CN108728152B

CN108728152B - Synthetic oil production method and synthetic oil production system

Info

Publication number: CN108728152B
Application number: CN201710256682.6A
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: CN108728152A

Abstract

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

Description

Synthetic oil production method and synthetic oil production system

Technical Field

The invention relates to a synthetic oil production method and a synthetic oil production system.

Background

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 Fischer-Tropsch (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. The process for directly preparing clean oil from coal by synthetic gas includes such steps as converting coal or natural gas to synthetic gas (CO and H)₂) And then directly preparing the liquid fuel through F-T synthesis. The most important advantage of the synthetic oil prepared by F-T synthesis is that the synthetic oil does not contain non-ideal components such as sulfur, nitrogen, aromatic hydrocarbon and the like, belongs to clean fuel, and completely conforms to the strict requirements and increasingly harsh environmental regulations of modern engines.

Currently, iron-based catalysts are generally used in industry, and slurry bed or fixed bed processes are adopted and connected in series with hydrocracking units to crack the high-carbon hydrocarbon-wax product into related liquid fuels and chemicals such as gasoline, diesel oil or lubricating oil. The technological process of preparing the synthetic oil by the process is shown in a figure 1, and comprises a coal water slurry preparation unit I, a coal gasification unit II, a water gas conversion unit III, a synthetic gas purification unit IV, a Fischer-Tropsch synthesis unit V and a synthetic oil separation unit VI which are sequentially connected. The specific process comprises the steps of preparing coal water slurry C from pulverized coal A and water B in a coal water slurry preparation unit I, conveying the coal water slurry C into a coal gasification unit II, reacting the coal water slurry C with oxygen D 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 through a water gas conversion unit III to enable the coal gasification crude synthesis gas E to meet the requirements of the Fischer-Tropsch synthesis reaction to obtain converted crude synthesis gas F, removing acid gas and sulfide M from the converted crude synthesis gas F through a synthesis gas purification unit IV to obtain purified synthesis gas J, conveying the obtained purified synthesis gas J into a Fischer-Tropsch synthesis unit V to perform the Fischer-Tropsch synthesis reaction to generate a Fischer-Tropsch reaction product N containing olefin, separating the Fischer-Tropsch reaction product N through a low-carbon olefin separation unit VI to obtain synthesis oil K, discharging carbon dioxide H and methane G generated by the Fischer-Tropsch synthesis unit V, and circulating a part of, another portion of the unreacted synthesis gas exits the system as purge gas Z.

The main problems of the above process 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 Fischer-Tropsch synthesis product distribution is limited by an Anderson-Schulz-Flory rule (the molar distribution of chain growth decreasing according to indexes), and is limited by the generation of a large amount of methane and carbon dioxide caused by strong exothermicity of reaction, so that the overall energy efficiency of the process is low, and the industrial process of the F-T synthesis process is seriously influenced. The coal gasification process uses a large amount of cooling water and external sewage to ensure that the water consumption is high.

Therefore, there is a need to optimize the fischer-tropsch process and select a system that is energy efficient and reduces greenhouse gas emissions.

Disclosure of Invention

The invention aims to provide a method for producing synthetic oil, 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 synthetic oil, 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 synthetic oil to obtain a Fischer-Tropsch synthesis product material flow;

s41, separating synthetic oil, 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 a synthetic oil production system comprising a steam reforming reaction unit, a dry reforming reaction unit, a syngas mixing unit, a fischer-tropsch synthesis reaction product separation unit, and a recycle unit,

the steam reforming reaction unit is used for contacting methane with steam to carry out steam reforming reaction to obtain steam reforming synthesis gas;

the dry reforming reaction unit is used for contacting methane and carbon dioxide to carry out dry reforming reaction to obtain dry reforming synthesis gas;

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

the Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor and is used for contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at the reaction temperature of producing the synthetic oil to obtain a Fischer-Tropsch synthesis product material flow containing the synthetic oil;

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

The synthetic oil production method and the synthetic oil production system can effectively reduce the energy consumption of the system and the emission of greenhouse gases (such as carbon dioxide).

Drawings

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

FIG. 2 is a schematic illustration of a synthetic oil production process and system according to the present invention.

FIG. 1 description of reference numerals

I: a coal water slurry preparation unit II: coal gasification unit

III: water gas shift unit IV: syngas purification unit

V: Fischer-Tropsch synthesis unit VI: synthetic oil separation unit

A: pulverized coal B: water (W)

C: coal water slurry D: oxygen gas

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

G: methane H: carbon dioxide

K: synthetic oil M: acid gases and sulfides

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: and (D) water: carbon dioxide

E: steam reforming syngas F: dry weight is integrated into finished gas

G: Fischer-Tropsch synthesis reaction feed H: Fischer-Tropsch synthesis product stream

L: hydrogen and carbon monoxide K for recycle: synthetic oil

M: methane N: carbon dioxide

Z: purge gas

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

In the present invention, the term "synthetic oil" means C₅-C₃₀The component (c).

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, preferably 1: 1-3. The methane may be contacted with the water vapor at a temperature of 700 ℃ to 950 ℃, preferably 800 ℃ to 900 ℃. The pressure in the reactor in which the methane is contacted with the steam may be 0.1 to 5MPa, preferably 1 to 3MPa, said pressure being a gauge pressure. The steam reforming reaction may be carried out in a common reactor. Preferably, the steam reforming reaction is carried out in a fixed bed reactor. The hourly space velocity of the feed gas may be 10000-^-1Excellence inIs selected from 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 hourly space velocity of the feed gas may be 10000-^-1Preferably 50000-100000 hours^-1。

In step S21, various dry reforming catalysts commonly used in the art for dry reforming reactions can be used. As an example, the dry reforming catalyst contains a carrier and an active component supported on the carrier. The carrier can be one or the combination of more than two of alumina, silica, zirconia and silicon carbide. Preferably, the support is alumina, in particularMay be gamma-Al₂O₃、θ-Al₂O₃、δ-Al₂O₃And alpha-Al₂O₃One or more than two of them. The active component can be a group VIII metal element, preferably a non-noble group VIII metal element, such as one or more of Fe, Co and Ni. More preferably, the active component is Ni. The loading of the active ingredient on the carrier may be conventionally selected. In general, the active component may be present in an amount of 1 to 30% by weight, preferably 5 to 25% by weight, more preferably 10 to 15% by weight, calculated as element, based on the total amount of the catalyst.

According to the process for producing synthetic oil 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 synthetic oil according to the present invention further includes the step of separating methane from the feed gas containing methane at S10, 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 may be separated from the 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 synthetic oil 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 synthetic oil 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 of 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 synthetic oil production method 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 formulate a fischer-tropsch synthesis reaction feed that meets the fischer-tropsch synthesis feed hydrogen-to-carbon ratio (i.e., the molar ratio of hydrogen to carbon monoxide). From the viewpoint of improving the selectivity of the synthetic oil, 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 the synthetic oil, in step S31, Fischer-Tropsch synthesis reaction feed is contacted with a Fischer-Tropsch synthesis catalyst to carry out Fischer-Tropsch synthesis reaction, so that a Fischer-Tropsch synthesis product material flow is obtained.

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

According to the fischer-tropsch synthesis catalyst of this preferred embodiment, the support is alumina, specific examples of which may include, but are not limited to: gamma-Al₂O₃、θ-Al₂O₃、δ-Al₂O₃And alpha-Al₂O₃One or more than two of them. Preferably, the carrier is γ -Al₂O₃。

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 canIs 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。

In the fischer-tropsch synthesis catalyst according to this preferred embodiment, the first metal element is one or two or more selected from group VIII metal elements. The group VIII metal element as an active component of the catalyst may be a group VIII noble metal element, a group VIII non-noble metal element, or a combination of the group VIII noble metal element and the group VIII non-noble metal element. 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 30% by weight or more, preferably 40% by weight or more, more preferably 50% by weight or more (e.g., 55% by weight or more), and still more preferably 60% by weight or more, based on the total amount of the group VIII metal element in the fischer-tropsch synthesis catalyst. Based on the total amount of the VIII group metal elements in the Fischer-Tropsch synthesis catalyst, the valence state of the VIII group metal elements is the maximum valence state of the VIII group metal elements which is lower than the maximum oxidation state of the VIII group metal elements in terms of elementsHigh levels may be 100% by weight, for example: 95 wt%, 90 wt%, 85 wt%. 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.

According to the Fischer-Tropsch synthesis catalyst of this preferred embodiment, in a particularly preferred example, the group VIII metal element is Fe, and the Fischer-Tropsch synthesis catalyst has an X-ray photoelectron spectrum in which peaks corresponding to FeO (typically found at 711.9eV and 724.4 eV) and peaks corresponding to Fe are present₅C₂The spectral peak of (usually occurs at 717.9 eV). The Fischer-Tropsch synthesis catalyst according to this particularly preferred example has more excellent catalytic performance. 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 content of Fe determined by the spectral peak of (a) may be 8 to 20: 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 10 to 15: 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 30% by weight or more, preferably 40% by weight or more, more preferably 50% by weight or more (e.g., 55% by weight or more), and still more preferably 60% by weight or more. From the peak corresponding to FeO and corresponding to Fe₅C₂The content of Fe determined by the peak of (a) is generally not higher than 95% by weight, preferably not higher than 90% by weight, more preferably not higher than 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 6 to 20 wt%, and even more preferably from 8 to 15 wt%, calculated as element, based on the total amount of the Fischer-Tropsch synthesis catalyst. In the present invention, the kind and content of each metal element in the catalyst and the catalyst precursor were measured by an X-ray fluorescence spectrum analysis method specified in RIPP 132-92 (compiled in methods for petrochemical engineering (RIPP experiments), Yanggui et al, science publishers, 1 st edition at 9/1990, p. 371-379).

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 supported on the carrier. The catalyst containing the second metal element exhibits more excellent catalytic activity.

The second metal element is one or two or more selected from a group IVB metal element, an optional alkali metal element, and an optional alkaline earth metal element. The group IVB metal element is preferably Zr and/or Ti. The alkali metal element is preferably one or two or more of Li, Na, and K, and more preferably Li and/or K. The alkaline earth metal element is preferably Mg and/or Ca, more preferably Mg. The content of the second metal element may be 0.5 to 10% by weight, preferably 1 to 8% by weight, more preferably 2 to 6% by weight, in terms of element, based on the total amount of the catalyst.

In a preferred embodiment, the second metal element is a group IVB metal element and one or more selected from alkali metal elements and alkaline earth metal elements, which can achieve a better catalytic effect. In this preferred embodiment, the content of the group IVB metal element is preferably 10 to 60% by weight, more preferably 20 to 55% by weight, and still more preferably 30 to 50% by weight, based on the total amount of the second metal element. More preferably, the second metal element is a group IVB metal element and an alkaline earth metal element. Further preferably, the second metal element is Zr and Mg.

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 profile, with a desorption peak (herein, this desorption peak is referred to as CO) in the temperature interval of 300-₂High temperature desorption peak). The CO is₂The peak area of the high-temperature desorption peak is generally 0.3 to 2.5a.u (arbitrary units), preferably 0.5 to 2a.u (arbitrary units). CO of the Fischer-Tropsch Synthesis catalyst according to this preferred embodiment₂In the desorption spectrum of TPD, another desorption peak (herein, the desorption peak is referred to as CO) exists in the temperature range of 100-200 ℃, preferably 140-180 ℃₂Low temperature desorption peak). The CO is₂The peak area of the low-temperature desorption peak is generally 0.5 to 3.5a.u (arbitrary units), preferably 1 to 2a.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 spectrum of the Fischer-Tropsch synthesis catalyst according to the preferred embodiment, a desorption peak (herein, the desorption peak is referred to as a CO high-temperature desorption peak) exists in a temperature interval of 300-600 ℃, preferably 400-500 ℃. The peak area of the CO high-temperature desorption peak is generally 1-5a.u (arbitrary unit), and preferably 2-4a.u (arbitrary unit). In the CO-TPD desorption spectrum of the Fischer-Tropsch synthesis catalyst according to the preferred embodiment, another desorption peak (herein, the desorption peak is referred to as a CO low-temperature desorption peak) exists in the temperature range of 100-200 ℃, preferably 160-180 ℃. The peak area of the CO low-temperature desorption peak is generally 0.5-2a.u (arbitrary unit).

In the present invention, CO₂the-TPD and the CO-TPD are detected on line by using a Michmark chemical adsorption instrument and an OMistar mass spectrometer as a detector, wherein the CO is detected₂TPD recorded signals for the nuclear to cytoplasmic ratio of 44 by the mass spectrometer and CO-TPD recorded signals for the nuclear to cytoplasmic ratio of 28 by the mass spectrometer. In the present invention, the position of the desorption peak is the position of the peak.

According to the production method of synthetic oil of the present invention, the fischer-tropsch synthesis catalyst may 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 contains a carrier, and a first metal element and a second metal element that are supported 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 are not described in detail herein.

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. Preferably, the inert gas is nitrogen and/or argon. When the first gas is a mixed gas of hydrogen and an inert gas, the molar ratio of the inert gas to the hydrogen may be 1 to 30: 1.

the contacting temperature of the Fischer-Tropsch synthesis catalyst precursor with the first gas is such that 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 (calculated by hydrogen) can be 1000-^-1Preferably 2000-10000 hours^-1. The pressure in the reactor in which the pre-reduction is carried out may be 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. 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. At the reduction activation temperatureSpecific examples of the hydrocarbons that are gaseous at a degree 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. 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 5 to 50: 1, more preferably 10 to 30: 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 hydrocarbon gaseous at the reduction activation temperature) may be 1000-10000 hours^-1Preferably 2000-^-1. In the process of carrying out the reduction activation, the pressure in the reactor for carrying out the reduction activation 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 reductive activation. Generally, the duration of the reductive activation may be 1 to 20 hours, preferably 2 to 8 hours, more preferably 4 to 6 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 a second metal element to obtain the catalyst precursor.

The second metal element may be supported on the carrier together with the group VIII metal element, or the second metal element may be supported on the carrier after the group VIII metal element is supported.

In a preferred embodiment, the precursor of the second metal element prior to the oxide of the group VIII metal element and the oxide of the group VIII metal element is supported on a carrier, which can significantly improve the catalytic activity of the catalyst prepared.

In the preferred embodiment, the alumina containing the second metal element can be obtained by a conventional method. In one example, the second metal element may be supported on alumina during the preparation of alumina, for example, by coprecipitation, while the alumina is prepared.

In another example, alumina loaded with a compound containing a second metal element is calcined, thereby obtaining alumina containing the second metal 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 320-800 deg.C, more preferably 350-700 deg.C, and the duration of the calcination may be selected depending on the calcination temperature, and may be generally 1 to 8 hours, preferably 1.5 to 6 hours. The firing is performed in an air atmosphere.

The second metal element may be supported on the alumina by means of impregnation. When the second metal element is supported on the alumina by the impregnation method, the alumina may be impregnated with an impregnation solution containing a compound containing the second metal element, and the alumina having the impregnation solution adsorbed thereon may be sequentially dried and calcined to obtain the alumina containing the second metal element.

The second metal element-containing compound may be a water-soluble salt and/or a water-soluble base containing the second metal element, 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 example, the impregnation may be by conventional impregnation methods, such as saturation impregnation or excess impregnation. The impregnation may be carried out at ambient temperature.

In this example, the drying may be carried out under conditions sufficient to remove volatile species (primarily solvent in the impregnating solution) from the alumina on which the impregnating solution is adsorbed. Specifically, the drying may be performed at a temperature of 50 to 300 ℃, preferably 80 to 250 ℃, more preferably 100 ℃ to 200 ℃, and the drying may be performed under normal pressure (i.e., 1 atm, the same applies) or under reduced pressure. The duration of the drying may be selected depending on the temperature of the drying and the pressure of the drying, and may be generally 1 to 15 hours, preferably 3 to 12 hours. The drying may be performed in an air atmosphere.

According to this preferred embodiment, the carrier for supporting the group VIII metal element may be all the alumina containing the second metal element, or may be a mixture of the alumina containing the second metal element and the alumina containing no second metal element. In general, the content of the alumina containing the second metal element may be 10% by weight or more, preferably 30% by weight or more, more preferably 50% by weight or more, and still more preferably 70% by weight or more, based on the total amount of the support. Particularly preferably, the carrier for supporting the group VIII metal element is all alumina containing the second metal element.

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, for example, one or two 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 80 to 250 ℃, more preferably 100 ℃ to 200 ℃, 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 drying and the pressure of drying, and may be generally 1 to 15 hours, preferably 4 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 300-700 deg.C, and the duration of the calcination may be selected depending on the calcination temperature, and may be generally 1 to 12 hours, preferably 2 to 8 hours. The firing is performed in an air atmosphere.

According to the synthetic oil production method of the present invention, in step S31, the fischer-tropsch synthesis reaction may be performed under conventional synthetic oil production conditions. Preferably, the Fischer-Tropsch synthesis reaction feed and the Fischer-Tropsch synthesis catalyst may be contacted at a temperature of from 200 ℃ to 300 ℃, preferably from 220 ℃ to 280 ℃, more preferably from 250 ℃ to 260 ℃. 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 production method of the synthetic oil, the Fischer-Tropsch synthesis reaction can be contacted in a fixed bed reactor, can also be contacted in a fluidized bed reactor, and can also be contacted in the combination of the fixed bed reactor and the fluidized bed reactor. Preferably, the hydrogen and carbon monoxide are contacted with the catalyst in a fixed bed reactor. When the hydrogen and the carbon monoxide are contacted with the Fischer-Tropsch synthesis catalyst in the fixed bed reactor, the volume space velocity of the feeding material of the Fischer-Tropsch synthesis reaction can be 2000-one 30000 hours^-1Preferably 4000-^-1。

According to the synthetic oil production method of the present invention, in step S41, synthetic oil, methane and carbon dioxide can be separated from the fischer-tropsch synthesis product stream by a conventional method. As an example, the Fischer-Tropsch synthesis product stream may be separated by cryogenic condensation to yield synthetic oil, methane and carbon dioxide, respectively.

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

According to a second aspect of the present invention, there is provided a synthetic oil 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 the reforming reaction of methane with carbon dioxide to give a dry reformed gas having hydrogen and carbon monoxide as the main components.

The synthesis gas mixing unit is respectively communicated with a steam reforming synthesis gas output port of the steam reforming unit and a dry reforming reaction unit dry weight integrated gas output port, and is used for mixing the steam reforming synthesis gas and the dry 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 the synthetic oil to obtain a Fischer-Tropsch synthesis product material flow containing the synthetic oil. 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 hydrogen 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 hydrogen to perform a pre-reduction reaction, so as to obtain a pre-reduction catalyst, and then input a second gas into the reduction activation reactor to contact the pre-reduction catalyst with the second gas to perform 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 synthetic oil 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 synthetic oil 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 a synthetic oil production system according to the present invention, which is described in detail below with reference to fig. 2.

As shown in fig. 2, the synthetic oil 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. And the Fischer-Tropsch synthesis reactor in the Fischer-Tropsch synthesis reaction unit IV operates at the temperature of producing the synthetic oil. And the Fischer-Tropsch synthesis product material flow H output by the Fischer-Tropsch synthesis reaction unit IV enters a Fischer-Tropsch synthesis product separation unit V for separation to obtain the synthetic oil K, unreacted hydrogen, carbon monoxide, methane M and carbon dioxide N. Wherein the synthetic oil 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) Diesel oil component (C)₉-C₁₈) Selectivity (S) of_{Diesel fuel component}) Selectivity of isomerous diesel oil (S)_{Heterogeneous diesel oil}) And synthetic oils (i.e., C)₅-C₃₀Component (b) 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_{diesel fuel component}To moles of diesel components formedThe number of moles;

n_{heterogeneous diesel oil}The number of moles of the isomerate diesel oil produced;

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

Preparation examples 1 to 19 were used for preparing Fischer-Tropsch synthesis catalysts and their properties were evaluated.

In the following preparation examples, the specific surface area, pore volume and average pore diameter were measured by a nitrogen adsorption method, specifically, N was used₂Measuring an adsorption isotherm at a constant temperature of 77K, calculating a specific surface area and a pore volume according to a BET formula, and calculating an average pore size distribution according to a BJH method; the particle size distribution was determined using a laser particle sizer.

In the following preparation examples, the 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 corrected by the C1s peak (284.6eV) of elemental carbon, data processing was performed on Thermo Avantage software, and sensitivity factor was used in the analytical modulesAnd (5) carrying out quantitative analysis.

Preparation example 1

(1) Preparation of the support

Dissolving zirconium nitrate pentahydrate in 75g of deionized water to prepare a modified zirconium solution, and adding 100.0g of gamma-Al into the modified zirconium solution₂O₃(the parameters of the properties are listed in Table 1), saturated and immersed for 2 hours at 25 ℃. 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 6% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric ammonium citrate into 17mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid. To the impregnation solution, 24g of the carrier was added, 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 150 ℃ for 6 hours under an air atmosphere at normal pressure. The dried substance was calcined at 420 ℃ for 5 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 2000 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 synthetic oils

After the reduction activation is finished, introducing synthesis gas into the reactor, adjusting the temperature of the reactor to 260 ℃ and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 6000 hour^-1The pressure is 1.5MPa, and the composition of the synthesis gas is H₂: 50 of CO: 50 (molar ratio). The composition of the gas mixture output from the 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 5.

Preparation example 2

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 1, except that, in step (1), γ -Al was used₂O₃After being calcined at 980 ℃ for 2 hours, the mixture is added into the impregnation liquid. Wherein, the roasted product is subjected to X-ray diffraction analysis, and theta-Al is determined to be obtained₂O₃The property parameters are shown in Table 1.

Preparation example 3

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 1, except that step (1) was not performed, and γ -Al was used₂O₃Directly used in step (2) to prepare the catalyst precursor.

Preparation example 4

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 1, except that, in step (2), instead of step (1), zirconium nitrate pentahydrate and ferric ammonium citrate, which were in the same weight as in step (1) of preparation example 1, were used to prepare an impregnation liquid, that is, the impregnation liquid for impregnating the carrier was dispersed with zirconium nitrate pentahydrate and ferric ammonium citrate, thereby obtaining a catalyst precursor.

Preparation example 5

A catalyst was prepared and a synthetic oil 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 6

The catalyst was prepared and synthetic oil 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, and ethane was not used.

Preparation example 7

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 1, except that in step (3), ethane was used in an equal volumeThe CO replacement is carried out, namely after the prereduction of the introduced hydrogen is finished, the reactor is cooled to 200 ℃, the hydrogen is switched into CO, and the volume space velocity of the CO is 2000 hours^-1And maintained for 4 hours.

Preparation example 8

A catalyst was prepared and synthetic oil 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 with hydrogen was completed, the reactor was cooled to 200 ℃, hydrogen was switched to a mixed gas of CO and nitrogen, and the volume ratio of the mixed gas of CO and 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 9

Preparing a catalyst and synthetic oil 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 increased from 25 ℃ to 200 ℃ and maintained at the temperature for 4 hours, wherein the volumetric space velocity of ethane was 2000 hours^-1。

Preparation example 10

The catalyst was prepared and synthetic oil 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, and ethane was not used.

Preparation example 11

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

Preparation example 12

(1) Preparation of the support

Dissolving zirconium nitrate pentahydrate in 68g of deionized water to prepare a modified zirconium solution, and adding 100.0g of gamma-Al into the modified zirconium solution₂O₃(the parameters of the properties are listed in Table 1), saturated and immersed for 2 hours at 25 ℃. 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 700 ℃ for 1.5 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 3% by weight in terms of element based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric ammonium citrate into 15mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid. To the impregnation solution, 24g of the carrier was added, 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 180 ℃ for 4 hours under atmospheric pressure in an air atmosphere. The dried substance was calcined at 700 ℃ for 2 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 2000 hours^-1The temperature of the reactor was increased from 25 ℃ to 450 ℃ and maintained 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 5000 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 synthetic oils

After the reduction activation is finished, introducing synthesis gas into the reactor, adjusting the temperature of the reactor to 240 ℃ and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 5000 hours^-1The pressure is 1.5MPa, and the composition of the synthesis gas is H₂: CO 55: 45 (molar ratio). The composition of the gas mixture output from the 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 5.

Preparation example 13

(1) Preparation of the support

Dissolving zirconium nitrate pentahydrate and potassium nitrate in 81g of deionized water to prepare a modified solution, and adding 100.0g of gamma-Al into the modified solution₂O₃(the parameters of the properties are listed in Table 1), saturated and immersed for 2 hours at 25 ℃. 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 at 350 ℃ in an air atmosphere for 6 hours to obtain the carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis, and it was determined that the content of Zr was 3 wt% and the content of K was 3 wt% in terms of the element, based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric ammonium citrate into 18mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid. To the impregnation solution, 24g of the carrier was added, 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. 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 hydrogen is 5000 hours^-1The temperature of the reactor was raised from 25 ℃ to 480 ℃ and maintained at this temperature for 4 hours. Then, the reactor was cooled to 260 ℃, hydrogen was switched to a mixed gas of ethane and argon (in which the molar ratio of ethane to argon was 1: 20), and the volume space velocity of ethane was 8000 hours^-1After 6 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 synthetic oils

After the reduction activation is finished, introducing synthesis gas into the reactor, adjusting the temperature of the reactor to 260 ℃ and carrying out Fischer-Tropsch synthesis reaction, whereinThe volume space velocity of the synthetic gas is 5000 hours^-1The pressure is 1MPa, and the composition of the synthesis gas is H₂: CO 60: 40 (molar ratio). The composition of the gas mixture output from the 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 5.

Preparation example 14

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 13, except that zirconium nitrate pentahydrate was not used in step (1).

Preparation example 15

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 13, except that potassium nitrate was not used in step (1).

Preparation example 16

(1) Preparation of the support

Dissolving zirconium nitrate pentahydrate and magnesium nitrate in 81g of deionized water to prepare a modified solution, and adding 100.0g of gamma-Al into the modified solution₂O₃(same as example 6), 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. The dried material was calcined at 460 ℃ for 3 hours in an air atmosphere to obtain a carrier. The prepared carrier was subjected to X-ray fluorescence spectroscopic analysis, and it was determined that the content of Zr was 3 wt% and the content of Mg was 3 wt% in terms of the element, based on the total amount of the carrier.

(2) Preparation of the catalyst precursor

Adding ferric ammonium citrate into 18mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid. To the impregnation solution, 24g of the carrier was added, 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. The dried substance was calcined at 400 ℃ for 5 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 gas mixture (whichWherein the molar ratio of argon to hydrogen is 10: 1) the pressure of the reactor is adjusted to be 0.15MPa, and the volume space velocity of the hydrogen is 6000 hours^-1The temperature of the reactor was raised from 25 ℃ to 350 ℃ and maintained at this temperature for 10 hours. Then, the reactor was cooled to 250 ℃, hydrogen was switched to a mixed gas of ethane and argon (in which the molar ratio of ethane to argon was 1: 10), and the volumetric space velocity of ethane was 4000 hours^-1After 6 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 synthetic oils

After the reduction activation is finished, introducing synthesis gas into the reactor, adjusting the temperature of the reactor to 260 ℃ and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 5000 hours^-1The pressure is 1.5MPa, and the composition of the synthesis gas is H₂: 50 of CO: 50 (molar ratio). The composition of the gas mixture output from the 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 5.

Preparation example 17

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 16, except that zirconium nitrate pentahydrate was not used in step (1).

Preparation example 18

A catalyst was prepared and a synthetic oil was prepared in the same manner as in preparation example 16, except that magnesium nitrate was not used in step (1).

Preparation example 19

(1) Preparation of the support

The carrier was prepared in the same manner as in preparation example 1.

(2) Preparation of the catalyst precursor

Adding cobalt nitrate hexahydrate into 18mL of deionized water, heating in a water bath at 50 ℃, stirring and mixing uniformly to obtain a steeping fluid. To the impregnation solution, 24g of the carrier was added, 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. 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 prepared in step (2) was reductively activated in the same manner as in example 1 to prepare a Fischer-Tropsch synthesis catalyst, the composition of which is shown in tables 2 and 4, and CO₂The results of the-TPD and CO-TPD tests are listed in Table 3.

(4) Preparation of synthetic oils

After the reduction activation is finished, introducing synthesis gas into the reactor, adjusting the temperature of the reactor to 280 ℃ and carrying out Fischer-Tropsch synthesis reaction, wherein the volume space velocity of the synthesis gas is 1000 hours^-1The pressure is 5MPa, and the composition of the synthesis gas is H₂: CO 40: 60 (molar ratio). The composition of the gas mixture output from the 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 5.

TABLE 1

TABLE 2 (based on the total amount of catalyst)

Numbering	Group VIII Metal element/content (wt%)	Second metal element/content (wt%)	Alumina type
				Preparation example 1	Fe/8	Zr/5.5	γ-Al₂O₃
Preparation example 2	Fe/8	Zr/5.5	θ-Al₂O₃
				Preparation example 3	Fe/8	Is free of	γ-Al₂O₃
Preparation example 4	Fe/8	Zr/5.5	γ-Al₂O₃
				Preparation example 5	Fe/8	Zr/5.5	γ-Al₂O₃
Preparation example 6	Fe/8	Zr/5.5	γ-Al₂O₃
				Preparation example 7	Fe/8	Zr/5.5	γ-Al₂O₃
Preparation example 8	Fe/8	Zr/5.5	γ-Al₂O₃
				Preparation example 9	Fe/8	Zr/5.5	γ-Al₂O₃
Preparation example 10	Fe/8	Zr/5.5	θ-Al₂O₃
				Preparation example 11	Fe/8	Zr/5.5	θ-Al₂O₃
Preparation example 12	Fe/11	Zr/2.7	γ-Al₂O₃
				Preparation example 13	Fe/12	Zr/2.6+K/2.6	γ-Al₂O₃
Preparation example 14	Fe/12	K/2.6	γ-Al₂O₃
				Preparation example 15	Fe/12	Zr/2.6	γ-Al₂O₃
Preparation example 16	Fe/13	Zr/2.6+Mg/2.6	γ-Al₂O₃
				Preparation example 17	Fe/13	Mg/2.6	γ-Al₂O₃
Preparation example 18	Fe/13	Zr/2.6	γ-Al₂O₃
				Preparation example 19	Co/11	Zr/5.3	γ-Al₂O₃

TABLE 3

TABLE 4

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

TABLE 5

¹: the diesel oil component comprises normal paraffin, isoparaffin and olefin, and the isomerate refers to the isoparaffin in the diesel oil component.

Comparing preparation examples 1 and 6 to 9, and preparation example 2 and 10 and 11, it can be seen that the catalytic activity of the finally formed reduction-activated catalyst can be significantly improved by pre-reducing the catalyst precursor with hydrogen and then performing reduction activation with hydrocarbon which is gaseous at the reduction activation temperature, and the selectivity to diesel components, particularly to isomerate diesel can be significantly improved. As can be seen by comparing preparation example 1 with preparation example 2, theta-Al was used₂O₃Can obviously improve the catalytic activity of the Fischer-Tropsch synthesis catalyst.

Examples 1 to 6 are intended to illustrate the synthetic oil production process and production system of the present invention.

Example 1

The present embodiment adopts the synthetic oil production system shown in fig. 2, which includes a raw material gas separation unit I, a steam reforming reaction unit II, a dry reforming reaction unit III, a fischer-tropsch synthesis reaction unit IV, a fischer-tropsch synthesis product separation unit V, and a circulation unit. The specific process flow is as follows.

(1) And (2) sending shale gas with the flow rate of 220kmol/h and the pressure of 2.0MPa as a raw material gas A into a raw material gas separation unit I for low-temperature condensation separation, and removing sulfur, carbon and other impurities to obtain methane B with the sulfur mass content of less than 1 ppm.

And dividing the methane B into two parts by a flow divider, and respectively sending the two parts into a steam reforming reaction unit II and a dry reforming reaction unit III.

(2) Mixing the first stream of methane with medium-pressure steam C with the flow rate of 120kmol/h, the temperature of 370 ℃ and the pressure of 3MPa, raising the temperature of the mixture to 600 ℃, and then entering a fixed bed reactor of a steam reforming reaction unit II for reforming reaction to obtain steam reforming synthesis gas E. Wherein the molar ratio of methane to water vapor is 1: 3, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content 15 wt% in terms of element, based on the total amount of the catalyst, Al₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed layer is 900 ℃, the pressure in the reactor is 3MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 50000h^-1。

(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 100kmol/h, the temperature of 370 ℃ and the pressure of 2MPa, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content is 10% by weight, calculated as element, based on the total amount of the catalyst; Al)₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed 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.

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 IVThe catalyst for Fischer-Tropsch synthesis (the catalyst prepared in preparation example 1) was contacted to conduct the Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 250 ℃, the pressure in the reactor is 2.0MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly volume space velocity is 4000h^-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 synthetic oil K and gas products; 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 synthetic oil 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 a synthetic oil separation unit VI which are sequentially connected. The specific process flow is as follows.

Preparing coal water slurry C from pulverized coal A (pulverized coal (particle size of 10mm) obtained by pulverizing and sieving solid raw material coal (brown coal produced in inner Mongolia)) in a coal water slurry preparation unit I at a flow rate of 360t/h and water B at a flow rate of 360t/h, and conveying the coal water slurry C into a coal gasification unitII, reacting with oxygen D at 1300 ℃ and 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, wherein the molar ratio of hydrogen to carbon monoxide is 2.1: 1. the purified synthesis gas obtained is conveyed into a Fischer-Tropsch synthesis unit V to carry out Fischer-Tropsch synthesis reaction in a 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 250 ℃, the pressure in the reactor is 2.0MPa, the total amount of the synthetic gas is taken as a reference, and the gas hourly volume space velocity is 4000h^-1. And separating the Fischer-Tropsch reaction product N to obtain synthetic oil K through a synthetic oil separation unit VI, discharging carbon dioxide H and methane G generated by the Fischer-Tropsch synthesis unit V, circulating a part of unreacted synthetic gas (the content is 98 percent based on the total amount of the separated synthetic gas) Y to the Fischer-Tropsch synthesis unit V, and discharging the other part of unreacted synthetic gas serving as purge gas Z out of the system.

Comparative example 2

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

Comparative example 3

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

Example 2

A synthetic oil was prepared in the same manner as in example 1, except that the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst prepared in preparation example 3.

Example 3

A synthetic oil was prepared in the same manner as in example 1, except that the Fischer-Tropsch synthesis catalyst used was the Fischer-Tropsch synthesis catalyst prepared in preparation example 6.

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 850 ℃, the pressure in the reactor is 3MPa, and the gas hour 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, raising the temperature of the mixture to 600 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1.5, 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 700 ℃, 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 prepareThe hydrogen-carbon ratio is 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 12) to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 260 ℃, 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 10000h^-1。

Discharging the synthetic oil 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, raising the temperature of the mixture to 700 ℃, and then entering a fixed bed reactor of a steam reforming reaction unit II for reforming reaction to obtain steam reforming synthesis gas E. Wherein the molar ratio of methane to water vapor is 1: 1, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content 12 wt% in terms of element, based on the total amount of the catalyst, Al₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed is 900 ℃, the pressure in the reactor is 1MPa, and the gas-time volume space velocity based on the total amount of methane and water vapor is 100000h^-1。

(3) And mixing the second strand of methane with carbon dioxide D with the flow rate of 150kmol/h, the temperature of 450 ℃ and the pressure of 3MPa, raising the temperature of the mixture to 700 ℃, and then feeding the mixture into a fixed bed reactor of a dry reforming reaction unit III for reforming reaction to obtain dry integrated syngas F. Wherein the molar ratio of methane to carbon dioxide is 1: 1, the catalyst filled in the reactor is Ni/Al₂O₃(Ni content 12 wt% in terms of element, based on the total amount of the catalyst, Al₂O₃Is alpha-Al₂O₃) The temperature in the catalyst bed layer is 750 ℃, the pressure in the reactor is 2MPa, 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 13) to carry out Fischer-Tropsch synthesis reaction. Wherein the temperature in the reactor is 240 ℃, 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 synthetic oil 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

The same system and method as in example 5 were used to produce synthetic oil, except that the Fischer-Tropsch synthesis catalyst was the Fischer-Tropsch synthesis catalyst prepared in preparation example 16, the temperature in the Fischer-Tropsch synthesis reactor was 240 ℃, the pressure in the Fischer-Tropsch synthesis reactor was 2.5MPa, and the gas hourly volume space velocity was 20000h based on the total amount of the synthesis gas^-1。

TABLE 7

¹: the diesel oil component comprises normal paraffin, isoparaffin and olefin, the isomerate diesel oil refers to the isomerate in the diesel oil component, and the isomerate diesel oil selectivity is calculated by taking the total amount of the diesel oil component as a reference.

TABLE 8

	Water consumption (ton/ton)_{Synthetic oil})	Carbon dioxide emission (ton/ton)_{Synthetic oil})	Energy efficiency (%)
				Example 1	12	0.7	59
Comparative example 1	22	6.4	35
				Comparative example 2	18	2.5	35
Comparative example 3	25	0.7	43
				Example 4	15	1.1	51
Example 5	14	1.0	55
				Example 6	13	0.9	53

Note: the energy efficiency is the sum of the calorific value of the synthetic oil finally discharged from the device and the calorific value of raw materials such as the coal-electric steam catalyst solvent entering the device, namely the calorific value of the synthetic oil obtained/the comprehensive energy consumption required for producing the synthetic oil. 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.

Comparing the embodiment 1 with the comparative example 1, the 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 obviously 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 synthetic oil, 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, and contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at a synthesis oil production reaction temperature to obtain a Fischer-Tropsch synthesis product material flow, wherein the Fischer-Tropsch synthesis catalyst comprises a carrier, and a first metal element and a second metal element which are loaded on the carrier, the first metal element is Fe, the second metal element is one or more selected from IVB group metal elements, optional alkali metal elements and optional alkaline earth metal elements, and the carrier is gamma-Al₂O₃At least part of Fe is lower than the highest oxidation valence of Fe, and in the X-ray photoelectron spectrum 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 Fe, calculated as element₅C₂The ratio of Fe content determined by the spectrum peak of (1) is 10-15: 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 (a) is 60% or more, the fischer-tropsch synthesis catalyst is obtained by subjecting a catalyst precursor to reductive activation, and the method of reductive activation includes:

(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, a first metal element and a second metal element, the first metal element is loaded on the carrier in the form of an oxide, and the valence of Fe in the oxide is the highest oxidation valence of Fe;

(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 an oxide of Fe and/or a precursor of the oxide of Fe and a compound containing a second metal element to obtain a catalyst precursor, wherein the second metal element is loaded on the carrier before the oxide of Fe and the precursor of the oxide of Fe;

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 process as claimed in any one of claims 1 to 3, 6 and 7, wherein the Fischer-Tropsch synthesis reaction temperature for producing synthetic oil is 200 ℃ to 300 ℃.

10. The method as claimed in claim 9, wherein the Fischer-Tropsch synthesis reaction temperature for producing synthetic oil is 220-280 ℃.

11. The process of any one of claims 1-3, 6, and 7, wherein the contacting in S31 is performed in a fixed bed reactor.

12. The process as claimed in claim 11, wherein the gas hourly space velocity of the Fischer-Tropsch synthesis reaction feed in S31 is 2000-30000 h^-1。

13. The process as claimed in claim 12, wherein the gas hourly space velocity of the Fischer-Tropsch synthesis reaction feed in S31 is 4000-^-1。

14. The process of any one of claims 1 to 3, 6 and 7, 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 0.8 to 3MPa, the pressure being in gauge pressure.

15. The process of claim 14, 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.

16. The process of any one of claims 1 to 3, 6 and 7, 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 500 DEG C₂High temperature desorption peak.

19. The process of claim 18, wherein the fischer-tropsch synthesis catalyst has CO₂In the TPD desorption diagram, CO is present in the temperature interval of 320 ℃ and 400 DEG C₂High temperature desorption peak.

20. The process of any one of claims 1, 18 and 19, wherein the fischer-tropsch synthesis catalyst has CO₂In the desorption spectrum of TPD, CO is also present in the temperature range of 100-₂Low temperature desorption peak.

21. The process of claim 20, wherein the fischer-tropsch synthesis catalyst has CO₂In the desorption spectrum of TPD, CO is also present in the temperature range of 140 ℃ and 180 DEG C₂Low temperature desorption peak.

22. The process as claimed in any one of claims 1, 18 and 19, wherein the fischer-tropsch synthesis catalyst has a CO high temperature desorption peak in the CO-TPD desorption profile within the temperature range of 300-600 ℃.

23. The method as claimed in claim 22, wherein the fischer-tropsch synthesis catalyst has a CO high temperature desorption peak in the CO-TPD desorption profile within the temperature range of 400-500 ℃.

24. The process as claimed in any one of claims 1, 18 and 19, wherein the fischer-tropsch synthesis catalyst has a CO-TPD desorption profile in which a CO low temperature desorption peak is also present in the temperature range of 100-200 ℃.

25. The process as claimed in claim 24, wherein the Fischer-Tropsch synthesis catalyst has a CO-TPD desorption spectrum in which a CO low-temperature desorption peak is present in a temperature range of 160-180 ℃.

26. The process of any one of claims 1, 18 and 19, wherein the amount of Fe having a valence below its maximum oxidation state is greater than or equal to 30% by element, based on the total amount of Fe in the fischer-tropsch catalyst.

27. The process of claim 26, wherein the amount of Fe having a valence below its maximum oxidation state is greater than 60% by element, based on the total amount of Fe in the fischer-tropsch catalyst.

28. The method of claim 1, wherein the group IVB metal element is Zr and/or Ti.

29. The method according to claim 1, wherein the alkali metal element is one or two or more of Li, Na, and K.

30. The method of claim 29, wherein the alkali metal element is Li and/or K.

31. The method according to claim 1, wherein the alkaline earth metal element is Mg and/or Ca.

32. The method of claim 31, wherein the alkaline earth element is Mg.

33. The method according to claim 1, wherein the second metal element is a group IVB metal element and one or two or more selected from an alkali metal element and an alkaline earth metal element.

34. The process of any one of claims 1 and 28 to 33, wherein the second metal element is present in an amount of from 0.5 to 10 wt% calculated as element on the total fischer-tropsch synthesis catalyst.

35. The process of claim 34, wherein the second metal element is present in an amount of 1 to 8 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 to 6 wt.% on an elemental basis, based on the total amount of the fischer-tropsch synthesis catalyst.

37. The process of claim 1 wherein the Fe content is from 3 to 30 wt% calculated as element on a fischer-tropsch synthesis catalyst basis.

38. The process of claim 37, wherein the Fe is present in an amount of from 6 to 20 wt% calculated as element on a fischer-tropsch synthesis catalyst basis.

39. The process of claim 38, wherein the Fe content is from 8 to 15 wt% calculated as element on the total fischer-tropsch synthesis catalyst.

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

41. The method as claimed in claim 40, wherein the pre-reduction is carried out at a temperature of 300-500 ℃.

42. The method as claimed in claim 1, wherein the volume space velocity of the first gas is 1000-^-1。

43. The method as claimed in claim 42, wherein the volume space velocity of the first gas is 2000-10000 h on a hydrogen basis^-1。

44. 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.

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

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

47. The method of claim 46, wherein the duration of the pre-reduction is 4-10 hours.

48. 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.

49. The method of claim 48, 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.

50. the method of claim 49, wherein the molar ratio of the inert gas to the hydrocarbon that is gaseous at the reductive activation temperature is from 10 to 30: 1.

51. 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.

52. The method of any one of claims 1 and 48-51, wherein said hydrocarbon that is gaseous at a reductive activation temperature is selected from C₁-C₄Alkane and C₂-C₄One or more than two kinds of olefins.

53. The method according to claim 52, 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.

54. The method as claimed in any one of claims 1 and 48 to 51, wherein the reductive activation is carried out at a temperature of 180-450 ℃.

55. The method of claim 54, wherein the reductive activation is carried out at a temperature of 200-400 ℃.

56. The method of claim 55, wherein the reductive activation is carried out at a temperature of 200-300 ℃.

57. The method as claimed in any one of claims 1 and 48 to 51, wherein the volume space velocity of the second gas is 1000-10000 hours in terms of the hydrocarbon which is gaseous at the reduction activation temperature^-1。

58. The method as claimed in claim 57, wherein the second gas has a volumetric space velocity, based on the hydrocarbon which is gaseous at the reductive activation temperature, of 2000-8000 hours^-1。

59. The process of any one of claims 1 and 48 to 51, wherein the pressure in the reactor in which the reductive activation is carried out is 0 to 3MPa gauge.

60. A process as claimed in claim 59, in which the pressure in the reactor at which the reductive activation is effected is from 0.1 to 1MPa by gauge.

61. The method of any one of claims 1 and 48-51, wherein the duration of reductive activation is from 1 to 20 hours.

62. The method of claim 61, wherein the duration of said reductive activation is from 2 to 8 hours.

63. The method of claim 62, wherein the duration of reductive activation is 4-6 hours.

64. The method according to any one of claims 1, 42, 43, and 48 to 51, 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.

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

66. The method as claimed in claim 1, wherein the calcination is carried out at a temperature of 300-900 ℃ and the duration of the calcination is 1-12 hours.

67. The method of claim 66, wherein the method of supporting the second metal element on the alumina comprises: and roasting the compound alumina loaded with the second metal element.

68. The method as claimed in claim 67, wherein in the method of supporting the second metal element on the alumina, the calcination is carried out at a temperature of 300-900 ℃ and the duration of the calcination is 1-8 hours.

69. 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.

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

71. The method of claim 70, wherein the feed gas is one or more selected from shale gas, coal bed gas, natural gas and coke oven gas.

72. The method of claim 70, wherein the feed gas is one or more selected from shale gas, coal bed gas, natural gas and refinery gas.

73. The process of claim 70, wherein methane is separated from said feed gas using a cryogenic condensation process.

74. The method of any of claims 1 and 69-73, wherein the weight ratio of methane employed in S11 to methane employed in S21 is 1: 0.5-2.5.

75. A synthetic oil production system comprises a steam reforming reaction unit, a dry reforming reaction unit, a synthetic gas mixing unit, a Fischer-Tropsch synthesis reaction product separation unit and a circulation unit,

the Fischer-Tropsch synthesis reaction unit is provided with a Fischer-Tropsch synthesis reactor and is used for contacting the Fischer-Tropsch synthesis reaction feed with a Fischer-Tropsch synthesis catalyst at the reaction temperature of producing the synthetic oil to obtain a Fischer-Tropsch synthesis product flow containing the synthetic oil, wherein the Fischer-Tropsch synthesis catalyst is the Fischer-Tropsch synthesis catalyst described in any one of claims 1 and 18-68;

76. The system of claim 75, 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 into the steam reforming reaction unit and the dry reforming reaction unit, respectively.

77. The system of claim 76, wherein the feed gas separation unit is provided with a cryogenic condenser for condensing the feed gas to separate out methane in the feed gas.

78. The system of any one of claims 75-77, wherein the Fischer-Tropsch synthesis reaction unit further comprises a reductive activation subunit for reductive activation of the Fischer-Tropsch synthesis catalyst precursor.

79. The system of claim 78, 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 the reduction activation reaction.

80. The system of claim 79, wherein the hydrocarbon that is gaseous at the reduction activation temperature is one 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.

81. The system of claim 80, 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.

82. The system of claim 81, 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.

83. The system according to claim 79, 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.

84. The system of claim 83, wherein the inert gas in the first gas and the second gas is the same or different, each being nitrogen and/or argon.

85. The system of claim 79, 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.

86. The system of any one of claims 75-77, wherein the Fischer-Tropsch synthesis reactor is a fixed bed reactor.