CN108728151B

CN108728151B - Alpha-olefin production method and alpha-olefin production system

Info

Publication number: CN108728151B
Application number: CN201710256681.1A
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: CN108728151A

Abstract

The invention discloses a production method and a production system of alpha-olefin, wherein the production 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; the method comprises the steps of 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 reaction temperature for producing alpha-olefin 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 alpha-olefin can effectively reduce the energy consumption of the system and the emission of greenhouse gases (such as carbon dioxide).

Description

Alpha-olefin production method and alpha-olefin production system

Technical Field

The invention relates to a method for producing alpha-olefin and also relates to an alpha-olefin production system.

Background

Linear alpha-olefins are important organic feedstocks and intermediates for the production of comonomers, lubricant base oils, surfactants, polyolefin resins, plasticizers, dyes, pharmaceutical formulations, and the like. The south Africa Sasol company has built a set of production device for separating 1-pentene and 1-hexene from Fischer-Tropsch (F-T) synthetic products (rich in alpha-olefin) and successfully put into production, and the process has the greatest advantages that coal is used as a raw material, the 1-pentene and 1-hexene are used as by-products for recovery, the industrial production cost is low, and higher benefits can be obtained.

The most widely used method for producing alpha-olefins at present is olefin oligomerization, but the method has too high production cost and cannot produce linear alpha-olefins with odd carbon numbers, which have the same market value. The cost of extracting linear 1-hexene from a crude product by adopting a high-temperature F-T Fischer-Tropsch synthesis technology by south Africa Sasol company is less than one third of the cost of producing the linear 1-hexene by adopting an ethylene trimerization method by Philips company, and meanwhile, high-temperature F-T synthesis can also obtain high-value-added products such as 1-pentene, 1-heptene and the like with odd carbon numbers based on the Anderson-Schulz-Flory distribution rule (the chain growth is in exponentially decreased molar distribution) of F-T synthesis products. Therefore, the separation of alpha-olefins from products of the Fischer-Tropsch synthesis is of significant commercial value.

The energy sources in China are in the resource distribution situation of rich coal, much natural gas and oil shortage, the indirect conversion of coal or natural gas into clean and efficient liquid fuel through F-T synthesis is an important aspect of reasonably utilizing resources, and a main technical approach for relieving the contradiction between supply and demand of petroleum in China can be realized. In recent years, in the field of coal chemical industry in China, coal rapidly rises up through the preparation of alpha-olefin from methanol, and the direct preparation of the alpha-olefin from the coal through synthesis gas (FTO process) is another process for preparing the alpha-olefin from the coal. The process first converts coal or natural gas to syngas (CO and H)₂) And then directly preparing alpha-olefin through F-T synthesis.

The technological process for preparing alpha-olefin by adopting FTO technology 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 an alpha-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 a 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 requirements of the Fischer-Tropsch synthesis reaction, removing acid gas and sulfide M from the converted crude synthetic gas F 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 for synthesis reaction, generating a Fischer-Tropsch reaction product N containing olefin, separating the Fischer-Tropsch reaction product N to obtain alpha-olefin K through an alpha-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 major problems with the FTO process described above are: 1. the energy consumption is high, and the utilization rate of carbon atoms is low; 2. the emission of carbon dioxide is 5-6 times of that of the traditional petroleum route; 3. the distribution of the Fischer-Tropsch synthesis product is limited by the Anderson-Schulz-Flory rule, and is limited by the generation of a large amount of methane and carbon dioxide caused by strong exothermicity of reaction, so that the overall energy efficiency of the FTO process is low, and the industrial process of the FTO process is seriously influenced. The FTO process uses a large amount of cooling water and discharged sewage to keep the water consumption high.

Therefore, there is a need to optimize the FTO process and select a system with high energy efficiency and reduced greenhouse gas emission.

Disclosure of Invention

The invention aims to provide a method for producing alpha-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 the production of alpha-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 alpha-olefin to obtain a Fischer-Tropsch synthesis product material flow;

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

According to a second aspect of the present invention, there is provided an α -olefin production system comprising a steam reforming reaction unit, a dry reforming reaction unit, a synthesis gas mixing unit, a fischer-tropsch synthesis reaction product separation unit, and a circulation 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 at the reaction temperature of producing alpha-olefin to obtain a Fischer-Tropsch synthesis product material flow containing the alpha-olefin;

the Fischer-Tropsch synthesis reaction product separation unit is used for separating the Fischer-Tropsch synthesis product material flow to obtain methane, carbon dioxide, alpha-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.

The method and the system for producing alpha-olefin can effectively reduce the energy consumption of the system and the emission of greenhouse gases (such as carbon dioxide).

Drawings

Fig. 1 is a diagram for illustrating a typical process flow of the prior art for directly preparing alpha-olefins from coal via synthesis gas (FTO process).

Fig. 2 is a diagram illustrating an alpha-olefin production process and system 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 II: water gas shift unit

IV: synthesis gas purification unit V: Fischer-Tropsch synthesis unit VI: alpha-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: alpha-olefins

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

Z: purge gas J: purifying synthesis 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: alpha-olefins

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.

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

In step S11, the molar ratio of methane to water vapor may be 1: 0.5 to 4, preferablyIs 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 pressure being expressed as 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 gas hourly volume space velocity of the feed may be 100 based on the total amount of methane and carbon dioxide00-100000 hours^-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 method for producing α -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 α -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 method for producing α -olefins of the present invention, the purity of methane, which is one of the raw materials for steam reforming and dry reforming, is generally 90% by weight or more. 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 alpha-olefins of the present invention, the raw material utilization rate of the method of the present invention can be further improved by controlling the amount of methane fed to the steps S11 and S21 according to the reaction properties of steam reforming and dry reforming and the requirements of the feed for the fischer-tropsch synthesis reaction. Preferably, the weight ratio of the methane used in step S11 to the methane used in step S21 is 1: 0.5-2.5.

According to the method for producing alpha-olefin of the present invention, in step S31, at least a part of the steam reformed synthesis gas and at least a part of the dry weight integrated synthesis gas are mixed to prepare a fischer-tropsch synthesis reaction feed meeting the hydrogen-carbon ratio (i.e., the molar ratio of hydrogen to carbon monoxide) of the fischer-tropsch synthesis feed. From the viewpoint of improving the selectivity of alpha-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 2.8: 1, more preferably 0.8 to 2.6: 1, more preferably 1.5 to 2.5: 1.

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

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, supported on the support, a first metal element, a second metal element and optionally a third metal element.

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/g。

Preferably, the carrier contains theta-Al₂O₃. By introducing theta-Al into the support₂O₃Therefore, the Fischer-Tropsch synthesis catalyst can obtain higher catalytic activity. Generally, the content of the θ -alumina may be 10% by weight or more, preferably 20% by weight or more, more preferably 30% by weight or more, further preferably 40% by weight or more, and further preferably 50% by weight or more, based on the total amount of alumina in the catalyst. 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 in terms of the element may be 40% by weight or more, preferably 50% by weight or more, more preferably 55% by weight or more, further preferably 60% by weight or more, and further preferably 70% by weight or more, based on the total amount of the group VIII metal element in the fischer-tropsch synthesis catalyst. The content of the group VIII metal element having a valence lower than its maximum oxidation valence may be up to 100% by weight, for example, 95%, 90%, 85% by weight in terms of the element, based on the total amount of the group VIII metal element in the catalyst. According to this preferred embodiment, the Fischer-Tropsch synthesis catalyst can be used directly to catalyze reactions without additional reductive 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.

The Fischer-Tropsch synthesis catalyst according to this preferred embodiment, in a particularly preferred example, the VI < th > catalystThe element of the group II metal is Fe, and in the X-ray photoelectron spectrum of the Fischer-Tropsch synthesis catalyst, peaks corresponding to FeO (generally appearing at 711.9eV and 724.4 eV) and peaks corresponding to Fe exist₅C₂The 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. 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, the content of Fe determined from the peak corresponding to FeO and the content of Fe determined from the peak corresponding to Fe₅C₂The ratio of the content of Fe determined by the spectral peak of (A) is preferably 8.5 to 20: 1, more preferably 9 to 18: 1, more preferably 9.5 to 16: 1, particularly preferably 9.5 to 12: 1. according to the Fischer-Tropsch synthesis catalyst of this particularly preferred example, from the viewpoint of further improving the catalytic activity, the total amount of Fe determined by X-ray photoelectron spectroscopy is measured as the reference of the elements by the peak corresponding to FeO and the peak corresponding to Fe₅C₂The content of Fe determined by the peak of (a) may be 40% by weight or more, preferably 50% by weight or more, more preferably 55% by weight or more, still more preferably 60% by weight or more, and still more preferably 70% by weight or more. Based on the total amount of Fe determined by X-ray photoelectron spectroscopy, the total amount of Fe is determined by the peak corresponding to FeO and the peak corresponding to Fe₅C₂The content of Fe determined by the spectral peak of (a) may be up to 100% by weight, for example, 95%, 90%, 85% by weight.

In the present invention, the X-ray photoelectron spectroscopy was measured on an ESCALab250 type X-ray photoelectron spectrometer equipped with Thermo Avantage V5.926 software of Thermo Scientific, the excitation source was monochromatized Al Ka X-ray, the energy was 1486.6eV, the power was 150W, the transmission energy for narrow scan was 30eV, and the base vacuum during analysis and test was 6.5X 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), Yangshui et al, scientific Press, 1 st edition at 1990, 9 months, p. 371-.

The fischer-tropsch synthesis catalyst according to this preferred embodiment contains a carrier and a group VIII metal element supported on the carrier, and also contains a second metal element and optionally a third metal element supported on the carrier. The catalyst containing the second metal element and optionally the third metal element shows more excellent catalytic activity. In the present invention, "optional" means either with or without.

The second metal element is one or more than two selected from VIIB metal elements. Preferably, the second metal element is Mn. Preferably, the second metal element may be contained in an amount of 0.1 to 10% by weight, preferably 1 to 8% by weight, more preferably 2 to 6% by weight, and further preferably 2.5 to 4% by weight, in terms of element, based on the total amount of the catalyst.

The third metal element is one or more selected from alkali metal elements, alkaline earth metal elements and group IVB metal elements. Specific examples of the third metal element may include, but are not limited to: one or more of Li, Na, K, Mg, Ca, Zr and Ti. Preferably, the third metal element is one or two or more of Li, K, Mg, and Zr. More preferably, the third metal element is K and/or Zr. Further preferably, the third metal elements are K and Zr, and in this case, the content of Zr is preferably 40 to 70 wt%, more preferably 50 to 65 wt%, based on the total amount of the third metal elements. The content of the third metal element may be 0.1 to 15% by weight, preferably 1 to 12% by weight, more preferably 2 to 11% by weight, and still more preferably 4 to 9% by weight in terms of element, based on the total amount of the 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 of the catalyst. When the catalyst contains both the second metal element and the third metal element, the second metal element is more preferably Mn, and the third metal element is more preferably one or two or more of Mg, K, Li, and Zr, further preferably K and/Zr, and further preferably K and Zr, so that more excellent catalytic activity 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 spectrogram, the higher the temperature of the desorption peak is, the catalyst is favorable for desorption of the hydrocarbon molecules, and for a plurality of catalysts with desorption peaks at the same position, the higher the peak area is, the stronger the desorption capacity of the catalyst on the hydrocarbon molecules is. The Fischer-Tropsch synthesis catalyst according to the preferred embodiment exhibits unique CO₂TPD desorption spectrum, with a desorption peak (herein, the desorption peak is referred to as CO) in a temperature range of 300-600 ℃, preferably 350-500 ℃, and more preferably 380-450 ℃₂High temperature desorption peak). The CO is₂The peak area of the high-temperature desorption peak is generally 0.5 to 4a.u (arbitrary units), preferably 1 to 3a.u (arbitrary units). CO of the Fischer-Tropsch Synthesis catalyst according to this preferred embodiment₂In the desorption spectrum of TPD, another desorption peak (herein, this desorption peak is referred to as CO) exists in the temperature range of 90-200 deg.C, preferably 140-₂Low temperature desorption peak). The CO is₂The peak area of the low-temperature desorption peak is generally 1 to 4.5a.u (arbitrary units), preferably 2 to 4a.u (arbitrary units).

CO-TPD (namely, temperature programmed desorption CO) can be used for representing the dissociation capability of the catalyst for CO, and the higher the temperature of a CO desorption peak is, the higher the activity of the catalyst is, and the improvement of the selectivity of olefin is facilitated. For multiple catalysts with desorption peaks at the same location, a catalyst with a larger peak area favors CO dissociation. In the CO-TPD desorption diagram of the Fischer-Tropsch synthesis catalyst according to the preferred embodiment, a desorption peak (herein, the desorption peak is referred to as a CO low-temperature desorption peak) exists in a temperature interval of 300-700 ℃, preferably 320-650 ℃, more preferably 450-600 ℃ and further preferably 480-560 ℃. The peak area of the CO low-temperature desorption peak is generally 1-8a.u (arbitrary unit), preferably 3-7a.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 high-temperature desorption peak) exists in the temperature interval of 350-720 ℃, preferably 550-700 ℃. The peak area of the CO high-temperature desorption peak is generally 0.5-4a.u (arbitrary unit), and preferably 1.5-3.5a.u (arbitrary unit). And the peak position of the CO low-temperature desorption peak is lower than the peak position of the CO high-temperature desorption peak.

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, the position of the desorption peak is the peak position of the desorption peak.

According to the production method of α -olefins of the present invention, the fischer-tropsch synthesis catalyst can be obtained by subjecting a fischer-tropsch synthesis catalyst precursor to reductive activation, the reductive activation including 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 first metal element, a second metal element and an optional third metal element which are loaded on the carrier. The types and contents of the carrier, the first metal element and the second metal element can be found in the fischer-tropsch synthesis catalyst described above, and will not be described in detail here.

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.

In the reduction activation, the first gas is hydrogen gas or a mixed gas of hydrogen gas and an 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, as a carrier gas. 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 200: 1, preferably 1 to 100: 1, more preferably 2 to 50: 1, more preferably 5 to 20: 1.

the contacting temperature of the Fischer-Tropsch synthesis catalyst precursor with the first gas is such that 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 3MPa, preferably 0.1 to 1MPa, in terms of gauge pressure. The duration of the pre-reduction may be selected depending on the temperature of the pre-reduction and the pressure of the first gas. Generally, the duration of the pre-reduction may be 1 to 20 hours, preferably 2 to 15 hours, more preferably 4 to 10 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 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, as a carrier gas. 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 2 to 50: 1, more preferably 5 to 20: 1.

in the reduction activation method, the reduction activation may be performed at a temperature of 150-. The volume space velocity of the second gas (in terms of the hydrocarbon that is gaseous at the reduction activation temperature) may be 1000-30000 hours^-1Preferably 2000-10000 hours^-1. In carrying out the reduction activation, the pressure in the reactor may be 0 to 3MPa, preferably 0.1 to 1MPa, 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 from 1 to 20 hoursPreferably 2 to 15 hours, more preferably 3 to 8 hours.

In the reductive activation process, the fischer-tropsch synthesis catalyst precursor may be prepared by a process comprising the steps of: and roasting a carrier loaded with oxides of the VIII group metal elements and/or precursors of the oxides of the VIII group metal elements and a compound containing an auxiliary element to obtain a catalyst precursor, wherein the carrier is alumina.

According to the method for preparing the catalyst precursor of the present invention, 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 being modified. In a preferred embodiment, at least part of the alumina 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. Further preferably, the modifying element is Mg and/or Zr.

From the viewpoint of further improving the catalytic activity of the finally prepared catalyst, the content of the modifying 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 carrier.

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

In a preferred example, alumina supporting a compound containing a modifying element may be calcined to obtain alumina containing a modifying element. 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 300-800 deg.C, and the duration of the calcination may be selected depending on the calcination temperature, and may be generally 0.5-12 hours, preferably 1-8 hours. The firing is performed in an air atmosphere. Specifically, the modifying element may be supported on the alumina by means of impregnation. When the modifying element is supported on the alumina by impregnation, the alumina may be impregnated with an impregnation solution containing a compound containing the modifying element, and the alumina adsorbed with the impregnation solution may be sequentially dried and calcined to obtain the alumina containing the 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 carried out at a temperature of 50 to 300 ℃, preferably 100 ℃ to 300 ℃, and the drying may be carried out 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 20 hours, preferably 2 to 12 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 impregnation liquid adsorbed thereon may be dried under conventional conditions to remove the solvent from the impregnation liquid, thereby obtaining the support loaded with the oxide and/or precursor. Generally, the drying may be carried out at a temperature of 50 to 300 ℃, preferably 100 ℃ to 300 ℃, 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 20 hours, preferably 2 to 12 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 obtain a catalyst precursor. The group VIII metal element in the catalyst precursor is substantially in its highest oxidation state. Generally, the calcination may be carried out at a temperature of 300-900 deg.C, preferably 300-800 deg.C, and the duration of the calcination may be selected depending on the calcination temperature, and may be generally 0.5-12 hours, preferably 1-8 hours. The firing is performed in an air atmosphere.

From the viewpoint of further improving the catalytic activity of the finally prepared catalyst, the catalyst further comprises an auxiliary element loaded on the carrier, wherein the auxiliary element is a VIIB metal element, or the auxiliary element is a VIIB metal element and an alkali metal element. The group VIIB metal element is preferably Mn. The alkali metal element is preferably one or two or more of Li, Na, and K, and more preferably Li and/or K. Preferably, the auxiliary element is Mn and one or two selected from Li and K.

The supporting amount of the promoter element on the carrier may be such that the content of the promoter element may be 0.1 to 10% by weight, preferably 1 to 8% by weight, more preferably 4 to 7% by weight, in terms of the element, based on the total amount of the catalyst precursor. When the auxiliary element is a group VIIB metal element and an alkali metal element, the content of the group VIIB metal element is preferably 10 to 90 wt%, preferably 20 to 85 wt%, more preferably 30 to 80 wt%, and more preferably 40 to 75 wt% in terms of the element, based on the total amount of the auxiliary element.

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 nitrate, lithium carbonate, 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 auxiliary agent 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 of the finally produced catalyst, it is preferable to perform the 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 α -olefins of the present invention, in step S31, the fischer-tropsch synthesis reaction may be performed under conventional conditions for producing α -olefins. Preferably, the Fischer-Tropsch synthesis reaction feed and the Fischer-Tropsch synthesis catalyst may be contacted at a temperature of from 200 ℃ to 380 ℃, preferably from 250 ℃ to 350 ℃. The pressure at which the Fischer-Tropsch synthesis reaction feed is contacted with the Fischer-Tropsch synthesis catalyst may be in the range 0.8 to 3MPa, preferably 1 to 2.8MPa, expressed as gauge pressure.

According to the method for producing alpha-olefin, the Fischer-Tropsch synthesis reaction can be contacted in a fixed bed reactor, a fluidized bed reactor or a 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 fixed bed reactor. The volume space velocity of the Fischer-Tropsch synthesis reaction feed may be such that, when hydrogen and carbon monoxide are contacted with the Fischer-Tropsch synthesis catalyst in a fixed bed reactor, the volume space velocity of the Fischer-Tropsch synthesis reaction feed is2000-50000 hours^-1Preferably 5000-^-1Preferably 10000-^-1。

According to the method for producing α -olefins of the present invention, in step S41, α -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 alpha olefins, methane and carbon dioxide, respectively.

According to the method for producing alpha-olefins of the present invention, methane separated from the product stream of the Fischer-Tropsch synthesis is fed to step S11 and/or step S21 as a raw material for 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 method for producing the alpha-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 flow are recycled, so that the utilization rate of raw materials is effectively improved, and the emission of greenhouse gas carbon dioxide is obviously reduced.

The method for producing alpha-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 from the viewpoint of further improving the utilization rate of the raw material, and feeding at least part of the hydrogen and/or at least part of the carbon monoxide to step S31 for formulating the feed for the fischer-tropsch synthesis reaction. Preferably, part of the hydrogen and/or part of the carbon monoxide separated from the product stream of the Fischer-Tropsch synthesis is recycled to step S31 for use in formulating the Fischer-Tropsch synthesis reaction feed, while the remainder of the hydrogen and/or carbon monoxide is vented as purge gas to the system. Generally, the amount of hydrogen and carbon monoxide used for recycle may be in the range of from 10 to 98%, preferably from 15 to 98%, based on the total amount of hydrogen and carbon monoxide separated from the Fischer-Tropsch synthesis product stream.

According to a second aspect of the present invention, there is provided an α -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.

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 dry reforming reaction unit may be provided with a conventional dry reforming reactor and corresponding feed, discharge and control means to enable a reforming reaction of methane with carbon dioxide to obtain a dry reformed gas having hydrogen and carbon monoxide as main components.

The synthesis gas mixing unit is 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 reforming reaction unit 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. Or a pipeline mixer can be adopted to directly mix the steam regenerated synthetic gas and the dry weight integrated synthetic gas in a conveying pipeline so as to obtain the Fischer-Tropsch synthesis reaction feed. The synthesis gas mixing unit can be provided with various common control devices for controlling the mixing proportion of the steam reforming synthesis gas and the dry weight integrated synthesis gas, so that the Fischer-Tropsch synthesis reaction feeding material meeting the hydrogen-carbon ratio of the Fischer-Tropsch synthesis reaction is obtained.

The Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor, 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 at the reaction temperature of producing alpha-olefin to obtain a Fischer-Tropsch synthesis product material flow containing the alpha-olefin. The Fischer-Tropsch synthesis reactor can be various common reactor forms, and specifically, the Fischer-Tropsch synthesis reactor can be a fixed bed reactor, a fluidized bed reactor or a combination of the fixed bed reactor and the fluidized bed reactor. Preferably, the fischer-tropsch synthesis reactor is a fixed 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 reductive activation subunit may reductively activate the fischer-tropsch synthesis catalyst precursor 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 types of the first gas and the second gas have 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 the first gas for 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 for a reduction reaction. The reducing gas control means may employ conventional control elements such as various control valves to control the type of gas fed to the reduction reactor and the amount of gas fed thereto.

The reduction reactor is used for accommodating 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 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 alpha-olefin production system, the circulating unit is used for circularly sending the methane separated by the Fischer-Tropsch synthesis reaction product separating unit into one or both of the steam reforming reaction unit and the dry reforming reaction unit, circularly sending the carbon dioxide separated by the Fischer-Tropsch synthesis reaction product separating unit into the dry reforming reaction unit, and circularly sending the hydrogen and/or the carbon monoxide separated by the Fischer-Tropsch synthesis reaction product separating unit into the Fischer-Tropsch synthesis reaction unit.

The circulation unit can be provided with a methane conveying pipeline which is respectively used for communicating the Fischer-Tropsch synthesis reaction product separation unit with the steam reforming reaction unit and the dry reforming reaction unit, and a control valve arranged on the methane conveying pipeline, so that the methane separated by the Fischer-Tropsch synthesis reaction product separation unit is respectively conveyed into the steam reforming reaction unit and the dry reforming reaction unit. The circulation unit can be provided with a carbon dioxide conveying pipeline for communicating the Fischer-Tropsch synthesis reaction product separation unit and the dry reforming reaction unit and a control valve arranged on the carbon dioxide conveying pipeline so as to convey the carbon dioxide output by the Fischer-Tropsch synthesis reaction product separation unit into the dry reforming reaction unit.

When the Fischer-Tropsch synthesis reaction product separation unit further separates hydrogen and carbon monoxide, the circulation unit is preferably provided with a conveying pipeline for communicating the Fischer-Tropsch synthesis reaction product separation unit with the Fischer-Tropsch synthesis reaction unit and a control valve arranged on the conveying pipeline, so that the hydrogen and the carbon monoxide separated by the Fischer-Tropsch synthesis reaction product separation unit are sent into the Fischer-Tropsch synthesis reaction unit. Can send into during the ft synthesis reaction unit with hydrogen and carbon monoxide through same conveying pipeline, also can send into the ft synthesis reaction unit respectively with hydrogen and carbon monoxide through different conveying pipeline, can set up hydrogen conveying pipeline and set up the control flap on hydrogen conveying pipeline this moment respectively and with carbon monoxide conveying pipeline and the control flap of setting on carbon monoxide conveying pipeline.

The α -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 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 an alpha-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 α -olefin production system includes a raw material gas separation unit I, a steam reforming reaction unit II, a dry reforming reaction unit III, a fischer-tropsch synthesis reaction unit IV, a fischer-tropsch synthesis product separation unit V, and a circulation unit.

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. The steam reforming synthesis gas E and the dry weight integrated synthesis gas F are mixed (preferably by adopting a pipeline mixer) to prepare the Fischer-Tropsch synthesis reaction feed G which accords 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. The Fischer-Tropsch synthesis reactor in the Fischer-Tropsch synthesis reaction unit IV operates at the temperature for producing alpha-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 alpha-olefin K, unreacted hydrogen, carbon monoxide, methane M and carbon dioxide N. Wherein the alpha-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, the conversion of CO (X)_CO)、C₅-C₁₅Selectivity to alpha-olefin (S)_{Alpha-olefins}) And C₅Above (C)₅₊) Selectivity of hydrocarbons

Respectively 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 produce CO₂The number of moles of (a);

n_{alpha-olefins}To moles of alpha-olefin produced;

to generate CH₄、C₂Hydrocarbons, C₃Hydrocarbons and C₄The sum of the moles of hydrocarbons.

Preparation examples 1 to 21 were used to prepare 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 kind and content of each metal element in the catalyst and the catalyst precursor were measured by the X-ray fluorescence spectrum analysis method specified in RIPP 132-92 (compiled in methods for petrochemical engineering analysis (RIPP test method), Yangshui et al, science publishers, 1 st edition at 1990/9, p 371) -379). 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 5% 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 manganese 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 loaded into a fixed bed reactor, and H is introduced into the reactor₂The pressure of the reactor is adjusted to be 0.1MPa, and the volume space velocity of hydrogen is 10000 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 ℃ and hydrogen was switched to ethane at a volumetric space velocity of 2000 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 alpha-olefins

After the reduction activation is finished, introducing synthesis gas into the reactor, and heating the temperature of the reactor to 310 ℃ to carry out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 10000 hours^-1At a pressure of 1.5MPa (in gauge pressure) and a synthesis gas 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 obtained after 50 hours of the reaction are shown in table 5.

Preparation example 2

By following the same procedure as in preparation example 1The catalyst and the alpha-olefin were prepared in the same manner except that, in the step (1), gamma-Al was used₂O₃The carrier was prepared by impregnating it directly with a modified zirconium solution without calcination, wherein the Zr content was 5 wt% in terms of the element, based on the total amount of the carrier.

Preparation example 3

A catalyst was prepared and an alpha-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 α -olefin was prepared in the same manner as in preparation example 1, except that manganese nitrate hexahydrate was not used in preparing the impregnation liquid in step (2).

Preparation example 5

A catalyst was prepared and an α -olefin was prepared in the same manner 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 impregnated mixture was sequentially dried and calcined, thereby obtaining a catalyst precursor.

Preparation example 6

A catalyst was prepared and an alpha-olefin was prepared 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 an alpha-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

The catalyst was prepared and alpha-olefin was prepared in the same manner as in preparation example 1, except that in step (3), after hydrogen was introduced, ethane was not continuously introduced, but step (4) was directly performed, i.e., only hydrogen was used for reductive activation, instead of ethane.

Preparation example 9

A catalyst was prepared and alpha-olefin was prepared in the same manner as in preparation example 1, except that in step (3), ethane was replaced with CO of the same volume, i.e., after the prereduction with hydrogen was completed, the reactor was cooled to 200 ℃ and hydrogen was switched to CO, and the volume space velocity of CO was 2000 hours^-1And maintained for 4 hours.

Preparation example 10

A catalyst was prepared and α -olefin was prepared in the same manner as in preparation example 1, except that, in step (3), ethane was replaced with a mixed gas of CO and nitrogen, that is, after the prereduction of hydrogen was completed, the reactor was cooled to 200 ℃, hydrogen was switched to a mixed gas of CO and nitrogen, and the molar ratio of CO to nitrogen was 1: 1, the volume space velocity of the mixed gas of CO and nitrogen is 2000 hours^-1And maintained for 4 hours.

Preparation example 11

A catalyst was prepared and alpha-olefin was produced in the same manner as in preparation example 1, except that in step (3), ethane was introduced directly into the reactor without passing hydrogen, that is, the catalyst was charged into a fixed bed reactor, ethane was passed into the reactor, the reactor pressure was adjusted to 0.1MPa, the reactor temperature was raised from 25 ℃ to 200 ℃ and kept at the temperature for 4 hours, wherein the volumetric space velocity of ethane was 2000 hours^-1。

Preparation example 12

The catalyst was prepared and alpha-olefin was prepared in the same manner as in preparation example 2, except that in step (3), after hydrogen was introduced, ethane was not continuously introduced, but step (4) was directly performed, i.e., only hydrogen was used for reductive activation, instead of ethane.

Preparation example 13

The catalyst was prepared and alpha-olefins were prepared in the same manner as in preparation example 2, except that in step (3) ethane was replaced by CO, i.e. after prereduction with hydrogen was complete, the reactor was cooled to 200 ℃ and hydrogen was switched to CO, and the volumetric space velocity of CO was 2000 hours^-1And maintained for 4 hours.

Preparation example 14

(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 zirconium nitrate pentahydrate in 41g 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 200 ℃ and atmospheric pressure 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 the content of Zr based on the total amount of the carrier as 2.5 wt% in terms of element.

(2) Preparation of the catalyst precursor

Adding ferric nitrate, potassium carbonate and manganese 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 for 1 hour at 800 ℃ 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 200 ℃ under atmospheric pressure for 3 hours in an air atmosphere. The dried substance was calcined at 800 ℃ for 1 hour in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is loaded into a fixed bed reactor, and H is introduced into the reactor₂Adjusting the pressure of the reactor to 0.1MPa and hydrogenHas a volume space velocity of 20000 hours^-1The temperature of the reactor was raised from 25 ℃ to 500 ℃ and kept constant at this temperature for 6 hours. The reactor was then cooled to 250 ℃, hydrogen switched to ethane, and the volumetric space velocity of ethane was 10000 hours^-1After 5 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 alpha-olefins

After the reduction activation is finished, introducing synthesis gas into the reactor, and adjusting the temperature of the reactor to 310 ℃ to carry out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 10000 hours^-1The pressure is 1.5MPa, and the composition of the synthesis 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 obtained after 50 hours of the reaction are shown in table 5.

Preparation example 15

A catalyst was prepared and an alpha-olefin was prepared in the same manner as in preparation example 14, except that, in step (1), gamma-Al₂O₃The carrier was prepared without calcination by direct saturation impregnation with a modified zirconium solution, in which the Zr content was 2.5 wt% in terms of the element, based on the total amount of the carrier.

Preparation example 16

(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.

Dissolving zirconium nitrate pentahydrate in 53g 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 mixture obtained by impregnation was placed in an oven at 300 ℃ and atmospheric pressureDried in an air atmosphere for 2 hours. 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 Zr was 6% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

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 300 ℃ under atmospheric pressure for 2 hours 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 300 ℃ under atmospheric pressure for 2 hours 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 loaded into a fixed bed reactor, and H is introduced into the reactor₂And argon (wherein the molar ratio of argon to hydrogen is 15: 1), adjusting the pressure of the reactor to be 0.15MPa, and the volume space velocity of hydrogen to be 20000 hours^-1The temperature of the reactor was increased from 25 ℃ to 450 ℃ and maintained at this temperature for 10 hours. The reactor was then cooled to 250 ℃, hydrogen switched to ethane, and the volumetric space velocity of ethane was 8000 hours^-1After a maintenance time of 8 hours, 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 alpha-olefins

After the reduction activation is finished, introducing synthesis gas into the reactor, and adjusting the temperature of the reactor to 315 ℃ to carry out Fischer-Tropsch synthesis reactionWherein the volume space velocity of the synthetic gas is 12000 hours^-1The pressure is 1.2MPa, and the composition of the synthesis gas 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 obtained after 50 hours of the reaction are shown in table 5.

Preparation example 17

(1) Preparation of the support

Preparation of theta-Al by the same method as in preparation example 1₂O₃。

Dissolving magnesium nitrate in 43g of deionized water to prepare a modified magnesium solution, and adding 100.0g of the prepared theta-Al into the modified magnesium solution₂O₃The resulting mixture was saturated and immersed at 25 ℃ for 2 hours. Then, the impregnated mixture was placed in an oven and dried at 100 ℃ under atmospheric pressure for 12 hours in an air atmosphere. And roasting the dried substance in an air atmosphere at 300 ℃ for 8 hours to obtain the carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis to determine that the content of Mg was 5% 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 manganese nitrate hexahydrate into 12mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid.

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 100 ℃ under atmospheric pressure for 12 hours in an air atmosphere. And roasting the dried substance in an air atmosphere at 300 ℃ for 8 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 100 ℃ under atmospheric pressure for 12 hours in an air atmosphere. The dried substance was calcined at 300 ℃ for 8 hours in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is loaded into a fixed bed reactor, and H is introduced into the reactor₂The pressure of the reactor is adjusted to be 0.1MPa, and the volume space velocity of the hydrogen is 20000 hours^-1The temperature of the reactor was raised from 25 ℃ to 500 ℃ and maintained at this temperature for 4 hours. Then, the reactor was cooled to 350 ℃, hydrogen was switched to a mixed gas of ethane and argon (wherein the molar ratio of argon to ethane was 10: 1), and the volume space velocity of ethane was 10000 hours^-1After 3 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 alpha-olefins

After the reduction activation is finished, introducing synthesis gas into the reactor, and adjusting the temperature of the reactor to 305 ℃ to carry out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 8000 hours^-1The pressure is 2MPa, and the composition of the synthesis gas 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 obtained after 50 hours of the reaction are shown in table 5.

Preparation example 18

(1) Preparation of the support

Preparation of theta-Al by the same method as in preparation example 14₂O₃。

Potassium nitrate was dissolved in 41g of deionized water to prepare a modified potassium solution, and 100.0g of the thus-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 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 carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis, and it was determined that the content of K was 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 manganese nitrate hexahydrate into 11mL 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 500 ℃ for 8 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 500 ℃ for 8 hours in an air atmosphere to obtain a catalyst precursor.

(3) Reductive activation of catalyst precursor

The catalyst precursor is loaded into a fixed bed reactor, and H is introduced into the reactor₂And argon (among them, argon and H)₂In a molar ratio of 5: 1) the pressure of the reactor is adjusted to be 0.1MPa, and the volume space velocity of the hydrogen is 20000 hours^-1The temperature of the reactor was raised from 25 ℃ to 400 ℃ and maintained at this temperature for 8 hours. Then, the reactor was cooled to 250 ℃, hydrogen was switched to a mixed gas of ethane and argon (wherein the molar ratio of argon to ethane was 20: 1), and the volume space velocity of ethane was 10000 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 alpha-olefins

After the reduction activation is finished, introducing synthesis gas into the reactor, and adjusting the temperature of the reactor to 320 ℃ to carry out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 15000 hours^-1The pressure is 1MPa, and the composition of the synthesis gas 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 obtained after 50 hours of the reaction are shown in table 5.

Preparation example 19

A catalyst was prepared and alpha-olefin was produced in the same manner as in preparation example 18, except that in step (2), ferric nitrate was replaced with cobalt nitrate.

Preparation example 20

A catalyst was prepared and alpha-olefin was produced in the same manner as in preparation example 18, except that in step (2), the iron nitrate was replaced with nickel nitrate.

Preparation example 21

A catalyst was prepared and alpha-olefin was prepared in the same manner as in preparation example 18, except that in step (2), potassium carbonate was not used, and the amount of manganese nitrate hexahydrate 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

As can be seen from comparison of production examples 1 with production examples 8 to 11, and production example 2 with production examples 12 and 13, the catalytic activity of the finally formed reduction-activated catalyst, particularly the selectivity for α -olefin, can be significantly improved by subjecting the catalyst precursor to pre-reduction with hydrogen and then to reduction activation with a hydrocarbon which is gaseous at the reduction activation temperature. As can be seen by comparing preparation examples 1 and 2 and preparation examples 14 and 15, theta-Al was used₂O₃Can obviously improve the catalytic activity of the Fischer-Tropsch synthesis catalyst.

Examples 1-7 are intended to illustrate the alpha-olefin production process and production system of the present invention.

Example 1

In this embodiment, the α -olefin production system shown in fig. 2 is used, and 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, catalyst filling in the reactorThe chemical agent 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, then exchanging heat with a heat exchanger, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed layer is 750 ℃, the pressure in the reactor is 2MPa, and the gas hourly volume space velocity is 80000h based on the total amount of methane and steam^-1。

(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture meeting the hydrogen-carbon ratio of 2.1: 1 fischer-tropsch synthesis reaction feed G.

And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fixed bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, and contacting with a Fischer-Tropsch synthesis catalyst (the catalyst prepared in the preparation example 1) to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 320 ℃, the pressure in the reactor is 1.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 15000h^-1。

(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 alpha-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 alpha-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 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 an alpha-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 through a synthesis gas purification unit IV to obtain purified synthesis gas (the molar ratio of hydrogen to carbon monoxide is 2.1: 1).

The purified synthesis gas obtained is conveyed into a Fischer-Tropsch synthesis unit V to carry out Fischer-Tropsch synthesis reaction in a fixed bed reactor (by adopting the catalyst prepared in the preparation example 1), and a Fischer-Tropsch reaction product N containing olefin is generated. Wherein the temperature in the reactor is 320 ℃, the pressure in the reactor is 1.5MPa to synthesizeThe total amount of formed gas is taken as a reference, and the gas hourly volume space velocity is 15000h^-1。

And separating alpha-olefin K from the Fischer-Tropsch reaction product N through an alpha-olefin separation unit VI, discharging carbon dioxide H and methane G generated by the 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) Y back 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.

The composition of the gaseous product stream exiting the Fischer-Tropsch reactor during the reaction was analyzed by an on-line gas chromatograph and the results obtained after 50 hours of reaction are shown in Table 7. The overall water consumption, carbon dioxide emissions, and energy efficiency of the system are listed in table 8.

Comparative example 2

Alpha-olefins were 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 fed into the steam reforming reaction unit II to undergo reforming reaction.

Comparative example 3

Alpha-olefins were 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 to the dry reforming reaction unit III for the reforming reaction.

Example 2

Alpha-olefins were 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

Alpha-olefins were 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 8.

Example 4

In this example, the reaction system shown in FIG. 2 was used, and the specific process flow was as follows.

(1) And (2) taking the coke oven gas with the flow rate of 500kmol/h and the pressure of 3.0MPa as a raw material gas A, sending the raw material gas A into a raw material gas separation unit I for low-temperature condensation separation, and removing sulfur, carbon and other impurities to obtain methane B with the sulfur mass content of less than 1 ppm.

(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) And mixing the second strand of methane with carbon dioxide D with the flow rate of 200kmol/h, the temperature of 370 ℃ and the pressure of 2MPa, then exchanging heat with a heat exchange medium, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1.5, the catalyst filled in the reactor is Ni/Al₂O₃(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 2.5: 1 fischer-tropsch synthesis reaction feed G. Feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fixed bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, contacting with a Fischer-Tropsch synthesis catalyst (the catalyst prepared in the preparation example 14) and carrying out Fischer-Tropsch synthesis reactionShould be used. Wherein the temperature in the reactor is 310 ℃, the pressure in the reactor is 1.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 10000h^-1。

Discharging the alpha-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 5

(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, then exchanging heat with a heat exchanger, and mixingThe temperature of the compound is raised to 700 ℃, and then the compound enters a fixed bed reactor of a steam reforming reaction unit II to carry out 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 860 ℃, the pressure in the reactor is 1MPa, and the gas hourly space velocity is 100000h based on the total amount of methane and steam^-1。

(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 150kmol/h, the temperature of 450 ℃ and the pressure of 3MPa, then exchanging heat with a heat exchange medium, raising the temperature of the mixture to 700 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry weight integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content 12 wt% in terms of element, based on the total amount of the catalyst, Al₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed layer is 650 ℃, the pressure in the reactor is 1.5MPa, and the gas hourly volume space velocity is 60000h based on the total amount of methane and water vapor^-1。

(4) Mixing the steam reforming synthesis gas E and the dry weight integrated synthesis gas F to prepare a mixture meeting the hydrogen-carbon ratio of 1.5: 1 fischer-tropsch synthesis reaction feed G. And (3) feeding the Fischer-Tropsch synthesis reaction feed G into a Fischer-Tropsch synthesis reactor (a fixed bed reactor) of a Fischer-Tropsch synthesis reaction unit IV, and contacting the Fischer-Tropsch synthesis reaction feed G with a Fischer-Tropsch synthesis catalyst (the catalyst prepared in the preparation example 16) to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 290 ℃, the pressure in the reactor is 2.5MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly space velocity is 20000h^-1。

Discharging the alpha-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.

Example 6

An alpha-olefin was produced using the same system and method as in example 5, except that the fischer-tropsch synthesis catalyst was the fischer-tropsch synthesis catalyst prepared in preparation example 17, the temperature in the fischer-tropsch synthesis reactor was 330 ℃, the pressure in the reactor was 1.5MPa, and the gas hourly volume space velocity was 30000h, based on the total amount of synthesis gas^-1。

Example 7

An alpha-olefin was produced using the same system and method as in example 5, except that the fischer-tropsch synthesis catalyst was the fischer-tropsch synthesis catalyst prepared in preparation example 18, the temperature in the fischer-tropsch synthesis reactor was 280 ℃, the pressure in the reactor was 1.5MPa, and the gas hourly volume space velocity was 30000h, based on the total amount of synthesis gas^-1。

TABLE 7

TABLE 8

Numbering	Water consumption (t/t)_{Alpha-olefins})	Carbon dioxide emission (t/t)_{Alpha-olefins})	Energy efficiency (%)
				Example 1	14	0.6	55
Comparative example 1	21	6.8	34
				Comparative example 2	19	2.3	36
Comparative example 3	26	0.8	45
				Example 4	15	0.9	52
Example 5	15	1.0	50
				Example 6	16	1.1	49
Example 7	18	1.3	44

Note: the energy efficiency is the sum of the calorific value of the alpha-olefin finally discharged from the device/the calorific value of the raw material such as the coal-electric steam catalyst solvent fed into the device, namely the calorific value of the obtained alpha-olefin/the comprehensive energy consumption required for producing the alpha-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 process for producing alpha-olefins, the process comprising the steps of:

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, contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at a reaction temperature for producing alpha-olefin to obtain a Fischer-Tropsch synthesis product flow, wherein the Fischer-Tropsch synthesis catalyst comprises a carrier, and a first metal element, a second metal element and an optional third metal element which are loaded on the carrier, the first metal element is a VIII group metal element, the VIII group metal element is Fe, the second metal element is one or more selected from VIIB group metal elements, the third metal element is one or more selected from alkali metal elements, alkaline earth metal elements and IVB group metal elements, the carrier is alumina, and the alumina contains theta-alumina, based on the total amount of alumina in the Fischer-Tropsch synthesis catalyst, the content of the theta-alumina is more than 50 weight percent, 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, and in an X-ray photoelectron spectrum chart of the Fischer-Tropsch synthesis catalyst, a spectrum peak corresponding to FeO and a spectrum peak corresponding to Fe exist₅C₂The Fe content determined from the peak corresponding to FeO and the Fe content determined from the peak corresponding to Fe₅C₂The ratio of 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 peak of the Fischer-Tropsch synthesis catalyst is 55 to 95 percent, 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 first metal element, a second metal element and an optional third metal element which are loaded on the carrier, the first metal element is loaded on the carrier in the form of an oxide, and the valence state of a group VIII 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: roasting a carrier loaded with oxides of VIII group metals and/or precursors of oxides of VIII group metals and a compound containing an auxiliary element, wherein the auxiliary element is a VIIB group metal element, or the auxiliary element is a VIIB group metal element and an alkali metal element, and loading the VIII group metals and the auxiliary element on the carrier in an impregnation mode, wherein the impregnation frequency is 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 2, wherein in S11, the methane and the water vapor are contacted at a temperature of 700 ℃ and 950 ℃ and a pressure of 0.1 to 5MPa, the pressure being expressed as gauge pressure.

4. The process of any one of claims 1 to 3, 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 6, wherein in S21, the contacting of methane with carbon dioxide is carried out at a temperature of 600-800 ℃ and a pressure of 0.1-5MPa, said pressure being expressed as gauge pressure.

8. The process as claimed in any one of claims 1, 6 and 7, wherein, in S21, the dry reforming reaction is carried out in a fixed bed reactor, and the hourly space velocity of the gas fed is 10000-100000 hours based on the total amount of methane and carbon dioxide^-1。

9. The method as claimed in claim 1, wherein the Fischer-Tropsch synthesis reaction temperature for producing alpha-olefin in S31 is 200-380 ℃.

10. The method as claimed in claim 9, wherein the temperature of the Fischer-Tropsch synthesis reaction for producing α -olefins in S31 is 250-250 ℃.

11. The process of any one of claims 1, 9 and 10, wherein the fischer-tropsch synthesis reaction feed is contacted with the fischer-tropsch synthesis catalyst at a pressure of from 0.8 to 3MPa, expressed as gauge pressure, S31.

12. The process of claim 11, wherein in S31, the fischer-tropsch synthesis reaction feed is contacted with the fischer-tropsch synthesis catalyst at a pressure in the range of from 1 to 2.8MPa, expressed as a gauge pressure.

13. The process of any one of claims 1, 9 and 10, wherein in S31, the contacting is carried out in a fixed bed reactor.

14. The process as claimed in claim 13, wherein the gas hourly space velocity of the Fischer-Tropsch synthesis reaction feed in S31 is 2000-50000 h^-1。

15. The process as claimed in claim 14, wherein the gas hourly space velocity of the Fischer-Tropsch synthesis reaction feed in S31 is 5000-^-1。

16. The process of any one of claims 1, 9 and 10, wherein the molar ratio of hydrogen to carbon monoxide in the fischer-tropsch synthesis reaction feed in S31 is in the range of from 0.4 to 3: 1.

17. the process of claim 16, 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 2.5: 1.

18. the process of claim 1, 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.

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 500 DEG C₂High temperature desorption peak.

20. The process of claim 19, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is present in the temperature range of 380 ℃ to 450 DEG C₂High temperature desorption peak.

21. The process of any one of claims 1 and 18 to 20, wherein the fischer-tropsch synthesis catalyst has CO₂In the desorption spectrum of TPD, CO is also present in the temperature range of 90-200 DEG C₂Low temperature desorption peak.

22. The process of claim 21, wherein the fischer-tropsch synthesis catalyst has CO₂In the desorption spectrum of TPD, CO is also present in the temperature range of 140 ℃ and 180 DEG C₂Low temperature desorption peak.

23. The method as claimed in claim 1, wherein the CO-TPD desorption diagram of the Fischer-Tropsch synthesis catalyst has a CO high-temperature desorption peak in a temperature range of 350-720 ℃.

24. The method as claimed in claim 23, wherein the fischer-tropsch synthesis catalyst has a CO high temperature desorption peak in the temperature range of 550-700 ℃ in the desorption diagram.

25. The method as claimed in any one of claims 1, 23 and 24, wherein in the desorption diagram of CO-TPD from the Fischer-Tropsch synthesis catalyst, a CO low-temperature desorption peak exists in the temperature range of 300 ℃ to 700 ℃, and the peak position of the CO low-temperature desorption peak is lower than that of the CO high-temperature desorption peak.

26. The method as claimed in claim 25, wherein in the desorption diagram of CO-TPD from the Fischer-Tropsch synthesis catalyst, a CO low-temperature desorption peak exists in a temperature range of 450 ℃ and 600 ℃, and the peak position of the CO low-temperature desorption peak is lower than that of the CO high-temperature desorption peak.

27. The process as claimed in claim 26, wherein in the desorption diagram of CO-TPD from the Fischer-Tropsch synthesis catalyst, a CO low-temperature desorption peak is also present in the temperature interval of 480 ℃ and 560 ℃, and the peak position of the CO low-temperature desorption peak is lower than that of the CO high-temperature desorption peak.

28. The method according to claim 1, wherein the Fe content determined from the peaks corresponding to FeO and from the peaks corresponding to Fe, expressed in elemental terms₅C₂The ratio of Fe content determined by the spectrum peak of (1) is 9.5-16: 1.

29. the method of claim 28, wherein the Fe content determined from the peaks corresponding to FeO and from the peaks corresponding to Fe, on an elemental basis₅C₂The ratio of Fe content determined by the spectrum peak of (1) is 9.5-12: 1.

30. the method according to claim 1, wherein the total amount of Fe determined by X-ray photoelectron spectroscopy is represented by a peak corresponding to FeO and a peak corresponding to Fe on an elemental basis₅C₂The Fe content determined by the peak of (1) is 60 to 90% by weight.

31. The process as claimed in any one of claims 1, 18 to 20, 23, 24 and 28 to 30, wherein the content of the group VIII metal element having a valence lower than its maximum oxidation valence is 40% by weight or more in terms of the element based on the total amount of the group VIII metal element in the catalyst.

32. The process as claimed in claim 31, wherein the content of the group VIII metal element having a valence lower than its maximum oxidation valence is 50% by weight or more in terms of the element based on the total amount of the group VIII metal element in the catalyst.

33. The process as claimed in claim 32, wherein the content of the group VIII metal element having a valence lower than its maximum oxidation valence is 60% by weight or more in terms of the element based on the total amount of the group VIII metal element in the catalyst.

34. The method of any one of claims 1, 18-20, 23, 24, and 28-30, wherein the second metallic element is Mn.

35. The process of claim 34, wherein the second metal element is present in an amount of 0.1 to 10 wt.% on an elemental basis based on the total amount of the fischer-tropsch synthesis catalyst.

36. The process of claim 35, wherein the second metal element is present in an amount of 2.5 to 4 wt.% on an elemental basis based on the total amount of the fischer-tropsch synthesis catalyst.

37. The method according to any one of claims 1, 18 to 20, 23, 24 and 28 to 30, wherein the third metal element is one or two or more selected from Li, Na, K, Mg, Ca, Zr and Ti.

38. The method according to claim 37, wherein the third metal element is one or two or more of Li, K, Mg, and Zr.

39. The method of claim 38, wherein the third metallic element is K and Zr.

40. The process of claim 39, wherein the third metal element is present in an amount of 0.1 to 15 wt.% as an element, based on the total amount of the Fischer-Tropsch synthesis catalyst.

41. The process of claim 40, wherein the third metal element is present in an amount of 4 to 9 wt.% as an element, based on the total amount of the Fischer-Tropsch synthesis catalyst.

42. The process of any one of claims 1, 18 to 20, 23, 24 and 28 to 30, wherein the group VIII metal element is present in an amount of from 3 to 30 wt.% as element, based on the total amount of the fischer-tropsch synthesis catalyst.

43. The process of claim 42 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.

44. The process of claim 43, 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.

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

46. The method of claim 45 wherein the pre-reduction is carried out at a temperature of 300-550 ℃.

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

48. 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.

49. A process as claimed in claim 48, in which the pressure in the reactor at which the pre-reduction is carried out is in the range 0.1 to 1MPa gauge.

50. The method of any one of claims 1 and 45-49, wherein the duration of the pre-reduction is 1-20 hours.

51. The method of claim 50, wherein the pre-reduction is for a duration of 2-15 hours.

52. 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.

53. The method of claim 52, 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.

54. the method of claim 53, wherein the molar ratio of the inert gas to the hydrocarbon that is gaseous at the reductive activation temperature is from 5 to 20: 1.

55. 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.

56. The method of claim 55, 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.

57. The method according to claim 56, 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.

58. The method as claimed in any one of claims 1 and 52-57, wherein the reductive activation is carried out at a temperature of 180-400 ℃.

59. The method of claim 58, wherein the reductive activation is carried out at a temperature of 200-350 ℃.

60. The method as claimed in any one of claims 1 and 52 to 57, wherein the second gas has a volume space velocity, based on the hydrocarbon which is gaseous at the reduction activation temperature, of 1000-30000 hours^-1。

61. A process as claimed in any one of claims 1 and 52 to 57, in which the pressure in the reactor at which the reductive activation is carried out is in the range 0 to 3MPa gauge.

62. The process as set forth in claim 61, wherein the pressure in the reactor in which the reductive activation is carried out is from 0.1 to 1MPa by gauge.

63. The method of any one of claims 1 and 52-57, wherein the duration of the reductive activation is from 1 to 20 hours.

64. The method of claim 63, wherein the duration of said reductive activation is from 2 to 15 hours.

65. The method according to any one of claims 1, 47 and 52-54, 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.

66. The method of claim 65, wherein the inert gas in the first gas and the second gas is the same or different, each being nitrogen and/or argon.

67. The method according to claim 1, wherein the auxiliary element is Mn, or the auxiliary element is Mn and one or two or more selected from Li, Na and K.

68. The method according to claim 1, wherein the compound containing an auxiliary element is supported on the support simultaneously with the group VIII metal element.

69. The method as claimed in claim 68, wherein the calcination is carried out at a temperature of 300-900 ℃ and the duration of the calcination is 0.5-12 hours.

70. The method according to claim 1, wherein at least part of the alumina is alumina containing a modifying element which is 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 Li, Na, K, Mg, Ca, Zr, and Ti.

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

73. The method of claim 70, wherein the modifying element is present in an amount of 0.1 to 10 wt.% on an elemental basis, based on the total amount of the support.

74. A process as claimed in claim 73, in which the modifying element is present in an amount of from 2 to 6% by weight, calculated as element, based on the total amount of the support.

75. The method of claim 70, wherein the modifying element-containing alumina is prepared by a method comprising: and roasting the alumina loaded with the compound containing the modifying element to obtain the alumina containing the modifying element.

76. The method as claimed in claim 75, wherein, in the preparation method of the alumina containing the modifying element, the calcination is carried out at a temperature of 300-900 ℃ and the duration of the calcination is 0.5-12 hours.

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

78. The method of claim 1 or 77, wherein the alumina is prepared using a method comprising: mixing gamma-Al₂O₃At 700-1050 deg.CThe calcination is carried out at a temperature in an air atmosphere.

79. The method as claimed in claim 78, wherein, in the preparation method of the alumina, the duration of the calcination is 0.5-5 hours.

80. 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.

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

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

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

84. The process of claim 81, wherein methane is separated from said feed gas by cryocondensation.

85. The method of any of claims 1 and 80-84, wherein the weight ratio of methane employed in S11 to methane employed in S21 is 1: 0.5-2.5.

86. An alpha-olefin production system comprises a steam reforming reaction unit, a dry reforming reaction unit, a synthesis 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 at the reaction temperature of producing alpha-olefin to obtain a Fischer-Tropsch synthesis product material flow containing the alpha-olefin, wherein the Fischer-Tropsch synthesis catalyst is the Fischer-Tropsch synthesis catalyst described in any one of claims 1 and 18-84;

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 optionally 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.

87. The system of claim 86, 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.

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

89. The system of any one of claims 86-88, wherein the fischer-tropsch synthesis reaction unit further comprises a reductive activation subunit for reductive activation of the fischer-tropsch synthesis catalyst precursor.

90. The system of claim 89, wherein said 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 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 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 a first gas into the reduction activation reactor to enable the Fischer-Tropsch synthesis catalyst precursor to be in contact with 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 to enable the pre-reduction catalyst to be in contact with the second gas to carry out a reduction reaction.

91. The system of claim 90, 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.

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

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

94. The system according to claim 90, 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.

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

96. The system of claim 90, wherein the reductive activation reactor is the same reactor as the fischer-tropsch synthesis reactor, or

The reduction 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 86-88, wherein the Fischer-Tropsch synthesis reactor is a fixed bed reactor.