A method and separation system for removing hydrogen sulfide and ammonia from dry gas

Abstract

The present invention discloses a method and a separation system for removing hydrogen sulfide and ammonia from dry gas. The method comprises: a desulfurization absorption rich liquid, a deammoniation absorption rich liquid and an oxygen-containing gas enter a catalytic oxidation reactor, contact with a catalytic oxidation catalyst for reaction, and obtain a thiosulfate emulsion; the thiosulfate emulsion is subjected to gas-liquid separation, and the obtained liquid phase is decanted and layered to obtain a desulfurization absorption lean liquid and a crude deammoniation absorption lean liquid; the crude deammoniation absorption lean liquid is subjected to thiosulfate removal to obtain a deammoniation absorption lean liquid and a waste salt solution. The present invention solves the problem of low process matching when the acidic impurity hydrogen sulfide and the alkaline impurity ammonia are often removed separately when hydrogen sulfide and ammonia exist in the existing raw material dry gas at the same time. Through the self-absorption reaction between the impurities, the use of subsequent desulfurization agents or deammoniation agents is reduced, which is beneficial to reducing the production cost of ethylbenzene devices.

Description

Technical Field

[0001] The invention relates to the technical field of preparing ethylbenzene from dry gas, and particularly relates to a method for removing hydrogen sulfide and ammonia from dry gas. Background Art

[0002] Dry gas refers to the tail gas that cannot be liquefied in the process of oil refining. It mainly comes from the secondary processing of crude oil and is generally used as fuel. In order to make full use of the dilute ethylene in dry gas, improve the utilization rate of oil resources, and solve the current shortage of ethylbenzene / styrene market, various countries have developed the process of producing ethylbenzene from dry gas. Dry gas to produce ethylbenzene is generally carried out under the action of a catalyst at a reaction temperature of 200℃-500℃, so that the ethylene and a small amount of propylene and butene in the dry gas react with benzene to produce ethylbenzene, propylbenzene, butylbenzene and polyalkylbenzenes.

[0003] In addition to a large amount of hydrogen, refinery dry gas generally contains a certain amount of hydrogen sulfide and trace amounts of ammonia. Ethylbenzene alkylation catalysts generally use acidic molecular sieves. Since dry gas contains a high content of sulfur-containing impurities, due to its strong polarity, it is easy to adsorb on the active centers on the surface of the catalyst, resulting in the covering of the active centers of the catalyst; ammonia is an alkaline impurity, which will cause the acidic active centers of the catalyst to be poisoned and deactivated. Both will significantly shorten the operating cycle of the ethylbenzene alkylation catalyst. Hydrogen sulfide and ammonia in dry gas are generally removed separately, using a hydrogen sulfide removal tower to remove hydrogen sulfide and ammonia to remove ammonia using a water scrubber. For example, "Summary of Intermediate Tests on Selective Removal of H2S under Pressure with Methyldiethanolamine Aqueous Solution" (Petroleum and Natural Gas Chemical Industry, Issue 2, 1987) discloses a method for removing hydrogen sulfide-containing natural gas using an absorption tower using a methyldiethanolamine aqueous solution, and the hydrogen sulfide content in the purified gas is reduced to 10 mg / m 3 "Industrial Application of Supergravity Complex Iron Method for Removing H2S" (Chemical Industry and Engineering Technology, Vol. 35, No. 2, 2014) discloses a method for removing hydrogen sulfide-containing oil field associated gas using complex iron and a supergravity machine. After desulfurization, the hydrogen sulfide concentration of the associated gas is less than 20 mg / m 2 CN109574780A discloses a method for adsorbing ammonia in dry gas using modified activated carbon. CN111450662A discloses a combined deammoniation process of a water washing tower, a cyclone separator, a coalescer and an adsorption tower, wherein the washing liquid is selected from industrial water or an acidic aqueous solution, and the adsorbent is formed activated carbon, silicon dioxide, alumina, formed molecular sieve and a strong acidic cation exchange resin loaded with an inorganic acid or an organic acid.

[0004] In summary, the existing dry gas desulfurization and deamination technologies are mainly based on the removal of two types of impurities, hydrogen sulfide and ammonia, respectively, resulting in excessive consumption of desulfurizers or deaminizers when removing hydrogen sulfide or ammonia. Summary of the invention

[0005] The present invention provides a method for removing hydrogen sulfide and ammonia from dry gas. The method of the present invention realizes effective contact and timely conversion of hydrogen sulfide and ammonia, and when used to remove hydrogen sulfide and ammonia from dry gas, it can reduce the use of subsequent desulfurization agents or deammonification agents, which is beneficial to reducing the production cost of ethylbenzene devices.

[0006] A first aspect of the present invention provides a method for removing hydrogen sulfide and ammonia from dry gas, comprising:

[0007] S1, the raw dry gas contacts the desulfurization absorption lean liquid in the hydrogen sulfide absorption tower, the desulfurization purified dry gas is obtained at the top of the tower, and the desulfurization absorption rich liquid is obtained at the bottom of the tower;

[0008] S2, the desulfurized purified dry gas is contacted with the deammoniation absorption lean liquid in an ammonia absorption tower, and the deammoniation purified dry gas is obtained at the top of the tower, and the deammoniation absorption rich liquid is obtained at the bottom of the tower;

[0009] S3, the desulfurization absorption rich liquid and the deammoniation absorption rich liquid and the oxygen-containing gas (such as air) enter the catalytic oxidation reactor, contact with the catalytic oxidation catalyst to react, and obtain a thiosulfate emulsion;

[0010] S4, the thiosulfate emulsion is subjected to gas-liquid separation, and the obtained liquid phase is decanted and layered to obtain an upper oil phase, i.e., a regenerated desulfurization absorption lean liquid, and a lower water phase, i.e., a crude deammoniation absorption lean liquid;

[0011] S5, the crude deammonification absorption lean liquid is subjected to removal of thiosulfate (for example, by membrane filtration) to obtain a clear liquid, i.e., regenerated deammonification absorption lean liquid, and a concentrated liquid, i.e., a waste salt solution.

[0012] In the above technical solution, in step S4, the liquid phase after gas-liquid separation can be filtered to remove solid phase substances according to the situation before decantation and stratification.

[0013] In the above technical solution, the catalytic oxidation catalyst described in step S3 includes zinc oxide modified by resin and silanization agent.

[0014] In the above technical solution, in the catalytic oxidation catalyst described in step S3, the resin is a polymerized white ball, which is copolymerized by methyl methacrylate and ethylene glycol dimethacrylate, and the amount of pendant double bonds in the copolymer is 4 to 10 mmol / g.

[0015] In the above technical solution, in the catalytic oxidation catalyst described in step S3, the specific surface area of ​​the zinc oxide is 50 to 70 m 2 / g, pore volume is 0.25~0.4cm 3 / g, the average pore diameter is 10 to 20 angstroms, and the particle size is 10 to 20 nanometers.

[0016] In the above technical scheme, in the catalytic oxidation catalyst described in step S3, in the zinc oxide modified by the silanization agent, the silanization agent is selected from at least one of N-(3-triethoxysilylpropyl)ethylenediamine, 3-aminopropyltriethoxysilane, N-methyl-3-(trimethoxysilyl)propylamine, and 11-aminoundecyltriethoxysilane. Based on the mass of the zinc oxide modified by the silanization agent, the mass content of the silanization agent is 70% to 83%, and the mass content of the zinc oxide is 17% to 30%.

[0017] In the above technical solution, in the catalytic oxidation catalyst described in step S3, based on the weight of the catalyst, the mass content of the resin is 68% to 89%, and the mass content of the zinc oxide modified by the silanization agent is 11% to 32%.

[0018] In the above technical solution, in the catalytic oxidation catalyst described in step S3, the specific surface area of ​​the catalyst is 300 to 600 m 2 / g, the average pore size is 30-40 nanometers, and the particle size of at least 95% of the particles is 0.6 mm to 1.3 mm.

[0019] In the above technical solution, the method for preparing the catalytic oxidation catalyst described in step S3 comprises:

[0020] a. preparing zinc oxide modified by a silanization agent;

[0021] b. Treating the resin with zinc oxide modified with the silanization agent obtained in step a, and drying the resin to obtain the catalyst.

[0022] In the above technical scheme, in step a, the method for preparing zinc oxide modified by a silanization agent comprises: mixing a silanization agent modification solution with zinc oxide, heating and refluxing the mixture to obtain the zinc oxide modified by the silanization agent. The properties of the zinc oxide are as follows: the specific surface area is 50 to 70 m 2 / g, pore volume is 0.25~0.4cm 3 / g, the average pore size is 10-20 angstroms, and the particle size is 10-20 nanometers. The preparation method of the silanization agent modification liquid comprises: by mass fraction, silanization agent: low carbon alcohol: water: alkaline catalyst = (200-250): (200-300): (15-23): (0.6-1.3), silanization agent, low carbon alcohol (the low carbon alcohol is preferably selected from at least one of methanol, ethanol, propanol, isopropanol), water and alkaline catalyst (the alkaline catalyst is preferably selected from at least one of sodium hydroxide and potassium hydroxide) are mixed at 50-70°C for 2-5 hours to obtain the silanization agent modification liquid. The mass ratio of the silanization agent modification liquid and zinc oxide is (250-450): (100-200). The heating reflux reaction conditions are as follows: reflux reaction at 110-130°C for 4-7 hours.

[0023] In the above technical scheme, the resin described in step b is preferably prepared by the following method, including: preparing an aqueous phase and an oil phase separately; adding the oil phase to the aqueous phase for suspension polymerization to obtain a resin, i.e., a polymerized white ball. The aqueous phase includes polyvinyl alcohol (molecular weight 1500-2500), sodium chloride, desalted water, and azobisisobutyronitrile. In terms of mass fractions, polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile = (1-3): (4-9): (80-100): (0.5-1.2). The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and a solvent. The solvent is selected from at least one of toluene, p-xylene, m-xylene, and o-xylene. In terms of mass fractions, methyl methacrylate: ethylene glycol dimethacrylate: sorbitan oleate: solvent = (65-85): (66-87): (7-10): (120-160). The process of adding the oil phase to the water phase for suspension polymerization includes: first cooling the water phase to 18-30°C, then dropping the oil phase into the water phase, controlling the stirring speed of the water phase to control the oil phase droplet size to 0.6-1.5 mm, and after the dropwise addition is completed, heating to 70-90°C and reacting for 5-10 hours. Then, extracting with anhydrous ethanol at 50-70°C for 4-8 times in equal volume, extracting with desalted water at 50-70°C for 4-8 times in equal volume, and drying (drying temperature is 110-150°C, drying time is 1-4 hours) to obtain resin, i.e., polymer white balls.

[0024] In the above technical scheme, in step b, polymer white ball, triethylene glycol di(2-methacrylate) ester and solvent, the solvent is selected from at least one of toluene, p-xylene, m-xylene and o-xylene, in the ratio of polymer white ball: triethylene glycol di(2-methacrylate) ester: solvent = 50: (2-5): (150-200) by weight, swell at 60-80°C for 6-10 hours to obtain system I. Zinc oxide modified with silanization agent obtained in step a, diacyl peroxide and N,N-dimethylformamide, in the ratio of zinc oxide modified with silanization agent: diacyl peroxide: N,N-dimethylformamide = 20: (0.3-0.8): (160-230) by weight, stir at 70-90°C for 6-10 hours to obtain system II. The system II is added dropwise to the system I, with a volume ratio of system II: system I = 1-3, and the temperature is raised to 100-120° C., and refluxed for 3-6 hours. Then, anhydrous ethanol is used for extraction at 50-70° C. for 4-8 times of equal volume, and desalted water is used for extraction at 50-70° C. for 4-8 times of equal volume, and dried (drying temperature is 100-120° C., and drying time is 1-5 hours) to obtain the catalyst.

[0025] In the above technical solution, the desulfurization absorption lean liquid obtained in step S4 can be recycled in step S1.

[0026] In the above technical solution, the deammoniation absorption lean liquid obtained in step S5 can be recycled in step S2.

[0027] In the above technical scheme, in step S1, the raw material dry gas is derived from catalytic cracking, thermal cracking, delayed coking, and hydrocracking in the refinery. The raw material dry gas includes, but is not limited to, ethylene, methane, ethane, propane, propylene, isobutane, n-butane, trans-butene, n-butene, isobutene, cis-butene, oxygen, nitrogen, hydrogen, carbon monoxide, carbon dioxide, acetylene, 1,3-butadiene, alkanes or alkenes with carbon five or more, and also includes hydrogen sulfide and ammonia. In the raw material dry gas, the volume content of ethylene is 5% to 40%, preferably 10% to 20%, and the volume content of hydrogen is 15% to 25%. In the raw material dry gas, the volume content of hydrogen sulfide is not less than 10000ppm, preferably not less than 70000ppm, further, the volume content of hydrogen sulfide is not higher than 100000ppm, the volume content of ammonia is not less than 6000ppm, preferably not less than 30000ppm, further, the volume content of ammonia is not higher than 50000ppm.

[0028] In the above technical solution, in step S1, the hydrogen sulfide absorption tower is a conventional plate tower or packed tower in the art. The desulfurization absorption lean liquid contacts with the raw dry gas in countercurrent in the tower body.

[0029] In the above technical scheme, in step S1, the desulfurization absorption lean liquid used includes 120# solvent oil after light removal and p-xylene. By mass percentage, the mass content of 120# solvent oil after light removal is 75% to 86%, and the mass content of p-xylene is 10% to 20%. The 120# solvent oil after light removal is 120# solvent oil after light components with low boiling points of 40 to 90°C are removed by conventional distillation. The usage ratio of raw material dry gas to desulfurization absorption lean liquid is 80 to 200 by volume.

[0030] In the above technical solution, in step S1, the hydrogen sulfide absorption tower has 2 to 7 theoretical plates, an operating temperature of 8 to 40°C, and an operating pressure of 0.5 to 1.2 MPaG. Preferably, a demister is provided at the top of the hydrogen sulfide absorption tower, and the separation accuracy is 3 to 10 microns.

[0031] In the above technical solution, in step S2, the ammonia absorption tower is a conventional plate tower or packed tower in the art. The deammoniation absorption lean liquid and the desulfurization purified dry gas are in countercurrent contact in the tower body.

[0032] In the above technical solution, in step S2, the theoretical plate number of the ammonia absorption tower is 3 to 9, the operating temperature is 10 to 40°C, and the operating pressure is 0.5 to 1.2 MPaG. Preferably, a demister is provided on the top of the ammonia absorption tower, and the separation accuracy is 3 to 10 microns. The deammoniation absorption lean liquid is water (such as industrial fresh water, or desalted water). The dosage ratio of the desulfurization purified dry gas to the deammoniation absorption lean liquid is, by volume, 60 to 90.

[0033] In the above technical solution, in step S3, preferably, the feed amount of the desulfurization absorption rich liquid and the deammoniation absorption rich liquid into the catalytic oxidation reactor is controlled, that is, the molar ratio of ammonia in the deammoniation absorption rich liquid to hydrogen sulfide in the desulfurization absorption rich liquid is (0.8-1.8): 1. The ammonia in the deammoniation absorption rich liquid and the hydrogen sulfide in the desulfurization absorption rich liquid can be analyzed by offline and / or online solute concentration, and the flow ratio of the desulfurization absorption rich liquid to the deammoniation absorption rich liquid can be manually and / or automatically controlled. Preferably, the feed amount of the desulfurization absorption rich liquid and the deammoniation absorption rich liquid is controlled by online analysis of the ammonia concentration in the deammoniation absorption rich liquid and the hydrogen sulfide concentration in the desulfurization absorption rich liquid.

[0034] In the above technical solution, in step S3, the molar ratio of the amount of oxygen-containing gas used in terms of O2 to ammonia in the deamination absorption rich liquid is (100-200): 1. Preferably, the amount of oxygen-containing gas added is controlled in combination with the flow rate of the deamination absorption rich liquid and the ammonia solute concentration analyzed offline and / or online.

[0035] In the above technical solution, in step S3, the catalytic oxidation reactor is filled with the catalytic oxidation catalyst described in the present invention.

[0036] In the above technical solution, preferably, in step S3, the desulfurization absorption rich liquid and the deamination absorption rich liquid can be mixed first, and then enter the catalytic oxidation reactor after heat exchange with the reaction product of step S3.

[0037] In the above technical solution, preferably, in step S3, the oxygen-containing gas (such as air) is preheated before entering the catalytic oxidation reactor.

[0038] In the above technical scheme, in step S3, the reaction conditions of the reaction are as follows: the reaction temperature is 90-120°C, the reaction pressure is 0.5-1.0 MPaG, and the reaction volume space velocity is 0.5-1.0 h -1 .

[0039] In the above technical solution, in step S4, the gas-liquid separation temperature of the thiosulfate emulsion is 30-40° C., the filtration accuracy is 3-8 microns, and the liquid-liquid stratification temperature is 30-40° C. The upper oil phase obtained by liquid-liquid stratification, i.e., the desulfurization absorption lean liquid, is pressurized by 100-500 kPa by the desulfurization absorption lean liquid circulation pump and then circulated to the hydrogen sulfide absorption tower in step S1.

[0040] In the above technical scheme, in step S5, the membrane filtration is an organic membrane with a molecular weight cutoff of 100 to 300 Daltons, an operating pressure of 2000 to 4000 kPa, an operating temperature of 30 to 50°C, and a concentrated liquid discharge of 2% to 6% of the membrane feed volume. The obtained membrane filtration clear liquid has a desalination rate of 96% to 99% in terms of thiosulfate. The desalted membrane filtration clear liquid, i.e., the regenerated deammoniation absorption rich liquid, i.e., the deammoniation absorption lean liquid, is returned to the ammonia absorption tower of step S2 for continued reuse.

[0041] A second aspect of the present invention provides a separation system for removing hydrogen sulfide and ammonia from dry gas, comprising:

[0042] 1) Hydrogen sulfide absorption tower, used to absorb hydrogen sulfide using desulfurized absorption lean liquid to provide a sulfur source for subsequent reactions;

[0043] 2) Ammonia absorption tower, used to remove ammonia from dry gas by deammonification absorption of lean liquid and provide nitrogen source for subsequent reactions;

[0044] 3) a catalytic oxidation reactor, used to convert unstable ammonium sulfide into stable thiosulfate by catalytic oxidation;

[0045] 4) A separation system is used to remove thiosulfate so as to regenerate the desulfurization absorption liquid and the deammoniation absorption liquid.

[0046] Compared with the prior art, the present invention has the following advantages:

[0047] 1. The inventors of the present invention found in the research process that a certain amount of hydrogen sulfide and a small amount of ammonia generally exist in the raw dry gas at the same time. The acidic impurity hydrogen sulfide to be removed and the alkaline impurity ammonia to be removed will contact to form ammonium sulfide. However, due to the poor stability of ammonium sulfide, it will decompose into hydrogen sulfide and ammonia again. In particular, when there is a large amount of reducing gas (such as hydrogen, carbon monoxide, etc.) in the raw gas, the stability of ammonium sulfide is even worse. The existing process generally removes hydrogen sulfide and ammonia in the raw gas separately, using a desulfurization tower to remove hydrogen sulfide and a water scrubber or a water scrubber to remove ammonia. For safety reasons, it is difficult to oxidize and remove low-concentration ammonium sulfide in the reducing atmosphere. The inventors have found through further research that the weakly polar solvent with regulated polarity has a high solubility for hydrogen sulfide, and ammonia can be directly absorbed by circulating water. The catalytic oxidation catalyst of the present invention is used as a phase interface catalyst to solve the immiscible problem of organic solvent and water when they are in contact, and the contact of hydrogen sulfide and ammonia under the action of the catalyst to produce a catalytic oxidation reaction to obtain a relatively stable thiosulfate; through two parallel absorptions, the impurities to be removed are separated from the main body of the raw material dry gas, and the subsequent process is safe; the polarity of the desulfurization absorbent is regulated while taking into account the effective phase separation with water, so that the mixed emulsion after catalytic oxidation can be separated again by phase separation, simplifying the separation operation; at the same time, by controlling the degree of oxidation to generate thiosulfate, it is convenient to remove it by membrane separation. At the same time, the oxidation catalyst prepared by the method of the present invention, when used for catalytic oxidation of emulsions, has high activity stability and long service life relative to conventional supported catalysts.

[0048] 2. The present invention solves the problem of low process matching when the acidic impurity hydrogen sulfide and the alkaline impurity ammonia are often removed separately when a certain amount of hydrogen sulfide and a small amount of ammonia are simultaneously present in the existing raw dry gas under the condition of high hydrogen content. Ammonia and hydrogen sulfide are transferred to the outside of the raw dry gas system by a preferred absorbent, and then the absorbed hydrogen sulfide and ammonia are converted into relatively stable impurities by the catalytic oxidation catalyst of the present invention and removed from the absorption system. The process is safe, and the use of subsequent desulfurization agents or deamination agents is reduced, which is beneficial to reducing the production cost of ethylbenzene devices. BRIEF DESCRIPTION OF THE DRAWINGS

[0049] Figure 1 This is a schematic diagram of a separation system for jointly removing hydrogen sulfide and ammonia from dry gas according to the present invention;

[0050] The reference numerals are described as follows:

[0051] 101 is a hydrogen sulfide absorption tower, 102 is an ammonia absorption tower, 103 is a desulfurization absorption rich liquid buffer tank, 104 is a desulfurization absorption rich liquid reaction feed pump, 105 is a deammoniation absorption rich liquid buffer tank, 106 is a deammoniation absorption rich liquid reaction feed pump, 107 is a catalytic oxidation reactor inlet and outlet heat exchanger, 108 is an air preheater, 109 is a catalytic oxidation reactor, 110 is an oxidation product gas-liquid separation tank, 111 is a liquid product delivery pump, 112 is a liquid product filter, 113 is a liquid-liquid decanter, 114 is an aqueous phase feed pump, 115 is a membrane filter, 116 is Catalytic oxidation reactor feed heater, 117 is absorption liquid component analysis and flow controller, 118 is catalytic oxidation reactor discharge cooler, 119 is catalytic oxidation reactor feed gas-liquid mixer, 120 is membrane filter clear liquid buffer tank, 121 is deamination absorption lean liquid circulation pump, 122 is desulfurization absorption lean liquid circulation pump, 201 is raw material dry gas inlet, 202 is supplementary fresh water inlet, 203 is air inlet, 204 is oxidation tail gas outlet, 205 is filter drain outlet, 206 is membrane filter concentrate discharge outlet, 207 is deamination purified dry gas discharge outlet. DETAILED DESCRIPTION

[0052] The technical solutions in the embodiments of the present invention are described clearly and completely below. Obviously, the described embodiments are only part of the embodiments of the present invention, not all of them. Based on the embodiments of the present invention, all other embodiments obtained by those skilled in the art without creative work are within the protection scope of the present invention.

[0053] In the present invention, the analysis of low-carbon hydrocarbons, hydrogen sulfide, and ammonia components uses a gas chromatograph Agilent 7890AGC (Agilent, USA), equipped with an HP-PLOT Al2O3 KCl (50m×0.53mm×15μm) capillary chromatographic column, and the column temperature is maintained at 100°C for 10 minutes, then increased to 120°C at 30°C per minute, and maintained for 3 minutes. The carrier gas He flow rate is 3mL / min, the injection volume is 0.1mL (quantitative loop), the split ratio is 5:1, the injection port temperature is 250°C, and the detector (FID) temperature is 250°C.

[0054] In the present invention, the content of zinc oxide in the resin is determined by a calcination method in an air atmosphere at a calcination temperature of 600-700° C. to obtain the amount of zinc oxide loaded on a certain amount of the resin.

[0055] In the present invention, the amount of dangling double bonds in the resin is analyzed by titration: first, the resin is fully swollen with carbon tetrachloride, then further soaked with pyridine, sulfuric acid, bromine glacial acetic acid solution, carbon tetrachloride and mercuric acetate solution, then potassium iodide solution is added and titrated with sodium thiosulfate standard solution.

[0056] A separation system for jointly removing hydrogen sulfide and ammonia from dry gas according to the present invention (such as Figure 1101 is a hydrogen sulfide absorption tower, 102 is an ammonia absorption tower, 103 is a desulfurization absorption rich liquid buffer tank, 104 is a desulfurization absorption rich liquid reaction feed pump, 105 is a deammoniation absorption rich liquid buffer tank, 106 is a deammoniation absorption rich liquid reaction feed pump, 107 is a catalytic oxidation reactor inlet and outlet heat exchanger, 108 is an air preheater, 109 is a catalytic oxidation reactor, 110 is an oxidation product gas-liquid separation tank, 111 is a liquid product delivery pump, 112 is a liquid product filter, 113 is a liquid-liquid decanter, 114 is a water phase feed pump, 115 is a membrane filter, 116 is a catalytic oxidation reactor feed heater, 117 is an absorption liquid component analysis and flow controller, 118 is a catalytic oxidation reactor discharge cooler, 119 is a catalytic oxidation reactor feed gas-liquid mixer, 120 is a membrane filter clear liquid buffer tank, 121 is a deammonification absorption lean liquid circulation pump, and 122 is a desulfurization absorption lean liquid circulation pump. The separation process is as follows: the raw dry gas (for example, from refinery catalytic cracking, thermal cracking, delayed coking, and hydrocracking) enters the middle and lower part of the hydrogen sulfide absorption tower 101 from the raw dry gas inlet 201, and countercurrently contacts with the desulfurization absorption lean liquid in the tower body, the desulfurization purified dry gas at the top of the hydrogen sulfide absorption tower 101 enters the ammonia absorption tower 102, and countercurrently contacts with the deammonification absorption lean liquid in the tower body, and the deammonification purified dry gas is discharged from the top of the ammonia absorption tower 102. Afterwards, if the sulfur content in the deamination purified dry gas still does not meet the standard, it can enter the subsequent desulfurization system to continue the desulfurization treatment. The rich liquid in the bottom of the hydrogen sulfide absorption tower 101 (i.e., the desulfurization absorption rich liquid) and the rich liquid in the bottom of the ammonia absorption tower 102 (deamination absorption rich liquid) are analyzed by the absorption liquid component analysis and flow controller 117. According to the composition and flow of the deamination absorption rich liquid, the flow of the desulfurization liquid is calculated and controlled with reference to the composition of the desulfurization absorption rich liquid.The deammoniation absorption rich liquid and the calculated and adjusted desulfurization absorption rich liquid are heat exchanged with the outlet logistics of the catalytic oxidation reactor 109 through the catalytic oxidation reactor inlet and outlet heat exchanger 107, and heated to a certain temperature through the catalytic oxidation reactor feed heater 116. After being mixed with the air heated by the air preheater 108 in the catalytic oxidation reactor feed gas-liquid mixer 119, they enter the catalytic oxidation reactor and contact with the catalytic oxidation catalyst of the present invention to react. The obtained oxidation products are cooled in turn through the catalytic oxidation reactor inlet and outlet heat exchanger 107 and the catalytic oxidation reactor outlet cooler 118 and then enter the oxidation product gas-liquid separation tank 110. The gas phase part of the gas-liquid separation tank is discharged to the outside from the oxidation tail gas outlet 204, and the liquid phase part of the gas-liquid separation tank is discharged to the outside from the oxidation tail gas outlet 204. After passing through the liquid product delivery pump 111 and the liquid product filter 112, it enters the liquid-liquid decanter 113, in which the oil phase and the water phase are separated, the upper oil phase is circulated back to the hydrogen sulfide absorption tower 101 for continued use through the desulfurization absorption lean liquid circulation pump 122, and the lower water phase is pressurized by the water phase feed pump 114 and enters the membrane filter 115. The membrane filter 115 discharges a part of the concentrated liquid from the membrane filter concentrated liquid discharge port 206 according to a certain proportion, and replenishes an equal amount of fresh water through the fresh water inlet 202. The clear liquid of the membrane filter 115, that is, the regenerated deammoniation absorption lean liquid, is buffered by the membrane filter clear liquid buffer tank 120, and then circulated back to the ammonia absorption tower 102 for continued use through the deammoniation absorption lean liquid circulation pump 121.

[0057] [Example 1]

[0058] The raw dry gas described in this embodiment is derived from catalytic cracking in a refinery, and the composition of the dry gas is calculated by volume, including: 17.9901% methane, 9.7208% ethane, 20.7122% ethylene, 0.1761% propane, 0.8422% propylene, 0.0708% isobutane, 0.0182% normal butane, 0.0073% trans-butene, 0.0155% normal butene, 0.0361% isobutene, 0.0002% cis-butene, 0.1075% oxygen, 1.0856% nitrogen, 28.7099% hydrogen, 9.6384% carbon monoxide, 3.0546% carbon dioxide, 0.0038% acetylene, 0.0001% 1,3-butadiene, 1.4061% alkanes or alkenes with a carbon content of more than five, 3.8906% hydrogen sulfide, 1.8767% ammonia, and 0.6263% other components. The feed volume of raw dry gas is 30000Nm 3 / Hour.

[0059] The separation process of this embodiment is the same as Figure 1 shown.

[0060] The hydrogen sulfide absorption tower is a plate tower. The desulfurization absorption lean liquid contacts the raw dry gas in countercurrent in the tower body. The hydrogen sulfide absorption tower has 4 theoretical plates, an operating temperature of 25°C, and an operating pressure of 0.7MPaG. A demister is provided on the top of the hydrogen sulfide absorption tower with a separation accuracy of 5 microns. The desulfurization absorption lean liquid includes 120# solvent oil (distillation range of 90-120°C) and paraxylene after light removal. By mass percentage, the mass content of 120# solvent oil after light removal is 82.7%, the mass content of paraxylene is 16.8%, and the rest is water and trace amounts of hydrogen sulfide. The usage ratio of raw dry gas to desulfurization absorption lean liquid is, by volume, 140.

[0061] The ammonia absorption tower is a plate tower. The deammoniation absorption lean liquid and the desulfurization purified dry gas are in countercurrent contact in the tower body. The theoretical number of plates of the ammonia absorption tower is 5, the operating temperature is 25℃, and the operating pressure is 0.7MPaG. A defoamer is provided at the top of the ammonia absorption tower, and the separation accuracy is 5 microns. The deammoniation absorption lean liquid is desalted water. The dosage ratio of the desulfurization purified dry gas to the deammoniation absorption lean liquid is, by volume, 75.

[0062] The desulfurization absorption rich liquid is mixed with the deammoniation absorption rich liquid. By analyzing the solute concentration online, the flow ratio of the desulfurization absorption rich liquid to the deammoniation absorption rich liquid is automatically controlled. The molar ratio of ammonia and hydrogen sulfide during mixing is controlled to be 1.4:1.

[0063] When the mixed liquid phase stream of the desulfurization absorption rich liquid and the deammoniation absorption rich liquid is mixed with air, the molar ratio of ammonia and oxygen when the mixed liquid phase stream is mixed with air is controlled to be 1:150 based on the flow rate of the deammoniation absorption rich liquid and the ammonia solute concentration analyzed online.

[0064] A catalytic oxidation reactor is provided with a catalytic oxidation catalyst, and the preparation method thereof comprises:

[0065] a. Preparation of zinc oxide modified with silanization reagent:

[0066] Among them, the properties of zinc oxide are as follows: the specific surface area is 60m 2 / g, pore volume is 0.32cm 3 / g, the average pore size is 14 angstroms, and the particle size is 15 nanometers. The preparation method of the silanization agent modification liquid includes: by mass, the silanization agent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 225:250:19:0.9, the silanization agent, ethanol, water and sodium hydroxide are evenly mixed and reacted at 60°C for 3 hours to obtain the silanization agent modification liquid. The mass ratio of the silanization agent modification liquid and zinc oxide is 350:150, the silanization agent modification liquid and zinc oxide are evenly mixed and refluxed at 120°C for 5 hours to obtain zinc oxide modified by the silanization agent; wherein, based on the mass of the zinc oxide modified by the silanization agent, the mass content of the silanization agent is 76%, and the mass content of zinc oxide is 24%;

[0067] b. Treating the resin with zinc oxide modified with the silanization agent obtained in step a, and drying the resin to obtain the catalyst:

[0068] The resin is prepared by the following method, including: preparing water phase and oil phase separately; adding the oil phase to the water phase for suspension polymerization to obtain the resin, i.e., polymerized white balls. The water phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, desalted water, and azobisisobutyronitrile, and the weight ratio of polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile is 2:6.5:90:0.85. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene, and the weight ratio of methyl methacrylate: ethylene glycol dimethacrylate: sorbitan oleate: toluene is 75:76:8.5:140. The process of adding the oil phase to the water phase for suspension polymerization includes: cooling the water phase to 25°C, then dripping the oil phase into the water phase, controlling the stirring speed of the water phase, and controlling the oil phase droplet size to 1.1 mm. After the dripping is completed, the temperature is raised to 80°C and the reaction is performed for 7.5 hours. Then, anhydrous ethanol was used to extract 6 times at 60°C in equal volume, and desalted water was used to extract 6 times at 60°C in equal volume, and dried (drying temperature was 130°C, and drying time was 2.5h) to obtain polymer white balls (the amount of pendant double bonds in the copolymer was 7mmol / g). The polymer white balls were swollen at 70°C for 8 hours with triethylene glycol di(2-methacrylate) ester and toluene in the ratio of polymer white balls: triethylene glycol di(2-methacrylate) ester: toluene = 50:3.5:175 by mass to obtain system I. The zinc oxide modified with silanization reagent obtained in step a was stirred at 80°C for 8 hours with diacyl peroxide and N,N-dimethylformamide in the ratio of zinc oxide modified with silanization reagent obtained in step a: diacyl peroxide: N,N-dimethylformamide = 20:0.5:195 by mass to obtain system II. The system II was added dropwise to the system I, with a volume ratio of system II: system I = 2, and the temperature was raised to 110°C and refluxed for 4.5 hours. Then, anhydrous ethanol was used for extraction at 60°C for 6 times in equal volume, and desalted water was used for extraction at 60°C for 6 times in equal volume. The drying temperature was 110°C and the drying time was 3 hours to obtain the catalyst.

[0069] Based on the weight of the catalyst, the mass content of the resin was 79.5%, and the mass content of the zinc oxide modified by the silanization agent was 20.5%.

[0070] Catalytic oxidation reactor, reaction temperature is 105℃, reaction pressure is 0.7MPaG, based on mixed solution, reaction volume space velocity is 0.75h -1 The thiosulfate emulsion at the reactor outlet has a gas-liquid separation temperature of 35°C, a filtration accuracy of 5 microns, and a liquid-liquid stratification temperature of 35°C. The upper oil phase obtained by liquid-liquid stratification, i.e., the desulfurization absorption lean liquid, is pressurized by 300 kPa by the desulfurization absorption lean liquid circulation pump and then circulated to the hydrogen sulfide absorption tower; the lower water phase, i.e., the crude deammoniation absorption lean liquid.

[0071] The crude deammoniation absorption lean liquid is filtered through a membrane to obtain a clear liquid, i.e., deammoniation absorption lean liquid, and a concentrated liquid, i.e., a waste salt solution. The membrane specification is an organic membrane with a molecular weight cutoff of 200 Daltons, an operating pressure of 3000 kPa, an operating temperature of 40°C, and a concentrated liquid discharge volume of 4% of the membrane feed volume. The obtained membrane filtration clear liquid, calculated as thiosulfate, has a desalination rate of 97.5%. The desalted membrane filtration clear liquid, i.e., deammoniation absorption lean liquid, is returned to the ammonia absorption tower for continued reuse.

[0072] The volume content of ammonia in the deammoniation purified dry gas obtained in this embodiment is 0.0003%, which is 1.8764 percentage points lower than that of the dry gas raw material. The content of hydrogen sulfide in the deammoniation purified dry gas is 1.2096%, which is 2.6810 percentage points lower than that of the dry gas raw material. By matching the desulfurization and deammoniation processes, the use of desulfurizers and deammonifiers is significantly reduced.

[0073] [Example 2]

[0074] The raw dry gas described in this embodiment is the same as that in embodiment 1. The separation process of this embodiment is the same as that in embodiment 1.

[0075] The hydrogen sulfide absorption tower is a plate tower. The desulfurization absorption lean liquid contacts the raw dry gas in countercurrent in the tower body. The hydrogen sulfide absorption tower has 6 theoretical plates, an operating temperature of 10°C, and an operating pressure of 1.0MPaG. A demister is provided on the top of the hydrogen sulfide absorption tower with a separation accuracy of 5 microns. The desulfurization absorption lean liquid includes 120# solvent oil (distillation range of 90-120°C) after light removal and paraxylene. By mass percentage, the mass content of 120# solvent oil after light removal is 85.2%, the mass content of paraxylene is 14.7%, and the rest is water and trace amounts of hydrogen sulfide. The usage ratio of raw dry gas to desulfurization absorption lean liquid is, by volume, 90.

[0076] The ammonia absorption tower is a plate tower. The deammoniation absorption lean liquid and the desulfurization purified dry gas are in countercurrent contact in the tower body. The theoretical number of plates of the ammonia absorption tower is 8, the operating temperature is 12°C, and the operating pressure is 1.0MPaG. A defoamer is provided at the top of the ammonia absorption tower, and the separation accuracy is 5 microns. The deammoniation absorption lean liquid is desalted water. The dosage ratio of the desulfurization purified dry gas to the deammoniation absorption lean liquid is, by volume, 65.

[0077] The desulfurization absorption rich liquid is mixed with the deammoniation absorption rich liquid. By analyzing the solute concentration online, the flow ratio of the desulfurization absorption rich liquid to the deammoniation absorption rich liquid is automatically controlled. The molar ratio of ammonia and hydrogen sulfide during mixing is controlled to be 1.0:1.

[0078] When the mixed liquid phase stream of the desulfurization absorption rich liquid and the deammoniation absorption rich liquid is mixed with air, the molar ratio of ammonia and oxygen when the mixed liquid phase stream is mixed with air is controlled to be 1:180 based on the flow rate of the deammoniation absorption rich liquid and the ammonia solute concentration analyzed online.

[0079] A catalytic oxidation reactor is provided with a catalytic oxidation catalyst, and the preparation method thereof comprises:

[0080] a. Preparation of zinc oxide modified with silanization reagent:

[0081] Among them, the properties of zinc oxide are as follows: specific surface area is 65m 2 / g, pore volume is 0.38cm 3 / g, the average pore size is 11 angstroms, and the particle size is 12 nanometers. In terms of mass fractions, silanization agent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 210:280:17:0.7. After the silanization agent, ethanol, water and sodium hydroxide are evenly mixed, they are reacted at 65°C for 4 hours to obtain the silanization agent modification solution. The mass ratio of the silanization agent modification solution and zinc oxide is 300:180. After the silanization agent modification solution and zinc oxide are evenly mixed, they are refluxed at 125°C for 6 hours to obtain zinc oxide modified with a silanization agent; wherein, based on the mass of the zinc oxide modified with the silanization agent, the mass content of the silanization agent is 72%, and the mass content of zinc oxide is 28%;

[0082] b. Treating the resin with zinc oxide modified with the silanization agent obtained in step a, and drying the resin to obtain the catalyst:

[0083] The resin is prepared by the following method, including: preparing water phase and oil phase separately; adding the oil phase to the water phase for suspension polymerization to obtain the resin, i.e., polymerized white balls. The water phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, desalted water, and azobisisobutyronitrile, and the weight ratio of polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile is 2.5:8.5:95:1.0. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene, and the weight ratio of methyl methacrylate: ethylene glycol dimethacrylate: sorbitan oleate: toluene is 80:81:9.0:150. The process of adding the oil phase to the water phase for suspension polymerization includes: cooling the water phase to 20°C, then dripping the organic phase into the water phase, controlling the stirring speed of the water phase, and controlling the oil phase droplet size to 0.8 mm. After the dripping is completed, the temperature is raised to 85°C and the reaction is performed for 9.5 hours. Filter, extract with anhydrous ethanol at 65℃ for 7 times, extract with desalted water at 65℃ for 7 times, and dry (drying temperature is 145℃, drying time is 3.5h) to obtain polymer white balls (the amount of pendant double bonds in the copolymer is 9mmol / g). The polymer white balls are swollen at 75℃ for 9 hours with triethylene glycol di(2-methacrylate) ester and toluene in the ratio of polymer white balls: triethylene glycol di(2-methacrylate) ester: toluene = 50:4.5:195 by mass to obtain system I. The zinc oxide modified with silanization reagent obtained in step a is stirred at 75℃ for 7 hours with diacyl peroxide and N,N-dimethylformamide in the ratio of zinc oxide modified with silanization reagent obtained in step a: diacyl peroxide: N,N-dimethylformamide = 20:0.7:210 by mass to obtain system II. The system II was added dropwise to the system I, with a volume ratio of system II: system I = 2.7, and the temperature was raised to 105°C, and refluxed for 5 hours. Then, anhydrous ethanol was used for extraction at 65°C for 7 times with equal volume, and desalted water was used for extraction at 65°C for 7 times with equal volume, and the drying temperature was 115°C for 4 hours to obtain the catalyst.

[0084] Based on the weight of the catalyst, the mass content of the resin was 71.7%, and the mass content of the zinc oxide modified by the silanization agent was 28.3%.

[0085] Catalytic oxidation reactor, reaction temperature is 115℃, reaction pressure is 0.9MPaG, based on mixed solution, reaction volume space velocity is 0.65h -1 The thiosulfate emulsion at the reactor outlet has a gas-liquid separation temperature of 32°C, a filtration accuracy of 4 microns, and a liquid-liquid stratification temperature of 32°C. The upper oil phase obtained by liquid-liquid stratification, i.e., the desulfurization absorption lean liquid, is pressurized to 400 kPa by the desulfurization absorption lean liquid circulation pump and then circulated to the hydrogen sulfide absorption tower; the lower water phase, i.e., the crude deammoniation absorption lean liquid.

[0086] The crude deammoniation absorption lean liquid is filtered through a membrane to obtain a clear liquid, i.e., deammoniation absorption lean liquid, and a concentrated liquid, i.e., a waste salt solution. The membrane specification is an organic membrane with a molecular weight cutoff of 150 Daltons, an operating pressure of 3500 KPa, an operating temperature of 35°C, and a concentrated liquid discharge volume of 5% of the membrane feed volume. The obtained membrane filtration clear liquid, calculated as thiosulfate, has a desalination rate of 98.5%. The desalted membrane filtration clear liquid, i.e., deammoniation absorption lean liquid, is returned to the ammonia absorption tower for continued reuse.

[0087] The volume content of ammonia in the deammoniation purified dry gas obtained in this embodiment is 0.0001%, which is 1.8766 percentage points lower than that of the dry gas raw material. The content of hydrogen sulfide in the deammoniation purified dry gas is 0.1372%, which is 3.7534 percentage points lower than that of the dry gas raw material. By matching the desulfurization and deammoniation processes, the use of desulfurizers and deammonifiers is greatly reduced.

[0088] [Example 3]

[0089] The raw dry gas described in this embodiment is the same as that in embodiment 1. The separation process of this embodiment is the same as that in embodiment 1.

[0090] The hydrogen sulfide absorption tower is a plate tower. The desulfurization absorption lean liquid contacts the raw dry gas in countercurrent in the tower body. The hydrogen sulfide absorption tower has 3 theoretical plates, an operating temperature of 35°C, and an operating pressure of 0.6MPaG. A demister is provided on the top of the hydrogen sulfide absorption tower with a separation accuracy of 4 microns. The desulfurization absorption lean liquid includes 120# solvent oil (distillation range of 90-120°C) and paraxylene after light removal. By mass percentage, the mass content of 120# solvent oil after light removal is 80.9%, the mass content of paraxylene is 18.7%, and the rest is water and trace amounts of hydrogen sulfide. The usage ratio of raw dry gas to desulfurization absorption lean liquid is, by volume, 180.

[0091] The ammonia absorption tower is a plate tower. The deammoniation absorption lean liquid and the desulfurization purified dry gas are in countercurrent contact in the tower body. The theoretical number of plates of the ammonia absorption tower is 4, the operating temperature is 35°C, and the operating pressure is 0.6MPaG. Preferably, a demister is provided at the top of the ammonia absorption tower, and the separation accuracy is 4 microns. The deammoniation absorption lean liquid is desalted water. The dosage ratio of the desulfurization purified dry gas to the deammoniation absorption lean liquid is, by volume, 80.

[0092] The desulfurization absorption rich liquid is mixed with the deammoniation absorption rich liquid. By analyzing the solute concentration online, the flow ratio of the desulfurization absorption rich liquid to the deammoniation absorption rich liquid is automatically controlled. The molar ratio of ammonia and hydrogen sulfide during mixing is controlled to be 1.6:1.

[0093] When the mixed liquid phase stream of the desulfurization absorption rich liquid and the deammoniation absorption rich liquid is mixed with air, the molar ratio of ammonia and oxygen when the mixed liquid phase stream is mixed with air is controlled to be 1:130 based on the flow rate of the deammoniation absorption rich liquid and the ammonia solute concentration analyzed online.

[0094] A catalytic oxidation reactor is provided with a catalytic oxidation catalyst, and the preparation method thereof comprises:

[0095] a. Preparation of zinc oxide modified with silanization reagent:

[0096] Among them, the properties of zinc oxide are as follows: specific surface area is 55m 2 / g, pore volume is 0.29cm 3 / g, the average pore size is 17 angstroms, and the particle size is 18 nanometers. The preparation method of the silanization agent modification liquid includes: by mass, silanization agent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 240:220:20:1.1, the silanization agent, ethanol, water and sodium hydroxide are mixed evenly and reacted at 55°C for 2.5 hours to obtain the silanization agent modification liquid. The mass ratio of the silanization agent modification liquid and zinc oxide is 400:120, the silanization agent modification liquid and zinc oxide are mixed evenly and refluxed at 115°C for 4.5 hours to obtain zinc oxide modified by the silanization agent, wherein, based on the mass of the zinc oxide modified by the silanization agent, the mass content of the silanization agent is 79%, and the mass content of zinc oxide is 21%;

[0097] b. Treating the resin with zinc oxide modified with the silanization agent obtained in step a, and drying the resin to obtain the catalyst:

[0098] The resin is prepared by the following method, including: preparing water phase and oil phase separately; adding the oil phase to the water phase for suspension polymerization to obtain the resin, i.e., polymerized white balls. The water phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, desalted water, and azobisisobutyronitrile, and the weight ratio of polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile is 1.5:4.5:85:0.7. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene, and the weight ratio of methyl methacrylate: ethylene glycol dimethacrylate: sorbitan oleate: toluene is 70:71:8.0:130. The process of adding the oil phase to the water phase for suspension polymerization includes: cooling the water phase to 28°C, then dripping the organic phase into the water phase, controlling the stirring speed of the water phase, and controlling the oil phase droplet size to 1.3 mm. After the dripping is completed, the temperature is raised to 75°C and the reaction is performed for 6.0 hours. Filter, extract with anhydrous ethanol at 55°C for 5 times with equal volume, extract with desalted water at 55°C for 5 times with equal volume, dry (drying temperature is 116°C, drying time is 2h), and obtain polymer white balls (the amount of pendant double bonds in the copolymer is 5mmol / g). The polymer white balls are swollen at 65°C for 7 hours with triethylene glycol di(2-methacrylate) ester and toluene in the ratio of polymer white balls: triethylene glycol di(2-methacrylate) ester: toluene = 50:2.5:160 by mass to obtain system I. The zinc oxide modified with silanization reagent obtained in step a is stirred at 85°C for 9 hours with diacyl peroxide and N,N-dimethylformamide in the ratio of zinc oxide modified with silanization reagent obtained in step a: diacyl peroxide: N,N-dimethylformamide = 20:0.4:170 by mass to obtain system II. System II was added dropwise to system I, with a volume ratio of system II: system I = 1.2, and the temperature was raised to 115°C, and refluxed for 3.5 hours. Then, anhydrous ethanol was used for extraction at 55°C for 5 times with equal volume, and desalted water was used for extraction at 55°C for 5 times with equal volume, and the drying temperature was 106°C for 1.3 hours to obtain the catalyst.

[0099] Based on the weight of the catalyst, the mass content of the resin was 86.6%, and the mass content of the zinc oxide modified by the silanization agent was 13.4%.

[0100] Catalytic oxidation reactor, reaction temperature is 95℃, reaction pressure is 0.6MPaG, based on mixed solution, reaction volume space velocity is 0.9h -1 The gas-liquid separation temperature of the thiosulfate emulsion at the reactor outlet is 38°C, the filtration accuracy is 6 microns, and the liquid-liquid stratification temperature is 38°C. The upper oil phase obtained by liquid-liquid stratification, i.e., the desulfurization absorption lean liquid, is pressurized by 200kPa by the desulfurization absorption lean liquid circulation pump and then circulated to the hydrogen sulfide absorption tower; the lower water phase, i.e., the crude deammoniation absorption lean liquid.

[0101] The crude deammoniation absorption lean liquid is filtered through a membrane to obtain a clear liquid, i.e., deammoniation absorption lean liquid, and a concentrated liquid, i.e., a waste salt solution. The membrane specification is an organic membrane with a molecular weight cutoff of 250 Daltons, an operating pressure of 2500 kPa, an operating temperature of 45°C, and a concentrated liquid discharge volume of 3% of the membrane feed volume. The obtained membrane filtration clear liquid has a desalination rate of 96.5% in terms of thiosulfate. The desalted membrane filtration clear liquid, i.e., deammoniation absorption lean liquid, is returned to the ammonia absorption tower for continued reuse.

[0102] The volume content of ammonia in the deammoniation purified dry gas obtained in this embodiment is 0.0005%, which is 1.8762 percentage points lower than that of the dry gas raw material. The content of hydrogen sulfide in the deammoniation purified dry gas is 1.5448%, which is 2.3458 percentage points lower than that of the dry gas raw material. By matching the desulfurization and deammoniation processes, the use of desulfurizers and deammonifiers is significantly reduced.

[0103] [Example 4]

[0104] The present embodiment is different from the embodiment 2 only in the raw material dry gas used. The dry gas composition, by volume ratio, includes: methane 16.9039%, ethane 9.1338%, ethylene 19.4617%, propane 0.1655%, propylene 0.7914%, isobutane 0.0665%, normal butane 0.0171%, trans-butene 0.0069%, normal butene 0.0146%, isobutene 0.0339%, cis-butene 0.0002%, oxygen 0.1010%, nitrogen 1.0200%, hydrogen 26.9765%, carbon monoxide 9.9961%, carbon dioxide 2.8702%, acetylene 0.0036%, 1,3-butadiene 0.0001%, C5+ alkane or olefin 1.3212%, hydrogen sulfide 7.3706%, ammonia 3.1567%, and other components 0.5819%.

[0105] The ammonia content in the deammoniation purified dry gas obtained in this embodiment is 0.0004%, which is 3.1563 percentage points lower than that of the dry gas raw material. The hydrogen sulfide content in the deammoniation purified dry gas is 1.0572%, which is 6.3134 percentage points lower than that of the dry gas raw material. By matching the desulfurization and deammoniation processes, the use of desulfurizers and deammonifiers is greatly reduced.

[0106] [Example 5]

[0107] The present embodiment is different from the embodiment 2 only in the raw material dry gas used. The dry gas composition, by volume ratio, includes: ethane 10.0323%, ethylene 21.3761%, propane 0.1817%, propylene 0.8692%, isobutane 0.0731%, normal butane 0.0188%, trans-butene 0.0076%, normal butene 0.0160%, isobutene 0.0372%, cis-butene 0.0002%, oxygen 0.1109%, nitrogen 1.1204%, hydrogen 29.6302%, carbon monoxide 9.9473%, carbon dioxide 3.1526%, acetylene 0.0039%, 1,3-butadiene 0.0001%, alkanes or alkenes with carbon 5 or more than 1.4511%, hydrogen sulfide 2.1380%, ammonia 0.6201%, and other components 0.6464%.

[0108] The ammonia content in the deammoniation purified dry gas obtained in this embodiment is 0.000063%, which is 0.6200 percentage points lower than that of the dry gas raw material. The hydrogen sulfide content in the deammoniation purified dry gas is 0.8977%, which is 1.2403 percentage points lower than that of the dry gas raw material. By matching the desulfurization and deammoniation processes, the use of desulfurizers and deammonifiers is greatly reduced.

[0109] [Comparative Example 1]

[0110] This comparative example uses a desulfurization tower to remove hydrogen sulfide alone, and a water scrubber to remove ammonia alone. The desulfurization tower is a plate tower with a theoretical plate number of 16, the desulfurizer is dimethylethanolamine, and the gas-liquid volume ratio is 80. The water scrubber is a packed tower with a theoretical plate number of 18, the deamination agent is a sulfuric acid aqueous solution with a mass concentration of 0.5%, and the gas-liquid volume ratio is 60. The feed amount of raw dry gas is 30000Nm 3 / hour, when the volume content of hydrogen sulfide is 3.8906% and the volume content of ammonia is 1.8767%, the consumption of sulfuric acid is about 1.2 tons / hour and the consumption of dimethylethanolamine is 4.6 tons / hour.

[0111] In contrast, the hydrogen sulfide in the raw dry gas is used to absorb the ammonia in the raw dry gas. Taking Example 2 as an example, the ammonia content in the deammoniation purified dry gas directly meets the requirements of the subsequent process, and no separate deammoniation is required; the consumption of dimethylethanolamine is reduced to about 160 kg / hour.

[0112] The above describes the specific implementation of the present invention in detail, but the present invention is not limited thereto. Within the technical concept of the present invention, the technical solution of the present invention can be subjected to a variety of simple modifications, including the various technical features being combined in any other suitable manner, and these simple modifications and combinations should also be regarded as the contents disclosed by the present invention and belong to the protection scope of the present invention.

Claims

1. A method for removing hydrogen sulfide and ammonia from dry gas, comprising: S1, the raw dry gas contacts the desulfurization absorption lean liquid in the hydrogen sulfide absorption tower, the desulfurization purified dry gas is obtained at the top of the tower, and the desulfurization absorption rich liquid is obtained at the bottom of the tower; S2, the desulfurized purified dry gas is contacted with the deammoniation absorption lean liquid in an ammonia absorption tower, and the deammoniation purified dry gas is obtained at the top of the tower, and the deammoniation absorption rich liquid is obtained at the bottom of the tower; S3, the desulfurization absorption rich liquid and the deammoniation absorption rich liquid and the oxygen-containing gas enter the catalytic oxidation reactor, contact with the catalytic oxidation catalyst to react, and obtain a thiosulfate emulsion; S4, the thiosulfate emulsion is subjected to gas-liquid separation, and the obtained liquid phase is decanted and layered to obtain an upper oil phase, i.e., the regenerated desulfurization absorption lean liquid, and a lower water phase, i.e., the crude deammoniation absorption lean liquid; S5, removing thiosulfate from the crude deammonification absorption lean liquid to obtain a clear liquid, i.e., the regenerated deammonification absorption lean liquid, and a concentrated liquid, i.e., a waste salt solution; The catalytic oxidation catalyst in step S3 includes zinc oxide modified by resin and silanization agent; In the catalytic oxidation catalyst of step S3, the resin is a polymerized white ball, which is copolymerized by methyl methacrylate and ethylene glycol dimethacrylate, and the amount of pendant double bonds in the copolymer is 4 to 10 mmol / g.

2. The method according to claim 1, characterized in that In the catalytic oxidation catalyst of step S3, the specific surface area of ​​the zinc oxide is 50 to 70 m 2 / g, pore volume is 0.25~0.4cm 3 / g, the average pore diameter is 10 to 20 angstroms, and the particle size is 10 to 20 nanometers.

3. The method according to claim 1, characterized in that In the catalytic oxidation catalyst described in step S3, in the zinc oxide modified by the silanization agent, the silanization agent is selected from at least one of N-(3-triethoxysilylpropyl)ethylenediamine, 3-aminopropyltriethoxysilane, N-methyl-3-(trimethoxysilyl)propylamine, and 11-aminoundecyltriethoxysilane.

4. The method according to claim 1, characterized in that In the catalytic oxidation catalyst of step S3, based on the mass of zinc oxide modified by the silanization agent, the mass content of the silanization agent is 70% to 83%, and the mass content of zinc oxide is 17% to 30%; And / or, based on the weight of the catalyst, the mass content of the resin is 68% to 89%, and the mass content of the zinc oxide modified by the silanization agent is 11% to 32%.

5. The method according to any one of claims 1 to 4, characterized in that: The method for preparing the catalytic oxidation catalyst in step S3 comprises: a. preparing zinc oxide modified by a silanization agent; b. Treating the resin with zinc oxide modified with the silanization agent obtained in step a, and drying the resin to obtain the catalyst.

6. The method according to claim 5, characterized in that In step a, the method for preparing zinc oxide modified by a silanization agent comprises: mixing a silanization agent modification solution with zinc oxide, heating and refluxing to obtain the zinc oxide modified by the silanization agent; the preparation method of the silanization agent modification solution comprises: by mass fraction, silanization agent: low carbon alcohol: water: alkaline catalyst = (200-250): (200-300): (15-23): (0.6-1.3), mixing the silanization agent, low carbon alcohol, water and alkaline catalyst at 50-70° C. for reaction for 2-5 hours to obtain the silanization agent modification solution; the mass ratio of the silanization agent modification solution to zinc oxide is (250-450): (100-200); the heating and refluxing reaction conditions are as follows: refluxing at 110-130° C. for reaction for 4-7 hours.

7. The method according to claim 6, characterized in that In step a, the low-carbon alcohol is selected from at least one of methanol, ethanol, propanol and isopropanol.

8. The method according to claim 6, characterized in that In step a, the alkaline catalyst is selected from at least one of sodium hydroxide and potassium hydroxide.

9. The method according to claim 5, characterized in that The resin described in step b is prepared by the following method, including: preparing an aqueous phase and an oil phase respectively; adding the oil phase to the aqueous phase for suspension polymerization to obtain a resin, i.e., a polymerized white ball; the aqueous phase comprises polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile, and the weight ratio of polyvinyl alcohol, sodium chloride, desalted water, and azobisisobutyronitrile is (1-3): (4-9): (80-100): (0.5-1.2); the oil phase comprises methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and a solvent, and the solvent is selected from at least one of toluene, p-xylene, m-xylene, and o-xylene, and the weight ratio of methyl methacrylate: ethylene glycol dimethacrylate: sorbitan oleate: solvent is (65-85): (66-87): (7-10): (120-160).

10. The method according to claim 9, characterized in that In step b, the molecular weight of the polyvinyl alcohol is 1500-2500.

11. The method according to claim 9, characterized in that In step b, the process of adding the oil phase to the aqueous phase for suspension polymerization includes: first cooling the aqueous phase to 18-30° C., then dropping the oil phase into the aqueous phase, controlling the stirring speed of the aqueous phase to control the oil phase droplet size to 0.6-1.5 mm, and after the dropwise addition is completed, heating to 70-90° C. and reacting for 5-10 hours.

12. The method according to claim 5, characterized in that In step b, polymer white ball, triethylene glycol di(2-methacrylate) ester and solvent, wherein the solvent is selected from at least one of toluene, p-xylene, m-xylene and o-xylene, in the ratio of polymer white ball: triethylene glycol di(2-methacrylate) ester: solvent = 50: (2-5): (150-200) by weight, swell at 60-80° C. for 6-10 hours to obtain system I; zinc oxide modified with silanization agent obtained in step a, diacyl peroxide and N,N-dimethylformamide, in the ratio of zinc oxide modified with silanization agent: diacyl peroxide: N,N-dimethylformamide = 20: (0.3-0.8): (160-230) by weight, stir at 70-90° C. for 6-10 hours to obtain system II; system II is added dropwise to system I, in the ratio of system II: system I = 1-3 by volume, the temperature is raised to 100-120° C., and refluxed for 3-6 hours.

13. The method according to claim 12, characterized in that The drying conditions in step b are as follows: the drying temperature is 100-120° C. and the drying time is 1-5 hours.

14. The method according to claim 1, characterized in that The desulfurization absorption lean liquid obtained in step S4 is recycled in step S1; and / or, the deammoniation absorption lean liquid obtained in step S5 is recycled in step S2; And / or, in step S3, the desulfurization absorption rich liquid and the deamination absorption rich liquid are first mixed, and after heat exchange with the reaction product of step S3, enter the catalytic oxidation reactor; And / or, in step S3, the oxygen-containing gas is preheated before entering the catalytic oxidation reactor.

15. The method according to claim 1, characterized in that In step S1, the volume content of hydrogen sulfide in the raw dry gas is not less than 10000 ppm; and / or the volume content of ammonia in the raw dry gas is not less than 6000 ppm.

16. The method according to claim 15, characterized in that In step S1, the volume content of hydrogen sulfide in the raw dry gas is not less than 70000 ppm; and / or the volume content of ammonia in the raw dry gas is not less than 30000 ppm.

17. The method according to claim 15, characterized in that In step S1, the volume content of hydrogen sulfide in the raw dry gas is not higher than 100,000 ppm; and / or the volume content of ammonia in the raw dry gas is not higher than 50,000 ppm.

18. The method according to claim 1, characterized in that In step S1, the hydrogen sulfide absorption tower is a plate tower or a packed tower; the desulfurization absorption lean liquid and the raw material dry gas are in countercurrent contact in the tower body.

19. The method according to claim 18, characterized in that In step S1, the desulfurization absorption lean liquid includes 120# solvent oil after light removal and p-xylene. Calculated by mass percentage, the mass content of 120# solvent oil after light removal is 75% to 86%, and the mass content of p-xylene is 10% to 20%; the usage ratio of raw dry gas to desulfurization absorption lean liquid is 80 to 200 by volume.

20. The method according to claim 18, characterized in that In step S1, the hydrogen sulfide absorption tower has 2 to 7 theoretical plates, an operating temperature of 8 to 40°C, and an operating pressure of 0.5 to 1.2 MPaG.

21. The method according to claim 1 or 18, characterized in that In step S2, the ammonia absorption tower is a plate tower or a packed tower; the deammoniation absorption lean liquid and the desulfurization purified dry gas are in countercurrent contact in the tower body.

22. The method according to claim 21, characterized in that In step S2, the ammonia absorption tower has 3 to 9 theoretical plates, an operating temperature of 10 to 40°C, and an operating pressure of 0.5 to 1.2 MPaG.

23. The method according to claim 21, characterized in that The deamination absorption lean liquid is water; the usage ratio of the desulfurization purified dry gas to the deamination absorption lean liquid is 60 to 90 by volume.

24. The method according to claim 1, characterized in that In step S3, the molar ratio of ammonia in the deamination absorption rich liquid to hydrogen sulfide in the desulfurization absorption rich liquid is controlled to be (0.8-1.8):

1.

25. The method according to claim 24, characterized in that In step S3, the molar ratio of the amount of oxygen-containing gas used, calculated as O2, to ammonia in the deammoniation absorption rich liquid is (100-200):

1.

26. The method according to claim 24, characterized in that In step S3, the reaction conditions of the reaction are as follows: the reaction temperature is 90-120°C, the reaction pressure is 0.5-1.0 MPaG, and the reaction volume space velocity is 0.5-1.0 h -1 .

27. The method according to claim 1, characterized in that In step S4, the gas-liquid separation temperature of the thiosulfate emulsion is 30-40°C, the filtration accuracy is 3-8 microns, and the liquid-liquid stratification temperature is 30-40°C; the upper oil phase obtained by liquid-liquid stratification is the desulfurization absorption lean liquid; And / or, in step S5, the method for removing thiosulfate is membrane filtration, the membrane specification is an organic membrane, the molecular weight cutoff is 100-300 Daltons, the operating pressure is 2000-4000 kPa, the operating temperature is 30-50°C, the concentrated liquid discharge is 2%-6% of the membrane feed volume, and the obtained membrane filtration clear liquid, calculated as thiosulfate, has a desalination rate of 96%-99%.

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