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CN113428861B - Methane and hydrogen sulfide reforming hydrogen production process - Google Patents

Methane and hydrogen sulfide reforming hydrogen production process Download PDF

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CN113428861B
CN113428861B CN202110815280.1A CN202110815280A CN113428861B CN 113428861 B CN113428861 B CN 113428861B CN 202110815280 A CN202110815280 A CN 202110815280A CN 113428861 B CN113428861 B CN 113428861B
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CN113428861A (en
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李平
肖植煌
马哲杰
冯昊
李准
李宗鸿
裴淏
付凯豪
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East China University of Science and Technology
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    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B32/00—Carbon; Compounds thereof
    • C01B32/70—Compounds containing carbon and sulfur, e.g. thiophosgene
    • C01B32/72—Carbon disulfide
    • C01B32/75—Preparation by reacting sulfur or sulfur compounds with hydrocarbons
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B17/00—Sulfur; Compounds thereof
    • C01B17/02—Preparation of sulfur; Purification
    • C01B17/04—Preparation of sulfur; Purification from gaseous sulfur compounds including gaseous sulfides
    • C01B17/0404—Preparation of sulfur; Purification from gaseous sulfur compounds including gaseous sulfides by processes comprising a dry catalytic conversion of hydrogen sulfide-containing gases, e.g. the Claus process
    • C01B17/046—Preparation of sulfur; Purification from gaseous sulfur compounds including gaseous sulfides by processes comprising a dry catalytic conversion of hydrogen sulfide-containing gases, e.g. the Claus process without intermediate formation of sulfur dioxide
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
    • C01B3/50—Separation of hydrogen or hydrogen containing gases from gaseous mixtures, e.g. purification
    • C01B3/52—Separation of hydrogen or hydrogen containing gases from gaseous mixtures, e.g. purification by contacting with liquids; Regeneration of used liquids
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/02—Processes for making hydrogen or synthesis gas
    • C01B2203/0205—Processes for making hydrogen or synthesis gas containing a reforming step
    • C01B2203/0227—Processes for making hydrogen or synthesis gas containing a reforming step containing a catalytic reforming step
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/04—Integrated processes for the production of hydrogen or synthesis gas containing a purification step for the hydrogen or the synthesis gas
    • C01B2203/0415—Purification by absorption in liquids
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/12—Feeding the process for making hydrogen or synthesis gas
    • C01B2203/1205—Composition of the feed
    • C01B2203/1211—Organic compounds or organic mixtures used in the process for making hydrogen or synthesis gas
    • C01B2203/1235—Hydrocarbons
    • C01B2203/1241—Natural gas or methane

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  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Inorganic Chemistry (AREA)
  • Engineering & Computer Science (AREA)
  • Combustion & Propulsion (AREA)
  • Health & Medical Sciences (AREA)
  • General Health & Medical Sciences (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

The invention discloses a methane and hydrogen sulfide reforming hydrogen production process. The method comprises the steps of: (1) The mixture contains CH at 600-900 DEG C 4 And H 2 The raw material gas of S is subjected to catalytic reaction to obtain reaction gas I; (2) separating the reaction gas I to obtain sulfur and a reaction gas II; (3) Flash evaporating reaction gas II to obtain coarse CS 2 Liquid and reaction gas III; (4) Crude CS of rectification 2 The liquid CS with the mass fraction of not less than 99.5% is obtained 2 The method comprises the steps of carrying out a first treatment on the surface of the The reaction gas III is contacted with the absorption liquid to obtain hydrogen-rich gas and absorption-rich liquid; and (5) separating the hydrogen-rich gas to obtain hydrogen with a volume fraction of more than 99.99% and the desorption gas.

Description

Methane and hydrogen sulfide reforming hydrogen production process
Technical Field
The invention relates to hydrogen preparation by hydrogen sulfide conversion, in particular to a flow and a process for preparing hydrogen by reforming methane and hydrogen sulfide.
Background
The high sulfur natural gas resources in China are rich and are mainly distributed in Bohai Bay basin and Sichuan basin, and the content of hydrogen sulfide (H2S) is generally more than 5%. Meanwhile, in the modification and refining process of sulfur-containing crude oil, a large amount of by-products H are generated in the hydrotreating process 2 S, S. Worldwide H 2 The S yield has reached 1 million tons per year. H 2 S is a highly corrosive and highly toxic gas, which can cause a great deal of problems for exploitation, transportation, storage and the like of sulfur-containing petroleum and natural gas, and the emission thereofAnd also pose a hazard to the natural environment and human health, so that it must be recycled.
Currently, the H is mainly prepared by adopting a physical or chemical method in the industry 2 S separation and enrichment, then using Claus process or LO-CAT process to make H 2 S is processed and converted, H 2 S is converted into sulfur and water. Although these processes can recover H 2 Sulfur in S, but at the same time H 2 The hydrogen in the S molecule is oxidized into water, so that the waste of hydrogen resources is caused. Hydrogen is an emerging clean energy source, and is expected to play a great role in the future, so that hydrogen becomes an important strategic resource worldwide. By H 2 S decomposition technology is simple in process, but the reaction conversion rate and the hydrogen yield are very low due to the limitation of thermodynamic equilibrium, and the pure H 2 S generally requires high temperatures above 1300℃to achieve conversions above 50%. The high temperature causes an increase in the cost of the technology, making the technology non-competitive. And H is 2 S decomposes and produces sulfur with low added value while producing hydrogen, and sulfur is very easy to block a pipeline in the reaction process, causes the problems of equipment corrosion and the like, and increases the operation cost. Thus H 2 S decomposition technology has not been industrialized so far.
Methane (CH) 4 ) And H is 2 S reforming hydrogen production is a brand new H 2 S conversion and hydrogen production technical route, its reaction process is: CH (CH) 4 +H 2 S＝CS 2 +H 2 . And H is 2 S decomposition compared with CH 4 And H is 2 The reaction conversion rate of the S reforming hydrogen production process under the same condition is high, the hydrogen yield is high, and the byproduct carbon disulfide (CS) 2 ) The added value of (2) is also obviously higher than that of sulfur, and the market price of the sulfur is 5 times of that of the sulfur at present, so the sulfur is more competitive.
There is therefore an urgent need in the art to provide CH 4 And H is 2 S reforming hydrogen production process and technology.
Disclosure of Invention
The invention aims to design and develop the high-concentration H 2 S is used as raw material, and is utilized together with CH 4 Reaction, synchronous production of hydrogen and CS 2 Is the routeRealizing the technical proposal of industrialization.
The invention provides a method for producing hydrogen by reforming methane and hydrogen sulfide, which comprises the following steps:
(1) The mixture contains CH at 600-900 DEG C 4 And H 2 The raw material gas of S is subjected to catalytic reaction to obtain reaction gas I;
(2) Separating the reaction gas I to obtain sulfur and a reaction gas II;
(3) Flash evaporating reaction gas II to obtain coarse CS 2 Liquid and reaction gas III;
(4) Crude CS of rectification 2 The liquid CS with the mass fraction of not less than 99.5% is obtained 2 The method comprises the steps of carrying out a first treatment on the surface of the The reaction gas III is contacted with the absorption liquid to obtain hydrogen-rich gas and absorption-rich liquid; and
(5) Separating the hydrogen-rich gas to obtain hydrogen with the volume fraction of more than 99.99% and the resolved gas.
In another embodiment, H in the feed gas of step (1) 2 S and CH 4 The molar ratio of (2) is 1.0 to 5.0, preferably 2.0 to 3.0.
In another embodiment, the catalytic reaction temperature in step (1) is 800-1000 ℃, preferably 900-1000 ℃.
In another embodiment, reaction gas II in step (3) is subjected to a secondary flash distillation.
In another embodiment, the absorption liquid in step (4) is selected from one or more of the following: low temperature methanol, ethanolamine, diethanolamine and N-methyldiethanolamine, preferably low temperature methanol.
In another embodiment, the molar ratio of the reaction gas III to the absorption liquid in step (4) is above 2.18.
In another embodiment, the temperature of the absorption liquid in step (4) is-20 to-60 ℃.
In another embodiment, step (4) is at H 2 And (3) enabling the reaction gas III to contact with the absorption liquid in the S absorption tower, wherein the pressure of the absorption tower is 2-5 MPa.
In another embodiment, step (5) separates the hydrogen-rich gas using a Pressure Swing Adsorption (PSA) unit.
In another embodiment, the CH-containing stream is preheated by a waste heat recovery system 4 And H 2 The temperature of the raw material gas of S reaches 600-900 ℃.
Accordingly, the present invention provides a CH 4 And H is 2 S reforming hydrogen production process and technology.
Drawings
FIG. 1 is a schematic diagram of a hydrogen production process system and process flow of the present invention.
Detailed Description
The inventor has studied extensively and intensively, and found that when the acid tail gas and the natural gas with high sulfur content of the refinery are treated by a flow process system, not only the acid component H can be treated 2 S, also can fully recycle H 2 H and S elements in S to obtain hydrogen and CS with high added value 2 The product changes waste into valuable, and improves the utilization rate of resources. And H is 2 Compared with the direct S decomposition technology, the invention adds CH 4 Make H 2 S and CH 4 The catalytic reforming reaction is carried out, the free energy of the reaction Gibbs is reduced, the reaction condition is milder, the equilibrium conversion rate of the reaction is improved, and the H is led to 2 The S hydrogen production process is more economical. On this basis, the present invention has been completed.
In particular, the present invention provides a method for simultaneous production of hydrogen and CS 2 Comprises the steps of:
in the first step, CH mixed according to a certain proportion 4 Gas and H 2 S gas is preheated to 600-900 ℃ and then is subjected to catalytic reaction to obtain reaction gas I;
secondly, separating the reaction gas I to obtain byproduct sulfur and reaction gas II;
step three, flash evaporation is carried out on the reaction gas II to obtain reaction gas III and crude CS 2 A liquid; coarse CS 2 Rectifying the liquid to obtain liquid CS with mass fraction not less than 99.5% 2 Product, H obtained at the same time 2 S gas is circulated back to the raw material gas for re-reaction;
fourth step, the reaction gas III is H 2 S, countercurrent absorption is carried out by an absorption tower, hydrogen-rich gas is obtained at the tower top, and rich absorption liquid is obtained at the tower bottom; separating the rich absorption liquid to obtain non-condensable gas, and refluxing the non-condensable gas to H 2 S absorption tower and liquid material obtained by gas-liquid separationFeeding the waste gas into an absorption liquid regeneration tower to obtain regenerated absorption liquid and analysis gas;
and fifthly, separating the hydrogen-rich gas obtained in the fourth step to obtain hydrogen and resolved gas with volume fraction of more than 99.99%.
In the first step, H in the raw material gas 2 S/CH 4 The molar ratio is 1.0 to 5.0, preferably 2.0 to 3.0.H 2 A proper excess of S favors the production of more hydrogen while suppressing carbon deposition, but H 2 A large excess of S leads to H 2 S conversion rate is reduced, circulation ratio is overlarge, equipment investment is greatly increased finally, and operation cost is increased.
In one embodiment of the present invention, the raw material H 2 S gas, H 2 The S volume content is 10-100%, and can be natural gas with high sulfur content, acid gas of refinery and H content 2 S waste gas and the like, wherein the raw material pretreatment unit is mainly used for pretreating H2S raw materials and removing oxygen (O2), water (H2O) and carbon dioxide (CO) in the raw materials 2 ) And solid particulate impurities.
In one embodiment of the invention, the feed gas is preheated in a manner that provides waste heat; the waste heat of the high-temperature flue gas and the reaction gas is utilized to preheat the raw material gas, and the air and the fuel gas can be preheated, and the flue gas and the reaction gas are cooled.
The catalytic reaction temperature in the first step is kept at 800-1000 ℃ and the catalytic reaction pressure is 0.05-0.15MPa (normal pressure operation, namely, around one atmosphere); catalysts used include, but are not limited to, moS 2 /Al 2 O 3 。
The way to separate the by-product sulfur in the reaction gas I that can be used in the above-mentioned second step includes, but is not limited to, the use of a sulfur trap, a sulfur condenser, etc.
The byproduct sulfur in the second step may be liquid sulfur (i.e., liquid sulfur).
In one embodiment of the invention, the reaction gas I may be recycled with heat before entering the sulfur trap.
In one embodiment of the invention, the heat of the reaction gas and the flue gas is recovered for preheating the raw material gas, and air and fuel gas can be preheated, and the reaction gas and the flue gas are cooled.
In the third step, the reaction gas II is subjected to pressurized cooling to perform flash evaporation, and secondary flash evaporation is preferred; the rectification may use CS 2 The rectifying tower has operation pressure of 1.0-2.0 MPa, temperature of 10-50 deg.c and reflux ratio of 0.1-0.3.
In one embodiment of the invention, the coarse CS 2 Liquid channel CS 2 The rectifying tower obtains liquid CS with mass fraction not less than 99.5% 2 H obtained from the top of the column 2 S gas is recycled to the feed gas for re-reaction.
In the fourth step, H is produced by the first catalytic reaction 2 The conversion of S is generally less than 50% so that H is entered 2 H in S absorption tower gas 2 S concentration is high (generally higher than 5%, the rest is H) 2 、CH 4 、N 2 ) Therefore, H must be taken 2 S is separated from the mixed gas and returned to the reaction furnace for recycling. The absorption liquid can be one or more of low-temperature methanol, ethanolamine, diethanolamine and N-methyldiethanolamine, and preferably adopts a low-temperature methanol washing method to separate high-concentration H 2 S, S. Compared with chemical absorption liquids such as ethanolamine, diethanolamine, N-methyldiethanolamine and the like, the low-temperature methanol gas-liquid ratio can reach more than 2.18mol of acid gas/mol of methanol, the high gas-liquid ratio effectively reduces the volume of equipment and the consumption of the absorption liquid, thereby reducing the investment cost of equipment and the operation cost. On the other hand, since no one has yet used the low temperature methanol washing method for higher concentration H 2 S separation, therefore, the application of this method to the present invention presents challenges, and first requires high solubility to sufficiently absorb a large amount of H 2 S gas, secondly requires a small influence of the heat of solution to avoid a large amount of H 2 S is dissolved to cause temperature rise to influence absorption effect, therefore, the invention is based on H 2 S solubility in low-temperature methanol and influence analysis of heat of solution, optimizing temperature and pressure of absorption liquid, and determining high concentration H by low-temperature methanol under the determined operating condition 2 S has high absorption efficiency and large treatment capacity.
In one embodiment of the invention, in H 2 And (3) enabling the reaction gas III to contact with an absorption liquid in the S absorption tower, wherein the temperature of the absorption liquid can be between 20 ℃ below zero and 60 ℃ below zero, and the pressure of the absorption tower can be between 2 and 5MPa. Preferably at a temperature of-40℃and a pressure of 3MPa.
In one embodiment of the present invention, the rich absorption liquid obtained in the fourth step is flashed and then passed through an absorption liquid regeneration tower to obtain a resolved gas, wherein the resolved gas has a main component of H 2 S, the reaction system can be recycled for re-reaction.
The method for separating the hydrogen-rich gas that can be used in the fifth step includes, but is not limited to, cryogenic rectification, membrane separation, conventional adsorption separation, etc.
In one embodiment of the invention, an activated carbon material with good performance is selected (for example, but not limited to, a specific surface area of more than 2000m 2 Per gram, pore volume greater than 1.8cm 3 Activated carbon material per gram), separating hydrogen-rich gas by Pressure Swing Adsorption (PSA) device to obtain a desorption gas containing CH as the main component 4 Recycling the hydrogen into the raw materials for re-reaction, and obtaining the high-purity hydrogen with the volume fraction of more than 99.99%.
The steps provided by the invention provide a CS 2 And H is 2 S separation and recovery method. The CS-containing product obtained in the second step 2 、H 2 Unreacted H 2 S、CH 4 The mixed gas (i.e. reaction gas II) is cooled and compressed to form non-condensable H 2 And CH (CH) 4 Still in the gaseous state, CS 2 Is completely liquefied (the liquefaction capacity depends on the temperature and CS 2 Physical properties and partial pressure) of (C) and (C) due to CS 2 Liquid pair H 2 S has a large solubility, so that part H 2 S will dissolve in CS 2 (the dissolution power depends on the temperature, H) 2 S partial pressure and H 2 S and CS 2 Affinity between) to cause liquefaction of CS 2 Contains part H 2 S, result in CS 2 The purity of the liquid product is low. To separate CS 2 H in liquid 2 S, a rectification method is generally adopted. However, the direct rectification method has high energy consumption, especiallyIn the invention, due to unreacted H 2 Higher S concentration results in dissolution in CS 2 H in liquid 2 The S content is higher, so that the energy consumption of a direct rectification method can be increased. For this purpose, the invention innovatively adopts a low-temperature flash evaporation method, so that CS 2 Dissolved H in liquid 2 S escapes rapidly, thereby improving CS 2 The purity of the liquid reduces the energy consumption burden of the next rectification separation. At the same time, through secondary flash evaporation, CS is also improved 2 Is a recovery rate of (2).
The process for synchronously producing hydrogen and carbon disulfide provided by the invention takes high-concentration hydrogen sulfide as a raw material to react with methane to simultaneously produce hydrogen and carbon disulfide, can treat the acid tail gas of a refinery and the natural gas with high sulfur content, and synchronously produces green energy source H 2 And high value added product CS 2 。
The invention also provides a hydrogen production flow system, which comprises a start-up heating furnace, a waste heat recovery system, a catalytic reaction furnace, a sulfur catcher, a gas compressor, a secondary flash separator and CS 2 Rectifying column, H 2 S absorption tower, absorption liquid regeneration tower and PSA pressure swing adsorption device. The raw material gas feeding pipeline is connected with the waste heat recovery system after passing through the starting heating furnace, and the starting heating furnace is only used during starting; the raw material gas is connected with an inlet at the upper end of the tube side of the catalytic reaction furnace after passing through the waste heat recovery system; the reaction gas at the outlet of the lower end of the catalytic reaction furnace is connected with a gas compressor through a waste heat recovery system, and is connected with the inlet of a secondary flash separator after being compressed and cooled; lower liquid phase outlet and CS of two-stage flash separator 2 The inlet of the rectifying tower is connected, and the upper gas outlet is cooled and H is led to 2 The inlet at the lower end of the S absorption tower is connected; CS (circuit switching) 2 The top outlet of the rectifying tower is connected with a raw material gas pipeline, and is circulated back to the reaction system, and the bottom outlet of the rectifying tower is connected with CS 2 The finished tank connection achieved 99.5wt.% CS 2 A product; h 2 The rich gas outlet at the upper end of the S absorption tower is connected with the inlet of the PSA pressure swing adsorption device, the methane gas outlet of the PSA pressure swing adsorption device is connected with a raw gas pipeline and recycled to the reaction system, and the hydrogen outlet of the PSA pressure swing adsorption device obtains a hydrogen product with the purity of 99.99 percent; h 2 S absorption tower lower extreme rich solution exportThe device is connected with a gas-liquid separation tank and is used for removing non-condensable gas in the rich liquid, and an outlet at the lower end of the gas-liquid separation tank is connected with an inlet of a methanol regeneration tower through heat exchange and cooling; the gas outlet at the top of the methanol regeneration tower is connected with a raw gas pipeline, and is circulated back to the reaction system, and the outlet of the tower kettle is connected with the liquid phase inlet at the top of the methanol absorption tower through pressurization and temperature reduction.
In one embodiment of the invention, the catalytic reactor is a tube reactor, the raw material gas is passed through a tube side, the fuel gas is passed through a shell side, the catalytic reforming reaction in the tube side belongs to a strong endothermic reaction, and the fuel gas is combusted through the shell side to provide heat for the reaction. The high-temperature reaction gas at the outlet of the reaction furnace tube and the high-temperature flue gas at the outlet of the shell pass are used for preheating raw material gas, byproduct steam and the like, so that comprehensive energy utilization is realized, and the byproduct steam can provide heat for reboilers of various rectifying towers of the separation unit. Through calculation, the high-temperature reaction gas at the outlet of the tube side and the high-temperature flue gas at the outlet of the shell side can recycle waste heat, can be sufficiently used for preheating raw material gas, fuel gas and air, and can ensure proper heat exchange temperature difference.
In one embodiment of the invention, the hydrogen production flow system comprises a startup heating furnace F-101, a high temperature catalytic reaction furnace R-101, a sulfur catcher V-101, a carbon disulfide rectifying tower T-201, a hydrogen sulfide absorption tower T-301, an absorption liquid regeneration tower T-302, a PSA pressure swing adsorption unit and a waste heat recovery system, wherein the waste heat recovery system comprises a plurality of heat exchangers (for example, 11). See fig. 1.
In one embodiment of the present invention, the following process using the flow system is provided as follows:
(1) Raw material CH 4 And H is 2 S gas is mixed according to a certain proportion and enters a start-up heating furnace F-101 for pretreatment. Preheating raw material gas to 600-900 ℃, and then entering a high-temperature catalytic reaction furnace R-101;
(2) The temperature of the catalytic reaction furnace is kept at 800-1000 ℃, and reaction gas I is obtained through catalytic reaction;
(3) The reaction gas I obtained in the step (2) is subjected to sulfur trap V-101 to separate out byproduct sulfur in the reaction gas I, so as to obtain liquid sulfur and reaction gas II;
(4) The reaction gas obtained in the step (3)II is pressurized and cooled and then enters flash separation tanks V-201 and V-202, reaction gas III is obtained at the top of the tower, and crude CS is obtained at the bottom of the tower 2 A liquid; coarse CS 2 Liquid in CS 2 Rectifying in a rectifying tower T-201 to obtain liquid CS with mass fraction not less than 99.5% 2 Product, H obtained from the top of the column 2 S gas is circulated back to the raw material gas for re-reaction;
(5) The reaction gas III obtained in the step (4) enters H 2 S recovery unit, reaction gas III in H 2 S, countercurrent absorption is carried out on the T-301 absorption tower, hydrogen-rich gas is obtained at the tower top, and rich absorption liquid is obtained at the tower bottom; the rich absorption liquid is decompressed and then enters a gas-liquid separation tank V-301 to obtain non-condensable gas which flows back to H 2 S, the absorption tower T-301 is used for separating gas from liquid, and the liquid material obtained by gas-liquid separation enters an absorption liquid regeneration tower T-302 to obtain regenerated absorption liquid and analysis gas;
(6) The main component of the resolved gas obtained in the step (5) is H 2 S, circulating the heating furnace F-101 to enter the catalytic reaction furnace R-101 for re-reaction;
(7) The hydrogen-rich gas obtained in the step (5) enters a Pressure Swing Adsorption (PSA) device to obtain hydrogen and resolved gas with volume fraction more than 99.99%;
(8) The main component of the analysis gas obtained in the step (7) is CH 4 And recycling to the high-temperature catalytic reaction furnace R-101 for re-reaction.
The catalytic reaction furnace R-101 is characterized in that the heating fuel is natural gas, the fuel gas passes through a shell side, and the raw material gas passes through a tube side. The startup heating furnace F-101 is only used during startup, and a waste heat recovery system is used for heating raw material gas in normal operation of equipment. The preheating temperature of the raw material gas is 600-900 ℃. The preheating temperature is lower than 600 ℃ and is unfavorable for reaction conversion, but the preheating temperature is too high to cause difficult material selection of equipment, so the preheating temperature is not higher than 900 ℃.
In one embodiment of the invention, the high temperature catalytic reactor R-101 is a tube array reactor; feed gas enters the tube side from the top inlet of the reactor and fuel gas and air enter the shell side from the lower inlet. The startup heating furnace F-101 is only used during startup, and raw materials are preheated through the waste heat recovery system in normal operation of the equipment.
The waste heat recovery system comprises nine heat exchangers: e-101, E-102, E-103, E-104, E-105, E-001, E-002, E-003, E-004. E-001-E-004 are used for recovering waste heat of flue gas at the outlet of the shell of the R-101 high-temperature catalytic reaction furnace, and E-101-E-105 are used for recovering waste heat of the outlet of the tube side reaction gas. The waste heat recovery system can preheat the raw material gas to 900 ℃, the air to 450 ℃ and the fuel gas to 300 ℃. Simultaneously, the temperature of the reaction gas is reduced to 256 ℃, and the temperature of the flue gas is reduced to 148 ℃.
In one embodiment of the invention, air is subjected to heat exchange by E-104 and then subjected to heat exchange by E-003, and then enters the shell side R-001 of the high temperature catalytic reactor R-101, and fuel gas CH4 is subjected to heat exchange by E-105 and then directly enters the shell side R-001 of the high temperature catalytic reactor R-101, instead of passing air or fuel gas through other heat exchangers.
In one embodiment of the invention, raw gas sequentially passes through heat exchangers E-004, E-102, E-002, E-101 and E-001 to exchange heat and then enters a tube side in a high-temperature catalytic reaction furnace R-101 for catalytic reaction.
In one embodiment of the invention, the reaction gas at the tube side outlet of the high-temperature catalytic reaction furnace R-101 is cooled by the heat exchangers E-101-E-105 and then enters the sulfur trap V-101.
In one embodiment of the invention, the flue gas at the outlet of the shell side R-001 of the high-temperature catalytic reaction furnace R-101 is cooled by the heat exchangers E-001-E-004 and then is exhausted.
According to the invention, the energy of the cold and hot material flows is calculated to match the material flows in different temperature intervals, and the heat exchange scheme is finally determined through the optimization of the complex heat exchange network, so that the full and efficient utilization of the energy of the whole process is realized, and the economical efficiency of the system is improved.
In one embodiment of the invention, the high temperature catalytic reactor R-101 has a reaction temperature of 800-1000 ℃, a reaction pressure of 0.1-0.5 MPa, preferably a reaction temperature of 900-1000 ℃ and a reaction pressure of 0.1-0.3 MPa. The proportion of the raw material gas is H 2 S/CH 4 ＝1.0～5.0。
In one embodiment of the invention, the high-temperature catalytic reaction furnace R-101 belongs to a tube type reactor, the tube side is filled with a catalyst,the feed gas enters the tube side from the inlet at the top end of the reactor; fuel gas and air (O) 2 ) Enters the shell side from the inlet at the lower end, and provides heat and high-temperature environment for the reaction by combusting fuel gas.
In one embodiment of the invention, the outlet temperature of the startup heating furnace F-101 is 600 ℃, and the outlet temperature of the raw material gas after passing through the waste heat recovery system is 900 ℃.
In one embodiment of the invention, the sulfur trap V-101 outlet line and H 2 The waste heat recovery system is arranged between gas phase inlet pipelines of the S rectifying tower T-201, and is also provided with a compressor C-101, heat exchangers E-106, E-107 and E-108 and flash separation tanks V-201 and V-202. Flash separation tank V-202 gas phase outlet and H 2 S rectifying tower T-201 pipeline is connected, flash separation tank V-201 liquid phase outlet is connected with CS 2 The rectifying tower T-101 is connected through a pipeline.
CS-containing effluent from the upper part of sulfur trap V-101 2 、H 2 Unreacted H 2 S、CH 4 The mixed gas is subjected to heat exchange cooling through heat exchangers E-106 and E-107 and is compressed by a compressor C-101, and then is non-condensable H 2 And CH (CH) 4 Still in the gaseous state, CS 2 Then it is completely liquefied.
In one embodiment of the invention, compressor C-101 employs three stages of compression with a compressor outlet pressure of 2 to 5MPa, preferably 3MPa. The inlet of the compressor C-101 and the outlet pipeline of the sulfur catcher V-101 are provided with a heat exchanger E-106. The outlet of the compressor C-101 and the inlet pipeline of the flash separation tank V-201 are provided with a heat exchanger E-107.
In one embodiment of the invention, an inlet of a used carbon disulfide rectifying tower T-201 is connected with liquid outlets of two-stage flash separators V-201 and V-202, 90% of crude carbon disulfide is recovered by the first-stage flash separator V-201, 99% of crude carbon disulfide is recovered by the second-stage flash separator V-202, part of unreacted hydrogen sulfide is obtained at the top of the rectifying tower T-201, and the unreacted hydrogen sulfide is recycled to a reaction system, and a carbon disulfide product is obtained at the bottom of the rectifying tower. At a separation system pressure of 3MPa and CS 2 And H 2 The S accounts for 37.13wt.% and 52.93wt.% respectively, the reaction gas is cooled to 20-25 ℃ and then is separated by first stage flash evaporationObtaining CS 2 The recovery rate of (2) reaches 90 percent; further cooling the residual gas phase to-5 to-8 ℃, performing secondary flash separation, and finally CS 2 The recovery rate of (3) reaches 99 percent. Crude CS obtained by two-stage flash separation 2 Feeding into CS 2 Rectifying in rectifying tower T-201 to obtain CS in 99.5wt 2 And (5) a product.
In another preferred example, a heat exchanger E-201 is arranged between the outlets of the flash separators V-201 and V-202 and the inlet of the rectifying tower T-201, and the outlet temperature of the heat exchanger is-5 ℃. The heat exchanger is used for controlling the temperature of the feed entering the rectifying tower.
In another preferred embodiment, a compressor is arranged between the reaction gas II and the inlet of the flash separation tank V-201, and the outlet pressure of the compressor is 2-5 MPa, preferably 3MPa.
In another preferred embodiment, the connection line of the compressor and the flash separator V-201 is provided with a heat exchanger.
In one embodiment of the invention, the flash separation tank V-202 vapor phase outlet is in fluid communication with H 2 The connecting pipeline of the S absorption tower T-301 is provided with a heat exchanger E-301, and the outlet temperature of the heat exchanger E-104 can be between-7 ℃ and-30 ℃, and is preferably between-20 ℃.
In one embodiment of the invention, the H 2 The temperature of the absorption liquid of the S absorption tower T-301 is-20 to-70 ℃ and the pressure is 2 to 5MPa.
In another preferred embodiment, the H 2 The operating pressure of the S absorption tower T-301 is 3MPa, and the temperature of the absorption liquid is-40 ℃.
In one embodiment of the invention, a gas-liquid separation tank and a heat exchanger are arranged between the hydrogen sulfide absorption tower and the absorption liquid regeneration tower, the gas-liquid separation tank is connected with the absorption tower, a rich liquid outlet below the absorption tower is connected with an inlet of the absorption liquid regeneration tower, a circulating pump and a heat exchanger are arranged between a tower bottom connecting pipeline of the regeneration tower and a tower bottom connecting pipeline of the rectifying tower, and a tower top gas outlet of the regeneration tower is connected with a raw gas pipeline.
For example, H 2 The gas-liquid separation is arranged on the connecting pipeline between the rich liquid outlet of the S absorption tower T-301 and the inlet of the absorption liquid regeneration tower T-302The gas-liquid separation tank is used for removing non-condensable gas in the rich absorption liquid, and comprises a separation tank V-301 and a heat exchanger E-304. Gas phase outlet of gas-liquid separation tank circulates back to H 2 S the gas phase inlet of the absorption tower T301; the heat exchanger is used for recovering the cold energy of the rich liquid and improving the energy utilization rate.
In one embodiment of the invention, H 2 And the rich gas outlet of the S absorption tower T-301 enters the PSA unit for separating methane and hydrogen, and the separated methane is recycled to the reaction system for re-reaction, and meanwhile, the hydrogen with the product purity of 99.99% is obtained.
In one embodiment of the present invention, the absorber T-302 may be operated at a pressure of 0.05 to 0.5MPa, preferably 0.1 to 0.2MPa.
In one embodiment of the invention, the liquid phase outlet of the methanol regeneration tower T-302 is connected with H 2 The liquid phase inlet connecting pipeline of the S absorption tower T-301 is provided with a circulating pump P-301, a heat exchanger E-304 and a heat exchanger E-305. The outlet temperature of the heat exchanger is-20 to-60 ℃, preferably-40 ℃. Said H 2 The S absorption tower is also provided with an absorption liquid supplementing system for supplementing the absorption liquid lost in operation.
The invention has the particularity that the invention not only relates to high-temperature reaction (-1000 ℃), but also needs low-temperature operation (-60 ℃), has wide temperature range and multiple operation units, and is a great difficulty in realizing comprehensive and efficient utilization of energy. According to the invention, the flow heat exchange network is optimized, the economy or low energy consumption is used as a target, the temperature zone diagram of each operation unit is drawn by adopting a problem table method, the energy of cold and hot material flows is calculated to carry out the mutual matching of the material flows, and finally, the heat exchange scheme is determined, so that the full and efficient utilization of the energy of the whole flow is realized, and the economy of the system is improved.
So that those skilled in the art can appreciate the features and effects of the present invention, a general description and definition of the terms and expressions set forth in the specification and claims follows. Unless otherwise defined, all technical and scientific terms used herein have the same meaning as commonly understood by one of ordinary skill in the art to which this invention belongs, and in the event of a conflict, the present specification shall control.
The theory or mechanism described and disclosed herein, whether right or wrong, is not meant to limit the scope of the invention in any way, i.e., the present disclosure may be practiced without limitation to any particular theory or mechanism.
In this document, all features, such as values, amounts, and concentrations, are for brevity and convenience only, as defined in the numerical or percent range. Accordingly, the description of a numerical range or percentage range should be considered to cover and specifically disclose all possible sub-ranges and individual values (including integers and fractions) within the range.
The above-mentioned features of the invention, or of the embodiments, may be combined in any desired manner. All of the features disclosed in this specification may be used in combination with any combination of features, provided that the combination of features is not inconsistent and all such combinations are contemplated as falling within the scope of the present specification. The various features disclosed in the specification may be replaced by alternative features serving the same, equivalent or similar purpose. Thus, unless expressly stated otherwise, the disclosed features are merely general examples of equivalent or similar features.
The invention has the main advantages that:
1. when the acid tail gas or the natural gas with high sulfur content of the refinery is treated, not only H can be treated 2 S is discharged, and H can be efficiently recovered 2 H and S elements in S synchronously generate green energy H 2 And high value added product CS 2 。
2. The reaction is carried out under the high temperature condition, and the waste heat recovery system effectively utilizes the high temperature waste heat of the reaction gas and the flue gas, thereby fully utilizing the energy and effectively reducing the energy consumption.
3. Preliminary CS separation by secondary flash separation 2 Separating, rectifying to obtain refined CS 2 A product; the separation method reduces energy consumption and efficiently recovers CS 2 Effectively improve CS 2 Is a recovery rate of (2).
4. Separation of higher concentration H by low temperature methanol wash 2 S, the consumption of the absorption liquid is greatly reduced, and the volumes of the absorption tower and the regeneration tower are effectively reduced, thereby reducing the installationA standby fee and an operating fee; the low-temperature methanol washing belongs to physical absorption, and compared with chemical absorption, the low-temperature methanol washing can avoid foaming phenomenon in the absorption process.
The invention will be further illustrated with reference to specific examples. It is to be understood that these examples are illustrative of the present invention and are not intended to limit the scope of the present invention. The experimental procedures, which do not address the specific conditions in the examples below, are generally carried out under conventional conditions or under conditions recommended by the manufacturer. All percentages, ratios, proportions, or parts are by weight unless otherwise indicated. The units in weight volume percent are well known to those skilled in the art and refer, for example, to the weight of solute in 100 milliliters of solution (grams). Unless defined otherwise, all technical and scientific terms used herein have the same meaning as commonly understood by one of ordinary skill in the art. In addition, any methods and materials similar or equivalent to those described herein can be used in the methods of the present invention. The preferred methods and materials described herein are presented for illustrative purposes only.
The invention provides a hydrogen production process and a process thereof, wherein the process system of the process comprises a start-up heating furnace F-101, waste heat recovery systems E101-E105, E001-E-004, a high-temperature catalytic reaction furnace R-101 (R-001 in the attached figure 1 represents the shell side of the reaction furnace), a sulfur catcher V-101, a gas compressor C-101, a secondary flash evaporation separation tank V-201, V-202 and CS 2 Rectifying tower T-201, H 2 S absorption tower T-301, absorption liquid (methanol) regeneration tower T-302, PSA pressure swing adsorption device T-401. The raw material gas passes through a startup heating furnace F-101 and then sequentially passes through heat exchangers E-004, E-102, E-001 and E-101 in a waste heat recovery system, and then is connected with an inlet at the upper end of a tube side of a high-temperature catalytic reaction furnace R-101; the fuel gas is connected with an inlet at the lower end of the shell side of the high-temperature catalytic reaction furnace R-101 through a waste heat recovery system E-104; the air or oxygen is connected with the inlet at the lower end of the shell side of the high-temperature catalytic reaction furnace R-101 through waste heat recovery systems E-105 and E-003. The high-temperature reaction gas at the tube side outlet of the high-temperature catalytic reaction furnace R-101 is subjected to heat exchange and temperature reduction through a waste heat recovery system E-101-E-105 and then is connected with the inlet at the upper end of the sulfur catcher V-101; high-temperature flue gas at the outlet of the shell side of the high-temperature catalytic reactor R-101 passes through a waste heat recovery system E-001-E-004, heat exchange and cooling, and then evacuating. The outlet of the sulfur catcher V-101 is connected with the inlet of the three-stage compressor C-101 after E-106 heat exchange, the outlet of the compressor C-101 is connected with the inlet of the second-stage flash evaporation separation V-201 after E-107 water cooling heat exchange and E-108 chilled brine heat exchange, the gas outlet above the flash evaporation separator V-201 is connected with the inlet of the second-stage flash evaporation separator V-202, and the liquid outlet of the second-stage flash evaporation separation is connected with CS 2 CS of rectifying column T-201 2 An inlet connection; the gas outlet above the flash separator V-202 is connected with H after passing through the heat exchanger E-301 2 The inlet of the S absorption tower T-301 is connected. CS (circuit switching) 2 The top outlet of the rectifying tower T-201 is connected with a raw material gas pipeline to realize partial raw material H 2 S is recycled, and CS with purity of 99.5wt.% is obtained at the outlet of the tower kettle of the rectifying tower 2 And (5) a product. H 2 The tower top inlet of the S absorption tower T-301 is connected with a regenerated methanol pipeline and a supplemented methanol pipeline at the tower bottom outlet of the absorption liquid (methanol) regeneration tower T-302, and the H is 2 The rich gas outlet at the top of the S absorption tower T-301 is connected with the inlet of the PSA pressure swing adsorption unit; the H is 2 And the rich liquid outlet of the tower kettle of the S absorption tower T-301 is connected with the inlet of the gas-liquid separation tank V-301. The gas outlet above the gas-liquid separation tank V-301 and the H 2 The bottom inlet of the S absorption tower T-301 is connected, and the lower liquid outlet is connected with the inlet of the absorption liquid (methanol) regeneration tower T-302 after E-304 exchanges heat. The top outlet of the absorption liquid (methanol) regeneration tower T-302 is connected with a raw material gas pipeline to realize the residual unreacted H 2 S, recycling; the outlet of the absorption liquid (methanol) regeneration tower T-302 tower kettle is connected with the H through a circulating pump P-301, a heat exchanger E-304 and a heat exchanger E-305 2 The tower top liquid inlet of the S absorption tower T-301 is connected. The outlet of the PSA pressure swing adsorption unit is connected with a raw material gas pipeline to obtain a hydrogen product with purity of 99.99%, and the outlet of the desorption gas is connected with the raw material gas pipeline to realize unreacted CH 4 Recycling.
Example 1
Treatment of feed gas CH 4 Concentration 100%, flow 400kmol/H, H 2 S concentration is 100%, flow rate is 800kmol/H, so that H in raw material gas 2 S/CH 4 The molar ratio is 2. The feed gas was preheated to 600 ℃ and the reaction pressure was 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 2.66mol of acid gas (containing H 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 3MPa and low temperature methanol feed temperature of-40 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 797 ℃, CH 4 Conversion was 31.33%, H 2 S conversion was 31.45%; CS having a purity of 99.5wt.% was obtained 2 125.51kmol/H H with a purity of 99.99% 2 11,208Nm 3 /h。
Example 2
Treatment of feed gas CH 4 Concentration 100%, flow 400kmol/H, H 2 S concentration 10% (the remainder being N) 2 Gas), flow 4000kmol/h; the feed gas was preheated to 700 ℃, the reaction pressure was 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 2.18mol of acid gas (containing H) 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 3MPa and low temperature methanol feed temperature of-40 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 785 ℃, CH 4 Conversion was 26.46%, H 2 S conversion was 52.52%; CS having a purity of 99.5wt.% was obtained 2 120.55kmol/H H with a purity of 99.99% 2 10,656Nm 3 /h。
Example 3
Treatment of feed gas CH 4 Concentration of 10% (the balance being N) 2 Gas), flow rate 1000kmol/H, H 2 S concentration 50% (the remainder being N) 2 Gas), flow rate 800kmol/h; the feed gas was preheated to 900 ℃ and the reaction pressure was 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 2.31mol of acid gas (containing H) 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 3MPa and low temperature methanol feed temperature of-40 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 829 ℃, and CH 4 Conversion is 45.90%, H 2 S conversion was 23.05%; CS having a purity of 99.5wt.% was obtained 2 45.41kmol/H H with 99.99% purity 2 4,068Nm 3 /h。
Example 4
Treatment of feed gas CH 4 Concentration 50% (the remainder being N) 2 Gas), flow rate of 400kmol/H, H 2 S concentration is 100%, flow is 1000kmol/h; feed gasPreheated to 800 ℃, and the reaction pressure is 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 2.57mol of acid gas (containing H) 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 3MPa and low temperature methanol feed temperature of-40 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 886 ℃, CH 4 Conversion is 57.69%, H 2 S conversion was 23.1%; CS having a purity of 99.5wt.% was obtained 2 115.14kmol/H H with a purity of 99.99% 2 10,304Nm 3 /h。
Example 5
Treatment of feed gas CH 4 Concentration 100%, flow 400kmol/H, H 2 S concentration is 50%, flow rate is 2400kmol/h; the feed gas was preheated to 900 ℃ and the reaction pressure was 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 2.24mol of acid gas (containing H) 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 3MPa and low temperature methanol feed temperature of-40 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 916 ℃, CH 4 Conversion is 60.40%, H 2 S conversion was 40.54%; CS having a purity of 99.5wt.% was obtained 2 242.28kmol/H H with a purity of 99.99% 2 21,720Nm 3 /h。
Example 6
Treatment of feed gas CH 4 Concentration 100%, flow 400kmol/H, H 2 S concentration 10% (the remainder being N) 2 Gas), flow 4000kmol/h; the feed gas was preheated to 700 ℃, the reaction pressure was 1atm; the absorption tower adopts low-temperature methanol as absorbent, and the gas-liquid ratio of the absorption tower is 4.02mol of acid gas (containing H) 2 S:42.19 wt.%)/mol of low temperature methanol, operating at 5MPa and low temperature methanol feed temperature of-60 ℃. After the raw material gas passes through a catalytic reaction furnace, the outlet temperature of the reaction gas is 785 ℃, CH 4 Conversion was 26.46%, H 2 S conversion was 52.52%; CS having a purity of 99.5wt.% was obtained 2 120.55kmol/H H with a purity of 99.99% 2 10,656Nm 3 /h。
Comparative example 1
Other conditions were the same as in example 1, and the feed gas preheating temperature was changed to 400 ℃.
Because the preheating temperature of the raw material gas is too low, the temperature in the catalytic reaction furnace cannot reach the high temperature condition, the outlet temperature of the reaction gas is only 667 ℃, and the CH is not reached 4 The conversion is only 13.75%, H 2 The S conversion is only 13.77%; CS having a purity of 99.5wt.% was obtained 2 54.94kmol/H H with a purity of 99.99% 2 4,928Nm 3 /h。
Comparative example 2
The other conditions were the same as in example 1, and the reaction pressure was changed to 3MPa.
After the raw material gas passes through the catalytic reaction furnace, the outlet temperature of the reaction gas is 814 ℃. Although the reaction is carried out at high temperature, CH is generated due to excessive reaction pressure 4 The conversion is only 14.25%, H 2 The S conversion is only 14.31%; CS having a purity of 99.5wt.% was obtained 2 56.96kmol/H H with a purity of 99.99% 2 5,107Nm 3 /h。
Comparative example 3
Other conditions were the same as in example 3, changing the absorption liquid feed temperature to-10 ℃.
The absorber column uses low temperature methanol as the absorbent, and the operating pressure of the absorber column is still 3MPa. Because the temperature of the absorption liquid is higher, the consumption of the absorption liquid is greatly increased, the original 1124.83kmol/H is increased to 1991.78kmol/H, and the gas-liquid ratio of the absorption tower is reduced to 1.53mol of acid gas (H) 2 S:42.19 wt.%)/mol of low temperature methanol.
Comparative example 4
Other conditions were the same as in example 3, and the operating pressure of the absorption column was changed to 1MPa.
The absorption tower adopts low-temperature methanol as an absorbent, and the feeding temperature of the absorption tower is still-40 ℃. Because the operating pressure of the absorption tower is reduced, the consumption of the required absorption liquid is greatly increased from the original 1124.83kmol/H to 2257.30kmol/H, and the gas-liquid ratio of the absorption tower is reduced to 1.36mol of acid gas (H 2 S:42.19 wt.%)/mol of low temperature methanol.
The foregoing description is only illustrative of the preferred embodiments of the present invention and is not intended to limit the scope of the invention, which is defined broadly in the appended claims, and any person skilled in the art to which the invention pertains will readily appreciate that many modifications, including those that fall within the metes and bounds of the claims, or equivalence of such metes and bounds thereof.

Claims (11)

1. A process for producing hydrogen by reforming methane with hydrogen sulfide, the process comprising the steps of:
(1) The mixture contains CH at 600-900 DEG C 4 And H 2 The raw material gas of S is subjected to catalytic reaction to obtain reaction gas I;
(2) Separating the reaction gas I to obtain sulfur and a reaction gas II;
(3) Flash evaporating reaction gas II to obtain coarse CS 2 Liquid and reaction gas III;
(4) Crude CS of rectification 2 The liquid CS with the mass fraction of not less than 99.5% is obtained 2 The method comprises the steps of carrying out a first treatment on the surface of the The reaction gas III is contacted with the absorption liquid to obtain hydrogen-rich gas and absorption-rich liquid;
(5) Separating the hydrogen-rich gas to obtain hydrogen with the volume fraction of more than 99.99% and analytic gas;
in the step (3), the reaction gas II is subjected to secondary flash evaporation; the temperature of the first-stage flash evaporation is 20-25 ℃, and the temperature of the second-stage flash evaporation is-5-8 ℃;
the absorption liquid in the step (4) is low-temperature methanol; the temperature of the absorption liquid in the step (4) is-20 to-60 ℃; step (4) at H 2 S, enabling the reaction gas III to contact with the absorption liquid in an absorption tower, wherein the pressure of the absorption tower is 2-5 MPa; in the step (4), the mol ratio of the reaction gas III to the absorption liquid is more than 2.18.
2. The method of claim 1, wherein H in the feed gas of step (1) 2 S and CH 4 The molar ratio of (2) is 1.0-5.0.
3. The method of claim 1, wherein H in the feed gas of step (1) 2 S and CH 4 The molar ratio of (2) is 2.0-3.0.
4. The process of claim 1, wherein the catalytic reaction temperature in step (1) is 800-1000 ℃.
5. The process of claim 1, wherein the catalytic reaction temperature in step (1) is 900 to 1000 ℃.
6. The method of claim 1, wherein the temperature of the absorption liquid in step (4) is-20 to-40 ℃ or-40 to-60 ℃.
7. The method of claim 1, wherein the temperature of the absorption liquid in step (4) is-40 ℃.
8. The method of claim 1, wherein step (4) is at H 2 And (3) enabling the reaction gas III to contact with the absorption liquid in the S absorption tower, wherein the pressure of the absorption tower is 2-3 MPa or 3-5 MPa.
9. The method of claim 1, wherein step (4) is at H 2 And (3) enabling the reaction gas III to be in contact with the absorption liquid in the S absorption tower, wherein the pressure of the absorption tower is 3MPa.
10. The method of claim 1, wherein step (5) separates the hydrogen-rich gas using a pressure swing adsorption apparatus.
11. The method according to any one of claims 1 to 10, wherein the CH-containing stream is preheated by a waste heat recovery system 4 And H 2 The temperature of the raw material gas of S reaches 600-900 ℃.
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