CN111068745A - α Process for the production of olefins - Google Patents

α Process for the production of olefins Download PDF

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Publication number
CN111068745A
CN111068745A CN201811224401.XA CN201811224401A CN111068745A CN 111068745 A CN111068745 A CN 111068745A CN 201811224401 A CN201811224401 A CN 201811224401A CN 111068745 A CN111068745 A CN 111068745A
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gas
fischer
tropsch synthesis
reaction
methane
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CN111068745B (en
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晋超
吴玉
张荣俊
侯朝鹏
夏国富
孙霞
阎振楠
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/03—Catalysts comprising molecular sieves not having base-exchange properties
    • B01J29/035—Microporous crystalline materials not having base exchange properties, such as silica polymorphs, e.g. silicalites
    • B01J29/0358—Microporous crystalline materials not having base exchange properties, such as silica polymorphs, e.g. silicalites containing arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/02—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon from oxides of a carbon
    • C07C1/04—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon from oxides of a carbon from carbon monoxide with hydrogen
    • C07C1/0425—Catalysts; their physical properties
    • C07C1/043—Catalysts; their physical properties characterised by the composition
    • C07C1/0435—Catalysts; their physical properties characterised by the composition containing a metal of group 8 or a compound thereof
    • C07C1/044—Catalysts; their physical properties characterised by the composition containing a metal of group 8 or a compound thereof containing iron
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts

Abstract

The invention provides a production method of α -olefin, which comprises the steps of contacting methane with steam to carry out steam reforming reaction, contacting methane with carbon dioxide to carry out dry reforming reaction, mixing products of the two, preparing Fischer-Tropsch synthesis reaction feed, contacting the feed with a Fischer-Tropsch synthesis catalyst obtained through reduction activation to carry out Fischer-Tropsch synthesis reaction, and separating α -olefin, methane and carbon dioxide from product flow, wherein the Fischer-Tropsch synthesis catalyst precursor comprises a carrier, and an active metal oxide and an auxiliary agent metal oxide which are loaded on the carrier, and the reduction activation comprises the steps of sequentially carrying out pre-reduction and re-reduction on the Fischer-Tropsch synthesis catalyst precursor to obtain the Fischer-Tropsch synthesis catalyst.

Description

α Process for the production of olefins
Technical Field
The invention relates to the field of olefin synthesis, in particular to a production method of α -olefin.
Background
The linear chain α -olefin is an important organic raw material and intermediate, and is widely used for producing comonomers, lubricating oil base oil, surfactants, polyolefin resin, plasticizers, dyes, pharmaceutical preparations and the like, Sasol company in south Africa has built a set of production devices for separating 1-pentene and 1-hexene from Fischer-Tropsch (F-T) synthetic products (rich in α -olefin) and successfully put into production, and the process has the greatest advantages that coal is used as the raw material, the 1-pentene and 1-hexene are used as by-products for recovery, the industrial production cost is low, and higher benefits can be obtained.
The most widely used method for producing α -olefin at present is an olefin oligomerization method, but the production cost of the method is too high, and linear α -olefin with odd number of carbon atoms and the same market value cannot be produced, the cost of extracting linear 1-hexene from a crude product by adopting a high-temperature F-T Fischer-Tropsch synthesis technology by Sasol company in south Africa is less than one third of the cost of producing linear 1-hexene by adopting an ethylene trimerization method by Philips company, and meanwhile, based on the Anderson-Schulz-Flory distribution rule (the molar distribution of chain growth decreasing by an index), high-temperature F-T synthesis can also obtain high-value-added products such as 1-pentene, 1-heptene and the like with odd number of carbon atoms, and thus, α -olefin separated from the Fischer-Tropsch synthesis product has important commercial value.
In recent years, in the field of coal chemical industry in China, coal is rapidly grown up through preparing α -olefin from methanol, and preparing α -olefin directly from synthetic gas (FTO process) from coal is another process for preparing α -olefin from coal2) And then the α -olefin is directly prepared by F-T synthesis.
The technological process for preparing α -olefin by adopting FTO process is shown in figure 1, and comprises a coal water slurry preparation unit I ', a coal gasification unit II', a water gas conversion unit III ', a synthetic gas purification unit IV', a Fischer-Tropsch synthesis unit V 'and a α -olefin separation unit VI' which are connected in sequence, wherein the coal water slurry C 'is prepared from pulverized coal A' and water B 'in the coal water slurry preparation unit I', the coal water slurry C 'is conveyed into the coal gasification unit II' and reacts with oxygen D 'to generate coal gasification crude synthetic gas E', the coal gasification crude synthetic gas E 'is subjected to the water gas conversion unit III' to adjust the molar ratio of hydrogen and carbon monoxide to become converted crude synthetic gas F 'meeting the Fischer-Tropsch synthesis reaction requirement, the converted crude synthetic gas F' is subjected to acid gas and sulfide M 'removal by the synthetic gas purification unit IV' to obtain purified synthetic gas J ', the obtained purified synthetic gas J' is conveyed into the Fischer-Tropsch synthesis unit V 'to perform synthesis reaction to generate a Fischer-Tropsch reaction product N', the Fischer-Tropsch reaction product N 'is separated into α -Tropsch olefin K' by the synthetic unit V 'through the α -olefin separation unit VI', the purified synthetic gas J 'is conveyed into carbon dioxide H' and is discharged as a part of a Fischer-Tropsch synthesis system, and is discharged as a part of a.
The major problems with the FTO process described above are: 1. the energy consumption is high, and the utilization rate of carbon atoms is low; 2. the emission of carbon dioxide is 5-6 times of that of the traditional petroleum route; 3. the distribution of the Fischer-Tropsch synthesis product is limited by the Anderson-Schulz-Flory rule, and is limited by the generation of a large amount of methane and carbon dioxide caused by strong exothermicity of reaction, so that the overall energy efficiency of the FTO process is low, and the industrial process of the FTO process is seriously influenced. The FTO process uses a large amount of cooling water and discharged sewage to keep the water consumption high.
Therefore, there is a need to optimize the FTO process and select a system with high energy efficiency and reduced greenhouse gas emission.
It is noted that the information disclosed in the foregoing background section is only for enhancement of background understanding of the invention and therefore it may contain information that does not constitute prior art that is already known to a person of ordinary skill in the art.
Disclosure of Invention
The invention aims to provide a method for producing α -olefin, which can effectively reduce the system energy consumption and the emission of greenhouse gas.
In order to achieve the purpose, the invention adopts the following technical scheme:
the invention provides a production method of α -olefin, which comprises the following steps:
contacting methane with steam to carry out steam reforming reaction to obtain steam reforming synthesis gas;
contacting methane with carbon dioxide to carry out dry reforming reaction to obtain dry reforming syngas;
mixing at least part of the steam reforming synthesis gas and at least part of the dry weight integrated synthesis gas to prepare a Fischer-Tropsch synthesis reaction feed;
carrying out reduction activation on the Fischer-Tropsch synthesis catalyst precursor to obtain a Fischer-Tropsch synthesis catalyst;
contacting the Fischer-Tropsch synthesis reaction feed with the Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction to obtain a Fischer-Tropsch synthesis product material flow; and
α -olefins, methane and carbon dioxide are separated from the Fischer-Tropsch synthesis product stream,
wherein the Fischer-Tropsch synthesis catalyst precursor comprises a carrier, and an active metal oxide and an auxiliary agent metal oxide which are loaded on the carrier, the carrier is a manganese oxide molecular sieve, the active metal in the active metal oxide is a VIII group metal, the auxiliary agent metal in the auxiliary agent metal oxide is one or more of IIB group metal, alkaline earth metal and IB group metal,
wherein the reductive activation comprises:
pre-reducing the Fischer-Tropsch synthesis catalyst precursor in a first gas to obtain a pre-reduced catalyst; and
and reducing the pre-reduction catalyst in a second gas to obtain the Fischer-Tropsch synthesis catalyst.
According to one embodiment of the invention, the active metal is one or more of Fe, Co, Ni, preferably Fe; the assistant metal is one or more of Zn, Cd, Cu and Ag, preferably Zn and/or Cu, and more preferably Zn.
According to one embodiment of the present invention, the content of the active metal oxide is 3 to 70% by weight, preferably 5 to 50% by weight, more preferably 8 to 30% by weight, based on the metal element; the content of the assistant metal oxide is 1-60 wt%, preferably 2-50 wt%; the carrier is present in an amount of 12 to 94% by weight, preferably 35 to 91% by weight.
According to one embodiment of the invention, the catalyst is prepared by the following method: and loading the active metal and the auxiliary metal on the carrier, and then drying and roasting the active metal and the auxiliary metal in sequence to obtain the supported catalyst.
According to one embodiment of the invention, the supporting is carried out by impregnation or coprecipitation.
According to one embodiment of the invention, the impregnation method is an isometric impregnation method or a saturation impregnation method.
According to one embodiment of the invention, the impregnation method comprises: dissolving the soluble salt of the active metal and the soluble salt of the auxiliary metal in a solvent to obtain an impregnation solution; and immersing the support in the immersion liquid.
According to one embodiment of the invention, the impregnation is a single impregnation or a stepwise impregnation.
According to one embodiment of the invention, the soluble salt of the active metal and the soluble salt of the auxiliary metal are nitrates or hydrochlorides.
According to one embodiment of the invention, the metal molar ratio of the active metal and the promoter metal in the impregnation solution is 1: 0.2-5, preferably 1: 0.3-3.
According to one embodiment of the invention, the temperature of the impregnation is between 10 and 80 ℃, preferably between 20 and 60 ℃; the impregnation time is 0.1 to 3 hours, preferably 0.5 to 1 hour.
According to one embodiment of the present invention, the temperature of the drying is 80-350 ℃, preferably 100-300 ℃; the drying time is 1 to 24 hours, preferably 2 to 12 hours.
According to one embodiment of the invention, the temperature of the roasting is between 250 ℃ and 900 ℃, preferably between 300 ℃ and 850 ℃, more preferably between 350 ℃ and 800 ℃; the calcination time is 0.5 to 12 hours, preferably 1 to 8 hours, and more preferably 2 to 6 hours.
According to one embodiment of the present invention, the first gas is hydrogen or a mixture of hydrogen and an inert gas.
According to one embodiment of the invention, the temperature of the pre-reduction is 200-600 ℃, preferably 300-550 ℃; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; the volume space velocity of the first gas is 5000--1。
According to one embodiment of the invention, the second gas is a gaseous hydrocarbon or a mixture of a gaseous hydrocarbon and an inert gas.
According to one embodiment of the invention, the temperature of the re-reduction is 180-; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; the volume space velocity of the second gas is 1000-30000 hours based on the gaseous hydrocarbon-1。
According to one embodiment of the invention, the molar ratio of the inert gas to the gaseous hydrocarbon is between 1 and 200: 1, preferably 5 to 20: 1.
according to one embodiment of the invention, the gaseous hydrocarbon is a gaseous alkane and/or a gaseous alkene, preferably selected from C1-C4And/or C2-C4More preferably, the olefin(s) of (b) is/are one or more selected from methane, ethane, ethylene, propylene, propane, butane and butene.
According to one embodiment of the present invention, the steam reforming reaction is carried out in a fixed bed reactor at a reaction temperature of 700-: 0.5-4, the hourly space velocity of the gas fed is preferably 10000--1Preferably 50000-100000 h-1.
According to one embodiment of the present invention, the dry reforming reaction is carried out in a fixed bed reactor at a reaction temperature of 600-: 0.5-5, wherein the gas hourly volume space velocity of the feed is 10000--1Preferably 50000-100000 hours-1。
According to one embodiment of the invention, the Fischer-Tropsch synthesis reaction is carried out in a fixed bed reactor, the reaction temperature is 200--1More preferably 5000--1Most preferably 10000--1。
According to one embodiment of the invention, the molar ratio of hydrogen to carbon monoxide in the feed to the fischer-tropsch synthesis reaction is in the range of from 0.4 to 3: 1, preferably 0.6 to 2.8: 1, more preferably 0.8 to 2.6: 1, most preferably 1.5-2.5: 1.
according to an embodiment of the present invention, the method further comprises recycling the separated methane to the steam reforming reaction and/or the dry reforming reaction, and recycling the separated carbon dioxide to the dry reforming reaction.
The production method can improve the utilization rate of two greenhouse gases of carbon dioxide and methane in the production process of α -olefin, so that the greenhouse gases are converted into products with high added values, the greenhouse gas emission is reduced, the energy consumption of a system is reduced, and the resource and energy utilization rate of the whole process is obviously improved.
Drawings
FIG. 1 is a flow diagram of a typical prior art process for the direct production of α -olefins from coal via syngas;
FIG. 2 is a process flow diagram of a α -olefin production process according to one embodiment of the invention;
wherein the reference numerals are as follows:
i': coal water slurry preparation unit II': a coal gasification unit III': water gas shift unit
IV ' syngas purification unit V ' Fischer-Tropsch synthesis unit VI ' α -olefin separation unit
A': pulverized coal B': water C': coal water slurry
D': oxygen E': coal gasification raw synthesis gas F': shifted raw synthesis gas
G ' methane H ' carbon dioxide K ' α -olefins
M': acid gas and sulfide N': Fischer-Tropsch reaction product Y': unreacted synthesis gas
Z': purge gas J': purifying synthesis gas
I: a raw material gas separation unit II: steam reforming reaction unit III: dry reforming reaction unit
IV: Fischer-Tropsch synthesis reaction unit V: Fischer-Tropsch synthesis product separation unit A: raw material gas
B: methane C: water vapor D: carbon dioxide
E: steam reforming syngas F: dry weight integration gas G: Fischer-Tropsch synthesis reaction feed
H Fischer-Tropsch synthesis product stream L hydrogen and carbon monoxide K α -olefins for recycle
M: methane N: carbon dioxide Z: purge gas
Detailed Description
The technical solution of the present invention is further explained below according to specific embodiments. The scope of protection of the invention is not limited to the following examples, which are set forth for illustrative purposes only and are not intended to limit the invention in any way.
In the present invention, anything or matters not mentioned is directly applicable to those known in the art without any change except those explicitly described. Moreover, any embodiment described herein may be freely combined with one or more other embodiments described herein, and the technical solutions or ideas thus formed are considered part of the original disclosure or original description of the present invention, and should not be considered as new matters not disclosed or contemplated herein, unless a person skilled in the art would consider such combination to be clearly unreasonable.
All features disclosed in this invention may be combined in any combination and such combinations are understood to be disclosed or described herein unless a person skilled in the art would consider such combinations to be clearly unreasonable. The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.
Any terms not directly defined herein should be understood to have meanings associated with them as commonly understood in the art of the present invention. The following terms as used throughout this specification should be understood to have the following meanings unless otherwise indicated.
According to a first aspect of the present invention there is provided a supported catalyst precursor comprising a support and an active metal oxide and a promoter metal oxide supported on the support.
In the supported catalyst precursor of the present invention, the carrier used is a manganese oxide molecular sieve, preferably an octahedral manganese oxide molecular sieve, a manganite (OMS-1).
The octahedral manganese oxide molecular sieve (OMS-1) can be prepared by a hydrothermal synthesis method, and specifically comprises the following steps:
(1) preparing a mixed solution of magnesium chloride and manganese chloride, and preparing a potassium permanganate solution containing sodium hydroxide;
(2) mixing the two solutions, heating, stirring, aging, washing, and drying to obtain Na-OL-1;
(3) Na-OL-1 and a magnesium chloride solution with a certain concentration are transferred to a hydrothermal kettle for crystallization reaction to obtain pure phase octahedral manganese oxide molecular sieve (OMS-1) (JCPDS 38-475).
In the hydrothermal synthesis method, the aging temperature of the two solutions after mixing is 30-90 ℃, preferably 40-70 ℃; the aging time is 10-50h, preferably 15-40 h; the crystallization reaction temperature of the Na-OL-1 and the magnesium chloride in the hydrothermal kettle is 100-200 ℃, preferably 120-180 ℃; the time of the crystallization reaction is 12 to 72 hours, preferably 24 to 60 hours.
In the supported catalyst precursor of the present invention, the active metal and the promoter metal are supported on the carrier in the form of an oxide, wherein the active metal is a group VIII metal, the promoter metal is one or more of a group IIB metal, an alkaline earth metal and a group IB metal, and the valence states of the active metal and the promoter metal in the oxide are the highest oxidation valence states thereof.
More specifically, the active metal may be one or more of Fe, Co, Ni, preferably Fe; the promoter metal may be one or more of Zn, Cd, Cu, Ag, preferably Zn and/or Cu, more preferably Zn.
In the supported catalyst precursor of the present invention, the content of the active metal oxide is 3 to 70% by weight, preferably 5 to 50% by weight, more preferably 8 to 30% by weight, in terms of the metal element; the content of the auxiliary metal oxide is 1 to 60 wt%, preferably 2 to 50 wt%; the content of the carrier is 12 to 94% by weight, preferably 35 to 91% by weight.
The supported catalyst precursor of the present invention can be obtained by supporting an active metal and an auxiliary metal on a carrier, followed by drying and calcination in this order.
The method for supporting the active metal and the promoter metal according to the present invention may be a method conventionally used in the art, and for example, an impregnation method or a coprecipitation method, preferably an impregnation method, may be used.
The impregnation method used in the invention is an isometric impregnation method or a saturation impregnation method, and specifically comprises the following steps: dissolving soluble salt of active metal and soluble salt of auxiliary metal in a solvent to obtain an impregnation solution; and immersing the support in the impregnation fluid.
The soluble salts of the active metal and the soluble salts of the auxiliary metal used in the preparation of the impregnation solution may be nitrates or hydrochlorides (chlorides).
The carrier is soaked in the soaking solution, one-time soaking can be adopted, step-by-step soaking can be adopted, soluble salts of the active metal and soluble salts of the auxiliary metal respectively form the soaking solution, then the soaking solution is sequentially loaded on the OMS-1 carrier through soaking, or the soluble salts of the active metal and the soluble salts of the auxiliary metal are dissolved together to form the soaking solution, and then the soaking solution is soaked on the carrier for two times or multiple times.
The metal molar ratio of the active metal to the auxiliary metal in the impregnation liquid is 1: 0.2-5, preferably 1:0.3-3, and the total concentration of solutes in the impregnation solution may be 30-70 wt%.
The dipping temperature can be 10-80 ℃, and preferably 20-60 ℃; the impregnation time may be from 0.1 to 3h, preferably from 0.5 to 1 h.
After OMS-1 is loaded with the active metal and the promoter metal, it is dried by a method conventionally used in the art, for example, by heat drying. The drying temperature may be 80-350 deg.C, preferably 100-300 deg.C, and the drying time may be 1-24 hours, preferably 2-12 hours.
After drying, the carrier supporting the active metal and the promoter metal is calcined, and the calcination method may be a method conventionally used in the art as long as the active metal and the promoter metal are respectively converted into corresponding oxides, for example, the calcination method is a calcination method under an air atmosphere, and the calcination conditions include: the roasting temperature is 250-900 ℃, preferably 300-850 ℃, and more preferably 350-800 ℃; the calcination time is 0.5 to 12 hours, preferably 1 to 8 hours, and more preferably 2 to 6 hours.
In another aspect, the present invention also provides the use of the supported catalyst precursor described above in the reaction of making α -olefin from synthesis gas.
In still another aspect, the present invention further provides a method for producing α -olefin, the method comprises contacting syngas with the supported catalyst, and specifically comprises the following steps:
s11, under the condition of steam reforming reaction, enabling methane to contact with steam to carry out steam reforming reaction to obtain steam reforming synthesis gas;
s21, under the condition of dry reforming reaction, enabling methane to contact with carbon dioxide for dry reforming reaction to obtain dry reforming synthesis gas;
s31, mixing at least part of steam reforming synthesis gas and at least part of dry weight integrated synthesis gas to prepare a Fischer-Tropsch synthesis reaction feed;
s41, carrying out reduction activation on the supported catalyst precursor to obtain a Fischer-Tropsch synthesis catalyst;
s51, contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction to obtain a Fischer-Tropsch synthesis product material flow;
s61, separating α -olefin, methane and carbon dioxide from the Fischer-Tropsch synthesis product stream.
In step S11, the steam reforming reaction may be performed in a conventional reactor, preferably, a fixed bed reactor. The molar ratio of methane to water vapor may be 1: 0.5 to 4, preferably 1: 1-3. The methane may be contacted with the water vapor at a temperature of 700 ℃ to 950 ℃, preferably 800 ℃ to 900 ℃. The pressure in the reactor in which methane is contacted with steam may be 0.1 to 5MPa, preferably 1 to 3MPa, which is a gauge pressure. The hourly space velocity of the gas fed is preferably 10000--1Preferably 50000-100000 hours-1。
In step S11, various steam reforming catalysts commonly used in the art and suitable for steam reforming reactions may be used. As an example, the steam reforming catalyst contains a carrier, which may be one or a combination of two or more of alumina, silica, zirconia, and silicon carbide, and an active component supported on the carrier. Preferably, the support is alumina, in particular gamma-Al2O3、θ-Al2O3、δ-Al2O3And α -Al2O3One or more than two of them. The active component can be a group VIII metal element, preferably a non-noble group VIII metal element, such as one or more of Fe, Co and Ni. More preferably, the active component is Ni. The amount of active ingredient supported on the carrier may be conventionally selected. In general, the active component may be present in an amount of 1 to 30% by weight, preferably 5 to 25% by weight, more preferably 10 to 15% by weight, calculated as element, based on the total amount of the catalyst.
In step S21, the dry reforming reaction may be carried out in a conventional reactor, preferably a fixed bed reactor. The molar ratio of methane to carbon dioxide may be 1: 0.5 to 5, preferably 1: 0.8 to 3, more preferably 1: 1-2. The methane and carbon dioxide may be contacted at a temperature of 600-800 deg.C, preferably 650-750 deg.C. The pressure in the reactor in which the methane and carbon dioxide are contacted may be in the range of from 0.1 to 5MPa, preferably from 1 to 3MPa, as a gauge pressure. The gas hourly volume space velocity of the feed is 10000-100000 hours based on the total amount of the methane and the carbon dioxide-1Preferably 50000-100000 hours-1。
In step S21, various dry reforming catalysts commonly used in the art for dry reforming reactions can be used. As an example, a dry reforming catalyst contains a support and an active component supported on the support. The carrier may be oxygenOne or a combination of two or more of aluminum oxide, silicon oxide, zirconium oxide and silicon carbide. Preferably, the support is alumina, in particular gamma-Al2O3、θ-Al2O3、δ-Al2O3And α -Al2O3One or more than two of them. The active component can be a group VIII metal element, preferably a non-noble group VIII metal element, such as one or more of Fe, Co and Ni. More preferably, the active component is Ni. The amount of active ingredient supported on the carrier may be conventionally selected. In general, the active component may be present in an amount of 1 to 30% by weight, preferably 5 to 25% by weight, more preferably 10 to 15% by weight, calculated as element, based on the total amount of the catalyst.
In the α -olefin production method of the present invention, the methane, which is one of the raw materials for steam reforming and dry reforming of methane, may be methane of various sources, and preferably methane separated from a methane-rich raw material gas, in this case, the α -olefin production method of the present invention may further include a step S10 of separating methane from a methane-containing raw material gas in S10.
The methane may be separated from the feed gas by conventional means, such as by pressure swing adsorption. As an example, methane is separated from a feed gas by a cryogenic condensation process. The cryocondensation method is a method for separating and purifying methane by using a difference in boiling point, and can determine whether to obtain methane from a gas phase or from a liquid phase according to the boiling point of each component in a raw material gas.
In the process for producing α -olefin according to the present invention, the purity of methane, which is one of the raw materials for steam reforming and dry reforming, is generally 90% by weight or more, and the mass content of elemental sulfur therein is generally 20ppm or less, preferably 10ppm or less, more preferably 5ppm or less, and still more preferably 1ppm or less.
In the α -olefin production method of the present invention, the raw material utilization rate of the method of the present invention can be further improved by controlling the amount of methane fed to step S11 and step S21 according to the reaction properties of steam reforming and dry reforming and the requirements of the Fischer-Tropsch synthesis reaction feed, preferably, the weight ratio of the methane used in step S11 to the methane used in step S21 is 1: 0.5-2.5.
In step S31, at least a portion of the steam reformed syngas and at least a portion of the dry syngas are combined to form a Fischer-Tropsch reaction feed having a hydrogen to carbon ratio (i.e., a molar ratio of hydrogen to carbon monoxide) that is consistent with the Fischer-Tropsch reaction feed, the molar ratio of hydrogen to carbon monoxide in the Fischer-Tropsch reaction feed is preferably 0.4 to 3: 1, more preferably 0.6 to 2.8: 1, even more preferably 0.8 to 2.6: 1, and even more preferably 1.5 to 2.5: 1, from the viewpoint of increasing α -olefin selectivity.
In step S41, the supported catalyst precursor needs to be subjected to reduction activation to obtain a fischer-tropsch synthesis catalyst before being applied to a reaction for preparing α -olefin from synthesis gas, and the reduction activation includes pre-reducing the supported catalyst precursor in a first gas to obtain a pre-reduced catalyst, and re-reducing the pre-reduced catalyst in a second gas to obtain the fischer-tropsch synthesis catalyst.
The first gas used in the pre-reduction is hydrogen or a mixture of hydrogen and an inert gas. That is, the pre-reduction may be performed in a pure hydrogen atmosphere, or may be performed in a mixed atmosphere of hydrogen and an inert gas, for example, in a mixed atmosphere of hydrogen and nitrogen and/or argon. The temperature of the pre-reduction is 200-600 ℃, preferably 300-550 ℃; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; the volume space velocity of the first gas is 5000--1。
The second gas in the re-reduction is a gaseous hydrocarbon or a mixed gas of the gaseous hydrocarbon and an inert gas. The temperature for re-reduction is 180 ℃ and 400 ℃, preferably 200 ℃ and 350 ℃; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; the volume space velocity of the second gas is 1000--1。
When a mixture of gaseous hydrocarbon and inert gas is selected, the molar ratio of the inert gas to the gaseous hydrocarbon is 1-200: 1, preferably 5 to 20: 1.
the gaseous hydrocarbon is an alkane and/or an alkene which is gaseous at the above-mentioned reduction activation temperature, and is preferably selected from C1-C4And/or C2-C4More preferably, the olefin(s) of (b) is/are one or more selected from methane, ethane, ethylene, propylene, propane, butane and butene.
The inert gas in the first gas and the second gas may be the same or different and each is independently selected from one or more of nitrogen and a group zero element gas, and preferably each is nitrogen and/or argon.
In step S51, the Fischer-Tropsch synthesis reaction can be carried out under the condition of conventionally producing α -olefin, the Fischer-Tropsch synthesis reaction can be carried out in a fixed bed reactor, a fluidized bed reactor or a combination of the fixed bed reactor and the fluidized bed reactor, preferably, hydrogen and carbon monoxide are contacted with the Fischer-Tropsch synthesis catalyst in the fixed bed reactor, preferably, the Fischer-Tropsch synthesis reaction feed and the Fischer-Tropsch synthesis catalyst can be contacted at the temperature of 200-380 ℃ and preferably 250-350 ℃, the pressure of the contact of the Fischer-Tropsch synthesis reaction feed and the Fischer-Tropsch synthesis catalyst can be 0.8-3MPa and preferably 1-2.8MPa, the pressure is measured by gauge pressure, and when the hydrogen and the carbon monoxide are contacted with the Fischer-Tropsch synthesis catalyst in the fixed bed reactor, the volume space velocity of the Fischer-Tropsch synthesis reaction feed can be 2000-50000 hours-1Preferably 5000--1Preferably 10000--1。
α -olefins, methane and carbon dioxide can be separated from the fischer-tropsch synthesis product stream in step S61 using conventional methods as an example, the fischer-tropsch synthesis product stream can be separated using a low temperature condensation process to provide α -olefins, methane and carbon dioxide, respectively.
The production method of the α -olefin can recycle the separated methane for steam reforming reaction and/or dry reforming reaction, namely, the methane separated from the Fischer-Tropsch synthesis product flow is sent to the step S11 and/or the step S21 to be used as the raw material for the steam reforming reaction and/or the dry reforming reaction, and the separated carbon dioxide can also recycle the dry reforming reaction, namely, the carbon dioxide separated from the Fischer-Tropsch synthesis product flow is sent to the step S21 to be used as the feed for the dry reforming reaction.
From the viewpoint of further improving the utilization rate of the raw materials, it is preferable to further comprise separating unreacted hydrogen and/or carbon monoxide from the product stream of the Fischer-Tropsch synthesis, and feeding at least part of the hydrogen and/or at least part of the carbon monoxide to step S31 for formulating the feed for the Fischer-Tropsch synthesis reaction. Preferably, part of the hydrogen and/or part of the carbon monoxide separated from the product stream of the Fischer-Tropsch synthesis is recycled to step S31 for use in formulating the Fischer-Tropsch synthesis reaction feed, while the remainder of the hydrogen and/or carbon monoxide is vented as purge gas to the system. Generally, the amount of hydrogen and carbon monoxide used for recycle may be in the range of from 10 to 98%, preferably from 15 to 98%, based on the total amount of hydrogen and carbon monoxide separated from the Fischer-Tropsch synthesis product stream.
In the α -olefin production method of the invention, the step S11 and the step S21 can be carried out simultaneously, and the step S31 and the step S41 can be carried out simultaneously, thereby saving the total reaction time and improving the production efficiency.
The α -olefin production process of the invention can be carried out by a α -olefin production system that includes a steam reforming reaction unit, a dry reforming reaction unit, a syngas mixing unit, a fischer-tropsch synthesis reaction product separation unit, and a recycle unit.
The steam reforming reaction unit is used for contacting methane with steam to carry out steam reforming reaction to obtain steam reforming synthesis gas. The steam reforming reaction unit may be provided with a conventional steam reforming reactor and corresponding feed, discharge and control means to enable the reforming reaction of methane with steam to produce a steam reformed synthesis gas having hydrogen and carbon monoxide as the main components.
And the dry reforming reaction unit is used for contacting methane and carbon dioxide to carry out dry reforming reaction to obtain dry reforming syngas. The dry reforming reaction unit may be provided with a conventional dry reforming reactor and corresponding feed, discharge and control means to enable the reforming reaction of methane with carbon dioxide to obtain a dry reformed gas having hydrogen and carbon monoxide as main components.
The synthesis gas mixing unit is used for mixing the steam reforming synthesis gas with the dry weight synthesis gas to prepare Fischer-Tropsch synthesis reaction feed, and the Fischer-Tropsch synthesis reaction feed is sent into the Fischer-Tropsch synthesis reaction unit. The synthesis gas mixing unit may be provided with a vessel for receiving and mixing the steam reformed synthesis gas and the dry integrated syngas, in which vessel the steam reformed synthesis gas is mixed with the dry integrated syngas to obtain the fischer-tropsch synthesis feed. Or a pipeline mixer can be adopted to directly mix the steam regenerated synthetic gas and the dry weight integrated synthetic gas in a conveying pipeline so as to obtain the Fischer-Tropsch synthesis reaction feed. The synthesis gas mixing unit can be provided with various common control devices for controlling the mixing proportion of the steam reforming synthesis gas and the dry weight integrated synthesis gas, so that the Fischer-Tropsch synthesis reaction feeding material meeting the hydrogen-carbon ratio of the Fischer-Tropsch synthesis reaction is obtained.
The Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor, and is used for contacting Fischer-Tropsch synthesis reaction feeding with a Fischer-Tropsch synthesis catalyst at the reaction temperature of producing α -olefin to obtain a Fischer-Tropsch synthesis product material flow containing α -olefin.
The Fischer-Tropsch synthesis reaction unit is preferably further provided with a reduction activation subunit, and the reduction activation subunit is used for carrying out reduction activation on the supported catalyst precursor so as to convert the supported catalyst precursor into the Fischer-Tropsch synthesis catalyst with catalytic activity. The reduction-activating subunit may reduce-activate the supported catalyst precursor by contacting the supported catalyst precursor with a reducing gas.
In a preferred embodiment, the reduction activation sub-unit comprises a first gas storage and delivery device, a second gas storage and delivery device, a reduction gas control device, and a reduction activation reactor.
The first gas storage and conveying device is used for storing the first gas and conveying the first gas into the reduction activation reactor. The first gas is hydrogen or a mixed gas of hydrogen and inert gas. The first gas storage and delivery device is configured to sufficiently store and deliver the first gas. The first gas storage and delivery means may be arranged in accordance with the teachings of the prior art to enable it to store and deliver the first gas.
And the second gas storage and conveying device is used for storing a second gas and conveying the second gas into the reduction activation reactor, wherein the second gas is hydrocarbon which is gaseous at the reduction temperature or a mixed gas of the hydrocarbon which is gaseous at the reduction temperature and inert gas. The types of the first gas and the second gas have been described in detail above and will not be described in detail here.
A reducing gas storage and conveying device, a reducing gas control device and a reduction activation reactor. The reducing gas is hydrogen or a mixed gas of hydrogen and inert gas. The reducing gas storage and delivery device is configured to sufficiently store and deliver the reducing gas. The reducing gas storage and delivery means may be provided according to the teachings of the prior art to enable storage and delivery of the reducing gas.
The reducing gas control means is for controlling the type of gas fed to the reduction activation reactor and the amount of gas fed thereto. Specifically, when the reduction activation subunit is operated, the reduction gas control device is configured to firstly input a first gas into the reduction activation reactor, so that the supported catalyst precursor is contacted with the first gas to carry out a pre-reduction reaction, and a pre-reduction catalyst is obtained. And then feeding a second gas into the reduction activation reactor so as to contact the pre-reduction catalyst with the second gas to carry out a re-reduction reaction. The reducing gas control means may employ conventional control elements such as various control valves to control the type of gas fed to the reduction reactor and the amount of gas fed thereto.
The reduction reactor is used for accommodating the supported catalyst precursor and is communicated with the first gas storage and conveying device and the second gas storage and conveying device, so that the supported catalyst precursor is sequentially contacted with the first gas and the second gas to carry out reduction activation, and the catalyst with the Fischer-Tropsch synthesis catalytic activity is obtained.
The reduction activation reactor and the Fischer-Tropsch synthesis reactor can be the same reactor, namely, the reduction activation of the supported catalyst precursor is carried out in the Fischer-Tropsch synthesis reactor.
The reduction activation reactor and the Fischer-Tropsch synthesis reactor also can be different reactors, namely the Fischer-Tropsch synthesis reactor and the reduction activation reactor are independent reactors. At this time, the reduction activation catalyst output port of the reduction activation reactor is set to be communicated with the catalyst input port of the fischer-tropsch synthesis reactor, so that the reduction activation catalyst output by the reduction activation reactor is sent into the fischer-tropsch synthesis reactor. The reduction activation catalyst output port of the reduction activation reactor and the catalyst input port of the Fischer-Tropsch synthesis reactor can be communicated by adopting a conveying pipeline, a control valve is arranged on the conveying pipeline, when the reduction activation reactor outputs the reduction activation catalyst, the control valve is opened, the reduction activation catalyst output port of the reduction activation reactor and the catalyst input port of the Fischer-Tropsch synthesis reactor are communicated, and the reduction activation catalyst is sent into the Fischer-Tropsch synthesis reactor.
In the α -olefin production system, the circulating unit is used for circularly sending the methane separated by the Fischer-Tropsch synthesis reaction product separating unit into one or both of the steam reforming reaction unit and the dry reforming reaction unit, circularly sending the carbon dioxide separated by the Fischer-Tropsch synthesis reaction product separating unit into the dry reforming reaction unit, and circularly sending the hydrogen and/or the carbon monoxide separated by the Fischer-Tropsch synthesis reaction product separating unit into the Fischer-Tropsch synthesis reaction unit.
The circulation unit can be provided with a methane conveying pipeline which is respectively used for communicating the Fischer-Tropsch synthesis reaction product separation unit with the steam reforming reaction unit and the dry reforming reaction unit, and a control valve arranged on the methane conveying pipeline, so that the methane separated by the Fischer-Tropsch synthesis reaction product separation unit is respectively conveyed into the steam reforming reaction unit and the dry reforming reaction unit. The circulation unit can be provided with a carbon dioxide conveying pipeline for communicating the Fischer-Tropsch synthesis reaction product separation unit and the dry reforming reaction unit and a control valve arranged on the carbon dioxide conveying pipeline so as to convey the carbon dioxide output by the Fischer-Tropsch synthesis reaction product separation unit into the dry reforming reaction unit.
When the fischer-tropsch synthesis reaction product separation unit further separates hydrogen and carbon monoxide, the circulation unit is preferably provided with a conveying pipeline for communicating the fischer-tropsch synthesis reaction product separation unit and the fischer-tropsch synthesis reaction unit, and a control valve arranged on the conveying pipeline, so that the hydrogen and the carbon monoxide separated by the fischer-tropsch synthesis reaction product separation unit are sent into the fischer-tropsch synthesis reaction unit. Can send into during the ft synthesis reaction unit with hydrogen and carbon monoxide through same conveying pipeline, also can send into the ft synthesis reaction unit respectively with hydrogen and carbon monoxide through different conveying pipeline, can set up hydrogen conveying pipeline and set up the control flap on hydrogen conveying pipeline this moment respectively and with carbon monoxide conveying pipeline and the control flap of setting on carbon monoxide conveying pipeline.
In the α -olefin production system according to the present invention, it is preferable that the system further comprises a raw material gas separation unit for separating methane from a raw material gas containing methane, and a methane output port of the raw material gas separation unit is connected to a methane raw material input port of the steam reforming reaction unit and a methane raw material input port of the dry reforming reaction unit, respectively, so as to send the separated methane to the steam reforming reaction unit and the dry reforming reaction unit, respectively.
The feed gas separation unit may employ conventional separation methods to separate methane from the feed gas. In one embodiment, the feed gas separation unit employs a pressure swing adsorption process to separate methane from the feed gas. In a more preferred embodiment, the feed gas separation unit employs cryogenic condensation to separate methane from the feed gas. In this more preferred embodiment, a low-temperature condenser may be provided in the raw gas separation unit to condense the raw gas to separate methane from the raw gas. The low-temperature condenser may be a conventional condenser, and is not particularly limited.
Fig. 2 illustrates a preferred embodiment of α -olefin production system according to the present invention, described in detail below with reference to fig. 2. as shown in fig. 2, the α -olefin production system includes a raw material gas separation unit I, a steam reforming reaction unit II, a dry reforming reaction unit III, a fischer-tropsch synthesis reaction unit IV, a fischer-tropsch synthesis product separation unit V, and a circulation unit.
The method comprises the steps of enabling a raw material gas A to enter a raw material gas separation unit I to be separated to obtain methane B, respectively sending the methane B into a steam reforming reaction unit II and a dry reforming reaction unit III, simultaneously sending steam C into the steam reforming reaction unit II to enable the methane and the steam to carry out reforming reaction to obtain steam reforming synthetic gas E, sending carbon dioxide D into the dry reforming reaction unit III to enable the methane and the carbon dioxide to carry out reforming reaction to obtain dry reforming synthetic gas F, mixing the steam reforming synthetic gas E and the dry reforming synthetic gas F (preferably adopting a pipeline mixer), enabling a Fischer-Tropsch synthesis reaction feeding G, a Fischer-Tropsch synthesis reaction feeding material G, which is in accordance with the Fischer-Tropsch synthesis reaction hydrogen-carbon ratio, to enter a Fischer-Tropsch synthesis reaction unit IV, contacting with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction, enabling a Fischer-Tropsch synthesis reactor in the Fischer-Tropsch synthesis reaction unit IV to operate at the temperature of α -Tropsch synthesis reaction products to operate, enabling a Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV to enter a Fischer-Tropsch synthesis product separation unit V to be separated to obtain α -olefins, unreacted hydrogen and carbon monoxide, and methane M.
The separated hydrogen and carbon monoxide can be recycled for preparing the Fischer-Tropsch synthesis reaction feed, can also be discharged out of the system, and can also be recycled for preparing the Fischer-Tropsch synthesis reaction feed in a part of the system, and discharged out of the system in the other part of the system. Preferably, as shown in FIG. 2, the hydrogen and carbon monoxide L for recycle are mixed with the steam reforming synthesis gas E and the dry weight integrated synthesis gas F for formulating the Fischer-Tropsch synthesis reaction feed G; the remaining part of the hydrogen and carbon monoxide is discharged out of the system as purge gas Z.
The separated carbon dioxide N is sent to the dry reforming reaction unit III and recycled as one of the raw materials for the dry reforming reaction. The separated methane M is respectively sent into the steam reforming reaction unit II and the dry reforming reaction unit III to be used as one of the raw materials of the reforming reaction for recycling.
The present invention will be described in detail with reference to examples, but the scope of the present invention is not limited thereto.
Examples
In the following examples, preparations and comparative examples, the pressures were gauge pressures unless otherwise specified.
In the following examples, preparations and comparative examples, the conversion of CO (X)CO)、C5-C15α selectivity to olefin (S)α -olefins) And C5Above (C)5+) Selectivity of hydrocarbons
Figure BDA0001835574500000161
Respectively calculated by the following formula:
Figure BDA0001835574500000162
Figure BDA0001835574500000163
Figure BDA0001835574500000164
wherein, V1、V2Respectively representing the volume of feed gas entering the reaction system and the volume of tail gas flowing out of the reaction system in a certain time period under a standard condition;
C1,CO、C2,COrespectively representing the molar contents of CO in raw gas entering a reaction system and tail gas flowing out of the reaction system;
nconis the mole number of CO participating in the reaction;
Figure BDA0001835574500000165
to produce CO2The number of moles of (a);
nα -olefinsTo moles of α -olefin produced;
Figure BDA0001835574500000171
to generate CH4、C2Hydrocarbons, C3Hydrocarbons and C4The sum of the moles of hydrocarbons.
In the following examples, preparations and comparative examples, the specific surface area, pore volume and average pore diameter were measured by nitrogen adsorption method, specifically, N was used2Measuring an adsorption isotherm at a constant temperature of 77K, calculating a specific surface area and a pore volume according to a BET formula, and calculating an average pore size distribution according to a BJH method; the particle size distribution was determined using a laser particle sizer.
In the following examples, preparation examples and comparative examples, the kind and content of each metal element in the catalyst and the catalyst precursor were measured by the X-ray fluorescence spectrum analysis method specified in RIPP 132-92 (published in the protocols of petrochemical engineering analysis (RIPP Experimental method), Yankee, et al, science publishers, 1 st edition at 1990, p. 371-. When the catalyst was tested, a sample of the catalyst was stored under an argon atmosphere.
In the following examples, preparations and comparative examples, CO2the-TPD and the CO-TPD are detected on line by using a Michmark chemical adsorption instrument and an OMistar mass spectrometer as a detector, wherein the CO is detected2TPD recorded signals for the nuclear to cytoplasmic ratio of 44 by the mass spectrometer and CO-TPD recorded signals for the nuclear to cytoplasmic ratio of 28 by the mass spectrometer.
In the following examples, preparations and comparative examples, X-ray photoelectron spectroscopy was carried out on ESCALB model 250X equipped with Thermo Avantage V5.926 software from Thermo ScientificThe test is carried out on a ray photoelectron spectrometer, an excitation source is monochromatic Al K α X rays, the energy is 1486.6eV, the power is 150W, the penetrating energy used by narrow scanning is 30eV, and the basic vacuum in the analysis test is 6.5 multiplied by 10-10mbar, electron binding energy was corrected for the C1s peak (284.6eV) of elemental carbon, data processed on Thermo Avantage software, and quantified in the analytical module using the sensitivity factor method.
Preparation example 1
(1) Preparation of the carrier: OMS-1 preparation:
dissolving 2.014g of anhydrous manganese chloride and 0.636g of magnesium chloride hexahydrate in 70ml of deionized water, heating and stirring in a water bath at 50 ℃ to fully dissolve the anhydrous manganese chloride and the magnesium chloride hexahydrate; dissolving 8.2g of sodium hydroxide in 70g of deionized water, adding 1.012g of potassium permanganate into the solution, and heating and stirring the solution in a water bath at 50 ℃ to fully dissolve the potassium permanganate; dripping the former into the latter to obtain precipitate, stirring in water bath at 50 deg.C for 6 hr, filtering the precipitate, washing with 80 deg.C water for 3 times to obtain Na-OL-1;
adding 42.63g of magnesium chloride hexahydrate into 120g of deionized water, adding Na-OL-1 into the deionized water, fully stirring, transferring the mixture into a 250ml hydrothermal kettle, and carrying out hydrothermal crystallization at 160 ℃ for 36h to obtain the OMS-1 carrier.
(2) Preparation of the catalyst precursor
4.04g of ferric nitrate nonahydrate and 2.98g of zinc nitrate hexahydrate are dissolved in 10mL of deionized water, and the mixture is heated and stirred in a water bath at 50 ℃ and is uniformly mixed to obtain a steeping fluid.
And dispersing the impregnation liquid into 10g of OMS-1 carrier, fully stirring at room temperature, placing in an oven at 120 ℃ for drying for 5h, and then roasting at 400 ℃ for 3h to obtain a catalyst precursor, wherein the composition of the catalyst precursor A1 is 10% Fe-10% Zn/OMS-1 in terms of metal elements and based on the weight of the prepared catalyst precursor.
(3) Reduction of catalyst precursor
The catalyst precursor A1 was charged into a fixed-bed reactor, and H was fed into the fixed-bed reactor2Adjusting the pressure of the reactor to 0.1MPa, the airspeed to 10000h-1, raising the temperature of the reactor to 400 ℃ at the heating rate of 10 ℃/min, keeping the temperature for 4h, then reducing the temperature to 200 ℃, and switching the reducing atmosphere into a mixed gas of ethane and argon, wherein the mixed gas is the mixture of ethane and argonThe medium ethane accounts for 10 percent of the volume and the space velocity is 2000h-1And maintaining for 4 hours to obtain the Fischer-Tropsch synthesis catalyst.
(4) Production of α -olefins from synthesis gas
Heating the fixed bed reactor subjected to reduction activation in the step (3) to 320 ℃, introducing synthesis gas into the fixed bed reactor to start reaction, wherein the space velocity is 5000h-1The pressure is 1.5MPa, and the composition of the synthesis gas is H2:CO:N2The composition of the off-gas was analyzed by on-line gas chromatography at 56:28:16 (volume ratio). The results obtained after 50 hours of reaction are shown in Table 1.
Preparation example 2
(1) Preparation of the support
OMS-1 vector was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst A2 was prepared in the same manner as in example 1, except that the Zn and Fe contents of the catalyst were different, and the catalyst composition was 20% Fe-10% Zn/OMS-1.
(3) Reduction of catalyst precursor
The catalyst precursor A2 was reduced in the same manner as in example 1, except that the mixed gas of ethane and argon was changed to a reduction activation temperature of 250 ℃.
(4) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in example 1 to evaluate the catalyst precursor A2. the results obtained after 50 hours of reaction are shown in Table 1.
Preparation example 3
(1) Preparation of the support
OMS-1 vector was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor A3 was prepared in the same manner as in preparation example 1, except that the contents of the metals in the catalyst precursor were not uniform, and the composition of catalyst precursor A3 was 10% Fe-20% Zn/OMS-1.
(3) Reduction of catalyst precursor
Catalyst precursor A3 was reduced in the same manner as in example 1, except that the molar ratio of ethane to argon in the mixed gas of ethane and argon was 1: 20.
(4) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in preparation example 1 to evaluate the catalyst precursor A3. the results obtained after 50 hours of reaction are shown in Table 1.
Preparation example 4
(1) Preparation of the support
OMS-1 vector was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor A4 was prepared in the same manner as in preparation example 1, except that the contents of the metals in the catalyst precursor were not uniform, and the composition of catalyst precursor A4 was 20% Fe-20% Zn/OMS-1.
(3) Reduction of catalyst precursor
The catalyst precursor A4 was reduced in the same manner as in example 1, except that pure ethane gas was introduced for activation.
(4) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in preparation example 1 to evaluate the catalyst precursor A4. the results obtained after 50 hours of reaction are shown in Table 1.
Preparation example 5
(1) Preparation of the support
OMS-1 vector was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor A5 was prepared in the same manner as in preparation example 1, except that the contents of the metals in the catalyst precursor were not uniform, and the composition of catalyst precursor A5 was 10% Fe-5% Zn/OMS-1.
(3) Reduction of catalyst precursor
The catalyst precursor A5 was reduced in the same manner as in example 1, except that a mixed gas of ethylene and argon was fed, in which ethylene accounted for 10% by volume.
(4) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in preparation example 1 to evaluate the catalyst precursor A5. the results obtained after 50 hours of reaction are shown in Table 1.
Preparation example 6
(1) Preparation of the support
OMS-1 vector was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor A6 was prepared in the same manner as in preparation example 1, except that the contents of the metals in the catalyst precursor were not uniform, and the composition of catalyst precursor A6 was 5% Fe-10% Zn/OMS-1.
(3) Reduction of catalyst precursor
The catalyst precursor A5 was reduced in the same manner as in example 1 except that a mixed gas of ethane, ethylene and argon was fed, wherein both ethane and ethylene contained 5% by volume.
(3) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in preparation example 1 to evaluate the catalyst precursor A6. the results obtained after 50 hours of reaction are shown in Table 1.
Comparative example 1
(1) Preparation of the support
The carrier OMS-1 was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor D1 was prepared in the same manner as in preparation example 1.
(3) Reduction of catalyst precursor
Catalyst precursor D1 was reduced in the same manner as in example 1, except that the reduction was carried out without using hydrocarbons, and catalyst precursor D1 was reduced with hydrogen only.
(4) Production of α -olefins from synthesis gas
α -olefin was synthesized in the same manner as in preparation example 1 to evaluate the catalyst precursor D1. the results obtained after 50 hours of reaction are shown in Table 1.
Comparative example 2
(1) Preparation of the support
The carrier OMS-1 was prepared in the same manner as in preparation example 1.
(2) Preparation of the catalyst precursor
Catalyst precursor D2 was prepared in the same manner as in preparation example 1.
(3) Reduction of catalyst precursor
Catalyst precursor D2 was reduced in the same manner as in example 1, except that the reduction was carried out without using hydrocarbons, and catalyst precursor D2 was reduced with hydrogen only.
(4) Production of α -olefins from synthesis gas
The catalyst precursor D2 was reduced and evaluated in the same manner as in preparation example 1. The results obtained after 50 hours of reaction are shown in Table 1.
Table 1 evaluation test data
Figure BDA0001835574500000211
Figure BDA0001835574500000221
Note that the oil phase product is from C5+, and the oil phase product contains components such as alkane, α -alkene, isomeric hydrocarbon, oxygen-containing compound, etc.
The results show that after the supported catalyst precursor provided by the invention is subjected to two-step reduction activation (prereduction and reducition), when the supported catalyst precursor is used in the reaction of preparing α -olefin from synthesis gas, the conversion rate of carbon monoxide is high, the selectivity of α -olefin is high, the carbon number of the product is concentrated, and the reaction conditions are mild.
Example 1
The α -olefin production system shown in FIG. 2 is adopted in this embodiment, and comprises a raw material gas separation unit I, a steam reforming reaction unit II, a dry reforming reaction unit III, a Fischer-Tropsch synthesis reaction unit IV, a Fischer-Tropsch synthesis product separation unit V, and a circulation unit.
(1) And (2) sending shale gas with the flow rate of 220kmol/h and the pressure of 2.0MPa as a raw material gas A into a raw material gas separation unit I for low-temperature condensation separation, and removing sulfur, carbon and other impurities to obtain methane B with the sulfur mass content of less than 1 ppm.
And dividing the methane B into two parts by a flow divider, and respectively sending the two parts into a steam reforming reaction unit II and a dry reforming reaction unit III.
(2) Mixing the first stream of methane with medium-pressure steam C with the flow rate of 120kmol/h, the temperature of 370 ℃ and the pressure of 3MPa, raising the temperature of the mixture to 600 ℃, and then entering a fixed bed reactor of a steam reforming reaction unit II for reforming reaction to obtain steam reforming synthesis gas E. Wherein the molar ratio of methane to water vapor is 1: 3, the catalyst filled in the reactor is Ni/Al2O3(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)2O3Is α -Al2O3) The temperature in the catalyst bed layer is 900 ℃, the pressure in the reactor is 3MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 50000h-1。
(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 100kmol/h, the temperature of 370 ℃ and the pressure of 2MPa, then exchanging heat with a heat exchanger, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al2O3(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)2O3Is α -Al2O3) The temperature in the catalyst bed layer is 750 ℃, the pressure in the reactor is 2MPa, and the gas hourly volume space velocity is 80000h based on the total amount of methane and steam-1。
(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture meeting the hydrogen-carbon ratio of 2.1: 1 fischer-tropsch synthesis reaction feed G.
The fischer-tropsch synthesis reaction feed G is fed into a fischer-tropsch synthesis reactor (a fixed bed reactor) of the fischer-tropsch synthesis reaction unit IV, and is contacted with a fischer-tropsch synthesis catalyst (obtained after reduction activation of the catalyst precursor in preparation example 1) to perform a fischer-tropsch synthesis reaction. Wherein the temperature in the reactor is 320 ℃, the pressure in the reactor is 1.5MPa, and the total amount of the synthesis gas isReference, gas hour volume airspeed of 15000h-1。
(5) And (2) sending the Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV into a Fischer-Tropsch synthesis product separation unit V for separation, wherein the separation process comprises the steps of firstly carrying out gas-liquid separation to obtain α -olefin K and gas products, then carrying out cryogenic separation on the gas products to remove carbon dioxide in the gas products, and then carrying out cryogenic separation on the gas products from which the carbon dioxide is separated to obtain methane, unreacted hydrogen and carbon monoxide.
Discharging α -olefin K out of a system, circularly feeding separated carbon dioxide N into a dry reforming reaction unit III, respectively feeding separated methane M into a steam reforming reaction unit II and the dry reforming reaction unit III, circularly feeding a part L of separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest part of separated hydrogen and carbon monoxide out of the system as purge gas Z, wherein the amount of the circulated hydrogen and carbon monoxide L is 98% based on the total amount of the separated hydrogen and carbon monoxide.
The composition of the gaseous product stream exiting the reactor of the fischer-tropsch synthesis reaction unit was analyzed by an on-line gas chromatograph during the reaction and the results obtained after 50 hours of reaction are shown in table 2. The overall water consumption, carbon dioxide emissions, and energy efficiency of the system are listed in table 3.
Comparative example 1
The comparative example adopts the system shown in FIG. 1, and comprises a coal water slurry preparation unit I ', a coal gasification unit II', a water gas shift unit III ', a synthesis gas purification unit IV', a Fischer-Tropsch synthesis unit V 'and an α -olefin separation unit VI' which are connected in sequence.
The coal slurry C ' is conveyed into a coal gasification unit II ' and reacts with oxygen D ' under the conditions that the temperature is 1300 ℃ and the pressure is 3MPa to generate coal gasification crude synthesis gas E.
Adjusting the molar ratio of hydrogen to carbon monoxide of the coal gasification crude synthesis gas E to be 2: 1, and then removing acid gas and sulfide through a synthesis gas purification unit IV' to obtain purified synthesis gas (the molar ratio of hydrogen to carbon monoxide is 2.1: 1).
The purified synthesis gas obtained is conveyed into a Fischer-Tropsch synthesis unit V 'to carry out Fischer-Tropsch synthesis reaction in a fixed bed reactor (by adopting the catalyst prepared in the preparation example 1), and a Fischer-Tropsch reaction product N' containing olefin is generated. Wherein the temperature in the reactor is 320 ℃, the pressure in the reactor is 1.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 15000h-1。
α -olefin K ' is separated from the Fischer-Tropsch reaction product N ' through an α -olefin separation unit VI ', carbon dioxide H ' and methane G ' generated by the Fischer-Tropsch synthesis unit V ' are discharged outside, one part of unreacted synthesis gas (the content is 98 percent based on the total amount of the separated synthesis gas) Y ' is recycled to the Fischer-Tropsch synthesis unit V ', and the other part of unreacted synthesis gas is discharged out of the system as purge gas Z '.
The composition of the gaseous product stream exiting the Fischer-Tropsch reactor during the reaction was analyzed by an on-line gas chromatograph and the results obtained after 50 hours of reaction are shown in Table 2. The overall water consumption, carbon dioxide emissions, and energy efficiency of the system are listed in table 3.
Comparative example 2
α -olefin was produced in the same manner as in production example 1, except that the dry reforming reaction unit III was not provided, and methane (including fresh methane and recycled methane) was entirely fed into the steam reforming reaction unit II to carry out the reforming reaction.
Comparative example 3
α -olefin was produced in the same manner as in production example 1, except that the steam reforming reaction unit II was not provided and methane (including fresh methane and recycled methane) was entirely fed to the dry reforming reaction unit III to be subjected to reforming reaction.
Example 2
α -olefin was prepared by the same method as in preparation example 1, except that the Fischer-Tropsch synthesis catalyst was obtained by reductive activation of the catalyst precursor prepared in preparation example 2.
Example 3
α -olefin was prepared by the same method as in preparation example 1, except that the Fischer-Tropsch synthesis catalyst was obtained by reductive activation of the catalyst precursor prepared in preparation example 3.
Example 4
In this example, the reaction system shown in FIG. 2 was used, and the specific process flow was as follows.
(1) And (2) taking the coke oven gas with the flow rate of 500kmol/h and the pressure of 3.0MPa as a raw material gas A, sending the raw material gas A into a raw material gas separation unit I for low-temperature condensation separation, and removing sulfur, carbon and other impurities to obtain methane B with the sulfur mass content of less than 1 ppm.
And dividing the methane B into two parts by a flow divider, and respectively sending the two parts into a steam reforming reaction unit II and a dry reforming reaction unit III.
(2) Mixing the first stream of methane with medium-pressure steam C with the flow rate of 240kmol/h, the temperature of 370 ℃ and the pressure of 3MPa, raising the temperature of the mixture to 700 ℃, and then entering a fixed bed reactor of a steam reforming reaction unit II for reforming reaction to obtain steam reforming synthesis gas E. Wherein the molar ratio of methane to water vapor is 1:2, the catalyst filled in the reactor is Ni/Al2O3(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)2O3Is α -Al2O3) The temperature in the catalyst bed layer is 900 ℃, the pressure in the reactor is 3MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 50000h-1。
(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 200kmol/h, the temperature of 370 ℃ and the pressure of 2MPa, then exchanging heat with a heat exchange medium, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1.5, the catalyst filled in the reactor is Ni/Al2O3(based on the total amount of the catalyst, the content of Ni is10% by weight of Al2O3Is α -Al2O3) The temperature in the catalyst bed is 750 ℃, the pressure in the reactor is 2MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 100000h-1。
(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture meeting the hydrogen-carbon ratio of 2.5: 1 fischer-tropsch synthesis reaction feed G. The fischer-tropsch synthesis reaction feed G was fed into a fischer-tropsch synthesis reactor (a fixed bed reactor) of the fischer-tropsch synthesis reaction unit IV, and contacted with a fischer-tropsch synthesis catalyst (obtained by reduction activation of the catalyst precursor prepared in preparation example 4) to perform a fischer-tropsch synthesis reaction. Wherein the temperature in the reactor is 310 ℃, the pressure in the reactor is 1.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 10000h-1。
(5) And (2) sending the Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV into a Fischer-Tropsch synthesis product separation unit V for separation, wherein the separation process comprises the steps of firstly carrying out gas-liquid separation to obtain α -olefin K and gas products, then carrying out cryogenic separation on the gas products to remove carbon dioxide in the gas products, and then carrying out cryogenic separation on the gas products from which the carbon dioxide is separated to obtain methane, unreacted hydrogen and carbon monoxide.
Discharging α -olefin K out of a system, circularly feeding separated carbon dioxide N into a dry reforming reaction unit III, respectively feeding separated methane M into a steam reforming reaction unit II and the dry reforming reaction unit III, circularly feeding a part L of separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest part of separated hydrogen and carbon monoxide out of the system as purge gas Z, wherein the amount of the circulated hydrogen and carbon monoxide L is 20% based on the total amount of the separated hydrogen and carbon monoxide.
During the reaction, the composition of the off-gas was analyzed by an on-line gas chromatograph, and the results obtained after 50 hours of the reaction are shown in table 2. The overall water consumption, carbon dioxide emissions and energy efficiency of the plant are listed in table 3.
Example 5
The preparation example adopts the reaction system shown in fig. 2, and the specific process flow is as follows.
(1) And (3) taking the coke oven gas with the flow rate of 150kmol/h and the pressure of 1MPa as a raw material gas A, sending the raw material gas A into a raw material gas separation unit I for low-temperature condensation separation, and removing sulfur, carbon and other impurities to obtain methane B with the sulfur mass content of less than 1 ppm.
And dividing the methane B into two parts by a flow divider, and respectively sending the two parts into a steam reforming reaction unit II and a dry reforming reaction unit III.
(2) Mixing the first stream of methane with medium-pressure steam C with the flow rate of 300kmol/h, the temperature of 450 ℃ and the pressure of 3MPa, then exchanging heat with a heat exchanger, raising the temperature of the mixture to 700 ℃, and then entering a fixed bed reactor of a steam reforming reaction unit II for reforming reaction to obtain steam reforming synthesis gas E. Wherein the molar ratio of methane to water vapor is 1: 1, the catalyst filled in the reactor is Ni/Al2O3(Ni content 15 wt% in terms of element, based on the total amount of the catalyst, Al2O3Is α -Al2O3) The temperature in the catalyst bed is 860 ℃, the pressure in the reactor is 1MPa, and the gas hourly space velocity is 100000h based on the total amount of methane and steam-1。
(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 150kmol/h, the temperature of 450 ℃ and the pressure of 3MPa, then exchanging heat with a heat exchange medium, raising the temperature of the mixture to 700 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al2O3(Ni content 12 wt% in terms of element, based on the total amount of the catalyst, Al2O3Is α -Al2O3) The temperature in the catalyst bed layer is 650 ℃, the pressure in the reactor is 1.5MPa, and the gas hourly volume space velocity is 60000h based on the total amount of methane and water vapor-1。
(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture meeting the hydrogen-carbon ratio of 1.5: fee of 1Synthetic reaction feed G. The fischer-tropsch synthesis reaction feed G was fed into a fischer-tropsch synthesis reactor (a fixed bed reactor) of the fischer-tropsch synthesis reaction unit IV, and contacted with a fischer-tropsch synthesis catalyst (obtained by reduction activation of the catalyst precursor prepared in preparation example 5) to perform a fischer-tropsch synthesis reaction. Wherein the temperature in the reactor is 290 ℃, the pressure in the reactor is 2.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 20000h-1。
(5) And (2) sending the Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV into a Fischer-Tropsch synthesis product separation unit V for separation, wherein the separation process comprises the steps of firstly carrying out gas-liquid separation to obtain α -olefin K and gas products, then carrying out cryogenic separation on the gas products to remove carbon dioxide in the gas products, and then carrying out cryogenic separation on the gas products from which the carbon dioxide is separated to obtain methane, unreacted hydrogen and carbon monoxide.
Discharging α -olefin K out of a system, circularly feeding separated carbon dioxide N into a dry reforming reaction unit III, respectively feeding separated methane M into a steam reforming reaction unit II and the dry reforming reaction unit III, circularly feeding a part L of separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest part of separated hydrogen and carbon monoxide out of the system as purge gas Z, wherein the amount of the circulated hydrogen and carbon monoxide L is 15% based on the total amount of the separated hydrogen and carbon monoxide.
During the reaction, the composition of the off-gas was analyzed by an on-line gas chromatograph, and the results obtained after 50 hours of the reaction are shown in table 2. The overall water consumption, carbon dioxide emissions and energy efficiency of the plant are listed in table 3.
Example 6
α -olefin was produced by the same system and method as in preparation example 5, except that the Fischer-Tropsch synthesis catalyst was obtained by reductive activation of the catalyst precursor prepared in preparation example 6, and the temperature in the Fischer-Tropsch synthesis reactor was 330 ℃ and the pressure in the reactor was 1.5MPa, the gas hourly space velocity based on the total amount of synthesis gas was 30000h-1。
TABLE 2
Figure BDA0001835574500000281
TABLE 3
Numbering Water consumption (t/t)α -olefins) Carbon dioxide emission (t/t)α -olefins) Energy efficiency (%)
Example 1 14 0.6 57
Comparative example 1 21 6.8 34
Comparative example 2 19 2.3 36
Comparative example 3 26 0.8 45
Example 2 16 2.5 48
Example 3 17 2.3 47
Example 4 15 0.9 52
Example 5 15 1.0 50
Example 6 16 1.1 49
The energy efficiency is the sum of the heat value of α -olefin which is finally discharged from the device and the heat value of raw materials such as a coal-electricity steam catalyst solvent which is fed into the device, namely the heat value of α -olefin which is obtained/the comprehensive energy consumption required for producing the α -olefin, wherein the comprehensive energy consumption comprises the heat value of the raw materials and the energy consumption of public engineering, and mainly comprises indirect energy consumption such as the heat value of fuel coal and raw material coal, the electric energy consumed by a motor pump for the device process, circulating cooling water, boiler make-up water, process air, instrument air, fresh water and the like.
As can be seen from Table 3, the invention combines the methane steam reforming process and the methane dry reforming process to simultaneously utilize two greenhouse gases, namely carbon dioxide and methane, so that the greenhouse gases are converted into products with high added values, the greenhouse gas emission is reduced, and the resource and energy utilization rate of the whole process is obviously improved.
In conclusion, the production method can improve the utilization rate of two greenhouse gases, namely carbon dioxide and methane, in the production process of α -olefin, so that the greenhouse gases are converted into products with high added values, the greenhouse gas emission is reduced, the system energy consumption is reduced, and the resource and energy utilization rate of the whole process is obviously improved.
It should be noted by those skilled in the art that the described embodiments of the present invention are merely exemplary and that various other substitutions, alterations, and modifications may be made within the scope of the present invention. Accordingly, the present invention is not limited to the above-described embodiments, but is only limited by the claims.

Claims (24)

1. A process for producing α -olefins, comprising:
contacting methane with steam to carry out steam reforming reaction to obtain steam reforming synthesis gas;
contacting methane with carbon dioxide to carry out dry reforming reaction to obtain dry reforming syngas;
mixing at least part of the steam reforming synthesis gas and at least part of the dry weight integrated synthesis gas to prepare a Fischer-Tropsch synthesis reaction feed;
carrying out reduction activation on the Fischer-Tropsch synthesis catalyst precursor to obtain a Fischer-Tropsch synthesis catalyst;
contacting the Fischer-Tropsch synthesis reaction feed with the Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction to obtain a Fischer-Tropsch synthesis product material flow; and
α -olefins, methane and carbon dioxide are separated from the Fischer-Tropsch synthesis product stream,
wherein the Fischer-Tropsch synthesis catalyst precursor comprises a carrier, and an active metal oxide and an auxiliary agent metal oxide which are loaded on the carrier, the carrier is a manganese oxide molecular sieve, the active metal in the active metal oxide is a VIII group metal, the auxiliary agent metal in the auxiliary agent metal oxide is one or more of IIB group metal, alkaline earth metal and IB group metal,
wherein the reductive activation comprises:
pre-reducing the Fischer-Tropsch synthesis catalyst precursor in a first gas to obtain a pre-reduced catalyst; and
and reducing the pre-reduction catalyst in a second gas to obtain the Fischer-Tropsch synthesis catalyst.
2. The production method according to claim 1, wherein the active metal is one or more of Fe, Co, Ni, preferably Fe; the assistant metal is one or more of Zn, Cd, Cu and Ag, preferably Zn and/or Cu, and more preferably Zn.
3. The production method according to claim 1, wherein the content of the active metal oxide is 3 to 70% by weight, preferably 5 to 50% by weight, more preferably 8 to 30% by weight, in terms of the metal element; the content of the assistant metal oxide is 1-60 wt%, preferably 2-50 wt%; the carrier is present in an amount of 12 to 94% by weight, preferably 35 to 91% by weight.
4. The production method according to claim 1, wherein the catalyst is prepared by: and loading the active metal and the auxiliary metal on the carrier, and then drying and roasting the active metal and the auxiliary metal in sequence to obtain the supported catalyst.
5. The production method according to claim 4, wherein the supporting is performed by an impregnation method or a coprecipitation method.
6. The production method according to claim 5, wherein the impregnation method is an equal-volume impregnation method or a saturation impregnation method.
7. The production method according to claim 6, wherein the impregnation method comprises:
dissolving the soluble salt of the active metal and the soluble salt of the auxiliary metal in a solvent to obtain an impregnation solution; and
immersing the carrier in the impregnation liquid.
8. The production method according to claim 7, wherein the impregnation is one-time impregnation or stepwise impregnation.
9. The production process according to claim 7, wherein the soluble salt of the active metal and the soluble salt of the auxiliary metal are nitrates or hydrochlorides.
10. The production method according to claim 7, wherein the metal molar ratio of the active metal and the auxiliary metal in the impregnation liquid is 1: 0.2-5, preferably 1: 0.3-3.
11. The production method according to claim 7, wherein the temperature of the impregnation is 10-80 ℃, preferably 20-60 ℃; the impregnation time is 0.1 to 3 hours, preferably 0.5 to 1 hour.
12. The production process according to claim 4, wherein the drying temperature is 80-350 ℃, preferably 100-300 ℃; the drying time is 1 to 24 hours, preferably 2 to 12 hours.
13. The production method according to claim 4, wherein the temperature of the roasting is 250-900 ℃, preferably 300-850 ℃, more preferably 350-800 ℃; the calcination time is 0.5 to 12 hours, preferably 1 to 8 hours, and more preferably 2 to 6 hours.
14. The production method according to claim 1, wherein the first gas is hydrogen or a mixed gas of hydrogen and an inert gas.
15. The production process according to claim 14, wherein the temperature of the pre-reduction is 200-600 ℃, preferably 300-550 ℃; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; calculated as hydrogen gas, theThe volume space velocity of the first gas is 5000--1。
16. The production method according to claim 1, wherein the second gas is a gaseous hydrocarbon or a mixed gas of a gaseous hydrocarbon and an inert gas.
17. The production process according to claim 16, wherein the temperature of the re-reduction is 180-400 ℃, preferably 200-350 ℃; the pressure is 0-3MPa, preferably 0.1-1 MPa; the duration is 1 to 20 hours, preferably 2 to 15 hours; the volume space velocity of the second gas is 1000-30000 hours based on the gaseous hydrocarbon-1。
18. The production method according to claim 16, wherein the molar ratio of the inert gas to the gaseous hydrocarbon is 1 to 200: 1, preferably 5 to 20: 1.
19. the production process according to claim 16, wherein the gaseous hydrocarbon is a gaseous alkane and/or a gaseous alkene, preferably selected from C1-C4And/or C2-C4More preferably, the olefin(s) of (b) is/are one or more selected from methane, ethane, ethylene, propylene, propane, butane and butene.
20. The production process according to any one of claims 1 to 19, wherein the steam reforming reaction is carried out in a fixed bed reactor at a reaction temperature of 700-: 0.5-4, the hourly space velocity of the gas fed is preferably 10000--1Preferably 50000-100000 hours-1。
21. The production method according to any one of claims 1 to 19, wherein the dry reforming reaction is carried out in a fixed bed reactor at a reaction temperature of600-800 ℃, preferably 650-750 ℃, the reaction pressure is 0.1-5MPa, preferably 1-3MPa, and the molar ratio of methane to carbon dioxide is 1: 0.5-5, wherein the gas hourly volume space velocity of the feed is 10000--1Preferably 50000-100000 hours-1。
22. The production process according to any one of claims 1 to 19, wherein the Fischer-Tropsch synthesis reaction is carried out in a fixed bed reactor at a reaction temperature of 200--1More preferably 5000--1Most preferably 10000--1。
23. The production process according to any one of claims 1 to 19, wherein the molar ratio of hydrogen to carbon monoxide in the fischer-tropsch synthesis reaction feed is in the range of from 0.4 to 3: 1, preferably 0.6 to 2.8: 1, more preferably 0.8 to 2.6: 1, most preferably 1.5-2.5: 1.
24. the production method according to any one of claims 1 to 19, further comprising recycling the separated methane to the steam reforming reaction and/or the dry reforming reaction, and recycling the separated carbon dioxide to the dry reforming reaction.
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