CN108728153B

CN108728153B - Production method of low-carbon olefin and low-carbon olefin production system

Info

Publication number: CN108728153B
Application number: CN201710256685.XA
Authority: CN
Inventors: 晋超; 吴玉; 张荣俊; 侯朝鹏; 孙霞; 阎振楠; 夏国富; 李明丰
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2017-04-19
Filing date: 2017-04-19
Publication date: 2021-05-14
Anticipated expiration: 2037-04-19
Also published as: CN108728153A

Abstract

The invention discloses a production method and a production system of low-carbon olefin, wherein the preparation method comprises the steps of contacting methane with water under the condition of steam reforming reaction to obtain steam reforming synthesis gas; under the condition of dry reforming reaction, contacting methane with carbon dioxide to obtain dry reforming syngas; integrating steam reforming synthesis gas and dry weight synthesis gas to prepare Fischer-Tropsch synthesis reaction feed, contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at a low carbon olefin production temperature to obtain a Fischer-Tropsch synthesis product material flow, separating methane and carbon dioxide in the Fischer-Tropsch synthesis product material flow, sending the methane into a steam reforming step and/or a dry weight reforming step, and sending the carbon dioxide into a dry reforming step. The method for producing the low-carbon olefin can effectively reduce the energy consumption of a system and the emission of greenhouse gases (such as carbon dioxide).

Description

Production method of low-carbon olefin and low-carbon olefin production system

Technical Field

The invention relates to a production method of low-carbon olefin and also relates to a production system of low-carbon olefin.

Background

Olefins are important basic chemical materials in the production of chemical industry, and are also a mark for measuring the development level of the national petrochemical industry. The existing methods for preparing low-carbon olefins can be divided into three main categories according to raw materials: oil routes, natural gas routes, and coal routes. The method for preparing low-carbon olefin by adopting light oil cracking, namely a method for preparing low-carbon olefin by using a petroleum route, is adopted by most countries in the world and accounts for about 65 percent of the yield of olefin. The natural gas is used as a raw material, the technology for preparing the low-carbon olefin by an oxidative coupling method or a Bensen method is adopted, ethylene is mainly used in the product, and the yield of propylene is low. The research of preparing olefin by using coal-based synthesis gas through methanol is rapidly developed, and a plurality of sets of process devices are built in China.

The energy sources in China are rich in coal, more natural gas and lack of oil, and the indirect conversion of the coal or the natural gas into clean and efficient liquid fuel through Fischer-Tropsch synthesis is an important aspect of reasonably utilizing the resources and a main technical approach for relieving the contradiction between supply and demand of petroleum in China. In recent years, in the field of coal chemical industry in China, coal rapidly rises through methanol to olefin, and the direct preparation of olefin from coal through synthetic gas (FTO process) is another process for preparing olefin from coal. The process comprises the steps of firstly converting coal or natural gas into synthesis gas (carbon monoxide and hydrogen), and directly preparing the low-carbon olefin with the carbon atom number less than or equal to 4 through Fischer-Tropsch synthesis.

The technological process for preparing olefin by the currently commonly used 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 low-carbon olefin separation unit VI which are connected in sequence, the specific process comprises the steps of preparing coal water slurry C from pulverized coal A and water B in the coal water slurry preparation unit I, conveying the coal water slurry C into the coal gasification unit II, reacting with oxygen D to generate coal gasification crude synthetic gas E, adjusting the molar ratio of hydrogen and carbon monoxide of the coal gasification crude synthetic gas E through the water gas conversion unit III to obtain converted crude synthetic gas F meeting the Fischer-Tropsch synthesis reaction requirement, removing acid gas and sulfide M through the synthetic gas purification unit IV to obtain purified synthetic gas J, conveying the obtained purified synthetic gas J into the Fischer-Tropsch synthesis unit V to perform Fischer-Tropsch synthesis reaction to generate a Fischer-Tropsch reaction product N containing olefin, and separating the low-carbon olefin K from the Fischer-Tropsch reaction product N through a low-carbon olefin separation unit VI, discharging carbon dioxide H and methane G generated by a Fischer-Tropsch synthesis unit V, circulating a part of unreacted synthesis gas Y back to the Fischer-Tropsch synthesis unit V, and discharging the other part of unreacted synthesis gas serving as purge gas Z out of the system. The main problems of the process for preparing the low-carbon olefin are as follows: 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 Fischer-Tropsch synthesis product distribution is limited by an Anderson-Schulz-Flory rule (the molar distribution of chain growth decreasing according to indexes), 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 process is low, and the industrial process of the Fischer-Tropsch synthesis process is seriously influenced. The coal gasification process uses a large amount of cooling water and discharged sewage to keep the water consumption high.

Therefore, there is still a need to optimize the fischer-tropsch synthesis process, reducing the system energy consumption and the emission of greenhouse gases.

Disclosure of Invention

The invention aims to provide a method for producing low-carbon olefin, which can effectively reduce the energy consumption of a system and the emission of greenhouse gases.

According to a first aspect of the present invention, there is provided a process for producing lower olefins, the process comprising the steps of:

s11, under the condition of steam reforming reaction, contacting methane with steam to obtain steam reforming synthesis gas;

s21, under the condition of dry reforming reaction, contacting methane with carbon dioxide to obtain dry reforming syngas;

s31, mixing at least part of steam reforming synthesis gas and at least part of dry weight synthesis gas to prepare Fischer-Tropsch synthesis reaction feed, and contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at the reaction temperature of producing low-carbon olefin to obtain a Fischer-Tropsch synthesis product material flow;

s41, separating low-carbon olefin, methane and carbon dioxide from the Fischer-Tropsch synthesis product stream, sending the separated methane into one or both of S11 and S21, and sending the separated carbon dioxide into S21.

According to a second aspect of the present invention, there is provided a low carbon olefin production system comprising 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 dry reforming reaction unit is used for contacting methane and carbon dioxide to carry out dry reforming reaction to obtain dry reforming synthesis gas;

the synthesis gas mixing unit is used for mixing the steam reforming synthesis gas with the dry weight integrated synthesis gas to prepare a Fischer-Tropsch synthesis reaction feed, and sending the Fischer-Tropsch synthesis reaction feed into the Fischer-Tropsch synthesis reaction unit;

the Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor and is used for contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst to obtain a Fischer-Tropsch synthesis product material flow containing low-carbon olefins;

the Fischer-Tropsch synthesis reaction product separation unit is used for separating the Fischer-Tropsch synthesis product material flow to obtain methane, carbon dioxide, low-carbon olefin, optional hydrogen and optional carbon monoxide;

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.

According to the method and the system for producing the low-carbon olefin, the energy consumption of the system and the emission of greenhouse gases (such as carbon dioxide) can be effectively reduced.

Drawings

Fig. 1 is used to illustrate a typical process flow of the prior art for directly preparing low carbon olefins from coal via synthesis gas (FTO process).

FIG. 2 is a schematic diagram illustrating a process and system for producing lower olefins according to the present invention.

FIG. 3 shows theta-Al prepared in preparation example 1₂O₃X-ray diffraction pattern of (a).

FIG. 1 description of reference numerals

I: a coal water slurry preparation unit II: a coal gasification unit III: water gas shift unit

IV: synthesis gas purification unit V: Fischer-Tropsch synthesis unit VI: low carbon olefin separation unit

A: pulverized coal B: and C, water C: coal water slurry

D: oxygen E: coal gasification crude synthesis gas F: shifted raw synthesis gas

G: methane H: carbon dioxide K: low carbon olefin

M: acid gas and sulfide N: Fischer-Tropsch reaction product Y: unreacted synthesis gas

Z: purge gas

FIG. 2 description of reference numerals

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 for recycle: low carbon olefin

M: methane N: carbon dioxide Z: purge gas

Detailed Description

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.

In the present invention, the term "lower olefin" means an olefin having a carbon number of 4 or less.

In step S11, 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 the methane is contacted with the steam may be 0.1 to 5MPa, preferably 1 to 3MPa, said pressure being a gauge pressure. The steam reforming reaction may be carried out in a common reactor. Preferably, the steam reforming reaction is carried out in a fixed bed reactor. The hourly space velocity of the feed gas may be 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 one example, the steam reforming catalyst contains a carrier and an active component supported on the carrier. The carrier can be one or the combination of more than two of alumina, silica, zirconia and silicon carbide. Preferably, the carrier is alumina, and may be gamma-Al in particular₂O₃、θ-Al₂O₃、δ-Al₂O₃And alpha-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 loading of the active ingredient 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 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, said pressureAs gauge pressure. The dry reforming reaction can be carried out in a common reactor. Preferably, the dry reforming reaction is carried out in a fixed bed reactor. The hourly space velocity of the feed gas may be 10000-^-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, the dry reforming catalyst contains a carrier and an active component supported on the carrier. The carrier can be one or the combination of more than two of alumina, silica, zirconia and silicon carbide. Preferably, the carrier is alumina, and may be gamma-Al in particular₂O₃、θ-Al₂O₃、δ-Al₂O₃And alpha-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 loading of the active ingredient 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.

According to the process for producing lower olefins of the present invention, methane, which is one of the raw materials for steam reforming of methane and dry reforming of methane, may be methane of various sources, preferably methane separated from a methane-rich raw material gas. At this time, the method for producing lower olefins according to the present invention further includes a step S10 of separating methane from the feed gas containing methane in S10. The feed gas may be a common methane-rich mixture. Specifically, the raw material gas may be one or more selected from shale gas, coal bed gas, natural gas, refinery gas and coke oven gas.

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 feed gas.

According to the process for producing lower olefins of the present invention, the content of elemental sulfur in methane, which is one of the raw materials for steam reforming and dry reforming, is generally 20ppm or less, preferably 10ppm or less, more preferably 5ppm or less, and still more preferably 1ppm or less, by mass.

According to the method for producing the low-carbon olefin, the raw material utilization rate of the method can be further improved by controlling the amount of the methane fed to the step S11 and the step S21 according to the reaction property of steam reforming and dry reforming and the requirement 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 weight integrated syngas are blended to formulate a fischer-tropsch synthesis reaction feed that meets the fischer-tropsch synthesis feed hydrogen to carbon ratio (i.e., the molar ratio of hydrogen to carbon monoxide). From the viewpoint of improving the selectivity of the low-carbon olefin, the molar ratio of hydrogen to carbon monoxide in the feed of the Fischer-Tropsch synthesis reaction is preferably 0.4-3: 1, more preferably 0.6 to 2.5: 1, more preferably 0.8 to 2.2: 1, more preferably 1.5 to 2.2: 1.

according to the production method of the low-carbon olefin, in the step S31, the Fischer-Tropsch synthesis reaction feed material is contacted with the Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction, so that a Fischer-Tropsch synthesis product material flow is obtained.

The Fischer-Tropsch synthesis catalyst can be a conventional catalyst with a catalytic effect on Fischer-Tropsch synthesis reaction. In a preferred embodiment, the fischer-tropsch synthesis catalyst comprises a support and a group VIII metal element supported on the support.

According to the fischer-tropsch synthesis catalyst of this preferred embodiment, the support is alumina, specific examples of which may include, but are not limited to: gamma-Al₂O₃、θ-Al₂O₃、δ-Al₂O₃And alpha-Al₂O₃One or more than two of them. The parameters of the specific surface area, the average pore diameter, the particle size distribution and the like of the alumina can be optimized according to the specific type of the alumina so as to further improve the catalytic performance of the catalyst. As an example, for γ -Al₂O₃The pore volume can be 0.6-1mL/g, preferably 0.65-0.9mL/g, more preferably 0.65-0.85 mL/g; the average pore diameter may be 8-35nm, preferably 12-30nm, more preferably 15-20 nm; the content of particles having a particle diameter in the range of 70 to 150 μm may be 80% by volume or more, preferably 85% by volume or more, more preferably 90% by volume or more; the specific surface area can be 100-300m²Per g, preferably 120-250m²(ii)/g, more preferably 150-²(ii) in terms of/g. As another example, for theta-Al₂O₃The pore volume can be 0.3-0.8mL/g, preferably 0.35-0.7mL/g, more preferably 0.4-0.6 mL/g; the average pore size may be 12-40nm, preferably 15-35nm, more preferably 18-25 nm; the content of particles having a particle diameter in the range of 70 to 150 μm may be 80% by volume or more, preferably 85% by volume or more, more preferably 90% by volume or more; the specific surface area can be 50-200m²A/g, preferably from 60 to 150m²A/g, more preferably from 65 to 100m²(ii) in terms of/g. In the present invention, the specific surface area, the pore volume and the average pore diameter are measured by a nitrogen adsorption method, specifically, N is 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.

Preferably, the carrier contains theta-Al₂O₃. By introducing theta-Al into the support₂O₃The Fischer-Tropsch synthesis catalyst can obtain higher catalytic activity and catalytic stability, shows higher CO conversion rate and low-carbon olefin selectivity, and has more excellent activity stability. Generally, the content of the theta alumina may be 10 wt% or more, preferably 20 wt% or more, more preferably 30 wt% or more, further preferably 40 wt% or more, and still further preferably 40 wt% or more, based on the total amount of alumina in the catalystIs 50% by weight or more. Particularly preferably, the support is theta alumina.

The theta-Al₂O₃Can be obtained commercially or by reacting gamma-Al₂O₃And baking to obtain the final product. Specifically, γ -Al may be added₂O₃The calcination is carried out at a temperature of 700-1050 deg.C, preferably 780-1050 deg.C. The duration of the firing may be selected based on the temperature of firing sufficient to convert the gamma-Al₂O₃Conversion to theta-Al₂O₃The standard is. In general, the duration of the calcination may be from 0.5 to 5 hours, preferably from 1 to 4 hours. The firing is performed in an air atmosphere.

In the fischer-tropsch synthesis catalyst according to this preferred embodiment, the group VIII metal element may be a group VIII noble metal element, a group VIII non-noble metal element, or a combination of a group VIII noble metal element and a group VIII non-noble metal element as an active component of the catalyst. In a preferred embodiment, the group VIII metal element is a group VIII non-noble metal element, and specific examples thereof may include, but are not limited to, one or two or more of Fe, Co, and Ni. More preferably, the group VIII metal element is Fe.

According to the Fischer-Tropsch synthesis catalyst of the preferred embodiment, at least a part of the group VIII metal element has a valence lower than the maximum oxidation state of the metal element. Generally, the content of the group VIII metal element having a valence lower than the maximum oxidation valence thereof may be 30% by weight or more, preferably 40% by weight or more, more preferably 45% by weight or more, further preferably 50% by weight or more (e.g., 55% by weight or more), further preferably 60% by weight or more, and particularly preferably 75% by weight or more, in terms of the element, based on the total amount of the group VIII metal element in the catalyst. The maximum content of the group VIII metal element having a valence lower than its maximum oxidation valence may be 100% by weight or less than 100% by weight, such as 99%, 96%, 93%, 90%, 87% by weight, based on the total amount of the group VIII metal element in the fischer-tropsch synthesis catalyst. According to the preferred embodimentThe Fischer-Tropsch synthesis catalyst can be directly used for catalytic reaction without additional reduction activation. In the present invention, the term "maximum oxidation state" refers to the valence of the metal element when it is completely oxidized, and in the case of Fe, the maximum oxidation state refers to iron oxide (Fe)₂O₃) The valence of the middle iron element is + 3. In the invention, the VIII group metal elements with different valence states and the content thereof are measured by an X-ray photoelectron spectroscopy.

According to the Fischer-Tropsch synthesis catalyst of this preferred embodiment, in a particularly preferred example, the group VIII metal element is Fe, and the Fischer-Tropsch synthesis catalyst has an X-ray photoelectron spectrum in which peaks corresponding to FeO (typically found at 711.9eV and 724.4 eV) and peaks corresponding to Fe are present₅C₂The spectral peak of (usually occurs at 717.9 eV). The fischer-tropsch synthesis catalyst according to this particularly preferred example has more excellent catalytic performance and activity stability. In this particularly preferred embodiment, the content of Fe determined from the peak corresponding to FeO and the content of Fe determined from the peak corresponding to Fe calculated on an elemental basis₅C₂The ratio of the Fe content determined by the spectral peaks of (a) may be 8 to 25: 1. from the viewpoint of further improving the catalytic activity and catalytic stability of the catalyst, Fe determined from the peak corresponding to FeO and Fe determined from the peak corresponding to Fe₅C₂The ratio of Fe determined by the spectral peaks of (a) is preferably 9 to 20: 1, more preferably 9.5 to 15: 1, more preferably 10 to 12: 1. from the viewpoint of further improving the catalytic activity and catalytic stability, the total amount of Fe determined by X-ray photoelectron spectroscopy is taken as a reference by the element, and the total amount of Fe is calculated by the peak corresponding to FeO and the peak corresponding to Fe₅C₂The content of Fe determined by the peak of the spectrum may be 30% or more (e.g., 30 to 99%), preferably not less than 40% (e.g., 40 to 96%), more preferably not less than 45% (e.g., 45 to 93%), further preferably not less than 50% (e.g., 55% or more), further preferably not less than 60% (e.g., 60 to 90%), particularly preferably not less than 75% (e.g., 75 to 87%).

In the present invention, X-ray photoelectron spectroscopy was carried out on an ESCALB 250 type X-ray photoelectron spectrometer equipped with Thermo Avantage V5.926 software, manufactured by Thermo ScientificThe test is carried out, wherein the excitation source is monochromatized Al K alpha X-ray, the energy is 1486.6eV, the power is 150W, the transmission energy for 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.

According to the fischer-tropsch synthesis catalyst of this preferred embodiment, the content of the group VIII metal element may be conventionally selected. Generally, the group VIII metal element may be present in an amount of from 3 to 30 wt%, preferably from 5 to 25 wt%, more preferably from 8 to 20 wt%, and even more preferably from 10 to 15 wt%, calculated as element, based on the total amount of the Fischer-Tropsch synthesis catalyst. In the present invention, the kind and content of each metal element in the catalyst and the catalyst precursor were measured by an X-ray fluorescence spectrum analysis method specified in RIPP 132-92 (compiled in methods for petrochemical engineering (RIPP experiments), Yanggui et al, science publishers, 1 st edition at 9/1990, p. 371-379).

According to the fischer-tropsch synthesis catalyst of this preferred embodiment, in addition to the group VIII metal element supported on the carrier, the catalyst may further contain a second metal element and/or a third metal element supported on the carrier, and preferably contains the third metal element and optionally the second metal element supported on the carrier. The catalyst containing the second metal element and/or the third metal element exhibits more excellent catalytic activity and catalytic stability. In the present invention, "optional" means either with or without.

The second metal element is one or more selected from alkali metal elements, alkaline earth metal elements and group IVB metal elements. Specific examples of the second metal element may include, but are not limited to: one or more of Li, Na, K, Mg, Ca, Zr and Ti. Preferably, the second metal element is one or two or more of Li, Zr, Mg, and K. The second metal element may be present in an amount of 0.1 to 15 wt%, preferably 1 to 12 wt%, more preferably 2 to 11 wt%, and still more preferably 5 to 7 wt%, calculated as element, based on the total amount of the fischer-tropsch synthesis catalyst.

From the viewpoint of further improving the catalytic activity and activity stability of the catalyst, the second metal element preferably contains a group IVB metal element (preferably Zr and/or Mg) and an alkali metal element (preferably K and/or Li), and the content of the group IVB metal element (preferably Zr and/or Mg) is preferably 0.5 to 8% by weight, more preferably 1 to 4% by weight, further preferably 2 to 3% by weight, and the content of the alkali metal element (preferably K and/or Li) is preferably 1 to 8% by weight, more preferably 2 to 6% by weight, further preferably 3 to 4% by weight, in terms of the elements, based on the total amount of the catalyst.

The third metal element is one or more selected from rare earth metal elements, and specific examples thereof may include, but are not limited to, one or more of lanthanum, cerium (Ce), praseodymium, and neodymium. Preferably, the third metal element is Ce. The content of the third metal element may be 0.1 to 10 wt%, preferably 0.5 to 6 wt%, more preferably 0.8 to 3 wt%, and still more preferably 1.2 to 2 wt% in terms of element based on the total amount of the fischer-tropsch synthesis catalyst.

The fischer-tropsch synthesis catalyst according to this preferred embodiment preferably contains the second metal element and the third metal element supported on the carrier, from the viewpoint of further improving the catalytic activity and catalytic stability of the catalyst. When the catalyst contains both the second metal element and the third metal element, the second metal element is more preferably a group IVB metal element and/or an alkali metal element, further preferably a group IVB metal element and an alkali metal element, and further preferably Zr and K, and the third metal element is more preferably Ce, so that more excellent catalytic activity and catalytic stability can be obtained.

CO₂TPD (i.e., temperature programmed desorption of CO)₂) Can be used for characterizing the desorption performance of the catalyst on hydrocarbon molecules, in CO₂In the TPD spectrum, the higher the temperature at which the desorption peak appears indicates that the catalyst is favorable for the desorption of the hydrocarbon molecules, and for a plurality of catalysts with desorption peaks at the same position, the larger the peak area of the catalyst is, the desorption of the hydrocarbon molecules is performedThe stronger the attachment capacity. The Fischer-Tropsch synthesis catalyst according to this preferred embodiment exhibits unique CO₂TPD desorption profile, with a desorption peak (herein, this desorption peak is referred to as CO) in the temperature range of 300-₂High temperature desorption peak). The CO is₂The peak area of the high-temperature desorption peak is generally 0.3 to 2.5a.u (arbitrary units), preferably 0.5 to 2a.u (arbitrary units). CO of the Fischer-Tropsch Synthesis catalyst according to this preferred embodiment₂In the desorption spectrum of TPD, another desorption peak (herein, the desorption peak is referred to as CO) exists in the temperature range of 100-200 ℃, preferably 150-190 ℃₂Low temperature desorption peak). The CO is₂The peak area of the low-temperature desorption peak is generally 0.5 to 3.5a.u (arbitrary units), preferably 1 to 3.2a.u (arbitrary units), and more preferably 2 to 3a.u (arbitrary units).

CO-TPD (i.e., temperature programmed desorption of CO) can be used to characterize the catalyst's ability to dissociate CO, with higher temperatures at which CO desorption peaks occur indicating higher activity of the catalyst. For multiple catalysts with desorption peaks at the same location, a catalyst with a larger peak area favors CO dissociation. According to the Fischer-Tropsch synthesis catalyst of the preferred embodiment, a desorption peak (herein, the desorption peak is referred to as a CO high-temperature desorption peak) exists in a temperature interval of 300-700 ℃, preferably 400-650 ℃, and more preferably 450-600 ℃. The peak area of the CO high-temperature desorption peak is generally 0.5-7a.u (arbitrary unit), preferably 1-6a.u (arbitrary unit), and more preferably 2-5.5a.u (arbitrary unit). In the CO-TPD desorption spectrum of the Fischer-Tropsch synthesis catalyst according to the preferred embodiment, another desorption peak (herein, the desorption peak is referred to as a CO low-temperature desorption peak) exists in the temperature interval of 100-200 ℃, preferably 150-190 ℃. The peak area of the CO low-temperature desorption peak is generally 0.5-2a.u (arbitrary unit), and preferably 0.8-1.6a.u (arbitrary unit).

In the present invention, 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 present invention, desorption is carried outThe position of the peak refers to the position at which the peak of the desorption peak is located.

The fischer-tropsch synthesis catalyst may be obtained by reductive activation of a fischer-tropsch synthesis catalyst precursor, the reductive activation comprising the steps of:

(1) carrying out pre-reduction on a Fischer-Tropsch synthesis catalyst precursor in a first gas to obtain a pre-reduction catalyst;

(2) and carrying out reduction activation on the pre-reduction catalyst in a second gas to obtain a reduction activation catalyst.

The Fischer-Tropsch synthesis catalyst precursor comprises a carrier and a VIII group metal element loaded on the carrier.

In the fischer-tropsch synthesis catalyst precursor, the group VIII metal element is supported on the carrier in the form of an oxide, and the valence of the group VIII metal element in the oxide is the highest oxidation valence of the metal element (herein, the oxide in which the valence of the metal element in the metal oxide is the highest oxidation valence is also referred to as a complete oxide). A typical example of the fischer-tropsch synthesis catalyst precursor is a catalyst precursor which has undergone drying and calcination (i.e. heat treatment in an oxygen atmosphere) during preparation without being subjected to reduction treatment. The VIII group metal element in the form of complete oxide needs to be subjected to reduction activation so as to have catalytic performance meeting the use requirement. The types and contents of the group VIII metal elements can be referred to the description of the Fischer-Tropsch synthesis catalyst part, and are not described in detail here.

The first gas is hydrogen or a mixed gas of hydrogen and inert gas. The inert gas may be one or two or more selected from nitrogen and a group zero element gas, and the group zero element gas may be, for example, argon. Preferably, the inert gas is nitrogen and/or argon. When the first gas is a mixed gas of hydrogen and an inert gas, the molar ratio of the inert gas to the hydrogen may be 1 to 30: 1.

the contacting temperature of the Fischer-Tropsch synthesis catalyst precursor with the first gas is such that at least part of the group VIII metal element in the Fischer-Tropsch synthesis catalyst precursor in the highest oxidation state is reduced (i.e. reduced in valence state).

In particular, the Fischer-Tropsch synthesis catalyst precursor and the first gas may be contacted at a temperature of 200-. The volume space velocity of the first gas (in terms of hydrogen) can be 5000-^-1Preferably 10000-^-1. The pressure in the reactor may be in the range of 0 to 2.5MPa, preferably 0.1 to 2MPa, in terms of gauge pressure. The duration of the pre-reduction may be selected depending on the temperature of the pre-reduction. Generally, the duration of the pre-reduction may be 1 to 20 hours, preferably 2 to 10 hours, more preferably 5 to 8 hours.

The second gas is a hydrocarbon that is gaseous at the reduction activation temperature, or a mixed gas of a hydrocarbon that is gaseous at the reduction activation temperature and an inert gas. The hydrocarbon which is gaseous at the reduction activation temperature may be one or two or more selected from an alkane which is gaseous at the reduction activation temperature and an alkene which is gaseous at the reduction activation temperature, and may be, for example, selected from C₁-C₄Alkane and C₂-C₄One or more kinds of olefins. Specific examples of the hydrocarbon that is gaseous at the reduction activation temperature may include, but are not limited to, one or two or more of methane, ethane, ethylene, propylene, propane, butane, and butene. From the viewpoint of further improving the catalytic activity and catalytic stability of the finally produced catalyst, the hydrocarbon which is gaseous at the reduction activation temperature is preferably one or more selected from alkanes which are gaseous at the reduction activation temperature, and more preferably selected from C₁-C₄One or two or more kinds of alkanes, and ethane is more preferable. The inert gas may be one or two or more selected from nitrogen and a group zero element gas, and the group zero element gas may be, for example, argon. Preferably, the inert gas is nitrogen and/or argon. When the second gas is a mixed gas of a hydrocarbon that is gaseous at the reduction activation temperature and an inert gas, the molar ratio of the inert gas to the hydrocarbon that is gaseous at the reduction activation temperature may be 1 to 200: 1, preferably 1 to 100: 1, more preferably 10 to 50:1, more preferably 15 to 30: 1.

the reduction activation may be carried out at a temperature of 150 ℃ to 500 ℃, preferably 180 ℃ to 450 ℃, more preferably 200 ℃ to 400 ℃. The volume space velocity of the second gas (in terms of hydrocarbon that is gaseous at the reduction activation temperature) may be 5000-^-1Preferably 10000-^-1. In carrying out the reduction activation, the pressure in the reactor may be 0 to 2.5MPa, preferably 0.1 to 2MPa, in terms of gauge pressure. The duration of the reductive activation may be selected according to the temperature of the reductive activation and the pressure of the second gas. Generally, the duration of the reductive activation may be 1 to 20 hours, preferably 2 to 15 hours, more preferably 4 to 12 hours.

The fischer-tropsch synthesis catalyst precursor may be prepared by a process comprising the steps of: loading a compound containing VIII group metal elements and an optional compound containing auxiliary elements on a carrier, and roasting the carrier to obtain a catalyst precursor, wherein the carrier is alumina.

The alumina may be used as a support without supporting an additional modifying element (i.e., pure alumina may be used as a support), or may be used as a support after modification. In a preferred embodiment, at least part of the support is alumina containing a modifying element. In general, the content of the alumina containing the modifying element may be 10% by weight or more, preferably 30% by weight or more, more preferably 50% by weight or more, further preferably 70% by weight or more, and further preferably 90% by weight or more, based on the total amount of the carrier. Particularly preferably, the support is alumina containing a modifying element.

The modifying element is one or more than two selected from alkali metal elements, alkaline earth metal elements and IVB group metal elements. Specific examples of the modifying element may include, but are not limited to, one or two or more of Li, Na, K, Mg, Ca, Zr, and Ti. More preferably, the modifying element is one or two or more of K, Mg and Zr.

From the viewpoint of further improving the catalytic activity and activity stability of the finally prepared catalyst, the content of the modifying element may be 0.1 to 15% by weight, preferably 0.5 to 12% by weight, more preferably 1 to 10% by weight, and still more preferably 1.5 to 8% by weight, in terms of the element, based on the total amount of the carrier. The modifying element is preferably supported on the alumina prior to the group VIII metal element and optionally the promoter element.

The alumina containing the modifying element can be obtained by a conventional method. For example, the modifying element may be supported on the alumina during the preparation of the alumina, such as by coprecipitation, while the alumina is being prepared.

In a preferred example, alumina loaded with a compound containing a modifying element is calcined to obtain a modifying element-containing alumina. The calcination may be carried out under conventional conditions, and generally, the calcination may be carried out at a temperature of 300-900 deg.C, preferably 400-800 deg.C, and the duration of the calcination may be selected depending on the calcination temperature, and may be generally 0.5 to 8 hours, preferably 1 to 6 hours. The firing is performed in an air atmosphere.

In this preferred embodiment, the compound containing the modifying element may be supported on the alumina by means of impregnation. Specifically, alumina may be impregnated with an impregnation solution containing a compound containing a modifying element, and the alumina having the impregnation solution adsorbed thereon may be sequentially dried and calcined to obtain alumina containing a modifying element.

In this preferred example, the modifying element-containing compound may be a modifying element-containing water-soluble salt and/or a water-soluble base, and specific examples thereof may include, but are not limited to: one or more of nitrate, oxalate, acetate, chloride, hydroxide, carbonate, bicarbonate and phosphate. In this preferred embodiment, the impregnation may be carried out by a conventional impregnation method such as saturation impregnation or excess impregnation. The impregnation may be carried out at ambient temperature.

In this preferred embodiment, the drying may be carried out under conditions sufficient to remove volatile species (primarily solvent in the impregnating solution) from the alumina on which the impregnating solution is adsorbed. Specifically, the drying may be performed at a temperature of 50 to 300 ℃, preferably 80 to 300 ℃, more preferably 120 ℃ to 300 ℃, and the drying may be performed under normal pressure (i.e., 1 atm, the same applies) or under reduced pressure. The duration of the drying may be selected depending on the temperature of the drying and the pressure of the drying, and may be generally 1 to 12 hours, preferably 2 to 6 hours. The drying may be performed in an air atmosphere.

The oxide of the group VIII metal element and/or the precursor of the oxide of the group VIII metal element may be supported on the carrier by a conventional method. For example, the co-precipitation method may be used to support the oxide of the group VIII metal element on the carrier during the preparation of the alumina (or the alumina containing the modifying element).

In a more preferred embodiment, a carrier is impregnated with an impregnation solution containing an oxide of a group VIII metal element and/or a precursor of an oxide of a group VIII metal element, and the carrier having the impregnation solution adsorbed thereon is dried to obtain a carrier having the oxide and/or the precursor supported thereon.

The type of the precursor of the group VIII metal element oxide may be selected depending on the solvent of the immersion liquid so that the precursor of the group VIII metal element oxide is soluble in the solvent, and may be one or more selected from the group consisting of an oxalate of the group VIII metal element, a nitrate of the group VIII metal element, a sulfate of the group VIII metal element, an acetate of the group VIII metal element, a chloride of the group VIII metal element, a carbonate of the group VIII metal element, a basic carbonate of the group VIII metal element, a hydroxide of the group VIII metal element, a phosphate of the group VIII metal element, a molybdate of the group VIII metal element, a tungstate of the group VIII metal element, and a water-soluble complex of the group VIII metal element. Specific examples of the precursor of the oxide of the group VIII metal element may include, but are not limited to: one or more of ferric nitrate, ferric sulfate, ferric acetate, nickel nitrate, nickel sulfate, nickel acetate, basic nickel carbonate, cobalt nitrate, cobalt sulfate, cobalt acetate, basic cobalt carbonate, cobalt chloride, nickel chloride and ferric ammonium citrate.

The support having the impregnating solution adsorbed thereon may be dried under conventional conditions to remove the solvent from the impregnating solution, thereby obtaining a support loaded with the oxide and/or precursor. Generally, the drying may be carried out at a temperature of 50 to 300 ℃, preferably 80 to 280 ℃, more preferably 120 to 280 ℃, and the drying may be carried out under normal pressure or under reduced pressure. The duration of the drying may be selected depending on the temperature of the drying and the pressure of the drying, and may be generally 1 to 12 hours, preferably 2 to 6 hours. The drying may be performed in an air atmosphere.

The support carrying the oxide and/or the precursor may be calcined under conventional conditions to provide a fischer-tropsch synthesis catalyst precursor. In the Fischer-Tropsch synthesis catalyst precursor, the VIII metal element is basically in the highest oxidation valence state. Generally, the calcination may be performed at a temperature of 300-900 deg.C, preferably 350-800 deg.C, more preferably 400-600 deg.C, and the duration of the calcination may be selected according to the calcination temperature, and may be generally 1-10 hours, preferably 2-6 hours. The firing is performed in an air atmosphere.

From the viewpoint of further improving the catalytic activity and activity stability of the finally prepared catalyst, it is preferable to further include supporting an auxiliary element on the carrier, the auxiliary element being one or more selected from alkali metal elements and rare earth metal elements (one or more of lanthanum, cerium, praseodymium and neodymium). Specific examples of the auxiliary element may include, but are not limited to: one or more than two of Li, Na, K and Ce. Preferably, the auxiliary element is one or more than two of Li, K and Ce. More preferably, the promoter element is K and/or Ce. Further preferably, the additive element is Ce. It should be noted that, although the kind of the auxiliary element and the modifying element used for modifying the alumina may be the same, when the auxiliary element and the modifying element are the same, even if the carrier is a carrier containing the modifying element, it is still necessary to additionally load the auxiliary element on the carrier containing the modifying element, and vice versa.

The supporting amount of the promoter element on the carrier is such that the content of the promoter element may be 0.1 to 10% by weight, preferably 1 to 8% by weight, more preferably 2 to 6% by weight, in terms of the element, based on the total amount of the catalyst precursor.

The auxiliary elements may be supported on the support by conventional methods, such as impregnation. The auxiliary element and the group VIII metal element may be simultaneously supported on the carrier, or the auxiliary element and the group VIII metal element may not be simultaneously supported on the carrier. Preferably, the promoter element and the group VIII metal element are simultaneously supported on the carrier, and in this case, the carrier may be impregnated with an impregnation solution containing an oxide of the group VIII metal element and/or a precursor of the oxide of the group VIII metal element and a compound containing the promoter element, and the carrier on which the impregnation solution is adsorbed may be successively dried and calcined to obtain the catalyst precursor.

The compound containing an auxiliary element may be a conventional compound capable of dissolving and dispersing in the impregnation solution, and may be one or two or more of nitrate, chloride, sulfate, acetate, oxalate, carbonate, bicarbonate, and hydroxide, for example. Specific examples of the compound containing an auxiliary element may include, but are not limited to: one or more of sodium nitrate, sodium chloride, sodium sulfate, sodium acetate, sodium oxalate, sodium carbonate, sodium bicarbonate, lithium carbonate, lithium nitrate, lithium chloride, potassium nitrate, potassium chloride, potassium sulfate, potassium acetate, potassium oxalate, potassium carbonate, potassium bicarbonate, manganese nitrate, and manganese chloride.

When the group VIII metal element and the optional auxiliary element are supported on the carrier by impregnation, the number of times of impregnation may be one or two or more. From the viewpoint of further improving the catalytic activity and activity stability of the finally prepared catalyst, it is preferable to perform impregnation twice or more. In the case where the impregnation is performed twice or more, it is preferable that the carrier having the impregnation liquid adsorbed thereon is dried and calcined in this order after each impregnation.

According to the method for producing low-carbon olefins of the present invention, in step S31, the fischer-tropsch synthesis reaction may be performed under conventional conditions. Preferably, the Fischer-Tropsch synthesis reaction feed and the Fischer-Tropsch synthesis catalyst may be contacted at a temperature of from 200 ℃ to 400 ℃, preferably from 300 ℃ to 380 ℃. The pressure at which the Fischer-Tropsch synthesis reaction feed is contacted with the Fischer-Tropsch synthesis catalyst may be in the range 0.5 to 3MPa, preferably 1 to 2.5MPa, expressed as gauge pressure.

According to the production method of the low-carbon olefin, the Fischer-Tropsch synthesis reaction can be contacted in a fixed bed reactor, can also be contacted in a fluidized bed reactor, and can also be contacted in the combination of the fixed bed reactor and the fluidized bed reactor. Preferably, the hydrogen and carbon monoxide are contacted with the Fischer-Tropsch synthesis catalyst in a fluidised bed reactor.

According to the method for producing the low-carbon olefin, when the Fischer-Tropsch synthesis catalyst of the preferred embodiment is adopted, a good catalytic reaction effect can be obtained even if hydrogen and carbon monoxide are contacted with the catalyst at a high space velocity. Specifically, the volume space velocity of the Fischer-Tropsch synthesis reaction feed can be 5000-¹Preferably 10000-^-1。

According to the method for producing low-carbon olefins, in step S41, low-carbon olefins, methane and carbon dioxide can be separated from the product stream of the fischer-tropsch synthesis by a conventional method. As an example, the Fischer-Tropsch synthesis product stream may be separated by cryocondensation to yield lower olefins, methane, and carbon dioxide, respectively.

According to the method for producing the low-carbon olefin, the methane separated from the Fischer-Tropsch synthesis product stream is sent to the step S11 and/or the step S21 to be used as the feed of the steam reforming reaction and/or the dry reforming reaction. Carbon dioxide separated from the product stream of the fischer-tropsch synthesis is fed to step S21 as a feed for the dry reforming reaction. According to the production method of the low-carbon olefin, the steam reforming and the dry reforming are combined for use, and the methane and the carbon dioxide separated from the Fischer-Tropsch synthesis product material flow are recycled, so that the utilization rate of the raw materials is effectively improved, and the emission of greenhouse gas carbon dioxide is obviously reduced.

From the viewpoint of further improving the utilization rate of the raw materials, the method for producing low-carbon olefins according to the present invention preferably further comprises separating unreacted hydrogen and/or carbon monoxide from the product stream of the fischer-tropsch synthesis, and sending 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.

According to a second aspect of the present invention, there is provided a low carbon olefin production system comprising 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 give a dry reformed gas having hydrogen and carbon monoxide as the main components.

The synthesis gas mixing unit is respectively communicated with a steam reforming synthesis gas output port of the steam reforming unit and a dry reforming reaction unit dry weight integrated gas output port, and is used for mixing the steam reforming synthesis gas and the dry weight integrated gas to prepare Fischer-Tropsch synthesis reaction feed, and feeding the Fischer-Tropsch synthesis reaction feed 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. The synthesis gas mixing unit can also adopt a pipeline mixer to directly mix the steam regenerated synthesis gas and the dry weight integrated synthesis 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, is communicated with a Fischer-Tropsch synthesis reaction feed outlet of the synthesis gas mixing unit, and is used for contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst to obtain a Fischer-Tropsch synthesis product material flow containing low-carbon olefins. The fischer-tropsch synthesis reactor may be in various common reactor forms, and specifically, the fischer-tropsch synthesis reactor may be a fixed bed reactor, a fluidized bed reactor, or a combination of a fixed bed reactor and a fluidized bed reactor, and is preferably a fluidized bed reactor.

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 Fischer-Tropsch synthesis catalyst precursor so as to convert the Fischer-Tropsch synthesis catalyst precursor into the Fischer-Tropsch synthesis catalyst with catalytic activity. The reduction activation subunit can reduce and activate the Fischer-Tropsch synthesis catalyst precursor into the Fischer-Tropsch synthesis catalyst by contacting the Fischer-Tropsch synthesis 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 be sufficient to store and deliver a 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. 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 composition of the first gas and the second gas has been described in detail above and will not be described in detail here.

The reducing gas control means is used 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 reducing gas control device is configured to firstly input a first gas into the reduction activation reactor to contact the fischer-tropsch synthesis catalyst precursor with hydrogen to perform a pre-reduction reaction, so as to obtain a pre-reduction catalyst, and then input a second gas into the reduction activation reactor to contact the pre-reduction catalyst with the second gas to perform the reduction activation 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 activation reactor and the amount of gas fed thereto.

The reduction activation reactor is used for accommodating a Fischer-Tropsch synthesis catalyst precursor and is communicated with the first gas storage and conveying device and the second gas storage and conveying device, so that the Fischer-Tropsch synthesis catalyst precursor is sequentially contacted with the first gas and the second gas to carry out reduction activation, and the Fischer-Tropsch synthesis catalyst with 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 Fischer-Tropsch synthesis catalyst precursor is carried out in the Fischer-Tropsch synthesis reactor.

The reduction activation reactor and the Fischer-Tropsch synthesis reactor also can not be the same reactor, namely the Fischer-Tropsch synthesis reactor and the reduction activation reactor are respectively 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.

According to the low-carbon olefin production system, the circulating unit is used for circularly conveying 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 conveying the carbon dioxide separated by the Fischer-Tropsch synthesis reaction product separating unit into the dry reforming reaction unit, and optionally circularly conveying 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 with the Fischer-Tropsch synthesis reaction unit and a control valve arranged on the conveying pipeline, so that the hydrogen and/or 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 pipelines, can set up hydrogen conveying pipeline and corresponding control valve this moment respectively and carbon monoxide conveying pipeline and corresponding control valve.

The low carbon olefin production system according to the present invention preferably 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 respectively communicated with 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, 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 shows a preferred embodiment of the low carbon olefin production system according to the present invention, which is described in detail below with reference to fig. 2. As shown in fig. 2, the low carbon 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.

And the raw material gas A enters a raw material gas separation unit I for separation to obtain methane B. And respectively feeding the methane B into the steam reforming reaction unit II and the dry reforming reaction unit III, and simultaneously feeding the steam C into the steam reforming reaction unit II so as to carry out reforming reaction on the methane and the steam to obtain the steam reforming synthesis gas E. And feeding carbon dioxide D into the dry reforming reaction unit III so as to carry out reforming reaction on methane and carbon dioxide to obtain dry reforming synthesis gas F. And (3) mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F (preferably adopting a pipeline mixer) to prepare the Fischer-Tropsch synthesis reaction feed G according with the hydrogen-carbon ratio of the Fischer-Tropsch synthesis reaction. And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reaction unit IV, and contacting with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction. And the Fischer-Tropsch synthesis reactor in the Fischer-Tropsch synthesis reaction unit IV operates at the temperature of producing the low-carbon olefin. And the Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV enters a Fischer-Tropsch synthesis product separation unit V for separation to obtain the low-carbon olefin K, unreacted hydrogen, carbon monoxide, methane M and carbon dioxide N. Wherein, the low carbon olefin K is sent out of the system.

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.

In the following examples, preparations and comparative examples, the pressures were gauge pressures unless otherwise specified.

In the following examples, preparations and comparative examples, CO conversion (X)_CO)、CH₄Selectivity is

And C₂-C₄The selectivity of the hydrocarbon (among them,

is represented by C₂-C₄The selectivity of the olefin is high,

is represented by C₂-C₄Alkane selectivity) was calculated by the following formula:

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;

C_1,CO、C_2,COrespectively representing the molar contents of CO in raw gas entering a reaction system and tail gas flowing out of the reaction system;

n_conis the mole number of CO participating in the reaction;

to generate CH₄The number of moles of (a);

to generate C₂-C₄Moles of hydrocarbons.

Preparation examples 1 to 24 were used for preparing Fischer-Tropsch synthesis catalysts and their properties were evaluated.

In the following preparation examples, the specific surface area, pore volume and average pore diameter were measured by a 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 preparation examples, the contents 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 (compiled by "analytical methods in petrochemical industry (RIPP test method)", Yanggui et al, scientific Press, 1 st edition (9 months, 1990), p. 371-. When the catalyst was tested, a sample of the catalyst was stored under an argon atmosphere.

In the following preparation example, 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 preparations, X-ray photoelectron spectroscopy was carried out on an ESCALB 250 type X-ray photoelectron spectrometer equipped with Thermo Avantage V5.926 software, manufactured by Thermo Scientific, with an excitation source of monochromated Al K.alpha.X rays, an energy of 1486.6eV, a power of 150W, a transmission energy for narrow scanning of 30eV, and a base vacuum of 6.5X 10 for analytical tests^-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 support

Taking gamma-Al₂O₃(Sasol product, its specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1) 200g, calcining in air atmosphere at 980 ℃ for 2 hours, subjecting the calcined product to X-ray diffraction analysis (shown in FIG. 1), and determining that theta-Al is obtained₂O₃The specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1.

Dissolving zirconium nitrate pentahydrate in 43g of deionized water to prepare a modified zirconium solution, and adding 100.0g of prepared theta-Al into the modified zirconium solution₂O₃The resulting mixture was saturated and immersed at 25 ℃ for 2 hours. Then, the impregnated mixture was placed in an oven and dried at 120 ℃ under atmospheric pressure (1 atm, the same applies hereinafter) for 5 hours in an air atmosphere. And roasting the dried substance in an air atmosphere at 400 ℃ for 3 hours to obtain the carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis to determine that the content of Zr was 3% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ammonium ferric citrate, potassium carbonate and cerous nitrate hexahydrate into 12mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain an impregnation liquid. A50 vol% of the impregnation solution was taken, 15g of the carrier was added to the impregnation solution, and the mixture was saturated and impregnated at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 120 ℃ under atmospheric pressure for 5 hours in an air atmosphere. And roasting the dried substance in an air atmosphere at 400 ℃ for 3 hours to obtain the catalyst after primary leaching.

The catalyst after the first impregnation was added to the remaining impregnation solution, and saturated impregnation was carried out at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 120 ℃ under atmospheric pressure for 5 hours in an air atmosphere. The dried substance was calcined at 400 ℃ for 3 hours in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is charged into a fluidized bed reactor, and H is introduced into the reactor₂Adjustment of the reverseThe reactor pressure is 0.1MPa, and the volume space velocity of hydrogen is 20000 hours^-1The temperature of the reactor was raised from 25 ℃ to 400 ℃ and maintained at this temperature for 8 hours. The reactor was then cooled to 200 ℃, hydrogen switched to ethane, and the volumetric space velocity of ethane was 15000 hours^-1After 12 hours of maintenance, a Fischer-Tropsch synthesis catalyst was obtained, the composition of which is shown in tables 2 and 4, CO₂The results of the-TPD and CO-TPD tests are listed in Table 3.

(4) Preparation of low-carbon olefin

After the reduction activation is finished, introducing synthesis gas into the reactor, raising the temperature of the reactor to 340 ℃, and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 30000 hours^-1At a pressure of 1.5MPa (in terms of gauge pressure), the synthesis gas is a mixture of hydrogen and carbon monoxide, the composition of which is H₂: 50 of CO: 50 (molar ratio). During the reaction, the composition of the reaction mixture gas outputted from the reactor was analyzed by an on-line gas chromatograph, and the results measured at 50 hours and 200 hours of the reaction were shown in tables 5 and 6, respectively.

Preparation example 2

A catalyst was prepared and a low-carbon olefin was prepared in the same manner as in preparation example 1, except that in step (1), γ -Al was used₂O₃The carrier was prepared without calcination and was directly impregnated with a modified zirconium solution, wherein the Zr content was 3 wt% in terms of the element based on the total amount of the carrier.

Preparation example 3

A catalyst was prepared and a low-carbon olefin was prepared in the same manner as in preparation example 1, except that in step (1), theta-Al₂O₃Is not contacted with the modified zirconium solution but is directly used in step (2) to prepare the catalyst precursor.

Preparation example 4

A catalyst was prepared and a low-carbon olefin was prepared in the same manner as in preparation example 1, except that cerium nitrate hexahydrate was not used in preparing the impregnation liquid in step (2).

Preparation example 5

A catalyst was prepared and a lower olefin was prepared by the same method as in preparation example 1, except that in step (2), impregnation was carried out once and the conditions of impregnation, drying and calcination were the same as in preparation example 1, that is, the carrier was impregnated with 6mL of the impregnation solution, and the mixture obtained by the impregnation was sequentially dried and calcined, thereby obtaining a catalyst precursor.

Preparation example 6

A catalyst was prepared and a lower olefin was produced in the same manner as in preparation example 1, except that in step (3), ethane was replaced with an equal volume of ethylene.

Preparation example 7

A catalyst was prepared and a low-carbon olefin was prepared in the same manner as in preparation example 1, except that in step (1), gamma-Al was brought into contact with the modified zirconium solution₂O₃And theta-Al₂O₃According to the weight ratio of 1: 1 mixing the resulting mixture.

Preparation example 8

A catalyst was prepared and a low carbon olefin was prepared in the same manner as in preparation example 1, except that in step (1), the amount of zirconium nitrate pentahydrate was changed to prepare a carrier in which the Zr content was 1.5% by weight in terms of element based on the total amount of the carrier.

Preparation example 9

A catalyst was prepared and a low carbon olefin was prepared in the same manner as in preparation example 1, except that in step (1), the amount of zirconium nitrate pentahydrate was changed to prepare a carrier in which the Zr content was 8% by weight in terms of element based on the total amount of the carrier.

Preparation example 10

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 1, except that in step (3), after the introduction of hydrogen, the introduction of ethane was not continued, but step (4) was directly performed, i.e., only hydrogen was used for the reduction activation, and no ethane was used.

Preparation example 11

A catalyst was prepared and a lower olefin was produced in the same manner as in preparation example 1, except that in step (3), ethane was replaced with an equal volume of CO.

Preparation example 12

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 1, except that in step (3), the operation of passing hydrogen was not performed, but ethane was directly passed through the reactor, i.e., only ethane was used for reductive activation, and hydrogen was not used.

Preparation example 13

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 2, except that in step (3), after the introduction of hydrogen, ethane was not continuously introduced, but step (4) was directly performed, i.e., only hydrogen was used for the reduction activation, and ethane was not used.

Preparation example 14

A catalyst was prepared and a lower olefin was prepared in the same manner as in preparation example 2, except that in step (3), ethane was replaced with an equal volume of CO.

Preparation example 15

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 3, except that in step (3), after the introduction of hydrogen, the introduction of ethane was not continued, but step (4) was directly performed, i.e., only hydrogen was used for the reduction activation, and no ethane was used.

Preparation example 16

A catalyst was prepared and a lower olefin was prepared in the same manner as in preparation example 3, except that in step (3), ethane was replaced with an equal volume of CO.

Preparation example 17

(1) Preparation of the support

Taking gamma-Al₂O₃(Sasol product, its specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1) 200g, calcining at 1050 deg.C for 1 hr in air atmosphere, subjecting the calcined product to X-ray diffraction analysis to determine that the product is theta-Al₂O₃The specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1.

Dissolving magnesium nitrate in 41g deionized water to obtain modified magnesium solution, and adding 100.0g of prepared theta-Al into the modified magnesium solution₂O₃The resulting mixture was saturated and immersed at 25 ℃ for 2 hours. Then, the mixture obtained by the impregnation was placed in an oven at 200 ℃ under atmospheric pressure in an air atmosphereAnd drying for 3 hours. And roasting the dried substance at 800 ℃ in an air atmosphere for 1 hour to obtain the carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis to determine that the content of Mg was 6% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric nitrate, lithium carbonate and cerous nitrate hexahydrate into 12mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain an impregnation liquid. A50 vol% of the impregnation solution was taken, 15g of the carrier was added to the impregnation solution, and the mixture was saturated and impregnated at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 200 ℃ under atmospheric pressure for 3 hours in an air atmosphere. And roasting the dried substance in an air atmosphere at 600 ℃ for 2 hours to obtain the catalyst after primary leaching.

The catalyst after the first impregnation was added to the remaining impregnation solution, and saturated impregnation was carried out at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 200 ℃ under atmospheric pressure for 3 hours in an air atmosphere. The dried substance was calcined at 600 ℃ for 2 hours in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is charged into a fluidized bed reactor, and H is introduced into the reactor₂And argon (wherein the molar ratio of argon to hydrogen is 10: 1), adjusting the pressure of the reactor to be 0.1MPa, and the volume space velocity of hydrogen to be 15000 hours^-1The temperature of the reactor was raised from 25 ℃ to 350 ℃ and maintained at this temperature for 8 hours. The reactor was then cooled to 250 ℃, hydrogen switched to ethane, and the volumetric space velocity of ethane was 10000 hours^-1After 4 hours of maintenance, the Fischer-Tropsch catalyst, the composition of which is shown in tables 2 and 4, CO₂The results of the-TPD and CO-TPD tests are listed in Table 3.

(4) Preparation of low-carbon olefin

After the reduction activation is finished, introducing synthesis gas into the reactor, raising the temperature of the reactor to 340 ℃, and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 30000 hours^-1At a pressure of 5MPa (gauge pressure), the synthesis gas is a mixture of hydrogen and carbon monoxide and has a composition of H₂: 50 of CO: 50 (molar ratio). During the reaction, the composition of the reaction mixture gas outputted from the reactor was analyzed by an on-line gas chromatograph, and the results measured at 50 hours and 200 hours of the reaction were shown in tables 5 and 6, respectively.

Preparation example 18

A catalyst was prepared and a low-carbon olefin was prepared in the same manner as in preparation example 17, except that in step (1), γ -Al₂O₃The carrier was prepared without calcination by direct impregnation with a modified magnesium solution, in which the Mg content was 6 wt% in terms of the element, based on the total amount of the carrier.

Preparation example 19

(1) Preparation of the support

Taking gamma-Al₂O₃(Sasol product, its specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1) 200g, roasting at 780 deg.C for 4 hr in air atmosphere, subjecting the roasted product to X-ray diffraction analysis to determine that the product is theta-Al₂O₃The specific surface area, pore volume, average pore diameter and particle size distribution are shown in Table 1.

Potassium nitrate was dissolved in 53g of deionized water to prepare a modified potassium solution, and 100.0g of the prepared theta-Al was added to the modified potassium solution₂O₃The resulting mixture was saturated and immersed at 25 ℃ for 2 hours. Then, the impregnated mixture was placed in an oven and dried at 300 ℃ under atmospheric pressure (1 atm, the same applies hereinafter) for 2 hours in an air atmosphere. The dried material was calcined at 500 ℃ for 6 hours in an air atmosphere to obtain a carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis to determine that the content of potassium was 2.5% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric nitrate, potassium carbonate and cerous nitrate hexahydrate into 15mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain an impregnation liquid. A50 vol% of the impregnation solution was taken, 15g of the carrier was added to the impregnation solution, and the mixture was saturated and impregnated at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 280 ℃ for 2 hours under atmospheric pressure in an air atmosphere. And roasting the dried substance for 6 hours at 500 ℃ in an air atmosphere to obtain the catalyst after primary leaching.

The catalyst after the first impregnation was added to the remaining impregnation solution, and saturated impregnation was carried out at ambient temperature (25 ℃ C.) for 1 hour. Then, the impregnated mixture was placed in an oven and dried at 280 ℃ for 2 hours under atmospheric pressure in an air atmosphere. The dried substance was calcined at 500 ℃ for 6 hours in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is charged into a fluidized bed reactor, and H is introduced into the reactor₂The pressure of the reactor is adjusted to be 0.15MPa, and the volume space velocity of hydrogen is 10000 hours^-1The temperature of the reactor was raised from 25 ℃ to 500 ℃ and kept constant at this temperature for 6 hours.

Then, the reactor was cooled to 350 ℃, hydrogen was switched to a mixed gas of ethane and argon (in which the molar ratio of ethane to argon was 1: 20), and the volume space velocity of ethane was 20000 hours^-1After 4 hours of maintenance, a Fischer-Tropsch synthesis catalyst was obtained, the composition of which is shown in tables 2 and 4, CO₂The results of the-TPD and CO-TPD tests are listed in Table 3.

(4) Preparation of low-carbon olefin

After the reduction activation is finished, introducing synthesis gas into the reactor, raising the temperature of the reactor to 340 ℃, and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 30000 hours^-1The pressure is 1MPa (measured by gauge pressure), and the synthesis gas is a mixed gas of hydrogen and carbon monoxide, and the composition of the mixed gas is H₂: CO 60: 40 (molar ratio). During the reaction, the composition of the reaction mixture gas outputted from the reactor was analyzed by an on-line gas chromatograph, and the results measured at 50 hours and 200 hours of the reaction were shown in tables 5 and 6, respectively.

Preparation example 20

A catalyst was prepared and a lower olefin was produced by the same method as in production example 19, except that in step (2), the iron nitrate was replaced with cobalt nitrate.

Preparation example 21

A catalyst was prepared and a lower olefin was produced by the same method as in production example 19, except that in step (2), the iron nitrate was replaced with nickel nitrate.

Preparation example 22

A catalyst was prepared and a lower olefin was produced in the same manner as in production example 19, except that in step (3), ethane was replaced with a mixed gas of CO and argon in an equal volume (wherein the molar ratio of CO to argon was 1: 20).

Preparation example 23

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 19, except that in step (2), potassium carbonate was not used in the preparation of the impregnation solution, and the amount of cerium nitrate hexahydrate was increased accordingly.

Preparation example 24

The catalyst was prepared and the lower olefins were prepared in the same manner as in preparation example 19, except that in step (2), cerium nitrate as a running water was not used in the preparation of the impregnation liquid, and the amount of potassium carbonate was increased accordingly.

TABLE 1

TABLE 2 (based on the total amount of catalyst)

TABLE 3

TABLE 4

¹: no Fe detected₅C₂ ²: FeO and Fe were not detected₅C₂

TABLE 5

*: O/P is C₂-C₄Olefin selectivity (S) of_C2 ^＝ _-C4 ^＝) And C₂-C₄Alkane selectivity (S)_C2°_-C4Deg.) ratio.

TABLE 6

*: based on the corresponding data for 50 hours

Comparing preparation examples 1 with preparation examples 10 to 12, preparation example 2 with preparation examples 13 and 14, preparation example 3 with preparation examples 15 and 16, and preparation example 19 with preparation example 22, it can be seen that, after the catalyst precursor is pre-reduced with hydrogen, reduction activation is performed with hydrocarbon which is gaseous at the reduction activation temperature, the catalytic activity of the finally formed reduction activation catalyst can be obviously improved, and particularly the selectivity for low-carbon olefins can be obviously improved. As can be seen from comparison of production examples 1 and 2 and production examples 17 and 18, theta-Al was used₂O₃Can obviously improve the catalytic activity of the Fischer-Tropsch synthesis catalyst.

Examples 1 to 6 are intended to illustrate the production process and production system of lower olefins according to the present invention.

Example 1

In this embodiment, the low-carbon olefin production system shown in fig. 2 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 specific process flow is as follows.

(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 alpha-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, 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 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 alpha-Al₂O₃) The temperature in the catalyst bed is 750 ℃, the pressure in the reactor is 2MPa, and the temperature is in the presence of methane and water vaporThe total meter and gas hour volume airspeed are 80000h^-1。

(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture with a hydrogen-carbon ratio of 2.1: 1 fischer-tropsch synthesis reaction feed G. And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fluidized bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, and contacting with a Fischer-Tropsch synthesis catalyst (the Fischer-Tropsch synthesis catalyst prepared in the preparation example 1) to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 340 ℃, the pressure in the reactor is 1MPa, and the gas hourly space velocity is 30000h based on the total amount of the synthesis gas^-1。

(5) And sending the Fischer-Tropsch synthesis product stream H output by the Fischer-Tropsch synthesis reaction unit IV into a Fischer-Tropsch synthesis product separation unit V for separation. The separation process comprises the following steps: firstly, carrying out gas-liquid separation to obtain low-carbon olefin K and a gas product; then, the gas product is subjected to cryogenic separation to remove carbon dioxide in the gas product; then, the gaseous product from which the carbon dioxide is separated is subjected to cryogenic separation to obtain methane, and unreacted hydrogen and carbon monoxide.

Discharging the low-carbon olefin K out of the system; the separated carbon dioxide N is circularly sent into a dry reforming reaction unit III; the separated methane M is respectively sent into a steam reforming reaction unit II and a dry reforming reaction unit III; and (3) circularly feeding a part of L of the separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest of L 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 reaction unit during the reaction was analyzed by an on-line gas chromatograph and the results obtained after 50 hours of reaction are shown in Table 7. The overall water consumption, carbon dioxide emissions, and energy efficiency of the system are listed in table 8.

Comparative example 1

The Fischer-Tropsch synthesis catalyst used in this comparative example was the same as the Fischer-Tropsch synthesis catalyst used in example 1.

The system shown in the figure 1 is adopted in the comparative example, 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 a low-carbon olefin separation unit VI which are sequentially connected. The specific process flow is as follows.

The coal water slurry C is prepared from pulverized coal A (pulverized coal (with the particle size of 10mm) obtained by crushing and screening solid raw material coal (brown coal produced by inner Mongolia)) in a coal water slurry preparation unit I at the flow rate of 360t/h and water B at the flow rate of 360t/h, and is conveyed into a coal gasification unit II to react 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: and 1, removing acid gas and sulfide M through a synthesis gas purification unit IV to obtain purified synthesis gas, wherein 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 fluidized bed reactor, and a Fischer-Tropsch reaction product N containing olefin is generated (the Fischer-Tropsch synthesis reaction conditions are the same as in example 1). And (3) separating the low-carbon olefin K from the Fischer-Tropsch reaction product N through a low-carbon olefin separation unit VI, discharging carbon dioxide H and methane G generated by a Fischer-Tropsch synthesis unit V, circulating a part of unreacted synthesis gas (the content is 98 percent based on the total amount of the separated synthesis gas) to the Fischer-Tropsch synthesis unit V, and discharging the other part of the unreacted synthesis gas as purge gas Z out of the system.

Comparative example 2

A low carbon olefin was produced in the same manner as in example 1, except that the dry reforming reaction unit III was not provided, and methane (including fresh methane and recycled methane) was entirely introduced into the steam reforming reaction unit II to carry out the reforming reaction.

Comparative example 3

A low carbon olefin was produced in the same manner as in example 1, except that the steam reforming reaction unit II was not provided, and methane (including fresh methane and recycled methane) was entirely fed into the dry reforming reaction unit III to undergo a reforming reaction.

Example 2

A low carbon olefin was produced by the same method as in example 1, except that the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst produced in preparation example 2.

Example 3

A low carbon olefin was produced by the same method as in example 1, except that the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst produced in production example 10.

Example 4

A low carbon olefin was produced by the same method as in example 1, except that the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst produced in production example 11.

Example 5

In this example, the reaction system shown in FIG. 2 was used, and the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst prepared in preparation example 17. The specific process flow is 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.

(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 alpha-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) Mixing the second stream of methane with carbon dioxide D at a flow rate of 200kmol/h, a temperature of 370 ℃ and a pressure of 2MPa, raising the temperature of the mixture to 600 ℃,then the gas enters 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(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)₂O₃Is alpha-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 2.2: 1 fischer-tropsch synthesis reaction feed G. And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fluidized bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, and contacting with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 340 ℃, 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 30000h^-1。

Discharging the low-carbon olefin K out of the system; the separated carbon dioxide N is circularly sent into a dry reforming reaction unit III; the separated methane M is respectively sent into a steam reforming reaction unit II and a dry reforming reaction unit III; and (3) circularly feeding a part of L of the separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest of L 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 7. The overall water consumption, carbon dioxide emissions, and energy efficiency of the plant are listed in table 8.

Example 6

In this example, the reaction system shown in FIG. 2 was used, and the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst prepared in preparation example 19. 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.

(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, 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 alpha-Al₂O₃) The temperature in the catalyst bed is 900 ℃, the pressure in the reactor is 1MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 100000h^-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, 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 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 alpha-Al₂O₃) The temperature in the catalyst bed was 750 ℃ and the pressure in the reactor was 2MPa, based on the total amount of the methane and the steam, the gas hourly volume space velocity is 80000h^-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: 1 fischer-tropsch synthesis reaction feed G. And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fluidized bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, and contacting with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 360 ℃, 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。

Discharging the low-carbon olefin K out of the system; the separated carbon dioxide N is circularly sent into a dry reforming reaction unit III; the separated methane M is respectively sent into a steam reforming reaction unit II and a dry reforming reaction unit III; and (3) circularly feeding a part of L of the separated hydrogen and carbon monoxide into a Fischer-Tropsch synthesis reaction unit IV, and discharging the rest of L 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 7. The overall water consumption, carbon dioxide emissions, and energy efficiency of the plant are listed in table 8.

TABLE 7

TABLE 8

	Water consumption (ton/ton)_{Low carbon olefin})	Carbon dioxide emission (ton/ton)_{Low carbon olefin})	Energy efficiency (%)
				Example 1	15	0.5	56
Comparative example 1	20	6.2	36
				Comparative example 2	21	4.2	41
Comparative example 3	19	0.6	46
				Example 5	19	1.2	48
Example 6	17	1.8	47

Note: the energy efficiency is the sum of the calorific value of the low-carbon olefin finally discharged out of the device/the calorific value of the raw materials such as the coal-electricity water vapor catalyst solvent entering the device, namely the calorific value of the obtained low-carbon olefin/the comprehensive energy consumption required for producing the low-carbon olefin. Wherein, the comprehensive energy consumption comprises raw material heat value and public engineering energy consumption, and mainly comprises: the heat value of fuel coal and raw material coal, the electric energy consumed by a motor pump for the device process, the indirect energy consumption of circulating cooling water, boiler make-up water, process air, instrument air, fresh water and the like.

The results in table 8 show that the present invention combines the methane steam reforming process and the methane dry reforming process to simultaneously utilize the two greenhouse gases of carbon dioxide and methane, so that the two 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 significantly improved.

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. A method for producing low-carbon olefins comprises the following steps:

s31, at least partially dryingSteam reforming synthesis gas and at least part of dry weight are integrated into synthesis gas to be mixed to prepare Fischer-Tropsch synthesis reaction feed, the Fischer-Tropsch synthesis reaction feed is contacted with a Fischer-Tropsch synthesis catalyst at the reaction temperature of producing low-carbon olefin to obtain a Fischer-Tropsch synthesis product material flow, the Fischer-Tropsch synthesis catalyst contains a carrier and a VIII group metal element loaded on the carrier, the carrier is alumina, the alumina contains theta-alumina, the theta-alumina content is more than 50 wt% based on the total amount of the alumina in the catalyst, the valence state of at least part of the VIII group metal element is lower than the highest oxidation valence state of the metal element, the VIII group metal element is Fe, and an X-ray photoelectron spectrum of the Fischer-Tropsch synthesis catalyst has a spectrum peak corresponding to FeO and a spectrum peak corresponding to Fe₅C₂The content of Fe determined from the peak corresponding to FeO and the content of Fe determined from the peak corresponding to Fe, calculated as elements₅C₂The ratio of the Fe content determined by the spectrum peak of (1) is 8-25: based on the total amount of Fe determined by X-ray photoelectron spectroscopy, the total Fe content is determined by the peak corresponding to FeO and the peak corresponding to Fe₅C₂The content of Fe determined by the spectrum peak is 30-99%,

the Fischer-Tropsch synthesis catalyst is obtained by carrying out reduction activation on a catalyst precursor, and the reduction activation method comprises the following steps:

(1) pre-reducing a catalyst precursor in a first gas to obtain a pre-reduced catalyst, wherein the first gas is hydrogen or a mixed gas of hydrogen and an inert gas, the catalyst precursor comprises a carrier and a VIII group metal element loaded on the carrier in the form of an oxide, and the valence state of the VIII group metal element in the oxide is the highest oxidation valence state of the metal element;

(2) reducing and activating the pre-reduction catalyst in a second gas to obtain a reduction activation catalyst, wherein the second gas is gaseous hydrocarbon at the reduction activation temperature or a mixed gas of the gaseous hydrocarbon and an inert gas at the reduction activation temperature, the reduction activation is carried out at the temperature of 150-,

the catalyst precursor is prepared by a method comprising the following steps: loading a compound containing a VIII group metal element and an optional compound containing an auxiliary agent element on a carrier, roasting the carrier to obtain a catalyst precursor, wherein the compound containing the VIII group metal element is an oxide of the VIII group metal element and/or a precursor of the oxide of the VIII group metal element, the auxiliary agent element is one or more selected from alkali metal elements and rare earth metal elements, and when the VIII group metal element and the optional auxiliary agent element are loaded on the carrier in a dipping mode, the dipping times are more than two times;

2. The method of claim 1, wherein in S11, the molar ratio of methane to water vapor is 1: 0.5-4.

3. The process as claimed in claim 1 or 2, wherein in S11, the methane is contacted with the steam at a temperature of 700 ℃ and 950 ℃ and at a pressure of 0.1 to 5MPa, the pressure being expressed as gauge pressure.

4. The process of claim 1 or 2, wherein in S11, the steam reforming reaction is carried out in a fixed bed reactor.

5. The method as claimed in claim 4, wherein the gas hourly space velocity of the feed in S11 is 10000-100000 hours in terms of the total amount of methane and steam^-1。

6. The method of claim 1, wherein the molar ratio of methane to carbon dioxide in S21 is 1: 0.5-5.

7. The process as claimed in claim 1 or 6, wherein in S21, the methane is contacted with the carbon dioxide at a temperature of 600-800 ℃ and a pressure of 0.1-5MPa, the pressure being expressed as gauge pressure.

8. The process of claim 1 or 6, wherein in S21, the dry reforming reaction is carried out in a fixed bed reactor.

9. The method as claimed in claim 8, wherein the gas hourly space velocity of the feed in S21 is 10000-100000 hours in terms of the total amount of methane and carbon dioxide^-1。

10. The method as claimed in claim 1, wherein the reaction temperature for producing the lower olefins in S31 is 200-400 ℃, and the pressure is 0.5-3MPa, wherein the pressure is measured by gauge pressure.

11. The process of claim 1 or 10, wherein in S31, 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.

12. the process of claim 1 or 10, wherein in S31, the molar ratio of hydrogen to carbon monoxide in the fischer-tropsch synthesis reaction feed is in the range of from 1.5 to 2.2: 1.

13. the process as claimed in claim 1 or 10, wherein, in S31, the contacting is carried out in a fluidized bed reactor, and the volume space velocity of the Fischer-Tropsch synthesis reaction feed is 5000-¹。

14. The method according to claim 1, wherein the content of Fe determined by the peak corresponding to FeO is related to the content of Fe determined by the peak corresponding to Fe, calculated on an elemental basis₅C₂The ratio of the determined Fe content of the peaks of (a) is 10-12: 1.

15. the method according to claim 1, wherein the total amount of Fe determined by X-ray photoelectron spectroscopy is measured as elemental by the peak and the couple corresponding to FeOIn response to Fe₅C₂The content of Fe determined by the peak of (1) is not less than 50%.

16. The method of claim 15, wherein the total amount of Fe, on an elemental basis as determined by X-ray photoelectron spectroscopy, is determined by the peaks corresponding to FeO and by the peaks corresponding to Fe₅C₂The content of Fe determined by the peak of (1) is not less than 75%.

17. The process of any one of claims 1 and 14 to 16, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is present in the temperature range of 300 ℃ and 600 DEG C₂High temperature desorption peak.

18. The process of claim 17, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is present in the temperature interval of 320 ℃ and 500 DEG C₂High temperature desorption peak.

19. The process of claim 18, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is present in the temperature range of 350 ℃ and 480 DEG C₂High temperature desorption peak.

20. The process of any one of claims 1 and 14 to 16, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is also present in the temperature range of 100-₂Low temperature desorption peak.

21. The process of claim 20, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is also present in the temperature range of 150 ℃ and 190 DEG C₂Low temperature desorption peak.

22. The process as claimed in any one of claims 1 and 14 to 16, wherein the Fischer-Tropsch synthesis catalyst has a CO-TPD desorption profile in which a CO high temperature desorption peak is present in a temperature range of 300 ℃ to 700 ℃.

23. The process as claimed in any one of claims 1 and 14 to 16, wherein the fischer-tropsch synthesis catalyst has a CO low temperature desorption peak in the CO-TPD desorption profile within the temperature range of 100 ℃ and 200 ℃.

24. The process according to any one of claims 1 and 14 to 16, wherein the content of the group VIII metal element having a valence lower than its maximum oxidation valence is 30% by weight or more in terms of the element based on the total amount of the group VIII metal element in the fischer-tropsch synthesis catalyst.

25. The process according to claim 24, wherein the content of the group VIII metal element having a valence lower than the maximum oxidation valence thereof is 50% by weight or more in terms of the element based on the total amount of the group VIII metal element in the Fischer-Tropsch synthesis catalyst.

26. The process of claim 25, wherein the amount of the group VIII metal element having a valence lower than its maximum oxidation valence is 55% by weight or more in terms of the element based on the total amount of the group VIII metal element in the fischer-tropsch catalyst.

27. The process according to claim 26, wherein the content of the group VIII metal element having a valence lower than the maximum oxidation valence thereof is 60% by weight or more in terms of the element based on the total amount of the group VIII metal element in the Fischer-Tropsch synthesis catalyst.

28. The process according to any one of claims 1 and 14 to 16, wherein the fischer-tropsch synthesis catalyst contains a second metal element and/or a third metal element supported on the carrier, the second metal element being one or two or more selected from alkali metal elements, alkaline earth metal elements, and group IVB metal elements, and the third metal element being one or two or more selected from rare earth metal elements.

29. The method according to claim 28, wherein the second metal element is one or two or more of Li, K, Mg, and Zr.

30. The method of claim 28, wherein the third metal element is Ce.

31. The method of claim 28, wherein the second metal element is present in an amount of 0 to 15 wt%, the third metal element is present in an amount of 0 to 10 wt%, and the second metal element and the third metal element are not present in an amount of 0 at the same time, calculated as elements based on the total amount of the fischer-tropsch synthesis catalyst.

32. The method of claim 31, wherein the second metal element is present in an amount of 2 to 11 wt%, the third metal element is present in an amount of 0.5 to 6 wt%, and the second metal element and the third metal element are not present in an amount of 0 on an elemental basis, based on the total amount of the fischer-tropsch synthesis catalyst.

33. The method of claim 32, wherein the second metal element is present in an amount of 5 to 7 wt%, the third metal element is present in an amount of 0.8 to 3 wt%, and the second metal element and the third metal element are not present in an amount of 0 on an elemental basis, based on the total amount of the fischer-tropsch synthesis catalyst.

34. The process of claim 28, wherein the fischer-tropsch synthesis catalyst comprises a second metallic element and a third metallic element supported on the support.

35. The method of claim 34, wherein the second metal element is a group IVB metal element and/or an alkali metal element, and the third metal element is Ce.

36. The method of claim 35, wherein the second metallic element is a group IVB metallic element and an alkali metallic element.

37. The method of claim 36, wherein the second metallic element is Zr and K.

38. The process of any one of claims 1 and 14 to 16, wherein the group VIII metal element is present in an amount of from 3 to 30 wt.% as the element, based on the total amount of the fischer-tropsch synthesis catalyst.

39. The process of claim 38 wherein the group VIII metal element is present in an amount of from 8 to 20 wt.% on an elemental basis based on the total amount of fischer-tropsch synthesis catalyst.

40. The process of claim 39, wherein the group VIII metal element is present in an amount of 10 to 15 wt.% on an elemental basis, based on the total amount of Fischer-Tropsch synthesis catalyst.

41. The method as claimed in claim 1, wherein the pre-reduction is carried out at a temperature of 200-600 ℃.

42. The method as claimed in claim 41, wherein the pre-reduction is carried out at a temperature of 300-550 ℃.

43. The process according to claim 1, wherein the pressure in the reactor in which the pre-reduction is carried out is 0 to 3MPa in gauge.

44. A process as claimed in claim 43, in which the pressure in the reactor at which the pre-reduction is carried out is in the range 0.1 to 1MPa gauge.

45. The method as claimed in claim 1, wherein the volume space velocity of the first gas is 5000-^-1。

46. The method as claimed in claim 45, wherein the volume space velocity of the first gas is 10000-^-1。

47. The method of claim 1, wherein the pre-reduction is for a duration of 1-20 hours.

48. The method of claim 47, wherein the duration of the pre-reduction is 2-10 hours.

49. The method of claim 1, wherein the second gas is a mixture of a hydrocarbon and an inert gas that is gaseous at a reduction activation temperature.

50. The method of claim 49, wherein the molar ratio of the inert gas to the hydrocarbon that is gaseous at the reductive activation temperature is from 1 to 200: 1.

51. the method of claim 50, wherein the molar ratio of the inert gas to the hydrocarbon that is gaseous at the reductive activation temperature is from 15 to 30: 1.

52. the method according to claim 1, wherein the hydrocarbon that is gaseous at the reduction activation temperature is one or two or more selected from an alkane that is gaseous at the reduction activation temperature and an alkene that is gaseous at the reduction activation temperature.

53. The method of claim 52, wherein the hydrocarbon that is gaseous at the reductive activation temperature is selected from C₁-C₄Alkane and C₂-C₄One or more than two kinds of olefins.

54. The method according to claim 53, wherein the hydrocarbon that is gaseous at the reduction activation temperature is one or two or more selected from methane, ethane, ethylene, propylene, propane, butane, and butene.

55. The method of claim 1, wherein the reductive activation is carried out at a temperature of 180-450 ℃.

56. The method of claim 55, wherein the reductive activation is carried out at a temperature of 200-400 ℃.

57. The process as claimed in claim 1, wherein the pressure in the reactor in which the reductive activation is carried out is 0 to 2.5MPa in gauge.

58. A process as claimed in claim 57, in which the pressure in the reactor at which the reductive activation is effected is from 0.1 to 2MPa gauge.

59. The method as claimed in claim 1, wherein the volume space velocity of the second gas is 5000-30000 hours in terms of the hydrocarbon which is gaseous at the reduction activation temperature^-1。

60. The method as claimed in claim 59, wherein the volume space velocity of the second gas is 10000-20000 hours based on the hydrocarbon which is gaseous at the reduction activation temperature^-1。

61. The method of claim 1, wherein the duration of the reductive activation is 1-20 hours.

62. The method of claim 61, wherein the duration of said reductive activation is from 4 to 12 hours.

63. The method according to claim 1, wherein the inert gas in the first gas and the second gas is the same or different and each is one or two or more selected from nitrogen and a group zero element gas.

64. The method of claim 63, wherein the inert gas in the first gas and the second gas is each nitrogen and/or argon.

65. The method according to claim 1, wherein the auxiliary element is one or two or more of Li, K and Ce.

66. The method of claim 65, wherein the promoter element is K and/or Ce.

67. The method according to claim 1, wherein the promoter element is supported on the carrier simultaneously with the group VIII metal element.

68. The method as claimed in claim 1, wherein the calcination is carried out at a temperature of 300-900 ℃ and the duration of the calcination is 1-10 hours.

69. The method as claimed in claim 68, wherein the calcination is carried out at a temperature of 400-600 ℃.

70. The method according to claim 1, wherein at least a part of the support is alumina containing a modifying element, the modifying element being one or two or more selected from the group consisting of alkali metal elements, alkaline earth metal elements and group IVB metal elements.

71. The method of claim 70, wherein the modifying element is one or more of K, Mg and Zr.

72. The method of claim 70, wherein the modifying element is present in an amount of 0.1 to 15 wt.% on an elemental basis, based on the total amount of the support.

73. The process of claim 72, wherein the modifying element is present in an amount of 1.5 to 8% by weight on an elemental basis based on the total amount of the support.

74. The method of any one of claims 70-73, wherein the modifying element-containing alumina is prepared by a method comprising: and (2) impregnating alumina with an impregnating solution containing a compound containing a modifying element, and drying and roasting the alumina adsorbed with the impregnating solution in sequence to obtain the alumina containing the modifying element.

75. The method as claimed in claim 74, wherein the drying is carried out at a temperature of 50-300 ℃, the duration of the drying is 1-12 hours, the baking is carried out at a temperature of 300-900 ℃, and the duration of the baking is 0.5-8 hours.

76. The method of claim 1, wherein the alumina is theta alumina.

77. The method of claim 1 or 76, wherein the alumina is prepared using a method comprising: mixing gamma-Al₂O₃The calcination is carried out in an air atmosphere at a temperature of 700-1050 ℃, and the duration time of the calcination is 0.5-5 hours.

78. The process of claim 1 further comprising separating unreacted hydrogen and/or carbon monoxide from the product stream of the fischer-tropsch synthesis and recycling at least part of the hydrogen and/or at least part of the carbon monoxide for use in formulating the feed to the fischer-tropsch synthesis reaction.

79. The process of claim 1, further comprising S10, in S10, separating methane from the methane-containing feed gas.

80. The method of claim 79, wherein the feed gas is one or more selected from shale gas, coal bed gas, natural gas and refinery gas.

81. The method of claim 79, wherein the feed gas is one or more selected from shale gas, coal bed gas, natural gas and coke oven gas.

82. The process of claim 79, wherein in S10, methane is separated from the feed gas using a cryogenic condensation process.

83. The method of claim 1, wherein the weight ratio of methane employed in S11 to methane employed in S21 is 1: 0.8-1.5.

84. The method of claim 83, wherein the weight ratio of methane employed in S11 to methane employed in S21 is 1: 0.9-1.3.

85. A low-carbon olefin production system comprises a steam reforming reaction unit, a dry reforming reaction unit, a synthetic gas mixing unit, a Fischer-Tropsch synthesis reaction product separation unit and a circulation unit,

the Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor and is used for contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst to obtain a Fischer-Tropsch synthesis product material flow containing low-carbon olefins, wherein the Fischer-Tropsch synthesis catalyst is the Fischer-Tropsch synthesis catalyst described in any one of claims 1 and 14-77;

86. The system of claim 85, further comprising a feed gas separation unit for separating methane from a methane-containing feed gas, wherein the methane output port of the feed gas separation unit is in communication with the methane feed input port of the steam reforming reaction unit and the methane feed input port of the dry reforming reaction unit, respectively, for feeding separated methane to the steam reforming reaction unit and the dry reforming reaction unit, respectively.

87. The system of claim 86, wherein the feed gas separation unit is provided with a cryogenic condenser for condensing the feed gas to separate methane from the feed gas.

88. The system of any one of claims 85-87, wherein the fischer-tropsch synthesis reaction unit further comprises a reductive activation subunit for reductive activation of the fischer-tropsch synthesis catalyst precursor.

89. The system of claim 88, wherein the reductive activation subunit comprises a first gas storage delivery device, a second gas storage delivery device, a reductive gas control device, and a reductive activation reactor,

the first gas storage and conveying device is used for storing a first gas and conveying the first gas into the reduction activation reactor, and the first gas is hydrogen or hydrogen and inert gas;

the second gas storage and conveying device is used for storing a second gas and conveying the second gas into the reduction activation reactor, 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,

and when the reduction activation subunit operates, the reduction gas control device is arranged to firstly input hydrogen into the reduction activation reactor so that the Fischer-Tropsch synthesis catalyst precursor contacts with the hydrogen to carry out a pre-reduction reaction to obtain a pre-reduction catalyst, and then input a second gas into the reduction activation reactor so that the pre-reduction catalyst contacts with the second gas to carry out the reduction activation reaction.

90. The system of claim 89, wherein the hydrocarbon that is gaseous at the reduction activation temperature is one or more than two selected from the group consisting of an alkane that is gaseous at the reduction activation temperature, and an alkene that is gaseous at the reduction activation temperature.

91. The system of claim 90, wherein the hydrocarbon that is gaseous at the reductive activation temperature is selected from C₁-C₄Alkane and C₂-C₄One or more than two kinds of olefins.

92. The system of claim 91, wherein the hydrocarbon that is gaseous at the reductive activation temperature is selected from one or more of methane, ethane, ethylene, propylene, propane, butane, and butene.

93. The system of claim 89, wherein the inert gas in the first gas and the second gas is the same or different.

94. The system of claim 93, wherein the inert gas in the first gas and the second gas is each one or more selected from nitrogen and a group zero gas.

95. The system of claim 94, wherein the inert gas in the first gas and the second gas is each nitrogen and/or argon.

96. The system of claim 89, wherein the reductive activation reactor is the same reactor as the fischer-tropsch synthesis reactor, or

The reduction activation reactor and the Fischer-Tropsch synthesis reactor are not the same reactor, and a reduction activation catalyst output port of the reduction activation reactor is communicated with a catalyst input port of the Fischer-Tropsch synthesis reactor so as to send the reduction activation catalyst output by the reduction activation reactor into the Fischer-Tropsch synthesis reactor.

97. The system of any one of claims 85 to 87, wherein the Fischer-Tropsch synthesis reactor is a fluidized bed reactor.