or less, preferably 13.5 or less, and more preferably 13 or less. By carrying out the first evaporation at the above pH, the evaporation of ammonia can be promoted, aqueous ammonia of higher concentration can be obtained, and high purity sodium sulfate and sodium chloride crystals can be conveniently obtained 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 the pH adjustment 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.

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 measuring device 61 and the fourth pH measuring 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 seventh heat exchange device 37 and the eighth heat exchange device 38, specifically, the ammonia-containing steam is sequentially passed through the eighth heat exchange device 38 and the seventh heat exchange device 37, and simultaneously the wastewater to be treated is heat-exchanged with the first ammonia-containing steam condensate by the seventh heat exchange device 37 and then heat-exchanged with the first ammonia-containing steam by the eighth heat exchange device 38.

According to a preferred embodiment of the present invention, a part of the wastewater to be treated exchanges heat with the first mother liquor in the fifth heat exchange device 35, a part of the wastewater to be treated exchanges heat with the first concentrated solution in the ninth heat exchange device 39 and then exchanges heat with the second ammonia-containing vapor condensate in the first heat exchange device 31, and the rest part of the wastewater to be treated exchanges heat with the first ammonia-containing vapor condensate in the seventh heat exchange device 37 and then is combined and introduced into the eighth heat exchange device 38 to exchange heat with the first ammonia-containing vapor.

Through the first heat exchange, the obtained first ammonia water is stored in the first ammonia water storage tank 52, and meanwhile, the temperature of the wastewater to be treated is raised to 75-137 ℃, preferably 95-117 ℃, so that the evaporation can be conveniently carried out.

The first heat exchange device 31, the fifth heat exchange device 35, the seventh heat exchange device 37, the ninth heat exchange device 39 and the eighth heat exchange device 38 are not particularly limited, and various heat exchangers conventionally used in the field can be used to achieve the purpose of 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 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 evaporated 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 conveying pipeline between the seventh heat exchange device 37 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 in the eighth heat exchange device 38, and finally the mixture is sent to the first MVR evaporation device 3. The first reflux ratio of the first evaporation is: 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. The reflux ratio may be set appropriately according to the evaporation amount to ensure that the first MVR evaporation device 3 can evaporate the required amount of water and ammonia at a given evaporation temperature. The first reflux ratio may be appropriately set as needed, and may be, for example, 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 102. Through to first containing ammonia steam compresses, for input energy among the MVR vaporization system, guarantees that waste water intensification-evaporation-cooling's process goes on in succession, needs input start-up steam when MVR vaporization process starts, only need pass through first compressor 102 energy supply after reaching continuous running state, no longer need input other energy. The first compressor 102 may be any one of various compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, a roots compressor, or the like. After compression by the first compressor 102, the temperature of the first ammonia-containing vapor increases by 5 ℃ to 20 ℃.

In the present invention, the second evaporation is performed to separate ammonia and sodium chloride in the wastewater by separating sodium chloride without separating sodium sulfate and by evaporating ammonia. According to the invention, by controlling the conditions of the second evaporation, sodium chloride is first precipitated as the solvent is continuously reduced, obtaining a second concentrated solution containing sodium chloride crystals. Controlling the second evaporation to precipitate only sodium chloride does not exclude sodium sulfate entrained by sodium chloride crystals or adsorbed on the surface. In the present invention, it is considered that only sodium chloride is precipitated (that is, the sodium sulfate is not crystallized by the second evaporation) when the total content of sodium sulfate, ammonium chloride and ammonium sulfate in the obtained sodium chloride crystal is 8% by mass or less (preferably 4% by mass or less, more preferably 3% by mass or less).

In the present invention, the second 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. The details of the device in the first evaporation are not repeated herein. In the present invention, the second evaporation is preferably performed by the second MVR evaporation device 2. 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 evaporation crystallizer.

In the present invention, the conditions of the second evaporation are not particularly limited, and may be appropriately selected as needed to achieve the purpose of precipitating crystals. In order to increase the efficiency of the second evaporation, the conditions of the second evaporation 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 45-60 ℃, and the pressure is-95 kPa to-87 kPa; preferably, the conditions of the second evaporation include: the temperature is 45-55 ℃, and the pressure is-95 kPa to-90 kPa.

In the present invention, the operation pressure of the second evaporation is preferably the saturated vapor pressure of the second evaporation 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 the second evaporation process can yield sodium chloride crystals of high purity, it is preferable that 1mol of SO contained in the first concentrated solution is used ₄ ^2- Cl contained in the first concentrated solution ^- Is 9.5mol or more, preferably 10mol or more, preferably 20mol or more, more preferably 30mol or more, further preferably 40mol or more, and may be, for example, 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, or the like. By reacting SO ₄ ^2- And Cl ^- The molar ratio of sodium sulfate to sodium chloride is controlled within the above range, and pure sodium chloride crystals can be obtained through the second evaporation, so that the separation of sodium sulfate and sodium chloride is realized.

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, a second concentrated solution containing only sodium chloride crystals cannot be obtained, and in this case, the crystals may be dissolved by adding water to the second concentrated solution, but the efficiency of wastewater treatment is impaired. Therefore, the second evaporation is preferably performed to such an extent that supersaturation of sodium sulfate in the second concentrated solution is not attained, that is, the second evaporation is performed so that the concentration of sodium sulfate in the second concentrated solution becomes Y or less (Y is a concentration of sodium sulfate at which both sodium sulfate and sodium chloride in the second concentrated solution become saturated under the conditions of the second evaporation). In view of precipitating sodium sulfate as much as possible while precipitating sodium chloride, the concentration of sodium sulfate in the second concentrated solution is preferably set to 0.9Y to 0.99Y, more preferably 0.95Y to 0.98Y by the second evaporation. By controlling the degree of the second evaporation within the above range, it is possible to ensure that sodium chloride is precipitated as much as possible during the second evaporation, and sodium sulfate is not precipitated, and pure sodium chloride crystals are finally 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 degree of progress of the second evaporation is performed by monitoring the concentration of the liquid obtained by the second evaporation, and specifically, by controlling the concentration of the liquid obtained by the second evaporation within the above range, the second evaporation does not cause crystallization of sodium sulfate. The concentration of the liquid resulting from the second evaporation is monitored by measuring the density, which may be carried out using a densitometer.

In the present invention, in order to increase the solid content in the second MVR evaporation device 2 and reduce the ammonia content in the liquid, it is preferable to return part of the liquid evaporated by the second MVR evaporation device 2 (i.e. the liquid located inside the second MVR evaporation device, hereinafter also referred to as the second circulation liquid) to the second MVR evaporation device 2 for evaporation, and preferably to the second MVR evaporation device 2 after heating for evaporation. The above-described process of returning the second circulation liquid to the second MVR evaporating device 2 may be returned to the second heat exchanging process by, for example, the seventh circulation 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 2 minus the amount of reflux. The reflux ratio may be set appropriately according to the evaporation amount to ensure that the second MVR evaporation device 2 can evaporate the required amount of water and ammonia at a given evaporation temperature. The second reflux ratio may be appropriately set as needed, and 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 first heat exchange. The compression of the second ammonia-containing vapor may be performed by a second compressor 101. Through compressing the second ammonia-containing steam, for input energy in the MVR evaporation system, guarantee that the process of waste water intensification-evaporation-cooling goes on in succession, need input start-up steam when MVR evaporation process starts, only through second compressor 101 energy supply after reaching continuous running state, no longer need input other energy. The second compressor 101 may employ various compressors conventionally used in the art, such as a centrifugal fan, a turbine compressor, or a roots compressor. After compression by the second compressor 101, 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 of the second ammonia-containing vapor obtained by the second evaporation, it is preferable to subject the first concentrated solution to the second heat exchange with the second ammonia-containing vapor before the first concentrated solution is sent to the second MVR evaporation device 2.

According to a preferred embodiment of the present invention, the second heat exchange of the first concentrate with the second ammonia-containing vapor is performed by second heat exchange means 32. Specifically, make the two heat transfer through second heat transfer device 32 respectively with the second including ammonia steam and first concentrate to make first concentrate intensification be convenient for the second evaporation, make the second including the ammonia steam cooling obtain the second aqua ammonia simultaneously. After the second heat exchange by the second heat exchange device 32, the temperature of the first concentrated solution is raised to 40 to 117 ℃, preferably 45 to 62 ℃.

The second heat exchange device 32 is not particularly limited, and various heat exchangers conventionally used in the art can be used to exchange heat between the second ammonia-containing steam and the wastewater to be treated. 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. Preferably, a duplex stainless steel plate heat exchanger is used.

According to the invention, the first concentrate is preferably adjusted to a pH greater than 9, preferably greater than 10.8, more preferably between 10.8 and 11.5, before it is passed to the second MVR evaporator 2. By adjusting the pH of the first concentrated solution to the above range, it can be ensured that ammonia is sufficiently evaporated in the second evaporation process, thereby improving the purity of the obtained sodium chloride. For the pH adjustment of the first concentrated solution, the method is the same as the pH adjustment of the wastewater to be treated, and is not described herein again.

In order to detect the pH value after the pH adjustment, it is preferable to provide a second pH measuring device 62 on the pipe for feeding the first concentrated solution to the second MVR evaporating device 2 to measure the pH value after the pH adjustment for the second time.

According to the present invention, the first solid-liquid separation may be performed by a first solid-liquid separation device (e.g., a centrifuge, a belt filter, a plate filter, etc.) 91. After the first solid-liquid separation, the first mother liquor obtained by the first solid-liquid separation device 91 (i.e., the liquid phase obtained by the first solid-liquid separation) is temporarily stored in the first mother liquor tank 53, and may be sent to the cooling crystallization device 1 for cooling crystallization, and specifically, the first mother liquor may be sent to the cooling crystallization device 1 by the sixth circulation pump 76. 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 first 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, preferably the sodium chloride crystals are washed with an aqueous solution of sodium chloride. More preferably, the concentration of the sodium chloride aqueous solution is preferably the concentration of sodium chloride in the aqueous solution at which sodium chloride and sodium sulfate reach saturation simultaneously at the temperature corresponding to the sodium chloride crystals to be washed.

The manner of the first solid-liquid separation and the first washing is not particularly limited, and may be carried out by using, for example, a combination of an elutriation apparatus and a solid-liquid separation apparatus which are conventional in the art, or may be carried out on a staged solid-liquid separation apparatus 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 first washing mode comprises elutriation and/or rinsing. 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 first washing can be used in a countercurrent mode for recycling 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 rinsed. 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. For the liquid resulting from the washing, it is preferably returned to the second MVR evaporation device 2.

According to a preferred embodiment of the present invention, after the initial solid-liquid separation by settling, the second concentrate containing sodium chloride crystals is subjected to elutriation in another elutriation tank using a liquid obtained in the subsequent washing of the sodium chloride crystals, the elutriated slurry is sent to a solid-liquid separation apparatus to be subjected to 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.

In the present invention, the purpose of the cooling crystallization is to precipitate sodium sulfate, but sodium chloride, ammonium sulfate, and the like are not precipitated, so that sodium sulfate can be separated from wastewater favorably. The cooling crystallization merely precipitates sodium sulfate, and sodium chloride and the like carried by the sodium sulfate crystals or adsorbed on the surface are not excluded. In the present invention, the content of sodium sulfate in the obtained sodium sulfate crystal is preferably 92% by mass or more, more preferably 96% by mass or more, and still more preferably 98% by mass or more. 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.

Before the first mother liquor is cooled and crystallized, the concentration of sodium chloride in the first mother liquor is preferably adjusted according to needs so that the concentration of sodium chloride in the crystal liquor obtained by cooling and crystallization is less than X, wherein X is the concentration of sodium chloride when sodium sulfate and sodium chloride in the crystal liquor are saturated under the condition of cooling and crystallization. By adjusting the concentration of sodium chloride in the liquid phase obtained by the first solid-liquid separation 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 0.95X-0.999X (wherein X is the concentration of sodium chloride when sodium sulfate and sodium chloride in the crystallization liquid reach saturation under the condition of cooling crystallization). The method of adjusting the sodium chloride concentration may be performed by mixing the first concentrated solution, 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 mother liquor is adjusted to be within the above range. To avoid introducing more liquid, it is preferred that the adjustment of the concentration of sodium chloride in the first mother liquor is performed by mixing the first concentrated solution, the catalyst production wastewater and/or the washing solution after washing the sodium sulfate crystals. By adjusting the concentration of sodium chloride in the first mother liquor to the range, the precipitation of sodium chloride in the cooling crystallization process is avoided, the purity of precipitated sodium sulfate in the cooling crystallization process is improved, and the cooling crystallization efficiency is improved.

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 ℃ to 17.5 ℃, preferably-20 ℃ to 10 ℃, more preferably-10 ℃ to 5 ℃, further preferably-5 ℃ to 0 ℃, and particularly preferably-4 ℃ to 0 ℃; the time (in terms of the residence time in the cooling crystallization apparatus 1) is 5min or more, preferably 60min to 180min, more preferably 90min to 150min, and still more preferably 130min to 140min. By controlling the cooling crystallization conditions within the above range, sodium sulfate can be sufficiently precipitated without precipitating sodium chloride, and the energy consumption in the cooling crystallization process can be reduced.

Specific examples of the temperature for cooling and crystallizing 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 invention, in order to ensure that the sodium sulfate crystals are obtained by cooling crystallization, SO is contained in the first mother liquor ₄ ^2- The concentration of (B) is preferably 0.01mol/L or more, more preferably 0.07mol/L or more, still more preferably 0.1mol/L or more, yet more preferably 0.2mol/L or more, and particularly preferably 0.3mol/L or more. According to the invention, in order to increase the purity of the sodium sulfate crystals obtained by cooling crystallization, the Cl in the first mother liquor ^- The concentration of (B) is preferably 5.2mol/L or less, more preferably 5mol/L or less, further preferably 4.5mol/L or less, and further preferably 4mol/L or less.

In the present invention, if SO is present in said first mother liquor ₄ ^2- 、Cl ^- The concentration of (b) is out of the above range, and concentration adjustment may be performed before the cooling crystallization is performed, and the concentration adjustment is preferably performed using the catalyst production wastewater, and specifically, the catalyst production wastewater may be mixed with the first mother liquor in the first mother liquor tank 53.

SO in the first mother liquor ₄ ^2- Specific examples of the content include: 0.01mol/L, 0.03mol/L, 0.05mol/L, 0.08mol/L, 0.1mol/L, 0.2mol/L, 0.3mol/L, 0.4mol/L, 0.5mol/L, 0.6mol/L, or 0.66mol/L, etc.

Additionally, cl is present in the first mother liquor ^- Specific examples of the content include: 0.01mol/L, 0.05mol/L, 0.1mol/L, 0.3mol/L, 0.6mol/L, 0.8mol/L, 1mol/L, 1.2mol/L, 1.4mol/L, 1.6mol/L, 1.8mol/L, 2.0mol/L, 2.2mol/L, 2.4mol/L, 2.6mol/L, 2.8mol/L, 3mol/L, 3.2mol/L, 3.4mol/L, 3.6mol/L, 3.8mol/L, 4mol/L, 4.5mol/L, or 5mol/L, etc.

According to the present invention, the cooling crystallization is carried out in a continuous or batch manner, and the temperature of the first mother liquor is lowered to precipitate sodium sulfate crystals, and the continuous cooling crystallization is preferably carried out. The cooling crystallization can be carried out by various cooling crystallization apparatuses 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 apparatus 1. The cooling part can lead the first mother liquor 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 blending part, such as a stirrer and the like, and the first mother liquor is blended to achieve the effect of uniform cooling, so that sodium sulfate in the first mother liquor 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 the 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 centrifugal pump with low rotating speed, and more preferably a guide pump impeller with large flow and low rotating speed or an axial pump with large flow, low lift and low rotating speed.

By crystallizing said cooled crystals at the above-mentioned temperature and Cl ^- The concentration is carried out, so that sodium sulfate can be fully precipitated in the cooling crystallization, and sodium chloride is not precipitated, thereby achieving the aim of separating and purifying the sodium sulfate.

In the present invention, in order to control the crystal size distribution in the cooling crystallization device 1 and reduce the content of fine crystal grains, it is preferable that a part of the liquid crystallized by the cooling crystallization device 1 (i.e., the liquid located inside the cooling crystallization device 1, hereinafter also referred to as cooling circulation liquid) is mixed with the catalyst production wastewater and returned to the cooling crystallization device 1 to be cooled and crystallized again. The above-mentioned process of returning the cooling circulation liquid to the cooling crystallization device 1 for crystallization may be, for example, by returning the cooling circulation liquid to the sixth heat exchanging device 36 by the second circulating pump 72, mixing with the first mother liquid, and then entering the cooling crystallization device 1 again for cooling crystallization. The return amount of the cooling circulation liquid can be defined by a cooling circulation ratio which is: the ratio of the circulating amount to the total amount of the liquid fed to the cooling crystallization apparatus 1 minus the circulating amount. The circulation ratio may be appropriately set according to the supersaturation degree of sodium sulfate in the cooling crystallization apparatus 1 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 second mother liquor (i.e. the liquid phase obtained by the second solid-liquid separation) are obtained after the second solid-liquid separation is carried out on the crystallization liquid containing the sodium sulfate crystal. The method of the second 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 second solid-liquid separation may be performed by using a second solid-liquid separation device 92 (for example, a centrifuge, a belt filter, a plate filter, or the like). After the second solid-liquid separation, the second mother liquor obtained by the second solid-liquid separation device 92 is temporarily stored in the second mother liquor tank 54, and may be returned to the first MVR evaporation device 2 by the ninth circulation pump 79 to perform the first evaporation. In addition, it is difficult to avoid that impurities such as chlorine ions and hydroxide ions are adsorbed on the obtained sodium sulfate crystals, and in order to remove the adsorbed impurities, reduce corrosiveness, and improve the purity of the crystals, the sodium sulfate crystals are preferably subjected to secondary washing with water or a sodium sulfate solution, and may be dried when it is necessary to obtain anhydrous sodium sulfate.

The manner of the above-mentioned second solid-liquid separation and second 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 on a staged solid-liquid separation device such as a belt filter. The washing is not particularly limited and may be carried out by a method conventional in the art. The second washing mode comprises elutriation and/or rinsing. The second washing method is preferably rinsing, and more preferably rinsing is performed after solid-liquid separation. 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 of higher purity. The second washing is preferably carried out using an aqueous sodium sulphate solution, the concentration of which is preferably such that the sodium chloride and the sodium sulphate reach simultaneously the concentration of sodium sulphate in a saturated aqueous solution at the temperature corresponding to the sodium sulphate 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 1, for example, may be returned to the cooling crystallization device 1 by the tenth circulation pump 80.

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 1 by the tenth circulation pump 80. By the above washing process, the purity of the obtained sodium sulfate crystal can be improved.

According to the present invention, in order to make full use of the refrigeration capacity of the second mother liquor, it is preferable that the first mother liquor and the second mother liquor are subjected to third heat exchange before the first mother liquor is cooled and crystallized.

According to a preferred embodiment of the present invention, the third heat exchange is performed by a third heat exchange device 33, and specifically, the first mother liquor and the second mother liquor are respectively passed through the third heat exchange device 33 and heat exchanged, so that the temperature of the first mother liquor is lowered to facilitate the cooling crystallization, and the temperature of the second mother liquor is raised to facilitate the second evaporation. After heat exchange by the third heat exchange device 33, the temperature of the first mother liquor is-19.7-15.5 ℃, preferably-19-9 ℃, more preferably-4-6 ℃, and is close to the temperature of cooling crystallization.

According to the present invention, in order to facilitate the cooling crystallization, it is preferable to further subject the first mother liquor to a third heat exchange with the refrigerating fluid. According to a preferred embodiment of the present invention, the third heat exchange between the first mother liquid and the refrigerating liquid is performed by the sixth heat exchange device 36, and specifically, the refrigerating liquid and the mixed liquid of the first mother liquid and the cooling circulation liquid are respectively passed through the sixth heat exchange device 36, and heat exchange is performed between the refrigerating liquid and the mixed liquid, so that the temperature of the mixed liquid of the first mother liquid and the cooling circulation liquid is lowered to facilitate the cooling crystallization. The refrigerating fluid can adopt a refrigerating medium which is conventionally used for reducing the temperature in the field, as long as the temperature of the first mother liquor can meet the requirement of cooling crystallization.

The third heat exchanger 33 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, 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. The third heat exchange device 33 and the sixth heat exchange device 36 are preferably heat exchangers made of plastic.

According to a preferred embodiment of the invention, the tail gas generated by cooling crystallization is discharged after ammonia removal; discharging the tail gas after ammonia removal of the tail gas left by the second heat exchange condensation; 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 1, and the second heat exchange condenses the residual tail gas which is the non-condensable gas discharged by the second heat exchange device 32; the first heat exchange condenses the remaining tail gas, i.e., the non-condensable 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 preferably 50g/L or more, further 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 30g/L or less.

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⁺ And the respective contents of inorganic salt ions such as rare earth element ionsPreferably 100mg/L or less, more preferably 50mg/L or less, still more preferably 10mg/L or less, and particularly preferably does not contain other inorganic salt ions. 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 the content of other inorganic salt ions in the catalyst production wastewater, the following impurity removal is preferably performed.

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, and may be, for example, 5 to 7.5 or 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, and potassium permanganate, can be used, 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 sequentially carrying out filtration, a weak acid cation exchange method and an ozone biological activated carbon adsorption oxidation method. 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 catalyst production wastewater of the present invention before the treatment 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 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 carrying out. 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 of the above-mentioned ozone biological activated carbon adsorption oxidation method are preferably: 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 bookAccording 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 evaporation and then cooling crystallization can be carried out according to the conditions of the invention; if the ion content of the catalyst production wastewater does not meet the conditions of the invention, cooling crystallization can be carried out firstly to obtain a crystallization liquid, solid-liquid separation is carried out to obtain sodium sulfate crystals and a second mother liquid, then the second mother liquid is mixed with the catalyst production wastewater to adjust the ion content of the wastewater to be treated to be in the range required by the invention, and evaporation is carried out to obtain sodium chloride 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 92g/L and Na) ₂ SO ₄ 39g/L、NH ₄ Cl 23g/L、(NH ₄ ) ₂ SO ₄ 9.9g/L, pH 6.7) at 5m ³ A speed of/h was fed into a pipeline of the treatment system, and a sodium hydroxide aqueous solution having a concentration of 45.16 mass% was introduced into the pipeline to perform a first pH adjustment, and the adjusted pH was monitored by a first pH measuring device 61 (pH meter) (measurement value was 8.6); a part of the catalyst production wastewater is sent to the seventh heat exchange device 37 through the first circulating pump 71 to exchange heat with the first ammonia-containing steam condensate, the other part of the catalyst production wastewater is sent to the first heat exchange device 31 to exchange heat with the second ammonia-containing steam condensate after being sent to the ninth heat exchange device 39 to exchange heat with a part of the first concentrated solution, and the rest part of the catalyst production wastewater and the ninth circulating pump 79 are returnedAnd mixing the returned second mother liquor, sending the mixed second mother liquor into a fifth heat exchange device 35 for heat exchange with the first mother liquor, combining the wastewater to obtain wastewater to be treated, introducing a sodium hydroxide aqueous solution with the concentration of 45.16 mass percent into a pipeline for sending the wastewater to be treated into an eighth heat exchange device 38 for second pH value adjustment, sending the wastewater to be treated into the eighth heat exchange device 38 for heat exchange with second ammonia-containing steam, heating to 107 ℃, monitoring the pH value (the measured value is 11) by using a fourth pH value measuring device 60 (a pH meter), and finally sending the wastewater to be treated into a first MVR evaporation device 3 (a falling film and forced circulation two-stage MVR evaporation crystallizer) for first evaporation to obtain first ammonia-containing steam and first concentrated solution. The first evaporation temperature is 100 deg.C, the pressure is-22.82 kPa, and the evaporation capacity is 0.90m ³ H is used as the reference value. After the first ammonia-containing steam is compressed by the first compressor 102 (the temperature is raised by 14 ℃), the first ammonia-containing steam exchanges heat with the wastewater to be treated and the catalyst production wastewater respectively in the eighth heat exchange device 38 and the seventh heat exchange device 37 in sequence to obtain first ammonia water, and the first ammonia water is stored in the first ammonia water storage tank 52. In addition, in order to increase the solids content in the first MVR evaporation device 3, part of the liquid evaporated in the first MVR evaporation device 3 was fed again as a first circulation liquid to the first MVR evaporation device 3 for evaporation by means of the fifth circulation pump 75 (first reflux ratio of 5.3). The degree of the first evaporation is monitored by a densimeter arranged on the first MVR evaporation device 3, and crystals are controlled not to be precipitated in the first concentrated solution. Measuring Cl in the first concentrated solution ^- Has a concentration of 4.24mol/L, SO ₄ ^2- Has a concentration of 0.198mol/L (Cl) ^- /SO ₄ ^2- Is 21.40).

A part of the first concentrated solution (6.12 m) was circulated by an eleventh circulating pump 70 ³ H) to a first mother liquor tank 53 to be mixed with the first mother liquor and the other part (7.48 m) ³ H) mixing with the second circulating liquid, introducing a sodium hydroxide aqueous solution with the concentration of 45.16 mass percent into the pipeline for third pH value adjustment, exchanging heat with second ammonia-containing steam through a second heat exchange device 32, monitoring the adjusted pH value (the measured value is 11.3) through a second pH value measuring device 62 (a pH meter), and finally sending the mixture into a second MVR evaporation device 2 (a falling film and forced circulation two-stage MVR evaporation crystallizer) for second evaporation to obtain a second circulating liquidA second concentrated solution containing ammonia vapor and sodium chloride crystals, the temperature of the second evaporation is 100 ℃, the pressure is-22.83 kPa, and the evaporation capacity is 3.44m ³ H is used as the reference value. After the second ammonia-containing steam is compressed by the second compressor 101 (the temperature is raised by 14 ℃), the second ammonia-containing steam exchanges heat with the first concentrated solution in the second heat exchange device 32 and the first heat exchange device 31 in sequence to obtain second ammonia water, and the second ammonia water is stored in the second ammonia water storage tank 51. In addition, in order to increase the solids content in the second MVR evaporation device 2, part of the liquid evaporated in the second MVR evaporation device 2 was fed again as a second circulation liquid to the second MVR evaporation device 2 for evaporation by means of the seventh circulation pump 77 (second reflux ratio of 45.3). The degree of the second evaporation is monitored by a mass flow meter arranged on the second MVR evaporation device 2, and the evaporation amount is controlled to be 3.44m ³ H (corresponding to the control of the sodium sulfate concentration in the second concentrated solution to 0.962Y (50.8 g/L)).

Sending the second concentrated solution containing sodium chloride crystals into a first solid-liquid separation device 91 (centrifugal machine) for solid-liquid separation and leaching to obtain 4.15m per hour ³ Contains NaCl307.1g/L and Na ₂ SO ₄ 50.8g/L、NaOH4.6g/L、NH ₃ 0.01g/L of the first mother liquor was temporarily stored in the first mother liquor tank 53. The obtained sodium chloride solid (680.45 kg of sodium chloride crystal cake with a water content of 14 mass% per hour, wherein the sodium sulfate content is 3 mass% or less) was washed with 307g/L of a sodium chloride solution having a dry mass equal to that of sodium chloride, and dried in a dryer to obtain 585.18kg of sodium chloride (purity: 99.6 mass%) per hour, and the second washing liquid obtained by washing was circulated to the second heat exchanger 32 by the eighth circulation pump 78.

6.12m as above ³ The first concentrated solution is subjected to heat exchange by the ninth heat exchange device 39 and then mixed with the first mother solution in the first mother solution tank 53 to obtain a crystallization mixed solution (the concentration of NaCl is measured to be 272.1g/L, na is measured to be contained in the crystallization mixed solution) ₂ SO ₄ Is 37.3 g/L), the crystallization mixed solution is respectively subjected to heat exchange with the catalyst production wastewater and the second mother solution through a fifth heat exchange device 35 and a third heat exchange device 33 in sequence by a sixth circulating pump 76, is cooled to-1 ℃, is mixed with the cooling circulating liquid, is subjected to further heat exchange with the refrigerating liquid through a sixth heat exchange device 36, and is finally sent toCooling and crystallizing in a cooling and crystallizing device 1 (continuous freezing and crystallizing tank) to obtain a crystallization liquid containing sodium sulfate crystals. Wherein the cooling crystallization temperature is-4 deg.C, the time is 120min, and the circulation amount of cooling crystallization is controlled to be 227.2m ³ And h, controlling the supersaturation degree of sodium sulfate in the cooling crystallization process to be not more than 1g/L.

The crystal liquid containing sodium sulfate crystals obtained from the cooling crystallization device 1 is sent to a second solid-liquid separation device 92 (centrifugal machine) for solid-liquid separation and leaching, and 9.49m per hour is obtained ³ Contains 294g/L NaCl and Na ₂ SO ₄ 14.2g/L、NaOH 3.7g/L、NH ₃ 0.2g/L of the second mother liquor was temporarily stored in the second mother liquor tank 54, and 995.57kg of a sodium sulfate decahydrate crystal cake having a purity of 98.6 mass% and a water content of 75 mass% was obtained per hour.

In this example, 0.90m of ammonia water having a concentration of 4.1 mass% was obtained per hour in the first ammonia water tank 52 ³ (ii) a The second ammonia water tank 51 was filled with 3.44m of 0.02 mass% ammonia water per hour ³ 。

In addition, the tail gas discharged from the cooling crystallization device 1, the second heat exchange device 32 and the eighth heat exchange device 38 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 a 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 further 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 40g/L, na ₂ SO ₄ 50g/L、NH ₄ Cl 24g/L、(NH ₄ ) ₂ SO ₄ 30.5g/L of catalyst production wastewater with the pH of 6.6 is treated, and the feeding amount is 5m ³ H is used as the reference value. Cl in the first concentrated solution ^- Has a concentration of 4.16mol/L, SO ₄ ^2- Has a concentration of 0.268mol/L, cl ^- /SO ₄ ^2- Is 15.52; a portion of the first concentrate (7.58 m) ³ H) after heat exchange by the ninth heat exchange device 39, mixing the heat-exchanged liquid with the first mother liquor in the first mother liquor tank 53 to obtain a crystallization mixed liquid, wherein the concentration of NaCl is 268.7g/L, and Na is contained ₂ SO ₄ The concentration of (2) was 46g/L.

The first evaporation temperature was 107 deg.C, the pressure was 0kPa, and the evaporation capacity was 1.02m ³ H; the second evaporation temperature is 75 deg.C, pressure is-72.74 kPa, and evaporation capacity is 2.76m ³ H; the cooling crystallization temperature is-2 deg.C, and the time is 120min.

The first solid-liquid separation device 91 obtained 384.19kg of a sodium chloride crystal cake having a water content of 15 mass% per hour, and finally obtained 326.56kg of sodium chloride (purity of 99.5 mass%) per hour; 5.51 m/hr ³ The concentration of NaCl is 303.2g/L and Na ₂ SO ₄ 56.8g/L、NaOH 3.8g/L、NH ₃ 0.02g/L of the first mother liquor.

The second solid-liquid separation device 92 yielded 1641.89kg (purity: 98.5% by mass) of a sodium sulfate decahydrate crystal cake containing 74.5% by mass of water per hour; yield 11.82m per hour ³ The concentration of NaCl is 297.6g/L and Na ₂ SO ₄ 15.9g/L、NaOH 3.4g/L、NH ₃ 0.2g/L of second mother liquor.

1.02m of ammonia water having a concentration of 4.8% by mass was obtained per hour in the first ammonia water tank 52 ³ (ii) a 2.76m of ammonia water having a concentration of 1.9% by mass per hour was obtained in the second ammonia water tank 51 ³ (ii) a The ammonia can be reused in the production process of the molecular sieve.

Example 3

The treatment of the catalyst production wastewater was carried out in the same manner as in example 1, except that: for NaCl 71g/L and Na ₂ SO ₄ 51g/L、NH ₄ Cl 26g/L、(NH ₄ ) ₂ SO ₄ 19g/L of catalyst production wastewater with pH of 6.9 is treated, and the feeding amount is 5m ³ H; cl in the first concentrated solution ^- Has a concentration of 4.24mol/L, SO ₄ ^2- In a concentration of 0.282mol/L, cl ^- /SO ₄ ^2- In a molar ratio ofIs 15.04; a portion of the first concentrate (5.33 m) ³ H) after heat exchange by the ninth heat exchange device 39, mixing the heat-exchanged liquid with the first mother liquor in the first mother liquor tank 53 to obtain a crystallization mixed liquid, wherein the concentration of NaCl is 268.6g/L, and Na is contained ₂ SO ₄ The concentration of (B) was 51.4g/L.

The first evaporation temperature was 105 deg.C, the pressure was-7.01 kPa, and the evaporation capacity was 1.15m ³ H; the second evaporation temperature is 50 deg.C, pressure is-92.67 kPa, and evaporation capacity is 2.88m ³ H; the cooling crystallization temperature is 0 deg.C, and the time is 120min.

The first solid-liquid separation device 91 obtained 576.37kg of a sodium chloride crystal cake having a water content of 14 mass% per hour, and finally 495.68kg of sodium chloride (purity of 99.5 mass%) per hour; yield 4.58m per hour ³ The concentration of NaCl is 292.2g/L and Na ₂ SO ₄ 64.6g/L、NaOH 4.1g/L、NH ₃ 0.03g/L of the first mother liquor.

The second solid-liquid separation device 92 yielded 1379.22kg (purity: 98.9 mass%) of a sodium sulfate decahydrate crystal cake containing 74 mass% of water per hour; 8.85m per hour ³ The concentration of NaCl is 300.9g/L and Na ₂ SO ₄ 17.2g/L、NH ₃ 0.2g/L of the second mother liquor.

1.15m of ammonia water having a concentration of 4.0 mass% is obtained per hour in the first ammonia water tank 52 ³ (ii) a 2.88m of ammonia water having a concentration of 1.7% by mass per hour was obtained in the second ammonia water tank 51 ³ (ii) a The ammonia can be reused in the production process of the molecular sieve.

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 catalyst production wastewater containing 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;

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

the second evaporation does not crystallize sodium sulfate out;

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;

the wastewater to be treated contains the catalyst production wastewater and a liquid phase obtained by the second solid-liquid separation.

2. The method according to claim 1, wherein the SO contained in the first concentrated solution is 1mol with respect to the SO contained in the first concentrated solution ₄ ^2- Cl contained in the first concentrated solution ^- 9.5mol or more.

3. The method according to claim 1, wherein the pH of the wastewater to be treated is adjusted to 10.8 or more before the wastewater to be treated is subjected to the first evaporation.

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

5. The process according to claim 1, wherein the second evaporation is carried out so that the concentration of sodium sulfate in the second concentrated solution is Y or less, wherein Y is the concentration of sodium sulfate at which both sodium sulfate and sodium chloride in the second concentrated solution are saturated under the conditions of the second evaporation.

6. The process of claim 5, wherein the second evaporation provides a sodium sulfate concentration in the second concentrated solution of 0.95Y to 0.99Y.

7. The process according to any one of claims 1 to 6, wherein the concentration of sodium chloride in the liquid phase obtained by the first solid-liquid separation is adjusted SO that SO is contained in the liquid phase obtained by the first solid-liquid separation before the cooling crystallization ₄ ^2- Has a concentration of 0.01mol/L or more and Cl ^- The concentration of (B) is 5.2mol/L or less.

8. The method of claim 7, wherein the cooling crystallization does not crystallize sodium chloride.

9. The process according to claim 8, wherein the adjustment of the concentration of sodium chloride in the liquid phase obtained by the first solid-liquid separation is carried out by mixing the first concentrated solution, the catalyst production wastewater and/or a washing solution for washing sodium sulfate crystals.

10. The method of any one of claims 1-6, wherein the conditions of the first evaporation comprise: the temperature is above 35 ℃ and the pressure is above-98 kPa.

11. The method of claim 10, wherein the conditions of the first evaporation comprise: the temperature is 75-130 ℃, and the pressure is-73-117 kPa.

12. The method of claim 11, wherein the conditions of the first evaporation comprise: the temperature is 85-130 ℃, and the pressure is-57.66 kPa-117 kPa.

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

14. The method of any one of claims 1-6, wherein the conditions of the second evaporation comprise: the temperature is above 17.5 ℃ and the pressure is above-101 kPa.

15. The method of claim 14, wherein the conditions of the second evaporation comprise: the temperature is 35-110 ℃, and the pressure is-98 kPa-12 kPa.

16. The method of claim 15, wherein the conditions of the second evaporation comprise: the temperature is 45-110 ℃, and the pressure is-95 kPa-12 kPa.

17. The method of claim 16, wherein the conditions of the second evaporation comprise: the temperature is 45-60 ℃, and the pressure is-95 kPa to-87 kPa.

18. The method of claim 17, wherein the conditions of the second evaporation comprise: the temperature is 45-55 ℃, and the pressure is-95 kPa to-90 kPa.

19. The method according to any one of claims 1 to 6, wherein the temperature of the cooling crystallization is from-21.7 ℃ to 17.5 ℃.

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

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

22. The method of claim 21, wherein the temperature of the cooling crystallization is from-5 ℃ to 5 ℃.

23. The method of claim 22, wherein the temperature of the cooling crystallization is from-5 ℃ to 0 ℃.

24. The method according to claim 19, wherein the cooling crystallization time is 5min or more.

25. The method according to claim 24, wherein the cooling crystallization time is 60min to 180min.

26. The method of claim 25, wherein the cooling crystallization time is 90min to 150min.

27. The method according to any one of claims 1 to 6, wherein the wastewater to be treated is subjected to a first heat exchange with a first ammonia-containing steam and a first aqueous ammonia is obtained before the wastewater to be treated is subjected to a first evaporation.

28. A process according to claim 27, wherein the first concentrate is subjected to a second heat exchange with the second ammonia-containing vapor and a second aqua ammonia is obtained before the first concentrate is subjected to a second evaporation.

29. The process as claimed in claim 28, wherein the second concentrated solution is subjected to a third heat exchange with a liquid phase obtained by a second solid-liquid separation before the second concentrated solution is subjected to cooling crystallization.

30. The method according to any one of claims 1 to 6, further comprising subjecting the second concentrated solution containing sodium chloride crystals to a first solid-liquid separation to obtain sodium chloride crystals.

31. The method of claim 30, further comprising washing the obtained sodium chloride crystals.

32. The method according to any one of claims 1 to 6, further comprising subjecting the sodium sulfate crystal-containing crystal liquid to a second solid-liquid separation to obtain sodium sulfate crystals.

33. The method of claim 32, further comprising washing the resulting sodium sulfate crystals.

34. The process of any one of claims 1 to 6, wherein the catalyst production wastewater is wastewater from a molecular sieve, alumina or refinery catalyst production process.

35. The method of claim 34, further comprising removing impurities and concentrating the catalyst process wastewater.