A catalytic oxidation catalyst, its preparation method and application

Abstract

The present invention discloses a catalytic oxidation catalyst, a preparation method thereof, and an application thereof in the combined removal of hydrogen sulfide and ammonia in a gas. The catalyst comprises a resin and zinc oxide modified by a silanizing reagent, and the preparation method comprises: a. preparing zinc oxide modified by a silanizing reagent; b. treating the resin with the zinc oxide modified by the silanizing reagent obtained in step a, and drying to obtain the catalyst. The catalyst of the present invention enables effective contact and timely conversion of hydrogen sulfide and ammonia. When used for removing hydrogen sulfide and ammonia in a gas, it can reduce the use of subsequent desulfurizing agents or deammoniating agents, which is beneficial to reducing the production cost of the device.

Description

Technical Field

[0001] The present invention relates to the technical field of catalyst preparation, and relates to a catalytic oxidation catalyst and its preparation method, as well as its application in removing hydrogen sulfide and ammonia from gas. Background Art

[0002] Refinery gas is various gases in the by-products of the petroleum processing process, including: thermal cracking gas, catalytic cracking gas, coking gas, reforming gas, etc. Coke oven gas is a by-product of coking production. In the coke oven, the coal is heated to a high temperature in an airtight environment to produce gas by coal dry distillation. Producer gas is made from coal or coke as raw materials and gasified in a gas producer. According to different gasifying agents, it can be divided into: air gas, air-steam gas, water gas, double water gas, semi-water gas, etc. Heavy oil cracking gas is made from heavy oil as raw material by partial oxidation to produce water gas or thermal cracking biogas. Anaerobic fermentation gas of organic matter includes biogas generated by fermentation and digestion of garbage, etc. During the generation process of such gases, due to the decomposition of nitrogen and sulfur components in petroleum, coal, and proteins in organic matter, a certain amount of hydrogen sulfide and trace amounts of ammonia are contained in the above gases.

[0003] In subsequent deep processing sections, various catalysts generally use acidic molecular sieves. Due to the high content of sulfur-containing impurities in the gas, because of its strong polarity, it is easily adsorbed on the active centers on the catalyst surface, causing the active centers of the catalyst to be covered; ammonia is a basic impurity, which will cause the acidic active centers of the catalyst to be poisoned and inactivated. Both will cause a significant shortening of the catalyst operation cycle. For hydrogen sulfide and ammonia in the gas, they are generally removed separately. A hydrogen sulfide removal tower is used to remove hydrogen sulfide, and a water washing tower is used to remove ammonia. For example, "Summary of the Intermediate Test on the Selective Removal of H2S by Aqueous Methyl Diethanolamine Solution under Pressure" (Chemical Engineering of Oil & Gas, No. 2, 1987) discloses a method for removing hydrogen sulfide-containing natural gas by using an aqueous methyl diethanolamine solution in an absorption tower, and the hydrogen sulfide content in the purified gas is reduced to 10mg / m 3 ."Industrial Application of the Removal of H2S by the Hypergravity Complex Iron Method" (Chemical Industry and Engineering Technology, Vol. 35, No. 2, 2014) discloses a method for removing hydrogen sulfide-containing associated gas from oilfields by using complex iron in a hypergravity machine, and the hydrogen sulfide concentration in the associated gas after desulfurization is less than 20mg / m 2 . CN109574780A discloses a method for adsorbing ammonia in dry gas by using modified activated carbon. CN111450662A discloses a combined ammonia removal process of a water washing tower - cyclone separator - coalescer - adsorption tower, the washing liquid is selected from industrial water or acidic aqueous solution, and the adsorbent is formed activated carbon, silica, alumina, formed molecular sieve, strongly acidic cation exchange resin, etc. loaded with inorganic acid or organic acid.

[0004] In summary, the existing gas desulfurization and denitrification technologies mainly remove hydrogen sulfide and ammonia impurities separately, resulting in excessive consumption of desulfurization agents or denitrification agents during the removal of hydrogen sulfide or ammonia. Summary of the Invention

[0005] The present invention provides a catalytic oxidation catalyst, a preparation method thereof, and an application in the combined removal of hydrogen sulfide and ammonia in gas. The catalyst of the present invention enables effective contact and timely conversion of hydrogen sulfide and ammonia. When used for removing hydrogen sulfide and ammonia in gas, it can reduce the use of subsequent desulfurization agents or denitrification agents, which is beneficial for the device to reduce production costs.

[0006] In a first aspect of the present invention, a catalytic oxidation catalyst is provided, wherein the catalyst comprises a resin and zinc oxide modified with a silanizing agent.

[0007] In the above technical solution, the resin is a polymeric white ball, copolymerized from methyl methacrylate and ethylene glycol dimethacrylate, and the amount of pendant double bonds in the copolymer is 4 - 10 mmol / g.

[0008] In the above technical solution, the specific surface area of the zinc oxide is 50 - 70 m 2 / g, the pore volume is 0.25 - 0.4 cm 3 / g, the average pore diameter is 10 - 20 Å, and the particle size is 10 - 20 nm.

[0009] In the above technical solution, in the zinc oxide modified with a silanizing agent, the silanizing agent is selected from at least one of N-(3-triethoxysilylpropyl)ethylenediamine, 3-aminopropyltriethoxysilane, N-methyl-3-(trimethoxysilyl)propylamine, 11-aminoundecyltriethoxysilane. Based on the mass of the zinc oxide modified with the silanizing agent, the mass content of the silanizing agent is 70% - 83%, and the mass content of the zinc oxide is 17% - 30%.

[0010] In the above technical solution, based on the weight of the catalyst, the mass content of the resin is 68% - 89%, and the mass content of the zinc oxide modified with the silanizing agent is 11% - 32%.

[0011] In the above technical solution, the specific surface area of the catalyst is 300 - 600 m 2 / g, the average pore diameter is 30 - 40 nm, and the particle size of at least 95% of the particles is 0.6 mm - 1.3 mm.

[0012] In a second aspect of the present invention, a preparation method of the above catalytic oxidation catalyst is provided, including:

[0013] a. Prepare zinc oxide modified with a silanizing agent;

[0014] b. Treat the resin with the zinc oxide modified by the silanizing reagent obtained in step a, and dry it to obtain the catalyst.

[0015] In the above technical solution, in step a, the method for preparing the zinc oxide modified by the silanizing reagent includes: mixing the zinc oxide with the silanizing reagent modification liquid and heating under reflux to obtain the zinc oxide modified by the silanizing reagent. Among them, the properties of the zinc oxide are as follows: the specific surface area is 50 - 70 m 2 / g, the pore volume is 0.25 - 0.4 cm 3 / g, the average pore diameter is 10 - 20 Å, and the particle size is 10 - 20 nm. The preparation method of the silanizing reagent modification liquid includes: by mass fraction, silanizing reagent: lower alcohol: water: base catalyst = (200 - 250):(200 - 300):(15 - 23):(0.6 - 1.3). Mix the silanizing reagent, lower alcohol (the lower alcohol is preferably at least one selected from methanol, ethanol, propanol, and isopropanol), water, and base catalyst (the base catalyst is preferably at least one selected from sodium hydroxide and potassium hydroxide) and react at 50 - 70 °C for 2 - 5 hours to obtain the silanizing reagent modification liquid. The mass ratio of mixing the silanizing reagent modification liquid with the zinc oxide is (250 - 450):(100 - 200). The conditions for the heating under reflux reaction are as follows: reflux reaction at 110 - 130 °C for 4 - 7 hours.

[0016] In the above technical solution, the resin in step b is preferably prepared by the following method, including: separately preparing an aqueous phase and an oil phase; adding the oil phase to the aqueous phase for suspension polymerization to obtain the resin, i.e., the polymerized white balls. The aqueous phase includes polyvinyl alcohol (molecular weight 1500 - 2500), sodium chloride, demineralized water, and azobisisobutyronitrile. By mass fraction, polyvinyl alcohol:sodium chloride:demineralized water: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. By mass fraction, 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 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 make the droplet size of the oil phase control within 0.6 - 1.5 mm. After the dropping is completed, raise the temperature to 70 - 90 °C and react for 5 - 10 hours. Then successively extract with absolute ethanol at 50 - 70 °C for 4 - 8 times with equal volume and extract with demineralized water at 50 - 70 °C for 4 - 8 times with equal volume, and dry (the drying temperature is 110 - 150 °C, and the drying time is 1 - 4 h) to obtain the resin, i.e., the polymerized white balls.

[0017] In the above technical solution, in step b, the polymerized white balls are combined with di(2-methylacrylic acid) triethylene glycol ester and a solvent. The solvent is selected from at least one of toluene, p-xylene, m-xylene, and o-xylene. By mass, the polymerized white balls: di(2-methylacrylic acid) triethylene glycol ester: solvent = 50: (2 - 5): (150 - 200). They are swollen at 60 - 80 °C for 6 - 10 hours to obtain System I. The zinc oxide modified by the silanization reagent, diacyl peroxide, and N,N-dimethylformamide obtained in step a, by mass, the zinc oxide modified by the silanization reagent: diacyl peroxide: N,N-dimethylformamide = 20: (0.3 - 0.8): (160 - 230). They are stirred at 70 - 90 °C for 6 - 10 hours to obtain System II. System II is added dropwise to System I. By volume ratio, System II: System I = 1 - 3. The temperature is raised to 100 - 120 °C and refluxed for 3 - 6 hours. Then, equal-volume extraction is performed 4 - 8 times with absolute ethanol at 50 - 70 °C and 4 - 8 times with demineralized water at 50 - 70 °C in sequence, and then dried (the drying temperature is 100 - 120 °C, and the drying time is 1 - 5 h) to obtain the catalyst.

[0018] The third aspect of the present invention provides an application of the catalytic oxidation catalyst provided in the first aspect or the catalytic oxidation catalyst prepared by the preparation method provided in the second aspect in the combined removal of hydrogen sulfide and ammonia in a gas.

[0019] In the above technical solution, the raw material gas can be a gas containing hydrogen sulfide and ammonia, especially a gas containing hydrogen, hydrogen sulfide, and ammonia simultaneously, such as at least one of refinery gas, coke oven gas, producer gas, heavy oil cracking gas, natural gas, and organic matter anaerobic fermentation gas.

[0020] The fourth aspect of the present invention provides a method for the combined removal of hydrogen sulfide and ammonia in a gas, including: contacting the desulfurization absorption lean liquid with the raw material gas to absorb and separate hydrogen sulfide therein to obtain the desulfurization absorption rich liquid and the desulfurization purified gas; contacting the deammoniation absorption lean liquid with the desulfurization purified gas to absorb and separate ammonia therein to obtain the deammoniation absorption rich liquid and the deammoniation purified gas; then contacting the desulfurization absorption rich liquid, the deammoniation absorption rich liquid, and an oxygen-containing gas with the catalytic oxidation catalyst for reaction to obtain a thiosulfate emulsion; and separating the thiosulfate emulsion subsequently to obtain the regenerated desulfurization absorption lean liquid and the deammoniation absorption lean liquid.

[0021] In the above technical solution, the desulfurization absorption lean liquid includes 120# solvent oil after light fraction removal and p-xylene. By mass percentage, the mass content of the 120# solvent oil after light fraction removal is 75% - 86%, and the mass content of p-xylene is 10% - 20%; the deammoniation absorption lean liquid is water.

[0022] In the above technical solution, in the feed of the reaction, the molar ratio of ammonia in the ammonia stripping absorption rich liquor to hydrogen sulfide in the desulfurization absorption rich liquor is (0.8 - 1.8):1; the amount of the oxygen-containing gas used, calculated as O2, and the ammonia in the ammonia stripping absorption rich liquor have a molar ratio of (100 - 200):1.

[0023] In the above technical solution, 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 based on the mixed solution, the reaction volume space velocity is 0.5 - 1.0 h -1 .

[0024] In the above technical solution, the subsequent separation process of the thiosulfate emulsion is as follows: first, after cooling, gas-liquid separation is carried out, and the obtained liquid phase is subjected to liquid-liquid decantation separation. The upper oil phase is the regenerated desulfurization absorption lean liquor, and the lower water phase is the crude ammonia stripping absorption lean liquor. The crude ammonia stripping absorption lean liquor is separated by membrane filtration to obtain the regenerated ammonia stripping absorption lean liquor.

[0025] In the above technical solution, the method for jointly removing hydrogen sulfide and ammonia from gas is specifically as follows:

[0026] S1, the desulfurization absorption lean liquor contacts with the raw material gas in the hydrogen sulfide absorption tower. The desulfurized purified gas is obtained at the top of the tower, and the desulfurization absorption rich liquor is obtained at the bottom of the tower;

[0027] S2, the ammonia stripping absorption lean liquor contacts with the desulfurized purified gas in the ammonia absorption tower. The ammonia-stripped purified gas is obtained at the top of the tower, and the ammonia stripping absorption rich liquor is obtained at the bottom of the tower;

[0028] S3, the desulfurization absorption rich liquor, the ammonia stripping absorption rich liquor, and the oxygen-containing gas (such as air) enter the catalytic oxidation reactor, contact with the catalytic oxidation catalyst for reaction, and a thiosulfate emulsion is obtained;

[0029] S4, the thiosulfate emulsion is cooled and then subjected to gas-liquid separation. The obtained liquid phase is decanted and stratified to obtain the upper oil phase, that is, the regenerated desulfurization absorption lean liquor, and the lower water phase, that is, the crude ammonia stripping absorption lean liquor;

[0030] S5, the crude ammonia stripping absorption lean liquor is desulfurized of thiosulfate (such as by membrane filtration) to obtain a clear liquid, that is, the regenerated ammonia stripping absorption lean liquor, and a concentrated solution, that is, the waste salt solution.

[0031] In the above technical solution, the desulfurization absorption lean liquor obtained in step S4 can be recycled in step S1.

[0032] In the above technical solution, the ammonia stripping absorption lean liquor obtained in step S5 can be recycled in step S2.

[0033] In the above technical solution, in step S4, before the liquid phase after gas-liquid separation undergoes decantation and stratification, solid substances can be filtered out according to the situation.

[0034] In the above technical solution, in step S1, the raw material gas comes from a gas containing hydrogen sulfide and ammonia, especially a gas containing a reducing gas (such as at least one of hydrogen, carbon monoxide, etc.), hydrogen sulfide and ammonia, such as at least one of refinery gas, coke oven gas, producer gas, heavy oil cracking gas, natural gas, anaerobic fermentation gas of organic matter, etc. The raw material gas may include, but is not limited to, at least one of 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 more than five carbon atoms, and also includes hydrogen sulfide and ammonia. In the raw material gas, the volume content of the reducing gas is 5% to 50%. In the raw material gas, the volume content of hydrogen sulfide is not less than 500 ppm, preferably not less than 70,000 ppm, and further preferably not more than 100,000 ppm; the volume content of ammonia is not less than 200 ppm, preferably not less than 30,000 ppm, and further preferably not more than 50,000 ppm.

[0035] 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 lean desulfurization absorption liquid and the raw material gas are in countercurrent contact in the tower body.

[0036] In the above technical solution, in step S1, the lean desulfurization absorption liquid used includes 120# solvent oil after light component removal and p-xylene. By mass percentage, the mass content of 120# solvent oil after light component removal is 75% to 86%, and the mass content of p-xylene is 10% to 20%. The 120# solvent oil after light component removal is the 120# solvent oil after removing low-boiling light components at 40 to 90°C by a conventional distillation method. The usage ratio of the raw material gas to the lean desulfurization absorption liquid is 80 to 200 by volume.

[0037] In the above technical solution, in step S1, the number of theoretical plates of the hydrogen sulfide absorption tower is 2 to 7, the operating temperature is 8 to 40°C, and the operating pressure is 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.

[0038] In the above technical solution, in step S2, the ammonia absorption tower is a conventional plate tower or packed tower in the art. The lean ammonia absorption liquid and the desulfurized and purified gas are in countercurrent contact in the tower body.

[0039] In the above technical solution, in step S2, the number of theoretical plates 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 at the top of the ammonia absorption tower, and the separation accuracy is 3 to 10 microns. The lean liquid for ammonia absorption and removal is water (such as industrial fresh water or demineralized water). The volume ratio of the desulfurized purified gas to the lean liquid for ammonia absorption and removal is 60 to 90.

[0040] In the above technical solution, in step S3, preferably, the feed rates of the rich liquid for desulfurization absorption and the rich liquid for ammonia absorption and removal entering the catalytic oxidation reactor are controlled, that is, the molar ratio of ammonia in the rich liquid for ammonia absorption and removal to hydrogen sulfide in the rich liquid for desulfurization absorption is (0.8 to 1.8):1. The ammonia in the rich liquid for ammonia absorption and removal and the hydrogen sulfide in the rich liquid for desulfurization absorption can analyze the solute concentration offline and / or online, and manually and / or automatically control the flow ratio of the rich liquid for desulfurization absorption to the rich liquid for ammonia absorption and removal. Preferably, the feed rates of the rich liquid for desulfurization absorption and the rich liquid for ammonia absorption and removal are controlled by online analyzing the ammonia concentration in the rich liquid for ammonia absorption and removal and the hydrogen sulfide concentration in the rich liquid for desulfurization absorption.

[0041] In the above technical solution, in step S3, the molar ratio of the amount of the oxygen-containing gas used, calculated as O2, to ammonia in the rich liquid for ammonia absorption and removal is (100 to 200):1. Preferably, the addition amount of the oxygen-containing gas is controlled in combination with the flow rate of the rich liquid for ammonia absorption and removal and the ammonia solute concentration analyzed offline and / or online.

[0042] In the above technical solution, in step S3, the catalytic oxidation reactor is filled with the catalytic oxidation catalyst of the present invention.

[0043] In the above technical solution, preferably, in step S3, the rich liquid for desulfurization absorption and the rich liquid for ammonia absorption and removal can be first mixed, heat-exchanged with the reaction product of step S3, and then enter the catalytic oxidation reactor.

[0044] In the above technical solution, preferably, in step S3, the oxygen-containing gas (such as air) is preheated and then enters the catalytic oxidation reactor.

[0045] In the above technical solution, in step S3, the reaction conditions of the reaction are as follows: the reaction temperature is 90 to 120 °C, the reaction pressure is 0.5 to 1.0 MPaG, and based on the mixed solution (rich liquid for ammonia absorption and removal and rich liquid for desulfurization absorption), the reaction volume space velocity is 0.5 to 1.0 h -1 。

[0046] In the above technical solution, in step S4, for the thiosulfate emulsion, the gas-liquid separation temperature is 30 to 40 °C, the filtration accuracy is 3 to 8 microns, and the liquid-liquid separation temperature is 30 to 40 °C. The upper oil phase obtained by liquid-liquid separation, i.e., the lean desulfurization absorption liquid, is pressurized by a lean desulfurization absorption liquid circulation pump by 100 to 500 kPa and then circulated to the hydrogen sulfide absorption tower in step S1.

[0047] In the above technical solution, in step S5, for the membrane filtration, the membrane specification is an organic membrane, the molecular weight cut-off is 100 to 300 daltons, the operating pressure is 2000 to 4000 kPa, the operating temperature is 30 to 50 °C, the concentrated liquid discharge amount is 2% to 6% of the membrane feed amount by volume, and for the obtained membrane filtration clear liquid, the desalination rate is 96% to 99% in terms of thiosulfate. The membrane filtration clear liquid after desalination, i.e., the regenerated rich ammonia absorption liquid, which is also the lean ammonia absorption liquid, returns to the ammonia absorption tower in step S2 for continued reuse.

[0048] The sixth aspect of the present invention provides a separation system for jointly removing hydrogen sulfide and ammonia from a gas, comprising:

[0049] 1) A hydrogen sulfide absorption tower for absorbing hydrogen sulfide using a lean desulfurization absorption liquid to provide a sulfur source for subsequent reactions;

[0050] 2) An ammonia absorption tower for removing ammonia from the gas using a lean ammonia absorption liquid and providing a nitrogen source for subsequent reactions;

[0051] 3) A catalytic oxidation reactor for catalytically oxidizing unstable ammonium sulfide into stable thiosulfate;

[0052] 4) A separation system for removing thiosulfate to regenerate the desulfurization absorption liquid and the ammonia absorption liquid.

[0053] Compared with the prior art, the present invention has the following advantages:

[0054] 1. During the research process, the inventors of the present invention found that there is a certain amount of hydrogen sulfide and ammonia in raw material gases such as refinery gas, coke oven gas, biogas, biomass gasification gas, water gas, and landfill anaerobic digestion gas during their direct application, deep processing, or domestic combustion. The acidic impurity hydrogen sulfide to be removed and the basic impurity ammonia to be removed will react to form ammonium sulfide. However, due to the poor stability of ammonium sulfide, it will decompose back into hydrogen sulfide and ammonia again. Especially when there are a large amount of reducing gases (such as hydrogen, carbon monoxide, etc.) in the raw material gas, the stability of ammonium sulfide is even worse. In the existing process, hydrogen sulfide and ammonia in the raw material gas are generally removed separately. A desulfurization tower is used to remove hydrogen sulfide, and a water washing tower or water washing tank is used to remove ammonia. For safety reasons, it is difficult to oxidize and remove low-concentration ammonium sulfide in a reducing atmosphere. The inventors of the present invention further found that a weakly polar solvent with adjusted polarity has a high solubility for hydrogen sulfide, and ammonia can be directly absorbed by circulating water. By using the catalytic oxidation catalyst of the present invention as an interfacial catalyst, the immiscibility problem when organic solvent and water contact is solved, and a catalytic oxidation reaction occurs when hydrogen sulfide and ammonia contact under the action of the catalyst 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 gas, and the subsequent process treatment is safe; while adjusting the polarity of the desulfurization absorbent and taking into account effective phase separation with water, it is convenient to separate the mixed emulsion after catalytic oxidation again through 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 through membrane separation. At the same time, the oxidation catalyst prepared by the method of the present invention has high activity stability and long service life when used for catalytic oxidation of emulsions compared with conventional supported catalysts.

[0055] 2. The present invention solves the problem of low process matching when acidic impurity hydrogen sulfide and basic impurity ammonia are often removed separately when there is a certain amount of hydrogen sulfide and a small amount of ammonia in the existing raw material gas under the condition of high content of reducing gases. By using a preferred absorbent, ammonia and hydrogen sulfide are transferred outside the raw material gas system, and then the absorbed hydrogen sulfide and ammonia are converted into relatively stable impurities and removed from the absorption system through the catalytic oxidation catalyst of the present invention. The process is safe, reduces the use of subsequent desulfurization agents or deammoniation agents, and is beneficial to reducing the production cost of the device. BRIEF DESCRIPTION OF THE DRAWINGS

[0056] Figure 1 It is a schematic diagram of a separation system for the combined removal of hydrogen sulfide and ammonia in a gas of the present invention;

[0057] Among them, the description of the reference numerals is as follows:

[0058] 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 phase product delivery pump, 112 is a liquid phase product filter, 113 is a liquid-liquid decanter, 114 is an aqueous phase feed pump, 115 is a membrane filter, 116 is a catalytic oxidation reactor feed heater, 117 is an absorption liquid component analyzer and flow controller, 118 is a catalytic oxidation reactor outlet cooler, 119 is a catalytic oxidation reactor feed gas-liquid mixer, 120 is a membrane filter clear liquid buffer tank, 121 is a deammoniation absorption lean liquid circulation pump, 122 is a desulfurization absorption lean liquid circulation pump, 201 is a raw material gas inlet, 202 is a supplementary fresh water inlet, 203 is an air inlet, 204 is an oxidation tail gas outlet, 205 is a filter blowdown port, 206 is a membrane filter concentrate discharge port, 207 is a deammoniation purified gas discharge port. Detailed implementation manners

[0059] The technical solutions in the embodiments of the present invention will be clearly and completely described below. Obviously, the described embodiments are only a part of the embodiments of the present invention, rather than all the embodiments. All other embodiments obtained by those skilled in the art based on the embodiments of the present invention without creative efforts shall fall within the protection scope of the present invention.

[0060] In the present invention, the analysis of low-carbon hydrocarbons, hydrogen sulfide, and ammonia components is carried out using a gas chromatograph Agilent 7890A GC (Agilent, USA), equipped with an HP-PLOT Al2O3 KCl (50m×0.53mm×15μm) capillary chromatographic column. The column temperature is maintained at 100°C for 10 minutes, and then increased to 120°C at a rate of 30°C per minute and maintained for 3 minutes. The carrier gas He flow rate is 3 mL / min, the injection volume is 0.1 mL (quantitative loop), the split ratio is 5:1, the injection port temperature is 250°C, and the detector (FID) temperature is 250°C.

[0061] In the present invention, the content of zinc oxide in the resin is determined by a roasting method in an air atmosphere at a roasting temperature of 600 - 700°C to obtain the amount of zinc oxide supported by a certain amount of resin.

[0062] In the present invention, the amount of pendant double bonds in the resin is analyzed by a titration method: First, the resin is fully swollen with carbon tetrachloride, and then further soaked with pyridine, sulfuric acid, bromine in glacial acetic acid solution, carbon tetrachloride, and mercury acetate solution, and then potassium iodide solution is added and titrated with a standard sodium thiosulfate solution for calculation.

[0063] A separation system for jointly removing hydrogen sulfide and ammonia from gas according to the present invention (such as Figure 1As shown in the figure, it includes: 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 phase product transfer pump, 112 is a liquid phase product filter, 113 is a liquid-liquid decanter, 114 is an aqueous 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 outlet cooler, 119 is a catalytic oxidation reactor feed gas-liquid mixer, 120 is a membrane filter clear liquid buffer tank, 121 is a deammoniation absorption lean liquid circulation pump, and 122 is a desulfurization absorption lean liquid circulation pump. The separation process is as follows: The raw gas (such as from refinery gas, coke oven gas, producer gas, heavy oil cracking gas, natural gas, anaerobic fermentation gas of organic matter, etc.) enters the middle and lower part of the hydrogen sulfide absorption tower 101 from the raw gas inlet 201, and countercurrently contacts with the desulfurization absorption lean liquid in the tower body. The desulfurized purified gas at the top of the hydrogen sulfide absorption tower 101 enters the ammonia absorption tower 102 and countercurrently contacts with the deammoniation absorption lean liquid in the tower body. The deammoniated purified gas is discharged from the top of the ammonia absorption tower 102. After that, if the sulfur content in the deammoniated purified gas still does not meet the standard, it can enter the subsequent hydrogen sulfide removal system for further desulfurization treatment. The tower bottom rich liquid of the hydrogen sulfide absorption tower 101 (i.e., the desulfurization absorption rich liquid) and the tower bottom rich liquid of the ammonia absorption tower 102 (the deammoniation absorption rich liquid) are analyzed for the absorption liquid components by the absorption liquid component analysis and flow controller 117. According to the composition and flow rate of the deammoniation absorption rich liquid, referring to the composition of the desulfurization absorption rich liquid, the flow rate of the desulfurization liquid is calculated and controlled.The deammoniation absorption rich liquid and the desulfurization absorption rich liquid after calculation adjustment are heat-exchanged with the outlet stream of the catalytic oxidation reactor through the feed and effluent heat exchanger 107 of the catalytic oxidation reactor 109, heated to a certain temperature by the feed heater 116 of the catalytic oxidation reactor, and then mixed with the air heated by the air preheater 108 in the feed gas-liquid mixer 119 of the catalytic oxidation reactor, and then enter the catalytic oxidation reactor to react with the catalytic oxidation catalyst of the present invention. The obtained oxidation products are cooled by the feed and effluent heat exchanger 107 of the catalytic oxidation reactor and the effluent cooler 118 of the catalytic oxidation reactor in sequence 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 of the boundary from the oxidation tail gas outlet 204. The liquid phase part of the gas-liquid separation tank passes through the liquid product transfer pump 111 and the liquid product filter 112 in sequence, and then enters the liquid-liquid decanter 113. In the decanter, the oil phase and the water phase are phase-separated. The upper oil phase is recycled back to the hydrogen sulfide absorption tower 101 through the desulfurization absorption lean liquid recycle pump 122 for continued use. The lower water phase is pressurized by the water phase feed pump 114 and then 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 at a certain ratio and supplements an equal amount of fresh water through the fresh water inlet 202. The clarified liquid of the membrane filter 115, that is, the regenerated deammoniation absorption lean liquid, is buffered by the membrane filter clarified liquid buffer tank 120 and then recycled back to the ammonia absorption tower 102 through the deammoniation absorption lean liquid recycle pump 121 for continued use.

[0064]

Example 1

[0065] In this example, the raw material gas is derived from refinery gas. The gas composition is by volume ratio, including: methane 17.9901%, ethane 9.7208%, ethylene 20.7122%, propane 0.1761%, propylene 0.8422%, isobutane 0.0708%, n-butane 0.0182%, trans-butene 0.0073%, n-butene 0.0155%, isobutene 0.0361%, cis-butene 0.0002%, oxygen 0.1075%, nitrogen 1.0856%, hydrogen 28.7099%, carbon monoxide 9.6384%, carbon dioxide 3.0546%, acetylene 0.0038%, 1,3-butadiene 0.0001%, alkanes or alkenes with more than five carbons 1.4061%, hydrogen sulfide 3.8906%, ammonia 1.8767%, and other components 0.6263%. The feed rate of the raw material gas is 30000 Nm 3 / h.

[0066] The separation process of this example is the same as that shown in Figure 1 shown.

[0067] The hydrogen sulfide absorption tower is a plate column. The lean desulfurization absorption liquid and the raw material gas are in countercurrent contact in the tower body. The theoretical number of plates of the hydrogen sulfide absorption tower is 4, the operating temperature is 25 °C, and the operating pressure is 0.7 MPaG. A demister is provided at the top of the hydrogen sulfide absorption tower, and the separation accuracy is 5 microns. The lean desulfurization absorption liquid includes 120# solvent oil (boiling range 90 - 120 °C) after light component removal and p-xylene. By mass percentage, the mass content of 120# solvent oil after light component removal is 82.7%, the mass content of p-xylene is 16.8%, and the rest is water and trace amounts of hydrogen sulfide. The volume ratio of the raw material gas to the lean desulfurization absorption liquid is 140.

[0068] The ammonia absorption tower is a plate column. The lean ammonia absorption liquid and the desulfurized purified 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 °C, and the operating pressure is 0.7 MPaG. A demister is provided at the top of the ammonia absorption tower, and the separation accuracy is 5 microns. The lean ammonia absorption liquid is demineralized water. The volume ratio of the desulfurized purified gas to the lean ammonia absorption liquid is 75.

[0069] The rich desulfurization absorption liquid and the rich ammonia absorption liquid are mixed. By online analyzing the solute concentration, the flow ratio of the rich desulfurization absorption liquid to the rich ammonia absorption liquid is automatically controlled to control the molar ratio of ammonia to hydrogen sulfide to be 1.4:1 during mixing.

[0070] When the mixed liquid phase stream of the rich desulfurization absorption liquid and the rich ammonia absorption liquid is mixed with air, in combination with the flow rate of the rich ammonia absorption liquid and the online analyzed ammonia solute concentration, the molar ratio of ammonia to oxygen during the mixing of the mixed liquid phase stream and air is controlled to be 1:150.

[0071] The catalytic oxidation reactor is equipped with a catalytic oxidation catalyst, and its preparation method includes:

[0072] a. Prepare zinc oxide modified with a silanization reagent:

[0073] Among them, the properties of zinc oxide are as follows: the specific surface area is 60 m 2 / g, and the pore volume is 0.32 cm 3 / g, with an average pore diameter of 14 Å and a particle size of 15 nm. The preparation method of the silanization reagent modification solution includes: by mass, silanization reagent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 225: 250: 19: 0.9. After mixing the silanization reagent, ethanol, water and sodium hydroxide evenly, react at 60 °C for 3 hours to obtain the silanization reagent modification solution. The mass ratio of the silanization reagent modification solution to zinc oxide is 350: 150. After mixing the silanization reagent modification solution and zinc oxide evenly, reflux and react at 120 °C for 5 hours to obtain zinc oxide modified with the silanization reagent; wherein, based on the mass of zinc oxide modified with the silanization reagent, the mass content of the silanization reagent is 76%, and the mass content of zinc oxide is 24%;

[0074] b. Treat the resin with the zinc oxide modified with the silanization reagent obtained in step a, and dry to obtain the catalyst:

[0075] Among them, the resin is prepared by the following method, including: separately preparing an aqueous phase and an oil phase; adding the oil phase to the aqueous phase for suspension polymerization to obtain the resin, i.e., the polymerized white balls. The aqueous phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, demineralized water, and azobisisobutyronitrile. By mass fraction, polyvinyl alcohol:sodium chloride:demineralized water:azobisisobutyronitrile = 2:6.5:90:0.85. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene. By mass fraction, methyl methacrylate:ethylene glycol dimethacrylate:sorbitan oleate:toluene = 75:76:8.5:140. The process of adding the oil phase to the aqueous phase for suspension polymerization includes: cooling the aqueous phase to 25°C, then dropping the oil phase into the aqueous phase, controlling the stirring speed of the aqueous phase to make the droplet size of the oil phase control at 1.1 mm. After the dropping is completed, the temperature is raised to 80°C and the reaction is carried out for 7.5 hours. Then, it is successively extracted 6 times with anhydrous ethanol in equal volume at 60°C and 6 times with demineralized water in equal volume at 60°C, and dried (the drying temperature is 130°C and the drying time is 2.5 h) to obtain the polymerized white balls (the amount of pendant double bonds in the copolymer is 7 mmol / g). The polymerized white balls, triethylene glycol di(2-methylacrylate), and toluene, by mass fraction, polymerized white balls:triethylene glycol di(2-methylacrylate):toluene = 50:3.5:175, are swollen at 70°C for 8 hours to obtain System I. The zinc oxide modified with the silanizing agent obtained in step a, benzoyl peroxide, and N,N-dimethylformamide, by mass fraction, the zinc oxide modified with the silanizing agent obtained in step a:benzoyl peroxide:N,N-dimethylformamide = 20:0.5:195, are stirred at 80°C for 8 hours to obtain System II. Drop System II into System I, and by volume ratio, System II:System I = 2, raise the temperature to 110°C and carry out reflux treatment for 4.5 hours. Then, it is successively extracted 6 times with anhydrous ethanol in equal volume at 60°C and 6 times with demineralized water in equal volume at 60°C, and the drying temperature is 110°C and the drying time is 3 h to obtain the catalyst.

[0076] Based on the weight of the catalyst, the mass content of the resin is 79.5%, and the mass content of the zinc oxide modified with the silanizing agent is 20.5%.

[0077] The catalytic oxidation reactor has a reaction temperature of 105°C, a reaction pressure of 0.7 MPaG, and a reaction volume space velocity of 0.75 h based on the mixed solution. -1 For the thiosulfate emulsion at the reactor outlet, the gas-liquid separation temperature is 35°C, the filtration accuracy is 5 microns, and the liquid-liquid separation temperature is 35°C. The upper oil phase obtained by liquid-liquid separation, i.e., the lean desulfurization absorption liquid, is pressurized by 300 kPa by the lean desulfurization absorption liquid circulation pump and then circulated to the hydrogen sulfide absorption tower; the lower aqueous phase is the lean crude deammoniation absorption liquid.

[0078] The lean solution of crude ammonia absorption is subjected to membrane filtration to obtain a clear liquid, i.e., the lean solution of ammonia absorption, and a concentrated liquid, i.e., the waste salt solution. The membrane specification is an organic membrane with a molecular weight cut-off of 200 Dalton. The operating pressure is 3000 kPa, the operating temperature is 40 °C, the discharge amount of the concentrated liquid is 4% of the volume of the membrane feed. For the clear liquid obtained by membrane filtration, the desalination rate is 97.5% calculated as thiosulfate. The clear liquid obtained by membrane filtration after desalination, i.e., the lean solution of ammonia absorption, is returned to the ammonia absorption tower for continued reuse.

[0079] In the deammoniated purified gas obtained in this example, the ammonia volume content is 0.0003%. Compared with the gas raw material, the ammonia content to be removed is reduced by 1.8764 percentage points. The hydrogen sulfide content in the deammoniated purified gas is 1.2096%. Compared with the gas raw material, the hydrogen sulfide content to be removed is reduced by 2.6810 percentage points. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is significantly reduced.

[0080]

Example 2

[0081] The raw material gas described in this example is the same as that in Example 1. The separation process in this example is the same as that in Example 1.

[0082] The hydrogen sulfide absorption tower is a plate tower. The lean solution of desulfurization absorption and the raw material gas are in countercurrent contact in the tower body. The number of theoretical plates of the hydrogen sulfide absorption tower is 6, the operating temperature is 10 °C, and the operating pressure is 1.0 MPaG. A demister is provided at the top of the hydrogen sulfide absorption tower with a separation accuracy of 5 microns. The lean solution of desulfurization absorption includes 120# solvent oil (boiling range 90 - 120 °C) after light component removal and p-xylene. By mass percentage, the mass content of 120# solvent oil after light component removal is 85.2%, the mass content of p-xylene is 14.7%, and the rest is water and trace hydrogen sulfide. The volume ratio of the raw material gas to the lean solution of desulfurization absorption is 90.

[0083] The ammonia absorption tower is a plate tower. The lean solution of deammoniation absorption and the desulfurized purified gas are in countercurrent contact in the tower body. The number of theoretical plates of the ammonia absorption tower is 8, the operating temperature is 12 °C, and the operating pressure is 1.0 MPaG. A demister is provided at the top of the ammonia absorption tower with a separation accuracy of 5 microns. The lean solution of deammoniation absorption is demineralized water. The volume ratio of the desulfurized purified gas to the lean solution of deammoniation absorption is 65.

[0084] The rich solution of desulfurization absorption and the rich solution of deammoniation absorption are mixed. By online analyzing the solute concentration, the flow ratio of the rich solution of desulfurization absorption to the rich solution of deammoniation absorption is automatically controlled to control the molar ratio of ammonia to hydrogen sulfide during mixing to be 1.0:1.

[0085] When the mixed liquid stream of the rich solution of desulfurization absorption and the rich solution of deammoniation absorption is mixed with air, in combination with the flow rate of the rich solution of deammoniation absorption and the ammonia solute concentration analyzed online, the molar ratio of ammonia to oxygen during the mixing of the mixed liquid stream and air is controlled to be 1:180.

[0086] A catalytic oxidation reactor equipped with a catalytic oxidation catalyst, and its preparation method includes:

[0087] a. Prepare zinc oxide modified with a silanization reagent:

[0088] Among them, the properties of zinc oxide are as follows: specific surface area is 65 m 2 / g, pore volume is 0.38 cm 3 / g, average pore diameter is 11 Å, and particle size is 12 nm. By mass fraction, silanization reagent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 210:280:17:0.7. After mixing the silanization reagent, ethanol, water and sodium hydroxide evenly, react at 65 °C for 4 hours to obtain the silanization reagent modification solution. The mass ratio of the silanization reagent modification solution to zinc oxide is 300:180. After mixing the silanization reagent modification solution and zinc oxide evenly, reflux and react at 125 °C for 6 hours to obtain zinc oxide modified with a silanization reagent; among them, based on the mass of zinc oxide modified with a silanization reagent, the mass content of the silanization reagent is 72%, and the mass content of zinc oxide is 28%;

[0089] b. Treat the resin with the zinc oxide modified with the silanization reagent obtained in step a, and dry it to obtain the catalyst:

[0090] Among them, the resin is prepared by the following method, including: separately preparing an aqueous phase and an oil phase; adding the oil phase to the aqueous phase for suspension polymerization to obtain the resin, i.e., polymerized white balls. The aqueous phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, demineralized water, and azobisisobutyronitrile. By mass, polyvinyl alcohol:sodium chloride:demineralized water:azobisisobutyronitrile = 2.5:8.5:95:1.0. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene. By mass, methyl methacrylate:ethylene glycol dimethacrylate:sorbitan oleate:toluene = 80:81:9.0:150. The process of adding the oil phase to the aqueous phase for suspension polymerization includes: cooling the aqueous phase to 20 °C, then dropping the organic phase into the aqueous phase, controlling the stirring speed of the aqueous phase to make the droplet size of the oil phase control at 0.8 mm. After the dropping is completed, the temperature is raised to 85 °C and the reaction is carried out for 9.5 hours. Filter, and sequentially extract with anhydrous ethanol at 65 °C in equal volume for 7 times and demineralized water at 65 °C in equal volume for 7 times, and dry (drying temperature is 145 °C, drying time is 3.5 h) to obtain polymerized white balls (the amount of pendant double bonds in the copolymer is 9 mmol / g). The polymerized white balls, triethylene glycol di(2-methylacrylate), and toluene, by mass, polymerized white balls:triethylene glycol di(2-methylacrylate):toluene = 50:4.5:195, are swollen at 75 °C for 9 hours to obtain System I. The zinc oxide modified with the silanizing reagent obtained in step a, diacyl peroxide, and N,N-dimethylformamide, by mass, the zinc oxide modified with the silanizing reagent obtained in step a:diacyl peroxide:N,N-dimethylformamide = 20:0.7:210, are stirred at 75 °C for 7 hours to obtain System II. Drop System II into System I, by volume ratio, System II:System I = 2.7, raise the temperature to 105 °C, and carry out reflux treatment for 5 hours. Then sequentially extract with anhydrous ethanol at 65 °C in equal volume for 7 times and demineralized water at 65 °C in equal volume for 7 times, and the drying temperature is 115 °C and the drying time is 4 h to obtain the catalyst.

[0091] Based on the weight of the catalyst, the mass content of the resin is 71.7%, and the mass content of the zinc oxide modified with the silanizing reagent is 28.3%.

[0092] The catalytic oxidation reactor has a reaction temperature of 115 °C, a reaction pressure of 0.9 MPaG, and a reaction volume space velocity of 0.65 h based on the mixed solution. -1 . For the thiosulfate emulsion at the reactor outlet, the gas-liquid separation temperature is 32 °C, the filtration accuracy is 4 microns, and the liquid-liquid separation temperature is 32 °C. The upper oil phase obtained by liquid-liquid separation, i.e., the lean desulfurization absorption liquid, is pressurized by a lean desulfurization absorption liquid circulation pump by 400 kPa and then circulated to the hydrogen sulfide absorption tower; the lower aqueous phase is the lean crude deammoniation absorption liquid.

[0093] The lean solution of crude ammonia absorption is filtered through a membrane to obtain a clear solution, namely the lean solution of ammonia absorption, and a concentrated solution, namely the waste salt solution. The membrane specification is an organic membrane with a cut-off molecular weight of 150 Dalton. The operating pressure is 3500 kPa and the operating temperature is 35 °C. The discharge amount of the concentrated solution is 5% of the volume of the membrane feed. The obtained clear solution after membrane filtration has a desalination rate of 98.5% in terms of thiosulfate. The clear solution after membrane filtration and desalination, namely the lean solution of ammonia absorption, is returned to the ammonia absorption tower for reuse.

[0094] In the deammoniated purified gas obtained in this example, the ammonia volume content is 0.0001%, which is 1.8766 percentage points lower than that of the gas raw material in terms of the ammonia content to be removed. The hydrogen sulfide content in the deammoniated purified gas is 0.1372%, which is 3.7534 percentage points lower than that of the gas raw material in terms of the hydrogen sulfide content to be removed. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is significantly reduced.

[0095]

Example 3

[0096] The raw material gas described in this example is the same as that in Example 1. The separation process in this example is the same as that in Example 1.

[0097] The hydrogen sulfide absorption tower is a plate tower. The lean solution of desulfurization absorption contacts the raw material gas countercurrently in the tower body. The number of theoretical plates of the hydrogen sulfide absorption tower is 3, the operating temperature is 35 °C, and the operating pressure is 0.6 MPaG. A demister is provided at the top of the hydrogen sulfide absorption tower with a separation accuracy of 4 microns. The lean solution of desulfurization absorption includes the 120# solvent oil (boiling range 90 - 120 °C) after light component removal and p-xylene. By mass percentage, the mass content of the 120# solvent oil after light component removal is 80.9%, and the mass content of p-xylene is 18.7%. The rest is water and trace amounts of hydrogen sulfide. The volume ratio of the raw material gas to the lean solution of desulfurization absorption is 180.

[0098] The ammonia absorption tower is a plate tower. The lean solution of deammoniation absorption contacts the desulfurized purified gas countercurrently in the tower body. The number of theoretical plates of the ammonia absorption tower is 4, the operating temperature is 35 °C, and the operating pressure is 0.6 MPaG. Preferably, a demister is provided at the top of the ammonia absorption tower with a separation accuracy of 4 microns. The lean solution of deammoniation absorption is demineralized water. The volume ratio of the desulfurized purified gas to the lean solution of deammoniation absorption is 80.

[0099] The rich solution of desulfurization absorption and the rich solution of deammoniation absorption are mixed. By online analyzing the solute concentration, the flow ratio of the rich solution of desulfurization absorption to the rich solution of deammoniation absorption is automatically controlled to control the molar ratio of ammonia to hydrogen sulfide to be 1.6:1 during mixing.

[0100] When the mixed liquid stream of the rich solution of desulfurization absorption and the rich solution of deammoniation absorption is mixed with air, considering the flow rate of the rich solution of deammoniation absorption and the ammonia solute concentration analyzed online, the molar ratio of ammonia to oxygen during the mixing of the mixed liquid stream and air is controlled to be 1:130.

[0101] A catalytic oxidation reactor is equipped with a catalytic oxidation catalyst, and its preparation method includes:

[0102] a. Prepare zinc oxide modified with a silanization reagent:

[0103] Among them, the properties of zinc oxide are as follows: the specific surface area is 55 m 2 / g, the pore volume is 0.29 cm 3 / g, the average pore diameter is 17 Å, and the particle size is 18 nm. The preparation method of the silanization reagent modification solution includes: by mass, silanization reagent (3-aminopropyltriethoxysilane): ethanol: water: sodium hydroxide = 240:220:20:1.1. Mix the silanization reagent, ethanol, water, and sodium hydroxide evenly and react at 55 °C for 2.5 hours to obtain the silanization reagent modification solution. The mass ratio of the silanization reagent modification solution to zinc oxide for mixing is 400:120. Mix the silanization reagent modification solution and zinc oxide evenly and reflux at 115 °C for 4.5 hours to obtain zinc oxide modified with a silanization reagent. Among them, based on the mass of zinc oxide modified with a silanization reagent, the mass content of the silanization reagent is 79%, and the mass content of zinc oxide is 21%;

[0104] b. Treat the resin with the zinc oxide modified with the silanization reagent obtained in step a and dry it to obtain the catalyst:

[0105] Among them, the resin 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 the resin, i.e., the polymerized white balls. The aqueous phase includes polyvinyl alcohol (molecular weight 2000), sodium chloride, demineralized water, and azobisisobutyronitrile. By mass fraction, polyvinyl alcohol:sodium chloride:demineralized water:azobisisobutyronitrile = 1.5:4.5:85:0.7. The oil phase includes methyl methacrylate, ethylene glycol dimethacrylate, sorbitan oleate, and toluene. By mass fraction, methyl methacrylate:ethylene glycol dimethacrylate:sorbitan oleate:toluene = 70:71:8.0:130. The process of adding the oil phase to the aqueous phase for suspension polymerization includes: cooling the aqueous phase to 28 °C, then dropping the organic phase into the aqueous phase, controlling the stirring speed of the aqueous phase to make the droplet size of the oil phase controlled at 1.3 mm. After the dropping is completed, the temperature is raised to 75 °C and the reaction is carried out for 6.0 hours. Filter, successively extract 5 times with anhydrous ethanol at 55 °C in equal volume and extract 5 times with demineralized water at 55 °C in equal volume, and dry (the drying temperature is 116 °C and the drying time is 2 h) to obtain the polymerized white balls (the amount of pendant double bonds in the copolymer is 5 mmol / g). The polymerized white balls, triethylene glycol di(2-methylacrylate), and toluene, by mass fraction, polymerized white balls:triethylene glycol di(2-methylacrylate):toluene = 50:2.5:160, are swollen at 65 °C for 7 hours to obtain System I. The zinc oxide modified with the silanizing reagent obtained in step a, dibenzoyl peroxide, and N,N-dimethylformamide, by mass fraction, zinc oxide modified with the silanizing reagent obtained in step a:dibenzoyl peroxide:N,N-dimethylformamide = 20:0.4:170, are stirred at 85 °C for 9 hours to obtain System II. Drop System II into System I, by volume ratio, System II:System I = 1.2, raise the temperature to 115 °C, and carry out reflux treatment for 3.5 hours. Then successively extract 5 times with anhydrous ethanol at 55 °C in equal volume and extract 5 times with demineralized water at 55 °C in equal volume, the drying temperature is 106 °C, and the drying time is 1.3 h to obtain the catalyst.

[0106] Based on the weight of the catalyst, the mass content of the resin is 86.6%, and the mass content of the zinc oxide modified with the silanizing reagent is 13.4%.

[0107] The catalytic oxidation reactor has a reaction temperature of 95 °C, a reaction pressure of 0.6 MPaG, and a reaction volume space velocity of 0.9 h based on the mixed solution. -1 For the thiosulfate emulsion at the reactor outlet, the gas-liquid separation temperature is 38 °C, the filtration accuracy is 6 microns, and the liquid-liquid separation temperature is 38 °C. The upper oil phase obtained by liquid-liquid separation, i.e., the lean desulfurization absorption liquid, is pressurized by 200 kPa by the lean desulfurization absorption liquid circulation pump and then circulated to the hydrogen sulfide absorption tower; the lower aqueous phase is the lean crude deammoniation absorption liquid.

[0108] The lean liquid of rough deammoniation absorption is subjected to membrane filtration to obtain a clear liquid, i.e., the lean liquid of deammoniation absorption, and a concentrated liquid, i.e., a waste salt solution. The membrane specification is an organic membrane with a cut-off molecular weight of 250 Dalton, an operating pressure of 2500 kPa, an operating temperature of 45 °C, and the discharge amount of the concentrated liquid is 3% of the volume of the membrane feed. The obtained clear liquid after membrane filtration has a desalination rate of 96.5% calculated as thiosulfate. The clear liquid after membrane filtration and desalination, i.e., the lean liquid of deammoniation absorption, is returned to the ammonia absorption tower for continued reuse.

[0109] In the deammoniated purified gas obtained in this example, the ammonia volume content is 0.0005%, and the ammonia content to be removed is reduced by 1.8762 percentage points relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 1.5448%, and the hydrogen sulfide content to be removed is reduced by 2.3458 percentage points relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is significantly reduced.

[0110]

Example 4

[0111] This example is compared with Example 2 only in that the source of the refinery gas as the raw material gas used is different. Among them, the gas composition is by volume ratio, including: methane 16.9039%, ethane 9.1338%, ethylene 19.4617%, propane 0.1655%, propylene 0.7914%, isobutane 0.0665%, n-butane 0.0171%, trans-butene 0.0069%, n-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%, paraffins or olefins with more than five carbons 1.3212%, hydrogen sulfide 7.3706%, ammonia 3.1567%, and other components 0.5819%.

[0112] In the deammoniated purified gas obtained in this example, the ammonia content is 0.0004%, and the ammonia content to be removed is reduced by 3.1563 percentage points relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 1.0572%, and the hydrogen sulfide content to be removed is reduced by 6.3134 percentage points relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0113]

Example 5

[0114] This example is only different from Example 2 in that the refinery gas as the raw material gas used is different. Among them, the gas composition is by volume percentage, including: ethane 10.0323%, ethylene 21.3761%, propane 0.1817%, propylene 0.8692%, isobutane 0.0731%, n-butane 0.0188%, trans-butene 0.0076%, n-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%, paraffins or olefins with more than five carbons 1.4511%, hydrogen sulfide 2.1380%, ammonia 0.6201%, and other components 0.6464%.

[0115] The ammonia content in the deammoniated purified gas obtained in this example is 0.000063%, and the ammonia content to be removed is reduced by 0.6200 percentage points relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 0.8977%, and the hydrogen sulfide content to be removed is reduced by 1.2403 percentage points relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0116]

Example 6

[0117] This example is only different from Example 2 in that the raw material gas used is different. The raw material gas described in this example is from coke oven gas, and the composition is by volume percentage, including:

[0118] carbon monoxide 5.3%, hydrogen 56.1%, methane 22.6%, (C2-C5) hydrocarbons 2.2%, oxygen 0.9%, nitrogen 8.4%, carbon dioxide 1.8%, ammonia 0.0430%, hydrogen sulfide 0.1115%, and other components 2.5455%.

[0119] The ammonia volume content in the deammoniated purified gas obtained in this example is 0.0001%, and the ammonia content to be removed is reduced by 0.0429% relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 0.0257%, and the hydrogen sulfide content to be removed is reduced by 0.0858% relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0120]

Example 7

[0121] This example is only different from Example 2 in that the raw material gas used is different. The raw material gas described in this example is from producer gas, and the composition is by volume percentage, including:

[0122] Carbon monoxide 26.2%, hydrogen 14.3%, methane 1.4%, (C2-C5) hydrocarbons 0.3%, oxygen 0.2%, nitrogen 49.5%, carbon dioxide 4.6%, ammonia 0.0620%, hydrogen sulfide 0.1731%, other components 3.2649%.

[0123] In the deammoniated purified gas obtained in this example, the volume content of ammonia is 0.00011%, and the ammonia content to be removed is reduced by 0.0619% relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 0.0493%, and the hydrogen sulfide content to be removed is reduced by 0.1238% relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0124]

Example 8

[0125] This example is only different from Example 2 in that the raw material gas used is different. The raw material gas described in this example is from heavy oil cracking gas, and its composition by volume ratio includes:

[0126] Carbon monoxide 1.5%, hydrogen 18.7%, methane 39.8%, (C2-C5) hydrocarbons 35.3%, oxygen 0.3%, nitrogen 2.1%, carbon dioxide 0.4%, ammonia 0.0710%, hydrogen sulfide 0.1842%, other components 1.6448%.

[0127] In the deammoniated purified gas obtained in this example, the volume content of ammonia is 0.00014%, and the ammonia content to be removed is reduced by 0.0709% relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 0.0425%, and the hydrogen sulfide content to be removed is reduced by 0.1417% relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0128]

Example 9

[0129] This example is only different from Example 2 in that the raw material gas used is different. The raw material gas described in this example is from natural gas, and its composition by volume ratio includes:

[0130] Methane 71.8%, (C2-C5) hydrocarbons 15.2%, nitrogen 7.9%, carbon dioxide 2.3%, ammonia 0.0650%, hydrogen sulfide 0.1664%, other components 2.5686%.

[0131] In the deammoniated purified gas obtained in this example, the volume content of ammonia is 0.00013%, and the ammonia content to be removed is reduced by 0.0649% relative to the gas raw material. The hydrogen sulfide content in the deammoniated purified gas is 0.0367%, and the hydrogen sulfide content to be removed is reduced by 0.1297% relative to the gas raw material. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0132]

Example 10

[0133] This example is only different from Example 2 in that the raw material gas used is different. The raw material gas described in this example is derived from the anaerobic fermentation gas of organic matter. By volume, its composition includes:

[0134] Carbon monoxide 0.4%, methane 68.9%, (C2-C5) hydrocarbons 0.9%, nitrogen 0.8%, carbon dioxide 27.3%, ammonia 0.0380%, hydrogen sulfide 0.0940%, other components 1.5680%.

[0135] In the deammoniated purified gas obtained in this example, the volume content of ammonia is 0.00009%. Compared with the gas raw material, the ammonia content to be removed is reduced by 0.0379%. The hydrogen sulfide content in the deammoniated purified gas is 0.0182%. Compared with the gas raw material, the hydrogen sulfide content to be removed is reduced by 0.0758%. Through the matching of the desulfurization and deammoniation processes, the use of desulfurizing agent and deammoniating agent is greatly reduced.

[0136]

Comparative Example 1

[0137] In this comparative example, a desulfurization tower is used to remove hydrogen sulfide alone and a water washing tower is used to remove ammonia alone. The desulfurization tower is a plate tower with 16 theoretical plates. The desulfurizing agent is dimethylethanolamine and the gas-liquid volume ratio is 80. The water washing tower is a packed tower with 18 theoretical plates. The deammoniating agent is a sulfuric acid aqueous solution with a mass concentration of 0.5% and the gas-liquid volume ratio is 60. When the feed rate of the raw material gas is 30000 Nm 3 / h, 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 / h, and the consumption of dimethylethanolamine is 4.6 tons / h.

[0138] In contrast, when using hydrogen sulfide in the raw material gas to absorb ammonia in the raw material gas, taking Example 2 as an example, the ammonia content in the deammoniated purified gas directly meets the requirements of the subsequent process and there is no need to remove ammonia separately; the consumption of dimethylethanolamine is reduced to about 160 kg / h.

[0139] The specific embodiments of the present invention have been described in detail above. However, the present invention is not limited thereto. Within the scope of the technical concept of the present invention, various simple modifications can be made to the technical solution of the present invention, including any other suitable combination of each technical feature. These simple modifications and combinations should also be regarded as the content disclosed by the present invention and fall within the protection scope of the present invention.

Claims

1. A catalytic oxidation catalyst, wherein, The catalyst includes a resin and zinc oxide modified by a silanizing agent; the resin is a polymeric white ball copolymerized from methyl methacrylate and ethylene glycol dimethacrylate, and the amount of pendant double bonds in the copolymer is 4 - 10 mmol / g; in the zinc oxide modified by the silanizing agent, the silanizing 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 silanizing agent, the mass content of the silanizing agent is 70% - 83%, and the mass content of zinc oxide is 17% - 30%; Based on the weight of the catalyst, the mass content of the resin is 68% - 89%, and the mass content of the zinc oxide modified by the silanizing agent is 11% - 32%; The preparation method of the zinc oxide modified by the silanizing agent includes: mixing a zinc oxide with a silanizing agent modification solution and heating under reflux to obtain the zinc oxide modified by the silanizing agent; the preparation method of the silanizing agent modification solution includes: mixing a silanizing agent, a lower alcohol, water, and an alkali catalyst and reacting to obtain the silanizing agent modification solution; The preparation method of the catalyst includes: mixing a polymeric white ball with di(2-methylacrylic acid) triethylene glycol ester and a solvent, swelling to obtain System I; mixing the zinc oxide modified by the silanizing agent with a diacyl peroxide and N,N-dimethylformamide, stirring to obtain System II; dropping System II into System I and refluxing to obtain the catalyst.

2. The catalyst according to claim 1, characterized in that, The specific surface area of the zinc oxide is 50-70 m 2 / g, the pore volume is 0.25-0.4 cm 3 / g, the average pore diameter is 10-20 Å, and the particle size is 10-20 nm.

3. The preparation method of the catalytic oxidation catalyst according to any one of claims 1 - 2, including: a. Preparing zinc oxide modified by a silanizing agent; b. Treating the resin with the zinc oxide modified by the silanizing agent obtained in step a and drying to obtain the catalyst.

4. The preparation method according to claim 3, characterized in that, In step a, the method for preparing the zinc oxide modified by the silanizing agent includes: mixing a zinc oxide with a silanizing agent modification solution and heating under reflux to obtain the zinc oxide modified by the silanizing agent; the preparation method of the silanizing agent modification solution includes: by mass fraction, silanizing agent: lower alcohol: water: alkali catalyst = (200 - 250):(200 - 300):(15 - 23):(0.6 - 1.3), mixing the silanizing agent, the lower alcohol, water, and the alkali catalyst and reacting at 50 - 70 °C for 2 - 5 hours to obtain the silanizing agent modification solution; the mass ratio of mixing the silanizing agent modification solution with the zinc oxide is (250 - 450):(100 - 200); the conditions for the heating under reflux reaction are as follows: refluxing at 110 - 130 °C for 4 - 7 hours.

5. The preparation method according to claim 4, wherein In step a, the lower alcohol is selected from at least one of methanol, ethanol, propanol, and isopropanol, and the alkali catalyst is selected from at least one of sodium hydroxide and potassium hydroxide.

6. The preparation method according to claim 3, 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 the resin, i.e., the polymerized white balls; the aqueous phase includes polyvinyl alcohol, sodium chloride, demineralized water, and azobisisobutyronitrile. By mass fraction, polyvinyl alcohol:sodium chloride:demineralized water: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. By mass fraction, methyl methacrylate:ethylene glycol dimethacrylate:sorbitan oleate:solvent = (65-85):(66-87):(7-10):(120-160).

7. The preparation method according to claim 6, characterized in that In step b, the molecular weight of polyvinyl alcohol is 1500-2500.

8. The preparation method according to claim 6, 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 make the droplet size of the oil phase control within 0.6-1.5 mm. After the dropping is completed, raise the temperature to 70-90 °C and react for 5-10 hours.

9. The preparation method according to claim 3, characterized in that, In step b, the polymerized white balls, di(2-methacryloyloxyethyl) glycol and a solvent. The solvent is selected from at least one of toluene, p-xylene, m-xylene, and o-xylene. By mass fraction, polymerized white balls:di(2-methacryloyloxyethyl) glycol:solvent = 50:(2-5):(150-200), and swell at 60-80 °C for 6-10 hours to obtain System I; the zinc oxide modified by the silanization reagent, peroxydiacyl and N,N-dimethylformamide obtained in step a. By mass fraction, zinc oxide modified by the silanization reagent:peroxydiacyl:N,N-dimethylformamide = 20:(0.3-0.8):(160-230), and stir at 70-90 °C for 6-10 hours to obtain System II; drop System II into System I. By volume ratio, System II:System I = 1-3, raise the temperature to 100-120 °C, and carry out reflux treatment for 3-6 hours.

10. The preparation method according to claim 9, wherein, The drying conditions described in step b are as follows: the drying temperature is 100-120 °C, and the drying time is 1-5 h.

11. Application of the catalytic oxidation catalyst described in any one of claims 1-2 or the catalytic oxidation catalyst prepared by the preparation method described in any one of claims 3-10 in jointly removing hydrogen sulfide and ammonia in gas.

12. The application according to claim 11, wherein The lean liquid for desulfurization absorption contacts with the raw material gas to absorb and separate hydrogen sulfide therein, obtaining a rich liquid for desulfurization absorption and a purified gas for desulfurization; the lean liquid for ammonia absorption contacts with the purified gas for desulfurization to absorb and separate ammonia therein, obtaining a rich liquid for ammonia absorption and a purified gas for ammonia removal; then the rich liquid for desulfurization absorption, the rich liquid for ammonia absorption and an oxygen-containing gas contact with the catalytic oxidation catalyst for reaction to obtain a thiosulfate emulsion; the thiosulfate emulsion is separated subsequently to obtain the regenerated lean liquid for desulfurization absorption and the lean liquid for ammonia absorption.

13. The application according to claim 12, characterized in that, The lean desulfurization absorption liquid includes the 120# solvent oil after light component removal and p-xylene. By mass percentage, the mass content of the 120# solvent oil after light component removal is 75% - 86%, and the mass content of p-xylene is 10% - 20%; the lean ammonia absorption liquid is water.

14. The application according to claim 12, characterized in that, In the feed of the reaction, the molar ratio of ammonia in the rich ammonia absorption liquid to hydrogen sulfide in the rich desulfurization absorption liquid is (0.8 - 1.8):1; the dosage of the oxygen-containing gas calculated as O2 and the molar ratio of ammonia in the rich ammonia absorption liquid is (100 - 200):

1.

15. The application according to claim 12 or 14, characterized in that The reaction conditions of the said reaction are as follows: the reaction temperature is 90 to 120 °C, the reaction pressure is 0.5 to 1.0 MPaG, and based on the mixed solution, the reaction volume space velocity is 0.5 to 1.0 h -1 .

16. The application according to claim 12, wherein The subsequent separation process of the thiosulfate emulsion is as follows: First, cool down and then separate gas and liquid. The obtained liquid phase is subjected to liquid-liquid decantation separation. The upper oil phase is the regenerated lean desulfurization absorption liquid, and the lower water phase is the crude lean ammonia absorption liquid. The crude lean ammonia absorption liquid is separated by membrane filtration to obtain the regenerated lean ammonia absorption liquid.

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