CN113428861A - Methane and hydrogen sulfide reforming hydrogen production process - Google Patents
Methane and hydrogen sulfide reforming hydrogen production process Download PDFInfo
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- CN113428861A CN113428861A CN202110815280.1A CN202110815280A CN113428861A CN 113428861 A CN113428861 A CN 113428861A CN 202110815280 A CN202110815280 A CN 202110815280A CN 113428861 A CN113428861 A CN 113428861A
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- 229910052739 hydrogen Inorganic materials 0.000 title claims abstract description 56
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 title claims abstract description 55
- 239000001257 hydrogen Substances 0.000 title claims abstract description 55
- RWSOTUBLDIXVET-UHFFFAOYSA-N Dihydrogen sulfide Chemical compound S RWSOTUBLDIXVET-UHFFFAOYSA-N 0.000 title claims abstract description 42
- VNWKTOKETHGBQD-UHFFFAOYSA-N methane Chemical compound C VNWKTOKETHGBQD-UHFFFAOYSA-N 0.000 title claims abstract description 42
- 229910000037 hydrogen sulfide Inorganic materials 0.000 title claims abstract description 41
- 238000004519 manufacturing process Methods 0.000 title claims abstract description 18
- 238000002407 reforming Methods 0.000 title claims abstract description 10
- 239000007789 gas Substances 0.000 claims abstract description 144
- 238000010521 absorption reaction Methods 0.000 claims abstract description 112
- 239000007788 liquid Substances 0.000 claims abstract description 105
- 239000012495 reaction gas Substances 0.000 claims abstract description 64
- 239000002994 raw material Substances 0.000 claims abstract description 50
- 238000006555 catalytic reaction Methods 0.000 claims abstract description 42
- 238000000034 method Methods 0.000 claims abstract description 42
- 229910052717 sulfur Inorganic materials 0.000 claims abstract description 37
- 239000011593 sulfur Substances 0.000 claims abstract description 37
- NINIDFKCEFEMDL-UHFFFAOYSA-N Sulfur Chemical compound [S] NINIDFKCEFEMDL-UHFFFAOYSA-N 0.000 claims abstract description 34
- 238000001704 evaporation Methods 0.000 claims abstract description 11
- 230000008020 evaporation Effects 0.000 claims abstract description 8
- OKKJLVBELUTLKV-UHFFFAOYSA-N Methanol Chemical compound OC OKKJLVBELUTLKV-UHFFFAOYSA-N 0.000 claims description 147
- 239000002918 waste heat Substances 0.000 claims description 27
- 238000011084 recovery Methods 0.000 claims description 26
- 230000008569 process Effects 0.000 claims description 18
- 238000001179 sorption measurement Methods 0.000 claims description 13
- HZAXFHJVJLSVMW-UHFFFAOYSA-N 2-Aminoethan-1-ol Chemical compound NCCO HZAXFHJVJLSVMW-UHFFFAOYSA-N 0.000 claims description 8
- ZBCBWPMODOFKDW-UHFFFAOYSA-N diethanolamine Chemical compound OCCNCCO ZBCBWPMODOFKDW-UHFFFAOYSA-N 0.000 claims description 4
- CRVGTESFCCXCTH-UHFFFAOYSA-N methyl diethanolamine Chemical compound OCCN(C)CCO CRVGTESFCCXCTH-UHFFFAOYSA-N 0.000 claims description 4
- QGJOPFRUJISHPQ-UHFFFAOYSA-N Carbon disulfide Chemical compound S=C=S QGJOPFRUJISHPQ-UHFFFAOYSA-N 0.000 description 133
- 238000006243 chemical reaction Methods 0.000 description 67
- 238000000926 separation method Methods 0.000 description 37
- 230000008929 regeneration Effects 0.000 description 18
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- 239000000047 product Substances 0.000 description 15
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- 238000010438 heat treatment Methods 0.000 description 13
- 230000002829 reductive effect Effects 0.000 description 13
- UGFAIRIUMAVXCW-UHFFFAOYSA-N Carbon monoxide Chemical compound [O+]#[C-] UGFAIRIUMAVXCW-UHFFFAOYSA-N 0.000 description 12
- 239000002253 acid Substances 0.000 description 12
- 239000003546 flue gas Substances 0.000 description 12
- 239000002250 absorbent Substances 0.000 description 9
- 230000002745 absorbent Effects 0.000 description 9
- 239000003345 natural gas Substances 0.000 description 9
- 239000006227 byproduct Substances 0.000 description 8
- 239000004230 Fast Yellow AB Substances 0.000 description 7
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- 238000005406 washing Methods 0.000 description 5
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- 230000000052 comparative effect Effects 0.000 description 4
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- 238000000354 decomposition reaction Methods 0.000 description 4
- 239000000542 fatty acid esters of ascorbic acid Substances 0.000 description 4
- 239000004172 quinoline yellow Substances 0.000 description 4
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 description 3
- CURLTUGMZLYLDI-UHFFFAOYSA-N Carbon dioxide Chemical compound O=C=O CURLTUGMZLYLDI-UHFFFAOYSA-N 0.000 description 3
- 239000003570 air Substances 0.000 description 3
- 208000012839 conversion disease Diseases 0.000 description 3
- 238000004090 dissolution Methods 0.000 description 3
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- 239000004149 tartrazine Substances 0.000 description 3
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 description 3
- 239000004261 Ascorbyl stearate Substances 0.000 description 2
- 230000002378 acidificating effect Effects 0.000 description 2
- QVGXLLKOCUKJST-UHFFFAOYSA-N atomic oxygen Chemical compound [O] QVGXLLKOCUKJST-UHFFFAOYSA-N 0.000 description 2
- 229910052799 carbon Inorganic materials 0.000 description 2
- 239000003054 catalyst Substances 0.000 description 2
- 230000003197 catalytic effect Effects 0.000 description 2
- 238000001833 catalytic reforming Methods 0.000 description 2
- 238000007906 compression Methods 0.000 description 2
- 238000004821 distillation Methods 0.000 description 2
- 230000000694 effects Effects 0.000 description 2
- 239000011344 liquid material Substances 0.000 description 2
- 230000007246 mechanism Effects 0.000 description 2
- 229910052760 oxygen Inorganic materials 0.000 description 2
- 239000001301 oxygen Substances 0.000 description 2
- 239000002151 riboflavin Substances 0.000 description 2
- PPASLZSBLFJQEF-RKJRWTFHSA-M sodium ascorbate Substances [Na+].OC[C@@H](O)[C@H]1OC(=O)C(O)=C1[O-] PPASLZSBLFJQEF-RKJRWTFHSA-M 0.000 description 2
- 230000001502 supplementing effect Effects 0.000 description 2
- 239000002699 waste material Substances 0.000 description 2
- 239000004229 Alkannin Substances 0.000 description 1
- 239000004283 Sodium sorbate Substances 0.000 description 1
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 description 1
- 239000012267 brine Substances 0.000 description 1
- 238000004364 calculation method Methods 0.000 description 1
- 150000001721 carbon Chemical class 0.000 description 1
- 239000001569 carbon dioxide Substances 0.000 description 1
- 229910002092 carbon dioxide Inorganic materials 0.000 description 1
- QGJOPFRUJISHPQ-NJFSPNSNSA-N carbon disulfide-14c Chemical compound S=[14C]=S QGJOPFRUJISHPQ-NJFSPNSNSA-N 0.000 description 1
- 239000007795 chemical reaction product Substances 0.000 description 1
- 230000002860 competitive effect Effects 0.000 description 1
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- 229910052593 corundum Inorganic materials 0.000 description 1
- 239000010779 crude oil Substances 0.000 description 1
- 230000008021 deposition Effects 0.000 description 1
- 238000010586 diagram Methods 0.000 description 1
- 238000002474 experimental method Methods 0.000 description 1
- 238000005187 foaming Methods 0.000 description 1
- 239000000446 fuel Substances 0.000 description 1
- 230000014509 gene expression Effects 0.000 description 1
- 230000036541 health Effects 0.000 description 1
- 150000002431 hydrogen Chemical class 0.000 description 1
- 239000012535 impurity Substances 0.000 description 1
- 230000002401 inhibitory effect Effects 0.000 description 1
- 239000012263 liquid product Substances 0.000 description 1
- 239000012528 membrane Substances 0.000 description 1
- 238000002156 mixing Methods 0.000 description 1
- 239000000203 mixture Substances 0.000 description 1
- 229910052961 molybdenite Inorganic materials 0.000 description 1
- CWQXQMHSOZUFJS-UHFFFAOYSA-N molybdenum disulfide Chemical compound S=[Mo]=S CWQXQMHSOZUFJS-UHFFFAOYSA-N 0.000 description 1
- 229910052982 molybdenum disulfide Inorganic materials 0.000 description 1
- 239000003921 oil Substances 0.000 description 1
- 238000005457 optimization Methods 0.000 description 1
- 239000011148 porous material Substances 0.000 description 1
- 238000002360 preparation method Methods 0.000 description 1
- 238000007670 refining Methods 0.000 description 1
- 238000010992 reflux Methods 0.000 description 1
- HPALAKNZSZLMCH-UHFFFAOYSA-M sodium;chloride;hydrate Chemical compound O.[Na+].[Cl-] HPALAKNZSZLMCH-UHFFFAOYSA-M 0.000 description 1
- 239000007787 solid Substances 0.000 description 1
- 239000007858 starting material Substances 0.000 description 1
- 238000003860 storage Methods 0.000 description 1
- 230000001360 synchronised effect Effects 0.000 description 1
- 239000002341 toxic gas Substances 0.000 description 1
- 230000036967 uncompetitive effect Effects 0.000 description 1
- 239000002912 waste gas Substances 0.000 description 1
- 229910001845 yogo sapphire Inorganic materials 0.000 description 1
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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
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- 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
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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
- 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
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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
- 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
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- 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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- General Health & Medical Sciences (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The invention discloses a hydrogen production process by reforming methane and hydrogen sulfide. The method comprises the following steps: (1) making the temperature of 600-4And H2Carrying out catalytic reaction on the raw material gas of S to obtain a reaction gas I; (2) separating the reaction gas I to obtain sulfur and reaction gas II; (3) the reaction gas II is subjected to flash evaporation to obtain crude CS2Liquid and reaction gas III; (4) rectifying the crude CS2Obtaining liquid CS with the mass fraction of not less than 99.5 percent from the liquid2(ii) a Contacting the reaction gas III with an absorption liquid to obtain hydrogen-rich gas and an absorption liquid; and (5) separating the hydrogen-rich gas to obtain hydrogen with the volume fraction of more than 99.99 percent and the desorbed gas.
Description
Technical Field
The invention relates to hydrogen preparation by hydrogen sulfide conversion, in particular to a process and a process for preparing hydrogen by reforming methane and hydrogen sulfide.
Background
China is rich in high-sulfur natural gas resources, the natural gas resources are mainly distributed in Bohai Bay basins and Sichuan basins, and the content of hydrogen sulfide (H2S) generally exceeds 5%. Meanwhile, in the process of upgrading and refining the sulfur-containing crude oil, a large amount of by-product H is generated in the hydrotreating process2And S. In the world H2The S yield has reached 1 million tons per year. H2S is a highly corrosive and toxic gas, which brings a lot of problems to the exploitation, transportation, storage and the like of sulfur-containing oil and natural gas, and the discharge of S also causes harm to the natural environment and human health, so that the S must be recycled.
At present, H is prepared by a physical or chemical method mainly used in the industry2S is separated and enriched, and then H is treated by adopting Claus process or LO-CAT process2S is treated and converted to H2S is converted into sulfur and water. Although these processes can recover H2Sulfur in S but simultaneously H2The hydrogen in the S molecule is oxidized to water, resulting in waste of hydrogen resources. Hydrogen is a new clean energy source, and is expected to play a significant role in the future, so that it has become an important strategic resource in the world. By means of H2Although the S decomposition technology has simple process, the reaction conversion rate and the hydrogen yield are very low due to the limitation of thermodynamic equilibrium, and pure H2S generally requires high temperatures above 1300 ℃ to achieve conversion rates above 50%. The high temperature causes an increase in the cost of the technology, making the technology uncompetitive. And, H2Sulfur with low added value can be produced while hydrogen is produced by decomposing S, and the sulfur is easy to block pipelines in the reaction process and causes the problems of equipment corrosion and the like, thereby increasing the operation cost. Thus, H2The S decomposition technology has not been industrialized so far.
Methane (CH)4) And H2The hydrogen production by S reforming is a brand new H2The technical route of S conversion and hydrogen production comprises the following reaction processes: CH (CH)4+H2S＝CS2+H2. And H2S decomposition phase ratio, CH4And H2The reaction conversion rate of the S reforming hydrogen production process is high under the same condition, the hydrogen yield is high, and the by-product carbon disulfide (CS) is generated2) The added value of the sulfur is also obviously higher than that of sulfur, and the former is 5 times of the latter in the current market price, so the sulfur is more competitive.
There is therefore a great need in the art to provide CH4And H2S reforming hydrogen production process.
Disclosure of Invention
The present invention is designed and developed to have a high concentration of H2S is taken as a raw material and is utilized together with CH4Reaction, synchronous production of hydrogen and CS2The process and the technology provide a technical scheme for realizing industrialization of the route.
The invention provides a method for preparing hydrogen by reforming methane and hydrogen sulfide, which comprises the following steps:
(1) making the temperature of 600-4And H2Carrying out catalytic reaction on the raw material gas of S to obtain a reaction gas I;
(2) separating the reaction gas I to obtain sulfur and reaction gas II;
(3) the reaction gas II is subjected to flash evaporation to obtain crude CS2Liquid and reaction gas III;
(4) rectifying the crude CS2Obtaining liquid CS with the mass fraction of not less than 99.5 percent from the liquid2(ii) a Contacting the reaction gas III with an absorption liquid to obtain hydrogen-rich gas and an absorption liquid; and
(5) separating the hydrogen-rich gas to obtain hydrogen with the volume fraction of more than 99.99 percent and the desorbed gas.
In another embodiment, H is in the feed gas of step (1)2S and CH4Is 1.0 to 5.0, preferably 2.0 to 3.0.
In another embodiment, the catalytic reaction temperature in step (1) is 800-.
In another embodiment, in step (3), the reaction gas II is subjected to a secondary flash evaporation.
In another embodiment, the absorbing solution in step (4) is selected from one or two or more of the following: low temperature methanol, ethanolamine, diethanolamine and N-methyldiethanolamine, low temperature methanol being preferred.
In another embodiment, the molar ratio of the reaction gas III to the absorption liquid in step (4) is 2.18 or more.
In another embodiment, the temperature of the absorption liquid in step (4) is in the range of-20 to-60 ℃.
In another embodiment, step (4) is performed at H2And (S) enabling the reaction gas III to contact with the absorption liquid in the absorption tower, wherein the pressure of the absorption tower is 2-5 MPa.
In another embodiment, step (5) uses a Pressure Swing Adsorption (PSA) unit to separate the hydrogen-rich gas.
In another embodiment, the preheating of the CH-containing stream by the waste heat recovery system4And H2The temperature of the S raw material gas is 600-900 ℃.
Accordingly, the present invention provides a CH4And H2S reforming hydrogen production process.
Drawings
FIG. 1 is a schematic diagram of a hydrogen production process system and process flow according to the present invention.
Detailed Description
The inventor of the present invention has conducted extensive and intensive studies and found that when a process system is used for treating refinery acidic tail gas and natural gas with high sulfur content, acidic component H can be treated2S, can also fully recover H2H and S elements in S to obtain hydrogen and CS with high added value2The product changes waste into valuable, and improves the utilization rate of resources. And H2Compared with the direct S decomposition technology, the invention adds CH4Let H stand for2S and CH4The catalytic reforming reaction is carried out, so that the Gibbs free energy of the reaction is reduced, the reaction condition is milder, the reaction equilibrium conversion rate is improved, and the H is ensured2The S hydrogen production process is more economical. On the basis of this, the present invention has been completed.
Specifically, the invention provides a method for synchronously producing hydrogen and CS2The method comprises the following steps:
in the first step, CH is mixed according to a certain proportion4Gas and H2Preheating S gas to 600-900 ℃, and then carrying out catalytic reaction to obtain reaction gas I;
secondly, separating the reaction gas I to obtain a byproduct sulfur and a reaction gas II;
thirdly, the reaction gas II is subjected to flash evaporation to obtain reaction gas III and crude CS2A liquid; coarse CS2Rectifying the liquid to obtain liquid CS with the mass fraction of not less than 99.5%2Product, H obtained simultaneously2S gas is recycled to the raw material gas for re-reaction;
the fourth step, reacting gas III in H2S, carrying out countercurrent absorption by an absorption tower to obtain hydrogen-rich gas at the tower top and obtain absorption-rich liquid at the tower bottom; the rich absorption liquid is separated to obtain noncondensable gas which flows back to H2S, an absorption tower, wherein liquid materials obtained by gas-liquid separation enter an absorption liquid regeneration tower to obtain regenerated absorption liquid and stripping gas;
and fifthly, separating the hydrogen-rich gas obtained in the fourth step to obtain hydrogen with the volume fraction of more than 99.99% and the desorbed gas.
In the first step, H is contained in the raw material gas2S/CH4The molar ratio is 1.0 to 5.0, preferably 2.0 to 3.0. H2A proper excess of S favors the production of more hydrogen while inhibiting carbon deposition, but H2A large excess of S leads to H2The conversion rate of S is reduced, the circulation ratio is too large, the equipment investment is greatly increased, and the operation cost is increased.
In one embodiment of the present invention, the feedstock H2S gas, H2The volume content of S is 10-100%, and the S can be high-sulfur natural gas, refinery acid gas and H-containing natural gas2S waste gas and the like, wherein the raw material pretreatment unit is mainly used for pretreating H2S raw material to remove oxygen (O2), water (H2O) and carbon dioxide (CO)2) And solid particulate impurities.
In one embodiment of the invention, the feed gas is preheated in a way that can provide waste heat; the waste heat of the high-temperature flue gas and the reaction gas is utilized to preheat the raw material gas, air and fuel gas, and the flue gas and the reaction gas are cooled.
The temperature of the catalytic reaction in the first step is kept at 800-1000 ℃, and the pressure of the catalytic reaction is keptA force of 0.05-0.15MPa (normal pressure operation, i.e. around one atmosphere); catalysts used include, but are not limited to, MoS2/Al2O3。
The manner of separating the by-product sulfur in the reaction gas I which can be employed in the above-mentioned second step includes, but is not limited to, the use of a sulfur trap, a sulfur condenser, etc.
The sulfur as a by-product in the second step may be liquid sulfur (i.e., liquid sulfur).
In one embodiment of the invention, the reaction gas I can be recycled into the sulfur trap after heat is recovered.
In one embodiment of the invention, the heat of the reaction gas and the flue gas is recovered and used 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 pressure and temperature reduction for flash evaporation, and preferably secondary flash evaporation is carried out; the rectification may use CS2The rectifying tower has an operating pressure of 1.0-2.0 MPa, a temperature of 10-50 ℃ and a reflux ratio of 0.1-0.3.
In one embodiment of the present invention, the coarse CS2Passing the liquid through CS2The rectifying tower obtains liquid CS with the mass fraction of not less than 99.5 percent2H obtained from the top of the column2And recycling the S gas back to the raw material gas for re-reaction.
In the fourth step, since H is generated in the first catalytic reaction2The conversion of S is generally less than 50%, leading to the entry of H2H in S absorption tower gas2The S concentration is higher (generally higher than 5%, and the rest is H)2、CH4、N2) Therefore, H must be substituted2And S is separated from the mixed gas and then returns 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 a low-temperature methanol washing method is adopted to separate H with higher concentration2And S. Compared with chemical absorption liquid such as ethanolamine, diethanolamine, N-methyldiethanolamine and the like, the absorption performance is better, the low-temperature methanol gas washing liquid ratio can reach more than 2.18mol of acid gas/mol of methanol, and the high gas-liquid ratio is effectiveThe equipment volume is reduced, and the consumption of absorption liquid is reduced, so that the equipment investment cost and the operation cost are reduced. On the other hand, since no low-temperature methanol washing method has been used for higher concentrations of H2Separation of S, and therefore applying this method to the present invention, presents challenges, first of all requiring high solubility to sufficiently absorb large amounts of H2S gas, the second requirement is that the heat of solution influence is small to avoid large amounts of H2The temperature rise caused by S dissolution affects the absorption effect, for which reason the invention is based on H2The influence of the solubility and the dissolution heat of S in the low-temperature methanol is analyzed, the temperature and the pressure of the absorption liquid are optimized, and the low-temperature methanol has higher concentration H within the range of the determined operation conditions at present2The absorption efficiency of S is high, and the treatment capacity is large.
In one embodiment of the invention, in H2And (3) contacting the reaction gas III with an absorption liquid in the S absorption tower, wherein the temperature of the absorption liquid can be-20 to-60 ℃, and the pressure of the absorption tower can be 2 to 5 MPa. The preferred temperature is-40 ℃ and the pressure is 3 MPa.
In an 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 desorption gas, wherein the desorption gas comprises H as a main component2And S, recycling the S to the reaction system for re-reaction.
The above-mentioned fifth step may be carried out by a method for separating a hydrogen-rich gas including, but not limited to, cryogenic rectification, membrane separation, conventional adsorption separation, etc.
In one embodiment of the present invention, a well-performing activated carbon material is selected (e.g., without limitation, a specific surface area greater than 2000m is selected2G, pore volume is more than 1.8cm3Activated carbon material/g), separating the hydrogen-rich gas using a Pressure Swing Adsorption (PSA) apparatus to obtain a desorption gas having CH as a major component4And the hydrogen is recycled to the raw material for re-reaction, and high-purity hydrogen with the volume fraction of more than 99.99 percent is obtained.
The steps provided by the invention provide a CS2And H2S separation and recovery method. Containing CS obtained in the above second step2、H2And unreacted H2S、CH4Mixed gases (i.e. reaction gases)II) H) which is not condensable after cooling and compression2And CH4Still in the gaseous state, CS2Is completely liquefied (the liquefaction capacity depends on the temperature and the CS2Physical properties and partial pressure) of the composition and due to CS2Liquid pair H2S has greater solubility, so part of H2S will dissolve in CS2Medium (dissolving power depends on temperature, H)2S partial pressure and H2S and CS2Affinity between) that causes liquefaction of CS2In which the moiety H2S, resulting in CS2The purity of the liquid product is low. To separate CS2H in liquids2S, a rectification method is generally adopted. However, the direct distillation process is energy-intensive, in particular in the case of the invention owing to unreacted H2Higher S concentration, resulting in dissolution in CS2H in liquids2The S content is higher, and further the energy consumption of the direct distillation method is increased. Therefore, the invention innovatively adopts a low-temperature flash evaporation method, so that the CS2H dissolved in liquid2S escapes rapidly, thereby improving CS2The purity of the liquid reduces the energy consumption burden of the next step of rectification and separation. Simultaneously, CS is also improved through secondary flash evaporation2The recovery rate of (1).
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, simultaneously generates hydrogen and carbon disulfide, can treat acid tail gas of a refinery and natural gas with high sulfur content, and synchronously produces green energy H2And high value-added product CS2。
The invention also provides a hydrogen production process system which comprises a start-up heating furnace, a waste heat recovery system, a catalytic reaction furnace, a sulfur trap, a gas compressor, a secondary flash separator, a CS2Rectifying column, H2S absorption tower, absorption liquid regenerator and PSA pressure swing adsorption equipment. The raw material gas feeding pipeline is connected with the waste heat recovery system after passing through the start-up heating furnace, and the start-up heating furnace is only used during start-up; the raw material gas passes through a waste heat recovery system and then is connected with an inlet at the upper end of the catalytic reaction furnace tube pass; the reaction gas at the outlet at the lower end of the catalytic reaction furnace is connected with a gas compressor through a waste heat recovery system and is compressed and reducedAfter the temperature is raised, the reaction kettle is connected with an inlet of a secondary flash separator; lower liquid phase outlet of secondary flash separator and CS2The inlet of the rectifying tower is connected, and the upper gas outlet is cooled and connected with the H2The lower end inlet of the S absorption tower is connected; CS2The outlet of the rectifying tower top is connected with a raw material gas pipeline and circulated back to the reaction system, and the outlet of the tower kettle is connected with the CS2Finished tank connection to obtain 99.5 wt.% CS2Producing a product; h2The 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 material gas pipeline and circulates back to the reaction system, and the hydrogen gas outlet of the PSA pressure swing adsorption device obtains a hydrogen product with the purity of 99.99 percent; h2A rich liquid outlet at the lower end of the S absorption tower is connected with a gas-liquid separation tank and 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 the methanol regeneration tower through heat exchange and temperature reduction; the gas outlet at the top of the methanol regeneration tower is connected with a raw material gas pipeline and circulated back to the reaction system, and the outlet of the tower kettle is connected with the liquid 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 tubular reactor, the raw material gas is led away from a tube side, the fuel gas is led away from a shell side, the catalytic reforming reaction in the tube side belongs to a strong endothermic reaction, and the heat is provided for the reaction by burning the fuel gas through the shell side. The high-temperature reaction gas at the outlet of the reaction furnace tube side and the high-temperature flue gas at the outlet of the shell side are used for preheating feed gas, byproduct steam and the like, so that comprehensive utilization of energy is realized, and the byproduct steam can provide heat for reboilers of rectifying towers of the separation unit. Through calculation, the high-temperature reaction gas at the tube pass outlet and the high-temperature flue gas at the shell pass outlet can recover waste heat, can be sufficiently used for preheating raw material gas, fuel gas and air, and ensure proper heat exchange temperature difference.
In one embodiment of the invention, the hydrogen production process system comprises a start-up heating furnace F-101, a high-temperature catalytic reaction furnace R-101, a sulfur trap 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 exchangers). See figure 1.
In one embodiment of the present invention, the following process using the flow system is provided as follows:
(1) starting material CH4And H2And mixing the S gas according to a certain proportion, and then, entering a start-up heating furnace F-101 for pretreatment. Preheating a raw material gas to 600-900 ℃, and then feeding the raw material gas into 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) separating a byproduct sulfur in the reaction gas I from the reaction gas I obtained in the step (2) through a sulfur trap V-101 to obtain liquid sulfur and a reaction gas II;
(4) pressurizing and cooling the reaction gas II obtained in the step (3), then feeding the reaction gas II into flash separation tanks V-201 and V-202, obtaining reaction gas III at the tower top, and obtaining crude CS at the tower bottom2A liquid; coarse CS2Liquid in CS2Rectifying in a rectifying tower T-201 to obtain liquid CS with the mass fraction not less than 99.5 percent in the tower bottom2Product, H obtained at the top of the column2S gas is recycled to the raw material gas for re-reaction;
(5) the reaction gas III obtained in the step (4) enters H2S recovery unit, reaction gas III in H2S, carrying out countercurrent absorption on the absorption tower T-301 to obtain hydrogen-rich gas at the tower top and obtain absorption-rich liquid at the tower bottom; the rich absorption liquid is decompressed and then enters a gas-liquid separation tank V-301 to obtain noncondensable gas which flows back to H2S, an absorption tower T-301, wherein liquid materials obtained by gas-liquid separation enter an absorption liquid regeneration tower T-302 to obtain regenerated absorption liquid and stripping gas;
(6) the main component of the desorption gas obtained in the step (5) is H2S, circulating the F-101 to the start-up heating furnace to enter the catalytic reaction furnace R-101 for re-reaction;
(7) allowing the hydrogen-rich gas obtained in the step (5) to enter a Pressure Swing Adsorption (PSA) device to obtain hydrogen and desorbed gas with volume fraction of more than 99.99%;
(8) the main component of the desorption gas obtained in the step (7) is CH4And the reaction product is circulated back to the high-temperature catalytic reaction furnace R-101 for re-reaction.
The heating fuel of the catalytic reaction furnace R-101 is natural gas, the fuel gas passes through the shell pass, and the raw material gas passes through the tube pass. The start-up heating furnace F-101 is only used during start-up, and a waste heat recovery system is used for heating the feed gas during normal operation of the equipment. The preheating temperature range of the feed gas is 600-900 ℃. The preheating temperature is lower than 600 ℃ to be unfavorable for reaction conversion, but the preheating temperature is too high to cause difficulty in 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 tubular reactor; the raw gas enters the tube side from the inlet at the top of the reactor, and the fuel gas and air enter the shell side from the inlet at the lower part. The start-up heating furnace F-101 is only used during start-up, and raw materials are preheated through a waste heat recovery system during 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 and E-004. E-001-E-004 is used for recovering the flue gas waste heat at the outlet of the furnace shell pass of the R-101 high-temperature catalytic reaction furnace, and E-101-E-105 is used for recovering the flue gas waste heat at the outlet of the tube pass reaction gas. The waste heat recovery system can preheat the raw material gas to 900 ℃, preheat the air to 450 ℃ and preheat the fuel gas to 300 ℃. Meanwhile, 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 firstly exchanges heat with E-104 and then enters the shell pass R-001 of the high-temperature catalytic reaction furnace R-101 after exchanging heat with E-003, and fuel gas CH4 directly enters the shell pass R-001 of the high-temperature catalytic reaction furnace R-101 after exchanging heat with E-105, instead of passing air or fuel gas through other heat exchangers.
In one embodiment of the invention, the raw material gas enters a tube pass in a high-temperature catalytic reaction furnace R-101 for catalytic reaction after sequentially passing through heat exchangers E-004, E-102, E-002, E-101 and E-001 for heat exchange.
In one embodiment of the invention, the reaction gas at the outlet of the tube pass 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 R-001 shell side of the high-temperature catalytic reaction furnace R-101 is cooled by the heat exchangers E-001-E-004 and then is evacuated.
According to the invention, the flows in different temperature intervals are matched by calculating the energy of cold and hot flows, and a heat exchange scheme is finally determined through complex heat exchange network optimization, so that the full and efficient utilization of the energy of the whole process is realized, and the system economy is improved.
In one embodiment of the invention, the reaction temperature of the high-temperature catalytic reaction furnace R-101 is 800-1000 ℃, the reaction pressure is 0.1-0.5 MPa, preferably the reaction temperature is 900-1000 ℃, and the reaction pressure is 0.1-0.3 MPa. The raw material gas ratio is H2S/CH4＝1.0～5.0。
In one embodiment of the invention, the high-temperature catalytic reaction furnace R-101 belongs to a tubular reactor, a tube side is filled with a catalyst, and a feed gas enters the tube side from an inlet at the top end of the reactor; fuel gas and air (O)2) Enters the shell pass from the inlet at the lower end, and provides heat and high-temperature environment for reaction by burning fuel gas.
In one embodiment of the invention, the outlet temperature of the start-up 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 H2The system for recovering the waste heat removed between the gas-phase inlet pipelines of the S rectifying tower T-201 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 H2The S rectifying tower is connected with a T-201 pipeline, and a V-201 liquid phase outlet of the flash separation tank is connected with the CS2The rectification tower T-101 is connected by pipelines.
CS-containing gas from the upper part of the sulfur trap V-1012、H2And unreacted H2S、CH4The mixed gas is cooled by heat exchange of heat exchangers E-106 and E-107 and compressed by a compressor C-101, and then is non-condensable H2And CH4Still in the gaseous state, CS2It is completely liquefied.
In one embodiment of the invention, the compressor C-101 adopts three-stage compression, and the outlet pressure of the compressor is 2-5 MPa, preferably 3 MPa. And a heat exchanger E-106 is arranged on a pipeline between the inlet of the compressor C-101 and the outlet of the sulfur trap V-101. And a heat exchanger E-107 is arranged on a pipeline between the outlet of the compressor C-101 and the inlet of the flash separation tank V-201.
In one embodiment of the invention, an inlet of a used carbon disulfide rectifying tower T-201 is connected with liquid outlets of secondary flash separators V-201 and V-202, 90% of crude carbon disulfide is recovered by the primary flash separator V-201, 99% of crude carbon disulfide is recovered by the secondary flash separator V-202, part of unreacted hydrogen sulfide is obtained at the top of the rectifying tower T-201 and 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 CS2And H2The S accounts for 37.13 wt.% and 52.93 wt.%, respectively, the temperature of the reaction gas is reduced to 20-25 ℃, and CS is separated by first-stage flash evaporation2The recovery rate of the product reaches 90 percent; further cooling the residual gas phase to-5-8 deg.C, performing secondary flash separation, and finally CS2The recovery rate of the product reaches 99 percent. Crude CS obtained by two-stage flash separation2Is sent into the CS2The rectifying tower T-201 is used for rectifying to finally obtain 99.5 wt.% of CS2And (5) producing the 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 example, 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 3 MPa.
In another preferred example, the pipeline connecting 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 has a gas phase outlet and H2The 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-7 to-30 ℃, preferably-20 ℃.
In one embodiment of the present invention, the H2The temperature of an absorption liquid of the S absorption tower T-301 is-20 to-70 ℃, and the pressure is 2 to 5 MPa.
In another preferred embodiment, the H2T-301 operating pressure selection of S absorption towerSelecting the pressure of 3MPa, and selecting the temperature of the absorption liquid to be-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 the tower bottom of the regeneration tower and the connecting pipeline at the tower bottom of the rectification tower, and a gas outlet at the tower top of the regeneration tower is connected with a raw material gas pipeline.
E.g. H2And a pipeline for connecting a rich liquid phase outlet of the S absorption tower T-301 with an inlet of the absorption liquid regeneration tower T-302 is provided with a gas-liquid separation tank V-301 and a heat exchanger E-304, and the gas-liquid separation tank is used for removing non-condensable gas in the rich absorption liquid. Gas phase outlet of the gas-liquid separation tank is recycled to H2A gas phase inlet of the S 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, H2And a rich gas outlet of the S absorption tower T-301 enters a PSA unit for separating methane and hydrogen, the separated methane is circulated back to the reaction system for re-reaction, and hydrogen with the product purity of 99.99% is obtained at the same time.
In one embodiment of the present invention, the operating pressure of the absorption liquid regeneration column T-302 may be 0.05 to 0.5MPa, preferably 0.1 to 0.2 MPa.
In one embodiment of the invention, the liquid phase outlet of the methanol regeneration tower T-302 and the H2The 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 ℃, and is preferably-40 ℃. Said H2And the S absorption tower is also provided with an absorption liquid supplementing system for supplementing the absorption liquid lost in the operation.
The invention has the particularity that the invention not only relates to high-temperature reaction (1000 ℃) but also needs low-temperature operation (minus 60 ℃), has wide temperature range and more operation units, and is a difficult problem on how to realize comprehensive and efficient utilization of energy. The invention optimizes the process heat exchange network, takes economy or low energy consumption as a target, adopts a problem table method to draw temperature area maps of each operation unit, performs mutual matching of streams by calculating the energy of cold and hot material flows, and finally determines a heat exchange scheme, thereby realizing the full and efficient utilization of the energy of the whole process and improving the system economy.
To make the features and effects of the present invention comprehensible to those skilled in the art, general description and definitions are made below with reference to terms and expressions mentioned in the specification and claims. 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 to which this invention belongs.
The theory or mechanism described and disclosed herein, whether correct or incorrect, should not limit the scope of the present invention in any way, i.e., the present disclosure may be practiced without limitation to any particular theory or mechanism.
All features defined herein as numerical ranges or percentage ranges, such as values, amounts, levels and concentrations, are for brevity and convenience only. Accordingly, the description of numerical ranges or percentage ranges should be considered to cover and specifically disclose all possible subranges and individual numerical values (including integers and fractions) within the range.
The features mentioned above with reference to the invention, or the features mentioned with reference to the embodiments, can be combined arbitrarily. All features disclosed in this specification may be combined in any combination, provided that there is no conflict between such features and the combination, and all possible combinations are to be considered within the scope of the present specification. Each feature disclosed in this specification may be replaced by an alternative feature serving the same, equivalent, or similar purpose. Thus, unless expressly stated otherwise, the features disclosed are merely generic examples of equivalent or similar features.
The main advantages of the invention are:
1. when the acid tail gas or the high-sulfur natural gas of the refinery is treated, H can be treated2S is discharged and H can be efficiently recovered2H and S elements in S synchronously generate green energy H2And high value-added product CS2。
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 separation of CS by secondary flash separation2Separating, and rectifying to obtain refined CS2Producing a product; the separation method not only reduces energy consumption, but also efficiently recovers CS2Effectively improve CS2The recovery rate of (1).
4. Separation of higher concentration H by low temperature methanol wash2S, the consumption of the absorption liquid is greatly reduced, and the volumes of the absorption tower and the regeneration tower are effectively reduced, so that the equipment cost and the operation cost are reduced; the low-temperature methanol washing belongs to physical absorption, and compared with chemical absorption, the low-temperature methanol washing can also avoid the foaming phenomenon in the absorption process.
The invention will be further illustrated with reference to the following specific examples. It should be understood that these examples are for illustrative purposes only and are not intended to limit the scope of the present invention. The experimental procedures, in which specific conditions are not noted in the following examples, are generally carried out according to conventional conditions or according to conditions recommended by the manufacturers. All percentages, ratios, proportions, or parts are by weight unless otherwise specified. The units in weight volume percent in the present invention are well known to those skilled in the art and refer to, for example, the weight (g) of solute in 100 ml of solution. 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 embodiments and materials described herein are intended to be exemplary only.
The invention provides a hydrogen production process and a hydrogen production process, wherein the process system comprises a start-up heating furnace F-101, a waste heat recovery system E101-E105, an E001-E-004, a high-temperature catalytic reaction furnace R-101 (R-001 in the attached figure 1 represents the shell pass of the reaction furnace), a sulfur trap V-101, a gas compressor C-101, a secondary flash separation tank V-201 and V-202, a CS2Rectifying tower T-201, H2S absorption tower T-301, absorption liquid (methanol) regenerationA tower T-302 and a PSA pressure swing adsorption device T-401. The raw material gas passes through a start-up 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 pass of a high-temperature catalytic reaction furnace R-101; the fuel gas is connected with an inlet at the lower end of the shell pass of the high-temperature catalytic reaction furnace R-101 through a waste heat recovery system E-104; and the air or the oxygen is connected with an 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 pass outlet of the R-101 of the high-temperature catalytic reaction furnace is subjected to heat exchange and temperature reduction by the waste heat recovery systems E-101-E-105 and then is connected with the inlet at the upper end of the sulfur trap V-101; high-temperature flue gas at the shell pass outlet of the high-temperature catalytic reaction furnace R-101 is subjected to heat exchange and temperature reduction by a waste heat recovery system E-001-E-004 and then is discharged. The outlet of the sulfur catcher V-101 is connected with the inlet of a three-stage compressor C-101 after heat exchange of E-106, the outlet of the compressor C-101 is connected with the inlet of a second-stage flash separation V-201 after water-cooling heat exchange of E-107 and heat exchange of E-108 frozen brine, the gas outlet above the flash separation V-201 is connected with the inlet of a second-stage flash separation V-202, and the liquid outlets of the second-stage flash separation are connected with the CS2CS of rectifying tower T-2012An inlet connection; the gas outlet above the flash separator V-202 passes through a heat exchanger E-301 and then is mixed with H2The inlet of the S absorption tower T-301 is connected. CS2The top outlet of the rectifying tower T-201 is connected with a raw material gas pipeline to realize partial raw material H2Recycling S, and obtaining CS with the purity of 99.5 wt.% at the tower bottom outlet of the rectifying tower2And (5) producing the product. H2The top inlet of the S absorption tower T-301 is connected with the regenerated methanol pipeline and the supplemented methanol pipeline of the absorption liquid (methanol) regeneration tower T-302 tower kettle outlet, and the H absorption tower is connected with the regenerated methanol pipeline and the supplemented methanol pipeline2The 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; said H2The rich liquid outlet of the tower bottom of the S absorption tower T-301 is connected with the inlet of the gas-liquid separation tank V-301. A gas outlet above the gas-liquid separation tank V-301 and the gas outlet H2The inlet of the bottom of the S absorption tower T-301 is connected, and the liquid outlet at the lower part is connected with the inlet of the absorption liquid (methanol) regeneration tower T-302 after heat exchange by the E-304. The outlet of the top of the absorption liquid (methanol) regeneration tower T-302 is connected with a raw material gas pipeline to realize the remaining unreacted H2S, recycling; the outlet of the tower bottom of the absorption liquid (methanol) regeneration tower T-302 is circulatedPump P-301, heat exchanger E-304, heat exchanger E-305 and the above-mentioned H2The liquid inlet at the top of the S absorption tower T-301 is connected. The PSA pressure swing adsorption unit outlet obtains a hydrogen product with the purity of 99.99 percent, and the desorption gas outlet is connected with a feed gas pipeline to realize the unreacted CH4And (5) recycling.
Example 1
Treatment of feed gas CH4Concentration 100%, flow 400kmol/H, H2The S concentration is 100 percent, and the flow rate is 800kmol/H, so that H in the feed gas2S/CH4The molar ratio is 2. The raw material gas is preheated to 600 ℃, and the reaction pressure is 1 atm; the absorption tower adopts low-temperature methanol as an absorbent, and the gas-liquid ratio of the absorption tower is 2.66mol of acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 3MPa, low temperature methanol feed temperature-40 ℃. After the raw material gas passes through the catalytic reaction furnace, the outlet temperature of the reaction gas is 797 ℃, and CH4Conversion was 31.33%, H2The S conversion was 31.45%; CS with a purity of 99.5 wt.% is obtained2125.51kmol/H H with a purity of 99.99%2 11,208Nm3/h。
Example 2
Treatment of feed gas CH4Concentration 100%, flow 400kmol/H, H2The concentration of S is 10% (the rest is N)2Gas), the flow rate is 4000 kmol/h; the raw material gas is preheated to 700 ℃, and the reaction pressure is 1 atm; the absorption tower adopts low-temperature methanol as an absorbent, and the gas-liquid ratio of the absorption tower is 2.18mol of acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 3MPa, low temperature methanol feed temperature-40 ℃. After the raw material gas passes through the catalytic reaction furnace, the outlet temperature of the reaction gas is 785 ℃, and CH4Conversion 26.46%, H2The S conversion was 52.52%; CS with a purity of 99.5 wt.% is obtained2120.55kmol/H H with a purity of 99.99%2 10,656Nm3/h。
Example 3
Treatment of feed gas CH4The concentration is 10 percent (the rest is N)2Gas), flow rate of 1000kmol/H, H2S concentration 50% (the remainder being N)2Gas), the flow rate is 800 kmol/h; preheating raw material gas to 900 ℃, wherein the reaction pressure is 1 atm; the absorption tower adopts a lower towerWarm methanol is taken as absorbent, the gas-liquid ratio of the absorption tower is 2.31mol acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 3MPa, low temperature methanol feed temperature-40 ℃. The outlet temperature of the reaction gas is 829 ℃ after the raw material gas passes through the catalytic reaction furnace, and CH4Conversion 45.90%, H2The S conversion rate is 23.05%; CS with a purity of 99.5 wt.% is obtained245.41kmol/H H with a purity of 99.99%2 4,068Nm3/h。
Example 4
Treatment of feed gas CH4The concentration is 50% (the rest is N)2Gas), flow rate of 400kmol/H, H2The concentration of S is 100 percent, and the flow rate is 1000 kmol/h; preheating raw material gas to 800 ℃, wherein the reaction pressure is 1 atm; the absorption tower adopts low-temperature methanol as an absorbent, and the gas-liquid ratio of the absorption tower is 2.57mol of acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 3MPa, low temperature methanol feed temperature-40 ℃. After the raw material gas passes through the catalytic reaction furnace, the outlet temperature of the reaction gas is 886 ℃, and CH4Conversion was 57.69%, H2The S conversion was 23.1%; CS with a purity of 99.5 wt.% is obtained2115.14kmol/H H with a purity of 99.99%2 10,304Nm3/h。
Example 5
Treatment of feed gas CH4Concentration 100%, flow 400kmol/H, H2The concentration of S is 50 percent, and the flow rate is 2400 kmol/h; preheating raw material gas to 900 ℃, wherein the reaction pressure is 1 atm; the absorption tower adopts low-temperature methanol as an absorbent, and the gas-liquid ratio of the absorption tower is 2.24mol of acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 3MPa, low temperature methanol feed temperature-40 ℃. The temperature of the reaction gas outlet of the raw material gas after passing through the catalytic reaction furnace is 916 ℃, and CH4Conversion was 60.40%, H2The S conversion rate is 40.54%; CS with a purity of 99.5 wt.% is obtained2242.28kmol/H H with a purity of 99.99%2 21,720Nm3/h。
Example 6
Treatment of feed gas CH4Concentration 100%, flow 400kmol/H, H2The concentration of S is 10% (the rest is N)2Gas), the flow rate is 4000 kmol/h; raw materialsThe gas is preheated to 700 ℃, and the reaction pressure is 1 atm; the absorption tower adopts low-temperature methanol as an absorbent, and the gas-liquid ratio of the absorption tower is 4.02mol of acid gas (containing H)2S: 42.19 wt.%)/mol low temperature methanol, operating pressure 5MPa, low temperature methanol feed temperature-60 ℃. After the raw material gas passes through the catalytic reaction furnace, the outlet temperature of the reaction gas is 785 ℃, and CH4Conversion 26.46%, H2The S conversion was 52.52%; CS with a purity of 99.5 wt.% is obtained2120.55kmol/H H with a purity of 99.99%2 10,656Nm3/h。
Comparative example 1
Compared with the example 1, the preheating temperature of the raw material gas is changed to 400 ℃ under the same other conditions.
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 temperature is4Conversion was only 13.75%, H2The S conversion was only 13.77%; CS with a purity of 99.5 wt.% is obtained254.94kmol/H H with a purity of 99.99%2 4,928Nm3/h。
Comparative example 2
The reaction pressure was changed to 3MPa in the same manner as in example 1.
The outlet temperature of the reaction gas is 814 ℃ after the raw material gas passes through the catalytic reaction furnace. Although the reaction is carried out at a high temperature, CH is generated due to an excessive reaction pressure4Conversion was only 14.25%, H2The S conversion was only 14.31%; CS with a purity of 99.5 wt.% is obtained256.96kmol/H H with a purity of 99.99%2 5,107Nm3/h。
Comparative example 3
The temperature of the absorbent feed was changed to-10 ℃ in the same manner as in example 3.
The absorption tower adopts low-temperature methanol as an absorbent, and the operating pressure of the absorption tower is still 3 MPa. Because the temperature of the absorption liquid is higher, the consumption of the required absorption liquid is greatly increased from original 1124.83kmol/H to 1991.78kmol/H, and the gas-liquid ratio of the absorption tower is reduced to 1.53mol of acid gas (H)2S: 42.19 wt.%)/mol low temperature methanol.
Comparative example 4
Compared with the example 3, the other conditions are the same, and the operating pressure of the absorption tower is changed to be 1 MPa.
The absorption tower adopts low-temperature methanol as an absorbent, and the feeding temperature of the absorption tower is still-40 ℃. The required absorption liquid dosage is greatly increased from 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)2S: 42.19 wt.%)/mol low temperature methanol.
The foregoing is merely a preferred embodiment of the invention and is not intended to limit the scope of the invention, which is defined by the claims appended hereto, and any other technical entity or method that is encompassed by the claims as broadly defined herein, or equivalent variations thereof, is contemplated as being encompassed by the claims.
Claims (10)
1. A method for producing hydrogen by reforming methane and hydrogen sulfide, which is characterized by comprising the following steps:
(1) making the temperature of 600-4And H2Carrying out catalytic reaction on the raw material gas of S to obtain a reaction gas I;
(2) separating the reaction gas I to obtain sulfur and reaction gas II;
(3) the reaction gas II is subjected to flash evaporation to obtain crude CS2Liquid and reaction gas III;
(4) rectifying the crude CS2Obtaining liquid CS with the mass fraction of not less than 99.5 percent from the liquid2(ii) a Contacting the reaction gas III with an absorption liquid to obtain hydrogen-rich gas and an absorption liquid;
(5) separating the hydrogen-rich gas to obtain hydrogen with the volume fraction of more than 99.99 percent and the desorbed gas.
2. The method of claim 1, wherein the feed gas in step (1) is H2S and CH4Is 1.0 to 5.0, preferably 2.0 to 3.0.
3. The process according to claim 1, wherein the catalytic reaction temperature in step (1) is 800-.
4. The process as claimed in claim 1, wherein in step (3), the reaction gas II is subjected to a secondary flash evaporation.
5. The method according to claim 1, wherein the absorbing solution in the step (4) is one or more selected from the group consisting of: low temperature methanol, ethanolamine, diethanolamine and N-methyldiethanolamine, low temperature methanol being preferred.
6. The process according to claim 1, wherein the molar ratio of the reaction gas III to the absorption liquid in step (4) is 2.18 or more.
7. The method of claim 1, wherein the temperature of the absorption liquid in step (4) is in the range of-20 to-60 ℃.
8. The method of claim 1, wherein step (4) is performed at H2And (S) enabling the reaction gas III to contact with the absorption liquid in the absorption tower, wherein the pressure of the absorption tower is 2-5 MPa.
9. The method of claim 1, wherein step (5) uses a Pressure Swing Adsorption (PSA) unit to separate the hydrogen-rich gas.
10. The method of any one of claims 1-9, wherein the preheating of the CH-containing stream by a waste heat recovery system4And H2The temperature of the S raw material gas is 600-900 ℃.
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