CN114437841A

CN114437841A - Method and device for co-producing natural gas, LPG and naphtha by using Fischer-Tropsch synthesis tail gas

Info

Publication number: CN114437841A
Application number: CN202210133450.2A
Authority: CN
Inventors: 贺树民; 高军虎; 郝栩; 董根全; 杨勇; 李永旺
Original assignee: Zhongke Synthetic Oil Technology Co Ltd
Current assignee: Zhongke Synthetic Oil Technology Co Ltd
Priority date: 2022-02-14
Filing date: 2022-02-14
Publication date: 2022-05-06

Abstract

The invention relates to a method for coproducing Synthetic Natural Gas (SNG), LPG and naphtha by utilizing Fischer-Tropsch synthetic tail gas and a device for implementing the method, wherein CO in the Fischer-Tropsch synthetic tail gas₂The content is adjusted to a specific range to obtain raw gas, and the raw gas is subjected to low-temperature separation treatment, methanation treatment and low-temperature oil washing treatment, so that synthetic natural gas is obtained and LPG and naphtha are co-produced. The method can realize the efficient utilization of various useful components in the Fischer-Tropsch synthesis tail gas, and improve the economic benefit of coal-to-liquids production enterprises.

Description

Method and device for co-producing natural gas, LPG and naphtha by using Fischer-Tropsch synthesis tail gas

Technical Field

The invention belongs to the field of Fischer-Tropsch synthesis, and relates to a Fischer-Tropsch synthesis tail gas treatment method. In particular, the invention relates to a catalyst which will have a specific CO₂The Fischer-Tropsch synthesis tail gas with the content is used as the raw material gasA process for the co-production of Synthetic Natural Gas (SNG), LPG and naphtha and a plant for carrying out the process.

Background

Natural gas is considered as a clean, convenient and safe high-quality energy source, and the main component of the natural gas is methane (CH)₄) The energy-saving device is widely applied to industries such as power generation, chemical industry, urban gas, automobile fuel and the like, and is one of main clean energy sources in the world. However, China is a country rich in coal, poor in oil and little in gas, and the natural gas reserves only account for 1.3% of the total reserves in the world. Moreover, the distribution of coal is extremely unbalanced, and the coal is mainly concentrated in inner Mongolia, Xinjiang, Shaanxi, Shanxi and the like, and the transportation is extremely inconvenient. The coal-based natural gas technology is used for converting high-carbon coal resources into low-carbon, hydrogen-rich and long-distance transportation-convenient substitute natural gas, and the heat value of the substitute natural gas is more than or equal to 34.6MJ/m³The heat energy utilization efficiency can reach 52.57%, but the coal consumption is in a considerable proportion by direct conversion through combustion, the heat energy efficiency is low, the resource waste is serious, and the environmental pollution is serious.

At present, the efficient and clean utilization of coal is realized by utilizing the Fischer-Tropsch synthesis reaction. The Fischer-Tropsch synthesis reaction refers to the steps of washing, converting and removing CO from the raw gas obtained after coal gasification₂And the synthesis gas (CO and H) obtained after desulfurization treatment₂) In a Fischer-Tropsch synthesis reactor, under the action of a Fischer-Tropsch synthesis catalyst, so as to produce crude products such as gaseous hydrocarbon, liquid hydrocarbon, synthetic wax and the like. The crude products are subjected to processes such as hydrofining, hydrocracking, catalytic cracking, pour point depression reaction and the like to produce high-quality products such as naphtha, gasoline, diesel oil, paraffin and the like. In the Fischer-Tropsch synthesis reaction, a large amount of Fischer-Tropsch synthesis tail gas is also generated, and if the tail gas is not utilized, the resource waste is caused.

At present, much research is carried out on recycling of fischer-tropsch synthesis tail gas, for example, a fischer-tropsch synthesis tail gas recycling system process is disclosed in chinese patent application No. cn201910952440, so as to obtain LNG products, propane, ethane, hydrogen products, carbon monoxide and the like. Chinese patent application CN201820639142.6 discloses a device for purifying fischer-tropsch synthesis tail gas and recovering light hydrocarbons, and LPG products are recovered. Chinese patent application CN201610342193.8 discloses a process and a system for preparing liquefied natural gas from fischer-tropsch synthesis tail gas, which are used to liquefy and separate liquefied natural gas, hydrogen-rich gas, carbon monoxide-rich gas, etc. from fischer-tropsch synthesis tail gas. The above process is concerned with recovering the components present in the tail gas. In addition, some methods not only recover components in the tail gas, but also recycle and convert part of the components in the tail gas, for example, chinese patent application CN201610695802.8 discloses a comprehensive utilization process of fischer-tropsch synthesis tail gas by selective catalytic oxidation conversion, in which hydrogen is separated and extracted from the tail gas, and low carbon hydrocarbons in the circulating tail gas and an oxidant undergo selective catalytic oxidation reaction to be converted into hydrogen and carbon monoxide under the action of a catalyst, so as to recycle and utilize the hydrogen and the carbon monoxide.

The inventor finds that in the actual production process, Fischer-Tropsch tail gas discharged by Fischer-Tropsch synthesis reaction contains a large amount of H₂、CO、CH₄Lower hydrocarbon (C)₈The following hydrocarbons), CO₂And N₂And the like. After the components of hydrogen, carbon monoxide, LNG and LPG are recovered, if the residual tail gas components are directly discharged or combusted, the environment is seriously polluted, and energy waste is caused, which is greatly incompatible with the expectation of high-efficiency clean utilization of energy at present. Therefore, the inventor finds that methanation treatment and the like of the Fischer-Tropsch synthesis tail gas is an efficient way for realizing clean utilization of energy sources based on control of specific composition of the Fischer-Tropsch synthesis tail gas.

Disclosure of Invention

In view of the above technical problems, the present inventors have found through research that CO is used therein₂The Fischer-Tropsch synthesis tail gas with the content adjusted to a specific range is used as a raw material gas to carry out low-temperature separation, methanation, low-temperature oil washing and other treatments, so that the synthetic natural gas SNG, LPG and naphtha can be co-produced.

In one aspect, the present invention relates to a process for co-producing natural gas, LPG and naphtha using fischer-tropsch synthesis tail gas, wherein the process comprises:

(1) will be provided withCO in Fischer-Tropsch synthesis tail gas₂The molar percentage of the raw material gas is adjusted to be 8 to 10 percent to obtain the raw material gas;

(2) carrying out low-temperature separation on the raw material gas to obtain absorption dry gas and absorption rich liquid, carrying out heat exchange on the absorption dry gas and the raw material gas to prepare a methanation raw material, wherein the methanation raw material is rich in H₂、CO、CO₂、N₂And C₃ ^-A hydrocarbon;

(3) carrying out methanation treatment on the methanation raw material to obtain synthetic natural gas SNG and condensate;

(4) carrying out flash evaporation on the absorption rich solution to obtain flash evaporation gas and flash evaporation liquid, and carrying out extraction separation on the flash evaporation liquid to obtain the CO-rich liquid₂And C₃ ⁺A mixture of hydrocarbons, enriching said CO₂And C₃ ⁺And carrying out low-temperature oil washing separation on the mixture of the hydrocarbons to obtain oil-washed dry gas, LPG and naphtha, and returning the oil-washed dry gas to be mixed with the absorption dry gas to obtain the methanation raw material.

In another aspect, the present invention relates to an apparatus for carrying out the above method, wherein the apparatus comprises:

the low-temperature separation unit comprises a first heat exchanger, a quencher, an absorption desorption tower, a flash tank and an extraction separation tower which are sequentially connected in a fluid communication manner;

a methanation unit coupled in fluid communication to the first heat exchanger and comprising a multi-stage methanation reactor, a second heat exchanger, a cooler, and a separator coupled in fluid communication in that order;

a low-temperature oil washing unit which is connected to the absorption desorption tower in a fluid communication manner and comprises an absorption tower, a stabilization tower and a regeneration tower which are sequentially connected in a fluid communication manner.

The method controls CO in the Fischer-Tropsch synthesis tail gas₂The raw material gas obtained with the content in a specific range is converted, so that the tail gas can be utilized as much as possibleAll useful components simplify the tail gas conversion and utilization process, not only can produce the synthetic natural gas SNG which is needed urgently, but also can obtain the CO-rich gas which is obtained after the flash evaporation liquid is extracted and separated₂And C₃ ⁺The mixture of products can be separated by oil washing to obtain LPG and naphtha.

Exemplary aspects of the present invention may be illustrated by the contents described in the following paragraphs, but the scope of the present invention is not limited thereto:

1. a process for co-producing natural gas, LPG and naphtha from fischer-tropsch synthesis tail gas, wherein the process comprises:

(1) CO in Fischer-Tropsch synthesis tail gas₂The molar percentage of the raw material gas is adjusted to 8 to 10 percent to obtain the raw material gas;

2. The method of paragraph 1 wherein in step (1) the CO in the Fischer-Tropsch synthesis tail gas is adjusted by₂The molar percentage of (A): removing part or all of CO from part of Fischer-Tropsch synthesis tail gas₂And obtaining decarbonized tail gas, and mixing the decarbonized tail gas with the other part of Fischer-Tropsch synthesis tail gas.

3. The method of paragraph 2 wherein in step (1) the tail is compared to the Fischer-Tropsch synthesisCO in gas₂The total amount of the catalyst is that more than 98 vol% of CO is removed from 20 vol% -25 vol% of Fischer-Tropsch synthesis tail gas₂And obtaining decarburization tail gas.

4. The method as claimed in any one of paragraphs 1 to 3, wherein, in step (2), the feed gas is first heat-exchanged with the absorption dry gas to be cooled to 0-10 ℃ and then cryogenically cooled to-60 ℃ to 0 ℃, and then the cryogenic separation is performed under the action of cold absorbent methanol.

5. The method as in any one of paragraphs 1 to 4, wherein, in the step (2), the raw material gas is subjected to the low-temperature separation in an absorption desorption tower, the number of theoretical plates of the absorption desorption tower is 5-50, the tower top temperature is-60 ℃ to 50 ℃, the tower bottom temperature is-50 ℃ to 50 ℃, and the operation pressure is 0-6 MPa.

6. The method as in any one of paragraphs 1-5, wherein, in step (3), the methanation treatment is performed after the methanation raw material is heated to 150-300 ℃.

7. The process of any of paragraphs 1-6, wherein, in step (3), the methanation feedstock is subjected to the methanation process under conditions such that: the method comprises the following steps of carrying out a first-stage reaction at 150-300 ℃ and 0-8.0 MPa, carrying out a second-stage reaction at 150-350 ℃ and 0-8.0 MPa, carrying out a third-stage reaction at 150-300 ℃ and 0-8.0 MPa, and carrying out a fourth-stage reaction at 150-300 ℃ and 0-8.0 MPa.

8. The process of paragraph 7, wherein, in step (3), the methanation feedstock is mixed with a primary return gas prior to the primary reaction.

9. The method of paragraph 7 or 8, wherein in step (3), the steam produced by the four-stage reaction is cooled to 30 ℃ to 50 ℃ to obtain the synthetic natural gas SNG.

10. A process as claimed in any one of paragraphs 1 to 9, wherein in step (4) the flash gas is compressed by a compressor and returned to step (1) and mixed with the fischer-tropsch synthesis tail gas as the feed gas.

11. The method of any of paragraphs 1-10, wherein in step (4), the flashing is carried out under the following conditions: the temperature is-40 ℃ to 50 ℃, and the pressure is 0MPa to 2 MPa.

12. The method as claimed in any one of paragraphs 1-11, wherein in step (4), said extraction separation is performed on said flashed liquid using water as extractant, thereby recovering methanol in the flashed liquid to obtain methanol aqueous solution, and separating said CO-enriched aqueous solution₂And C₃ ⁺A mixture of hydrocarbons.

13. The method of any of paragraphs 1-12, wherein in step (4), the extractive separation is carried out under the following conditions: the temperature is-40 ℃ to 50 ℃, the pressure is 0MPa to 2MPa, and the number of tower plates is 2 to 15.

14. The method of any of paragraphs 1-13, wherein the method further comprises: and carrying out desorption separation on the recovered methanol aqueous solution to obtain the extractant water and the cold absorbent methanol.

15. The method of paragraph 14 wherein said recovered aqueous methanol solution is subjected to desorption separation in an absorbent recovery column.

16. The method of paragraph 15, wherein the number of theoretical plates of the absorber recovery column is 5 to 100, the overhead temperature is 0 to 100 ℃, the kettle temperature is 50 to 150 ℃, and the operating pressure is 0 to 5 MPa.

17. The method as in any of paragraphs 14-16, wherein the cold absorbent separated by desorption is returned partially or completely to the absorption and desorption column.

18. The method as in any one of paragraphs 14-16, wherein the separated extractant is desorbed and returned to step (4) for recycling.

19. The method of any of paragraphs 1-18, wherein, in step (4), the cryogenic oil wash separation comprises the operations of: applying naphtha as absorbent to said CO rich fraction in an absorber₂And C₃₊Treating the mixture of hydrocarbons to obtain dry gas at the top of the tower and liquid at the bottom of the absorption tower; enabling the tower kettle liquid of the absorption tower to enter a stabilizing tower for treatment to obtain LPG and tower kettle liquid of the stabilizing tower; and (3) allowing the liquid in the tower bottom of the stabilizer to enter a regeneration tower for treatment to obtain oil-washed naphtha and absorbent naphtha.

20. The process of paragraph 19, wherein the overhead dry gas is recycled to step (2) and mixed with the absorption dry gas and oil wash dry gas as methanation feedstock.

21. The method of paragraph 19 or 20 wherein the absorbent naphtha is cooled to-20 to 20 ℃ and returned to the absorber for recycling.

22. The method of any one of paragraphs 19 to 21, wherein the number of theoretical plates of the absorber is 5 to 50, the overhead temperature is-60 ℃ to 50 ℃, the kettle temperature is-50 ℃ to 50 ℃, and the operating pressure is 0 to 6 MPa.

23. The method of any of paragraphs 19-22, wherein the number of theoretical plates in the stabilizer is 5-150, the overhead temperature is 0-100 ℃, the kettle temperature is 0-200 ℃ and the operating pressure is 0-4 MPa.

24. The method of any of paragraphs 19-23, wherein the number of theoretical plates in the regeneration column is 5-100, the overhead temperature is 0-100 ℃, the kettle temperature is 0-200 ℃, and the operating pressure is 0-3 MPa.

25. An apparatus for performing the method of any of paragraphs 1-24, wherein the apparatus comprises:

26. The apparatus of paragraph 25, wherein the apparatus further comprises a decarbonization unit coupled in fluid communication upstream of the first heat exchanger.

27. The apparatus of paragraph 25 or 26, wherein the apparatus further comprises an absorbent recovery column coupled in fluid communication to the absorption desorber and the extraction separation column.

28. The apparatus of any of paragraphs 25-27, wherein said multi-stage methanation reactor comprises a one-stage methanation reactor, a two-stage methanation reactor, and a four-stage methanation reactor.

29. The apparatus of paragraph 28, wherein the second heat exchanger comprises a heat exchanger coupled in fluid communication downstream of the primary, secondary and four-stage methanation reactors, respectively.

The method provided by the invention is not only beneficial to realizing comprehensive utilization of Fischer-Tropsch synthesis tail gas with relatively low cost, but also can improve the comprehensive utilization rate of coal energy efficiency and improve the economic benefit of coal-to-liquid production enterprises. By co-producing LPG, naphtha and SNG, various useful components in Fischer-Tropsch synthesis tail gas can be utilized to the maximum extent, and the products can relieve the energy structure of rich coal, little oil and gas in China, thereby bringing great convenience to the production life.

Drawings

FIG. 1 is an exemplary flow diagram of a process for co-producing natural gas, LPG and naphtha using Fischer-Tropsch synthesis tail gas.

Fig. 2 is an exemplary process schematic of feed gas entering cryogenic separation unit U1 for treatment.

FIG. 3 is a schematic diagram of an exemplary methanation process.

FIG. 4 is a schematic diagram of an exemplary low temperature oil wash process.

In the figure, S1 raw gas; s2 CO₂And C₃₊A mixture of hydrocarbons; s3 absorbing dry gas; s4 SNG; s5 LPG; s6 naphtha; s6' absorbent naphtha; s7 condensing; an S8 extractant; s9 methanol; s10 oil washes dry gas. A U1 cryogenic separation unit; a U2 methanation unit; u3 low temperature oil wash unit. 101 a first heat exchanger; 202. 203, 205, 211 second heat exchanger; 102 a chiller; 212 a cooler; 103 absorption desorption tower; 104 flash evaporationA tank; 105. 214, 215 compressor; 106 extraction separation tower; 107 absorbent recovery column; 207. 208, 213 separators; a stage 201 methanation reactor; 204 two-stage methanation reactor; 206 three sections of methanation reactors; 210 four-stage methanation reactor. 301 an absorption column; 302 a stabilizer column; 303 regenerating the column.

Detailed Description

While the illustrative embodiments of the invention are described below, it will be understood by those skilled in the art that the scope of the invention is not limited thereto, but various modifications and combinations can be made to the aspects of the invention without departing from the spirit and scope of the invention.

In the present invention, the term "C_n ^-"denotes a hydrocarbon having n or less carbon atoms, the term" C_n ⁺"represents a hydrocarbon having n or more carbon atoms, and n may be an integer of 1 or more, for example, a natural integer of 1, 2, 3, 4, 5, or the like. E.g. C₃ ⁺Represents a hydrocarbon having 3 or more carbon atoms.

In the present invention, unless otherwise specified, the terms "part" and "a portion" are used interchangeably to refer to some of the objects modified by the term, and may, for example, represent any value in the range of from greater than 0% to less than 100% relative to the total amount of the objects modified by the term.

Herein, methanation refers to H in syngas₂And CO is reacted chemically at certain temperature and pressure under the action of catalyst to produce methane. Methanation is a strongly exothermic, volume-reducing, reversible reaction and carbon evolution is possible during the reaction. The main chemical reactions in the methanation process are as shown in the following formulas (1), (2) and (3).

Methanation not only can utilize synthesis gas, but also can convert CO₂Conversion to useful fuels (CH)₄) And the method is recycled and has important commercial value. The temperature rise of 70-72 ℃ can be generated when every 1mol of CO is converted, and every 1mol of CO is converted₂A temperature rise of 60 ℃ can be generated.

CO+3H₂＝CH₄+H₂O-206.0 kJ/mol formula (1)

CO₂+4H₂＝CH₄+2H₂O-165.0 kJ/mol formula (2)

CO+H₂O＝H₂+CO₂-41.0 kJ/mol of formula (3)

The inventor finds that part of CO is removed from Fischer-Tropsch synthesis tail gas₂The obtained raw gas is subjected to methanation treatment, so that various useful components in the Fischer-Tropsch synthesis tail gas can be utilized to the maximum extent, and the synthetic natural gas SNG, LPG and naphtha are co-produced. Compared with the traditional technology for separating useful components from Fischer-Tropsch synthesis tail gas, the method disclosed by the invention can convert the most components into the components with economic values through simple treatment, so that the method is more environment-friendly and efficient, the economic benefits of coal-to-oil enterprises can be improved as much as possible, and the efficient clean utilization of resources and the maximization of economic benefits are realized.

In one embodiment, the invention provides a method for co-producing natural gas, LPG and naphtha from fischer-tropsch synthesis tail gas, wherein the method comprises:

In some preferred embodiments, in step (1), the CO in the Fischer-Tropsch synthesis tail gas is adjusted by₂The molar percentage of (A): removing part or all of CO from part of Fischer-Tropsch synthesis tail gas₂And obtaining decarbonized tail gas, and mixing the decarbonized tail gas with the other part of Fischer-Tropsch synthesis tail gas.

In some preferred embodiments, in step (1), the amount of CO in the Fischer-Tropsch synthesis tail gas is relative to the amount of CO in the Fischer-Tropsch synthesis tail gas₂The total amount of the catalyst is that more than 98 vol% of CO is removed from 20 vol% -25 vol% of Fischer-Tropsch synthesis tail gas₂And obtaining decarburization tail gas.

In some preferred embodiments, in step (2), the raw material gas is first subjected to heat exchange with the absorption dry gas to be cooled to 0-10 ℃ and then subjected to cryogenic cooling to-60 ℃ -0 ℃ (preferably-20 ℃ -0 ℃), and then subjected to the cryogenic separation under the action of cold absorbent methanol.

In some preferred embodiments, in the step (2), the raw material gas is subjected to the low-temperature separation in an absorption desorption tower, the number of theoretical plates of the absorption desorption tower is 5-50, the tower top temperature is-60-50 ℃, the tower bottom temperature is-50 ℃, and the operating pressure is 0-6 MPa.

In this context, the absorption rich solution absorbs C in the feed gas for the cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And (4) obtaining the final product.

In some preferred embodiments, in the step (3), the methanation treatment is performed after the methanation raw material is heated to 150-300 ℃.

In some preferred embodiments, in step (3), the methanation feedstock is subjected to the methanation process under the following conditions: performing a first-stage reaction at 150-300 ℃ and 0-8.0 MPa, performing a second-stage reaction at 150-350 ℃ and 0-8.0 MPa, performing a third-stage reaction at 150-300 ℃ and 0-8.0 MPa, and performing a fourth-stage reaction at 150-300 ℃ and 0-8.0 MPa.

In some preferred embodiments, in step (3), the methanation feedstock is mixed with a primary return gas prior to the primary reaction.

In some preferred embodiments, in step (3), the steam produced by the four-stage reaction is cooled to 30 to 50 ℃ (e.g., 40 ℃) to obtain the synthetic natural gas SNG.

In this context, as shown in fig. 3, for example, after the methanation raw material is mixed with the first-stage return gas and heated, the mixture enters the first-stage methanation reactor to perform a first-stage reaction, the generated tail gas is subjected to two-stage heat exchange (for example, the operating temperature of the two-stage heat exchanger is 0 to 400 ℃, the operating pressure is 0 to 5 MPa; for example, the operating temperature is 200 to 350 ℃, and the operating pressure is 2 to 4MPa), part of the tail gas is pressurized by a compressor (for example, the operating temperature of the compressor is 0 to 300 ℃, the operating pressure is 0 to 7 MPa; for example, the operating temperature is 150 to 280 ℃, the operating pressure is 2 to 4MPa) and is used as a raw material adjusting gas (i.e., "first-stage return gas") to return to the inlet of the first-stage methanation reactor, the rest of the tail gas enters the second-stage methanation reactor to perform a second-stage reaction, and the generated second-stage reaction tail gas is cooled by heat exchange, the reaction product enters a three-stage methanation reactor for three-stage reaction, after cooling and separating the generated three-stage reaction tail gas, a part of the reaction tail gas is further cooled and separated, a part of condensate (mainly water) is separated out, the condensate is pressurized by a compressor (for example, the operating temperature of the compressor is 0-300 ℃, the operating pressure is 0-7 MPa; for example, the operating temperature is 150-280 ℃, the operating pressure is 2-4 MPa) and then returns to the inlet of the two-stage methanation reactor, and the rest three-stage reaction tail gas enters a four-stage methanation reactor for four-stage reaction to carry out four-stage reaction on trace CO and CO in the tail gas₂Continuing the conversion until the generated four-stage reaction tail gas is basically free of H₂CO and CO₂And is mainly CH₄And (4) and the like. And after heat exchange is carried out on the four-stage reaction tail gas, further cooling and separating out condensate to obtain qualified synthetic natural gas.

In some preferred embodiments, in step (4), the flash gas is compressed by a compressor and then returned to step (1) and mixed with the fischer-tropsch synthesis tail gas to serve as the feed gas.

In some preferred embodiments, in step (4), the flashing is performed under the following conditions: the temperature is-40 ℃ to 50 ℃, and the pressure is 0MPa to 2 MPa.

In some preferred embodiments, in step (4), said extraction separation is performed on said flash liquid using water as an extractant, thereby recovering methanol in the flash liquid to obtain an aqueous methanol solution, and separating said CO-enriched water₂And C₃ ⁺A mixture of hydrocarbons.

In some preferred embodiments, in step (4), the extractive separation is carried out under the following conditions: the temperature is-40 ℃ to 50 ℃, the pressure is 0MPa to 2MPa, and the number of tower plates is 2 to 15.

In some preferred embodiments, the method further comprises: and carrying out desorption separation on the recovered methanol aqueous solution to obtain the extractant water and the cold absorbent methanol. In a further preferred embodiment, the recovered aqueous methanol solution is subjected to desorption separation in an absorbent recovery column. In a further preferred embodiment, the number of theoretical plates of the absorbent recovery tower is 5-100, the tower top temperature is 0-100 ℃, the tower kettle temperature is 50-150 ℃, and the operating pressure is 0-5 MPa. In a further preferred embodiment, part or all of the cold absorbent obtained by desorption separation is returned to the absorption desorption column. Preferably, part of the cold absorbent obtained by desorption and separation returns to the absorption and desorption tower, and the rest of the cold absorbent is mixed with the Fischer-Tropsch synthesis tail gas, so as to avoid freezing of water condensate in the heat exchange and cooling process of the feed gas. In some preferred embodiments, the extractant obtained by desorption and separation is returned to the step (4) for recycling.

In some preferred embodiments, in step (4), the low-temperature oil-wash separation comprises the following operations: applying naphtha as absorbent to said CO rich fraction in an absorber₂And C₃₊Treating the mixture of hydrocarbons to obtain dry gas at the top of the tower and liquid at the bottom of the absorption tower; enabling the tower kettle liquid of the absorption tower to enter a stabilizing tower for treatment to obtain LPG and tower kettle liquid of the stabilizing tower; and (3) allowing the liquid in the tower bottom of the stabilizer to enter a regeneration tower for treatment to obtain oil-washed naphtha and absorbent naphtha. In a further preferred embodiment, the overhead dry gas is passed to a columnAnd (3) recycling the obtained product to the step (2) and mixing the obtained product with the absorption dry gas and the oil-washed dry gas to obtain a methanation raw material. In a further preferred embodiment, the absorbent naphtha is cooled to-20 to 20 ℃ and then returned to the absorption tower for recycling.

Herein, the fraction of the oil-washed naphtha ranges from 0 to 150 ℃, preferably from 40 to 120 ℃. Herein, an LPG component, wherein C is produced at the top of the stabilizer column₃And C₄The hydrocarbon content is greater than 87%, preferably greater than 98%.

In some preferred embodiments, the number of theoretical plates of the absorption tower is 5-50, the tower top temperature is-60-50 ℃, the tower kettle temperature is-50 ℃, and the operation pressure is 0-6 MPa.

In some preferred embodiments, the number of theoretical plates of the stabilizer is 5-150, the temperature of the top of the stabilizer is 0-100 ℃, the temperature of the bottom of the stabilizer is 0-200 ℃, and the operating pressure is 0-4 MPa.

In some preferred embodiments, the number of theoretical plates of the regeneration tower is 5-100, the tower top temperature is 0-100 ℃, the tower kettle temperature is 0-200 ℃, and the operating pressure is 0-3 MPa.

In one embodiment, the present invention relates to an apparatus for carrying out the above method, wherein the apparatus comprises:

In some preferred embodiments, the apparatus further comprises a decarbonization unit coupled in fluid communication upstream of the first heat exchanger.

In some preferred embodiments, the apparatus further comprises an absorbent recovery column coupled in fluid communication to the absorption desorber and the extractive separation column.

Cryogenic separation unit

Herein, the first heat exchanger is one heat exchanger, or a plurality of heat exchangers having the same or different operating temperatures from each other. The quencher is one quencher or a plurality of quenchers with the same or different operating temperatures from each other.

In some preferred embodiments, the number of theoretical plates of the absorption desorption tower is 5-50, the tower top temperature is-60-50 ℃, the tower kettle temperature is-50-100 ℃, and the operating pressure is 0-6 MPa. In a further preferred embodiment, the number of theoretical plates of the absorption and desorption tower is 10-30, the tower top temperature is-40-20 ℃, the tower kettle temperature is-40 ℃, and the operating pressure is 2-4 MPa.

In some preferred embodiments, the flash tank is operated at a temperature of-40 ℃ to 50 ℃ and an operating pressure of 0 to 2 MPa. In a further preferred embodiment, the operating temperature of the flash tank is-30 ℃ to 20 ℃ and the operating pressure is 0MPa to 1 MPa.

In some preferred embodiments, the extraction separation tower has an operating temperature of-40 to 50 ℃, an operating pressure of 0 to 2MPa, and a number of plates of 2 to 15. In a further preferred embodiment, the operating temperature of the extraction separation tower is-10 ℃ to 40 ℃, the operating pressure is 0 to 1MPa, and the number of the tower plates is 8 to 12.

In some preferred embodiments, the number of theoretical plates of the absorbent recovery tower is 5-100, the tower top temperature is 0-100 ℃, the tower kettle temperature is 50-150 ℃, and the operating pressure is 0-5 MPa. In a further preferred embodiment, the number of theoretical plates of the absorbent recovery tower is 20-80, the tower top temperature is 30-80 ℃, the tower kettle temperature is 80-140 ℃, and the operating pressure is 0-1 MPa.

Methanation unit

In some preferred embodiments, the multi-stage methanation reactor comprises a one-stage methanation reactor, a two-stage methanation reactor, and a four-stage methanation reactor.

In some preferred embodiments, the operation temperature of the primary methanation reactor is 150-300 ℃, and the operation pressure is 0-8.0 MPa. In a further preferred embodiment, the operation temperature of the primary methanation reactor is 200-300 ℃, and the operation pressure is 2-4 MPa.

In some preferred embodiments, the second heat exchanger comprises a heat exchanger fluidly connected downstream of the primary, secondary, and fourth methanation reactors, respectively. Wherein, the methanation reactor connected with the first section in a fluid communication way can be a two-stage heat exchanger.

In some preferred embodiments, the operating temperature of the two-stage methanation reactor is 150-350 ℃, and the operating pressure is 0-8.0 MPa. In some preferred embodiments, the three-stage methanation reactor and the four-stage methanation reactor are operated at the temperature of 150-300 ℃ and the pressure of 0-8.0 MPa. In a further preferred embodiment, the three-stage methanation reactor and the four-stage methanation reactor have the operation temperature of 200-300 ℃ and the operation pressure of 2-4 MPa.

Low-temperature oil washing unit

In some preferred embodiments, the number of theoretical plates of the absorption tower is 5-50, the tower top temperature is-60-50 ℃, the tower kettle temperature is-50 ℃, and the operation pressure is 0-6 MPa. In a further preferred embodiment, the number of theoretical plates of the absorption tower is 20-40, the tower top temperature is-40-20 ℃, the tower kettle temperature is-40-20 ℃, and the operation pressure is 2-5 MPa.

In some preferred embodiments, the number of theoretical plates of the stabilizer is 5-150, the temperature of the top of the stabilizer is 0-100 ℃, the temperature of the bottom of the stabilizer is 0-200 ℃, and the operating pressure is 0-4 MPa. In a further preferred embodiment, the number of theoretical plates of the stabilizer is 20-80, the temperature of the top of the stabilizer is 20-90 ℃, the temperature of the bottom of the stabilizer is 80-180 ℃, and the operating pressure is 2-3 MPa.

In some preferred embodiments, the number of theoretical plates of the regeneration tower is 5-100, the tower top temperature is 0-100 ℃, the tower kettle temperature is 0-200 ℃, and the operating pressure is 0-3 MPa. In a further preferred embodiment, the number of theoretical plates of the regeneration tower is 20 to 60, the tower top temperature is 40 to 90 ℃, the tower kettle temperature is 60 to 160 ℃, and the operating pressure is 0.2 to 1 MPa.

Examples

The present invention will be described in further detail with reference to examples, but the present invention is not limited to these examples.

The measured compositions of the fischer-tropsch synthesis tail gas and the decarbonated tail gas used in the following examples are shown in table 1 below.

TABLE 1 Fischer-Tropsch Synthesis Tail gas and decarbonized Tail gas compositions in mole percent

Name (R)	H₂	N₂	CO	CH₄	CO₂	C₂H₆	C₂H₄	C₃+
									Synthetic tail gas/mol%	61.89	3.99	7.12	10.26	12	1.15	0.28	3.31
Decarburization tail gas/mol%	69.93	4.51	8.05	11.59	0.56	1.30	0.32	3.74

Example 1

The Fischer-Tropsch synthesis tail gas is divided into two streams, and one stream of the tail gas is decarbonized by a decarbonization unit with 20 vol% and then is mixed with the remaining Fischer-Tropsch tail gas with 80 vol% to obtain the feed gas. The raw material gas and the absorption dry gas are subjected to heat exchange in a heat exchanger to be cooled to 8 ℃, and then are subjected to deep cooling to-20 ℃ in a quencher to enter an absorption desorption tower (the temperature at the top of the tower is-10 ℃, the temperature at the bottom of the tower is-12 ℃, the operation pressure is 3.0MPa, and the number of theoretical plates is 8), and the C in the raw material gas is removed under the action of a cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And after the generated absorption dry gas is reheated in a heat exchanger, the absorption dry gas enters a downstream methanation unit as a part of the methanation raw material.

The methanation raw material is mixed with the first-stage return gas and heated to 280 ℃, and then the mixture enters a first-stage methanation reactorReacting at 280 ℃, wherein the reaction pressure is 3.4MPa, the outlet temperature of the reactor is about 600 ℃, the generated first-stage reaction tail gas generates 3.2MPa medium-pressure steam through two-stage heat exchange, the temperature is reduced to 320 ℃, part of the first-stage reaction tail gas is compressed to 3.5MPa by a compressor and then is used as raw material gas for regulating the gas and returning to the inlet, and the rest of the first-stage reaction tail gas enters a second-stage methanation reactor for reacting at 320 ℃, and the reaction pressure is 3.2 MPa; the outlet temperature of the second-stage methanation reactor is 600 ℃, the generated second-stage reaction tail gas generates 3.2MPa medium-pressure steam through heat exchange, the steam is cooled to 280 ℃ and enters the third-stage methanation reactor to react at 280 ℃, and the reaction pressure is 3.0 MPa; the outlet temperature of the three-section methanation reactor is about 450 ℃, the generated three-section reaction tail gas is cooled to 250 ℃, one part of the three-section reaction tail gas is further cooled to 100 ℃ for separation, partial condensate (mainly water) is separated out, the condensate is pressurized by a compressor and then returns to the two-section inlet, the rest three-section reaction tail gas enters the four-section methanation reactor for reaction at 250 ℃, the reaction pressure is 2.8MPa, and trace CO and CO in the tail gas are subjected to reaction₂Continuing to carry out conversion; the outlet temperature of the four-section methanation reactor is 300 ℃, and the generated four-section reaction tail gas (basically does not contain H)₂CO and CO₂And is mainly CH₄Equal component, CH₄The content is more than 87vol percent) to generate 1.0MPa low-pressure steam through heat exchange, the low-pressure steam is further cooled to about 40 ℃ in a cooler, and condensate is separated out through a separator to obtain qualified synthetic natural gas.

Absorbing rich liquid is flashed out in a flash tank and contains H₂、CO、N₂、CO₂、CH₄And a small amount of C₃ ^-Compressing the flash gas of gaseous hydrocarbons (mainly ethane and ethylene) by a compressor, returning to the feed gas stream, introducing the flash liquid obtained at-38 deg.C under 0.8MPa into an extraction separation tower, recovering methanol as cold absorbent to obtain a mixture of the extractant and the cold absorbent, and separating to obtain a C-rich gas₃ ⁺Hydrocarbons and CO₂As a feedstock for the low temperature oil wash. The operating temperature of the extraction separation tower is 20 ℃, the pressure is 0.8MPa, and the number of tower plates is 10.

The low-temperature oil-washing raw material is pressurized to 5MPa by a compressor and then enters the bottom of an absorption tower, and naphtha with the temperature of minus 20 ℃ is adopted as an absorbent at the top of the absorption tower. The temperature of the top of the absorption tower is-5 ℃, the temperature of the bottom of the absorption tower is-20 ℃, the operating pressure is 4.8MPa, the number of tower plates is 40, the absorbed dry gas at the top of the absorption tower returns to be mixed with the absorbed dry gas and the oil-washed dry gas to be used as a part of the methanation raw material, and the liquid at the bottom of the absorption tower enters a stabilizing tower; the temperature of the top of the stabilizer tower is 70 ℃, the temperature of the bottom of the stabilizer tower is 180 ℃, the operating pressure is 1.5MPa, the number of tower plates is 25, LPG is discharged from the top of the stabilizer tower, and liquid in the bottom of the stabilizer tower enters a regeneration tower; the temperature of the top of the regeneration tower is 80 ℃, the temperature of the bottom of the regeneration tower is 150 ℃, the operating pressure is 0.5MPa, the number of tower plates is 30, oil-washed naphtha is discharged from the top of the regeneration tower, absorbent naphtha is taken as the bottom of the regeneration tower, and the absorbent naphtha is cooled to-20 ℃ and then returned to the absorption tower for circulation.

And (3) carrying out desorption separation on the mixture of the extractant and the cold absorbent obtained by recycling in an absorbent recycling tower, wherein the tower top temperature of the absorbent recycling tower is 80 ℃, the tower bottom temperature is 120 ℃, the operating pressure is 1.0MPa, the number of tower plates is 20, one part of the cold absorbent obtained by recycling returns to an absorption desorption tower, and the rest part of the cold absorbent is mixed with Fischer-Tropsch synthesis tail gas so as to avoid freezing of a hydraulic liquid in the heat exchange and cooling process of the feed gas. The extractant returns to the extraction separation tower for recycling.

Example 2

The Fischer-Tropsch synthesis tail gas is divided into two streams, one stream is decarbonized by a decarbonization unit with 21 vol% and then is mixed with the remaining Fischer-Tropsch tail gas with 79 vol% to obtain the feed gas. The raw material gas and the absorption dry gas are subjected to heat exchange in a heat exchanger to be cooled to 8 ℃, and then are subjected to deep cooling to-20 ℃ in a quencher to enter an absorption desorption tower (the temperature at the top of the tower is-10 ℃, the temperature at the bottom of the tower is-12 ℃, the operation pressure is 3.0MPa, and the number of theoretical plates is 8), and the C in the raw material gas is removed under the action of a cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And after the generated absorption dry gas is reheated in a heat exchanger, the absorption dry gas enters a downstream methanation unit as a part of the methanation raw material.

Mixing the methanation raw material with the first-stage return gas, heating to 280 ℃, then feeding the mixture into a first-stage methanation reactor for reaction at the temperature of 280 ℃, wherein the reaction pressure is 3.4MPa, and the outlet temperature of the reactor is about 600 DEG CThe generated first-stage reaction tail gas generates 3.2MPa medium-pressure steam through two-stage heat exchange, the temperature is reduced to 320 ℃, part of the first-stage reaction tail gas is compressed to 3.5MPa by a compressor and then is returned to an inlet as a raw material gas-adjusting gas, and the rest of the first-stage reaction tail gas enters a two-stage methanation reactor to react at 320 ℃, wherein the reaction pressure is 3.2 MPa; the outlet temperature of the second-stage methanation reactor is 600 ℃, the generated second-stage reaction tail gas generates 3.2MPa medium-pressure steam through heat exchange, the steam is cooled to 280 ℃ and enters the third-stage methanation reactor to react at 280 ℃, and the reaction pressure is 3.0 MPa; the outlet temperature of the three-stage methanation reactor is about 450 ℃, the generated three-stage reaction tail gas is cooled to 250 ℃, one part of the three-stage reaction tail gas is further cooled to 100 ℃ for separation, partial condensate (mainly water) is separated out and returns to the two-stage inlet after being pressurized by a compressor, the rest three-stage reaction tail gas enters the four-stage methanation reactor to react at 250 ℃, the reaction pressure is 2.8MPa, and trace CO and CO in the tail gas are subjected to reaction₂Continuing to carry out conversion; the outlet temperature of the four-section methanation reactor is 300 ℃, the generated four-section reaction tail gas generates low-pressure steam of 1.0MPa through heat exchange, the low-pressure steam is further cooled to about 40 ℃ in a cooler, and condensate is separated out through a separator to obtain qualified synthetic natural gas.

The low-temperature oil-washing raw material is pressurized to 5MPa by a compressor and then enters the bottom of an absorption tower, and naphtha with the temperature of minus 20 ℃ is adopted as an absorbent at the top of the absorption tower. The temperature of the top of the absorption tower is-5 ℃, the temperature of the bottom of the absorption tower is-20 ℃, the operating pressure is 4.8MPa, the number of tower plates is 40, the absorbed dry gas at the top of the absorption tower returns to be mixed with the absorbed dry gas and the oil-washed dry gas to be used as a part of the methanation raw material, and the liquid at the bottom of the absorption tower enters a stabilizing tower; the temperature of the top of the stabilizing tower is 70 ℃, the temperature of the bottom of the stabilizing tower is 180 ℃, the operating pressure is 1.5MPa, the number of tower plates is 25, LPG is discharged from the top of the stabilizing tower, and liquid in the bottom of the stabilizing tower enters a regeneration tower; the temperature of the top of the regeneration tower is 80 ℃, the temperature of the bottom of the regeneration tower is 150 ℃, the operating pressure is 0.5MPa, the number of tower plates is 30, oil-washed naphtha is discharged from the top of the regeneration tower, absorbent naphtha is taken as the bottom of the regeneration tower, and the absorbent naphtha is cooled to-20 ℃ and then returned to the absorption tower for circulation.

Example 3

The Fischer-Tropsch synthesis tail gas is divided into two streams, one stream is decarbonized by a decarbonization unit with 22 vol% and then is mixed with the remaining Fischer-Tropsch tail gas with 78 vol% to obtain the feed gas. The raw material gas and the absorption dry gas exchange heat in a heat exchanger to be cooled to 8 ℃, are subjected to deep cooling to-20 ℃ in a quencher to enter an absorption desorption tower (the temperature of the top of the tower is-10 ℃, the temperature of the bottom of the tower is-12 ℃, the operating pressure is 3.0MPa, and the number of tower plates is 8), and the C in the raw material gas is removed under the action of cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And after generated absorption dry gas is reheated in a heat exchanger, the absorption dry gas enters a downstream methanation unit as a part of the methanation raw material.

Mixing the methanation raw material and the first-stage return gas, heating to 280 ℃, then feeding the mixture into a first-stage methanation reactor to react at 280 ℃, wherein the reaction pressure is 3.4MPa, the temperature of the outlet of the reactor is about 600 ℃, the generated first-stage reaction tail gas generates 3.2MPa medium-pressure steam through two-stage heat exchange, the temperature is reduced to 320 ℃, part of the first-stage reaction tail gas is pressurized to 3.5MPa by a compressor and then is used as raw material gas-regulating gas to return to the inlet, and the rest of the first-stage reaction tail gas enters a second-stage methane reactorThe reaction is carried out in the chemical reactor at 320 ℃, and the reaction pressure is 3.2 MPa; the outlet temperature of the second-stage methanation reactor is 600 ℃, the generated second-stage reaction tail gas generates 3.2MPa medium-pressure steam through heat exchange, the steam is cooled to 280 ℃ and enters the third-stage methanation reactor to react at 280 ℃, and the reaction pressure is 3.0 MPa; the outlet temperature of the three-section methanation reactor is about 450 ℃, the generated three-section reaction tail gas is cooled to 250 ℃, one part of the three-section reaction tail gas is further cooled to 100 ℃ for separation, partial condensate (mainly water) is separated out, the condensate is pressurized by a compressor and then returns to the two-section inlet, the rest three-section reaction tail gas enters the four-section methanation reactor for reaction at 250 ℃, the reaction pressure is 2.8MPa, and trace CO and CO in the tail gas are subjected to reaction₂Continuing to carry out conversion; the outlet temperature of the four-section methanation reactor is 300 ℃, the generated four-section reaction tail gas generates low-pressure steam of 1.0MPa through heat exchange, the low-pressure steam is further cooled to about 40 ℃ in a cooler, and condensate is separated out through a separator to obtain qualified synthetic natural gas.

Absorbing rich liquid is flashed out in a flash tank and contains H₂、CO、N₂、CO₂、CH₄And a small amount of C₃ ^-Compressing the flash gas of gaseous hydrocarbons (mainly ethane and ethylene) by a compressor, returning to the feed gas stream, introducing the flash liquid obtained at 38 ℃ and 0.8MPa into an extraction separation tower, recovering the cold absorbent methanol by using water as an extractant to obtain a mixture of the extractant and the cold absorbent, and separating the mixture to obtain the C-rich gas₃ ⁺Hydrocarbons and CO₂As a feedstock for the low temperature oil wash. The operating temperature of the extraction separation tower is 20 ℃, the pressure is 0.8MPa, and the number of tower plates is 10.

Pressurizing the low-temperature oil-washing raw material to 5MPa by a compressor, then feeding the raw material into the bottom of an absorption tower, and adopting naphtha with the temperature of minus 20 ℃ as an absorbent at the tower top. The temperature of the top of the absorption tower is-5 ℃, the temperature of the bottom of the absorption tower is-20 ℃, the operating pressure is 4.8MPa, the number of tower plates is 40, the absorbed dry gas at the top of the absorption tower returns to be mixed with the absorbed dry gas and the oil-washed dry gas to be used as a part of the methanation raw material, and the liquid at the bottom of the absorption tower enters a stabilizing tower; the temperature of the top of the stabilizing tower is 70 ℃, the temperature of the bottom of the stabilizing tower is 180 ℃, the operating pressure is 1.5MPa, the number of tower plates is 25, LPG is discharged from the top of the stabilizing tower, and liquid in the bottom of the stabilizing tower enters a regeneration tower; the temperature of the top of the regeneration tower is 80 ℃, the temperature of the bottom of the regeneration tower is 150 ℃, the operating pressure is 0.5MPa, the number of tower plates is 30, oil-washed naphtha is discharged from the top of the regeneration tower, absorbent naphtha is taken as the bottom of the regeneration tower, and the absorbent naphtha is cooled to-20 ℃ and then returned to the absorption tower for circulation.

Example 4

The Fischer-Tropsch synthesis tail gas is divided into two streams, and one stream of 23 vol% decarbonization unit is decarbonized and then is mixed with the remaining 77 vol% Fischer-Tropsch tail gas to obtain the feed gas. The raw material gas and the absorption dry gas exchange heat in a heat exchanger to be cooled to 8 ℃, are subjected to deep cooling to-20 ℃ in a quencher to enter an absorption desorption tower (the temperature of the top of the tower is-10 ℃, the temperature of the bottom of the tower is-12 ℃, the operating pressure is 3.0MPa, and the number of tower plates is 8), and the C in the raw material gas is removed under the action of cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And after generated absorption dry gas is reheated in a heat exchanger, the absorption dry gas enters a downstream methanation unit as a part of the methanation raw material.

Mixing the methanation raw material and the first-stage return gas, heating to 280 ℃, then feeding the mixture into a first-stage methanation reactor to react at 280 ℃, wherein the reaction pressure is 3.4MPa, the temperature of the outlet of the reactor is about 600 ℃, the generated first-stage reaction tail gas generates 3.2MPa medium-pressure steam through two-stage heat exchange, the temperature is reduced to 320 ℃, part of the first-stage reaction tail gas is pressurized to 3.5MPa by a compressor and then is used as raw material gas-regulating gas to be returned to the inlet, and the rest of the first-stage reaction tail gas enters a second-stage methanation reactor to react at 320 ℃, wherein the reaction pressure is 3.2 MPa; the outlet temperature of the second-stage methanation reactor is 600 ℃, the generated second-stage reaction tail gas generates 3.2MPa medium pressure steam through heat exchange, the steam is cooled to 280 ℃ and enters the third-stage methanation reactor to be at 280 DEG CThe reaction is carried out under the reaction pressure of 3.0 MPa; the outlet temperature of the three-section methanation reactor is about 450 ℃, the generated three-section reaction tail gas is cooled to 250 ℃, one part of the three-section reaction tail gas is further cooled to 100 ℃ for separation, partial condensate (mainly water) is separated out, the condensate is pressurized by a compressor and then returns to the two-section inlet, the rest three-section reaction tail gas enters the four-section methanation reactor for reaction at 250 ℃, the reaction pressure is 2.8MPa, and trace CO and CO in the tail gas are subjected to reaction₂Continuing to carry out conversion; the outlet temperature of the four-section methanation reactor is 300 ℃, the generated four-section reaction tail gas generates low-pressure steam of 1.0MPa through heat exchange, the low-pressure steam is further cooled to about 40 ℃ in a cooler, and condensate is separated out through a separator to obtain qualified synthetic natural gas.

Absorbing rich liquid is flashed out in a flash tank and contains H₂、CO、N₂、CO₂、CH₄And a small amount of C_3-Compressing the flash gas of gaseous hydrocarbons (mainly ethane and ethylene) by a compressor, returning to the feed gas stream, introducing the flash liquid obtained at-38 deg.C under 0.8MPa into an extraction separation tower, recovering methanol as cold absorbent to obtain a mixture of the extractant and the cold absorbent, and separating to obtain a C-rich gas₃₊Hydrocarbons and CO₂As a feedstock for the low temperature oil wash. The operating temperature of the extraction separation tower is 20 ℃, the pressure is 0.8MPa, and the number of tower plates is 10.

Example 5

The Fischer-Tropsch synthesis tail gas is divided into two streams, one stream of the tail gas is decarbonized by a decarbonization unit with 24 vol% and then is mixed with the remaining Fischer-Tropsch tail gas with 76 vol% to obtain the feed gas. The raw material gas and the absorption dry gas exchange heat in a heat exchanger to be cooled to 8 ℃, are subjected to deep cooling to-20 ℃ in a quencher to enter an absorption desorption tower (the temperature at the top of the tower is-10 ℃, the temperature at the bottom of the tower is-12 ℃, the operating pressure is 3.0MPa, and the number of tower plates is 8), and the C in the raw material gas is removed under the action of cold absorbent methanol₃ ⁺Hydrocarbons and part of CO₂And after generated absorption dry gas is reheated in a heat exchanger, the absorption dry gas enters a downstream methanation unit as a part of the methanation raw material.

Mixing the methanation raw material and the first-stage return gas, heating to 280 ℃, then feeding the mixture into a first-stage methanation reactor to react at 280 ℃, wherein the reaction pressure is 3.4MPa, the temperature of the outlet of the reactor is about 600 ℃, the generated first-stage reaction tail gas generates 3.2MPa medium-pressure steam through two-stage heat exchange, the temperature is reduced to 320 ℃, part of the first-stage reaction tail gas is pressurized to 3.5MPa by a compressor and then returns to the inlet as the raw material gas-regulating gas, and the rest of the first-stage reaction tail gas enters a second-stage methanation reactor to react at 320 ℃, wherein the reaction pressure is 3.2 MPa; the outlet temperature of the second-stage methanation reactor is 600 ℃, the generated second-stage reaction tail gas generates 3.2MPa medium-pressure steam through heat exchange, the steam is cooled to 280 ℃ and enters the third-stage methanation reactor to react at 280 ℃, and the reaction pressure is 3.0 MPa; the outlet temperature of the three-stage methanation reactor is about 450 ℃, the generated three-stage reaction tail gas is cooled to 250 ℃, one part of the three-stage reaction tail gas is further cooled to 100 ℃ for separation, partial condensate (mainly water) is separated out, and the condensate is returned after being pressurized by a compressorReturning to the second-stage inlet, and allowing the rest three-stage reaction tail gas to enter a four-stage methanation reactor for reaction at 250 ℃, wherein the reaction pressure is 2.8MPa, and trace amounts of CO and CO in the tail gas are treated₂Continuing to carry out conversion; the outlet temperature of the four-section methanation reactor is 300 ℃, the generated four-section reaction tail gas generates low-pressure steam of 1.0MPa through heat exchange, the low-pressure steam is further cooled to about 40 ℃ in a cooler, and condensate is separated out through a separator to obtain qualified synthetic natural gas.

The low-temperature oil washing raw gas is pressurized to 5MPa by a compressor and then enters the bottom of an absorption tower, and naphtha with the temperature of minus 20 ℃ is adopted as an absorbent at the top of the absorption tower. The temperature of the top of the absorption tower is-5 ℃, the temperature of the bottom of the absorption tower is-20 ℃, the operating pressure is 4.8MPa, the number of tower plates is 40, the absorbed dry gas at the top of the absorption tower returns to be mixed with the absorbed dry gas and the oil-washed dry gas to be used as a part of the methanation raw material, and the liquid at the bottom of the absorption tower enters a stabilizing tower; the temperature of the top of the stabilizing tower is 70 ℃, the temperature of the bottom of the stabilizing tower is 180 ℃, the operating pressure is 1.5MPa, the number of tower plates is 25, LPG is discharged from the top of the stabilizing tower, and liquid in the bottom of the stabilizing tower enters a regeneration tower; the temperature of the top of the regeneration tower is 80 ℃, the temperature of the bottom of the regeneration tower is 150 ℃, the operating pressure is 0.5MPa, the number of tower plates is 30, oil-washed naphtha is discharged from the top of the regeneration tower, absorbent naphtha is taken as the bottom of the regeneration tower, and the absorbent naphtha is cooled to-20 ℃ and then returned to the absorption tower for circulation.

TABLE 2 composition of the methanation feed in the examples

Name (R)	H₂	N₂	CO	CH₄	CO₂	C₂H₆	C₂H₄	C₃+	Total mol%
										Example 1	63.30	4.08	7.28	10.49	10.00	1.18	0.29	3.39	100.00
Example 2	63.58	4.10	7.31	10.54	9.60	1.18	0.29	3.40	100.00
										Example 3	63.90	4.12	7.35	10.59	9.14	1.19	0.29	3.42	100.00
Example 4	64.30	4.15	7.40	10.66	8.57	1.19	0.29	3.44	100.00
										Example 5	64.70	4.17	7.44	10.73	8.00	1.20	0.29	3.46	100.00

Table 3 composition of the synthetic natural gas obtained in each example

Name (R)	H₂	N₂	CO	CH₄	C₂H₆	Total mol%
							Example 1	3.40	11.89	0.00	80.47	4.24	100.00
Example 2	8.22	11.40	0.00	76.29	4.08	100.00
							Example 3	13.27	10.90	0.00	71.91	3.92	100.00
Example 4	18.98	10.32	0.00	66.96	3.74	100.00
							Example 5	24.12	9.80	0.00	62.51	3.57	100.00

The inventors of the present invention adopted the use of a catalyst having CO in a specific range₂The Fischer-Tropsch synthesis tail gas with the content is used as the raw material gas to carry out methanation treatment, LPG and naphtha can be separated, and the separated CH₄、H₂、CO、CO₂Synthesizing SNG. As can be seen from tables 2 and 3, the feed gas is an ideal methanation feed material, and various useful components in Fischer-Tropsch synthesis tail gas can be converted and utilized to the maximum extent by the method, so that synthetic natural gas SNG, LPG and naphtha with higher economic benefit can be obtained by co-production. The method is not only beneficial to realizing the comprehensive utilization of the Fischer-Tropsch synthesis tail gas with relatively low cost, but also can improve the comprehensive utilization rate of the coal energy efficiency and improve the economic benefit of coal-to-liquid production enterprises.

The above description is only a preferred embodiment of the present invention, and the scope of the present invention is not limited to the above embodiment, and it will be apparent to those skilled in the art that various modifications and substitutions can be made in the present invention, for example, direct methanation of the decarbonized fischer-tropsch synthesis tail gas. Therefore, any modification, equivalent replacement, improvement and the like made under the principle and spirit of the present invention shall fall within the protection scope of the present invention.

Claims

(4) carrying out flash evaporation on the absorption rich solutionObtaining flash gas and flash liquid, and extracting and separating the flash liquid to obtain the CO-rich liquid₂And C₃ ⁺A mixture of hydrocarbons, enriching said CO₂And C₃ ⁺And carrying out low-temperature oil washing separation on the mixture of the hydrocarbons to obtain oil-washed dry gas, LPG and naphtha, and returning the oil-washed dry gas to be mixed with the absorption dry gas to obtain the methanation raw material.

2. The method of claim 1, wherein in step (1), the CO in the Fischer-Tropsch synthesis tail gas is adjusted by₂The molar percentage of (A): removing part or all of CO from part of Fischer-Tropsch synthesis tail gas₂Obtaining decarbonized tail gas, and mixing the decarbonized tail gas with the other part of Fischer-Tropsch synthesis tail gas;

preferably, relative to the CO in the Fischer-Tropsch synthesis tail gas₂The total amount of the catalyst is that more than 98 vol% of CO is removed from 20 vol% -25 vol% of Fischer-Tropsch synthesis tail gas₂And obtaining decarburization tail gas.

3. The method according to claim 1 or 2, wherein in the step (2), the raw material gas and the absorption dry gas are subjected to heat exchange to be cooled to 0-10 ℃ and then are subjected to cryogenic cooling to-60-0 ℃, and then the cryogenic separation is carried out under the action of cold absorbent methanol;

preferably, in the step (2), the feed gas is subjected to low-temperature separation in an absorption desorption tower, the number of theoretical plates of the absorption desorption tower is 5-50, the temperature of the top of the absorption desorption tower is-60-50 ℃, the temperature of the bottom of the absorption desorption tower is-50 ℃, and the operating pressure is 0-6 MPa.

4. The method according to any one of claims 1 to 3, wherein in the step (3), the methanation treatment is carried out after the methanation raw material is heated to 150-300 ℃;

preferably, the methanation raw material is subjected to the methanation treatment under the following conditions: performing a first-stage reaction at 150-300 ℃ and 0-8.0 MPa, performing a second-stage reaction at 150-350 ℃ and 0-8.0 MPa, performing a third-stage reaction at 150-300 ℃ and 0-8.0 MPa, and performing a fourth-stage reaction at 150-300 ℃ and 0-8.0 MPa;

preferably, the methanation raw material is mixed with a primary return gas and then the primary reaction is carried out;

preferably, steam generated by the four-stage reaction is cooled to 30-50 ℃ to obtain the synthetic natural gas SNG.

5. The process of any one of claims 1 to 4 wherein in step (4) the flash gas is compressed by a compressor and returned to step (1) and mixed with the Fischer-Tropsch synthesis tail gas as the feed gas;

preferably, the flashing is carried out under the following conditions: the temperature is-40 ℃ to 50 ℃, and the pressure is 0MPa to 2 MPa.

6. A process according to any one of claims 1 to 5, wherein in step (4) said flash liquid is subjected to said extractive separation using water as extractant, whereby methanol in the flash liquid is recovered as aqueous methanol solution and said CO enriched CO is separated₂And C₃ ⁺A mixture of hydrocarbons;

preferably, the extractive separation is carried out under the following conditions: the temperature is-40 ℃ to 50 ℃, the pressure is 0MPa to 2MPa, and the number of tower plates is 2 to 15.

7. The method of any one of claims 1-6, wherein the method further comprises: carrying out desorption separation on the recovered methanol aqueous solution to obtain the extractant water and the cold absorbent methanol;

preferably, the recovered methanol aqueous solution is subjected to desorption separation in an absorbent recovery tower;

preferably, the number of theoretical plates of the absorbent recovery tower is 5-100, the temperature of the top of the tower is 0-100 ℃, the temperature of a tower kettle is 50-150 ℃, and the operating pressure is 0-5 MPa;

preferably, part or all of the cold absorbent obtained by desorption and separation is returned to the absorption and desorption tower;

preferably, the extractant obtained by desorption and separation is returned to the step (4) for recycling.

8. The method of any one of claims 1-7, wherein in step (4), the low temperature oil wash separation comprises the operations of: applying naphtha as absorbent to said CO rich fraction in an absorber₂And C₃₊Treating the mixture of hydrocarbons to obtain dry gas at the top of the tower and liquid at the bottom of the absorption tower; enabling the tower kettle liquid of the absorption tower to enter a stabilizing tower for treatment to obtain LPG and tower kettle liquid of the stabilizing tower; enabling the liquid in the tower bottom of the stabilizer to enter a regeneration tower for treatment to obtain oil-washed naphtha and absorbent naphtha;

preferably, the tower top dry gas is recycled to the step (2) and is mixed with the absorption dry gas and the oil washing dry gas to be used as a methanation raw material;

preferably, the absorbent naphtha is cooled to-20 ℃ and then returned to the absorption tower for recycling;

preferably, the number of theoretical plates of the absorption tower is 5-50, the temperature of the tower top is-60-50 ℃, the temperature of the tower kettle is-50 ℃, and the operating pressure is 0-6 MPa;

preferably, the number of theoretical plates of the stabilizing tower is 5-150, the temperature of the top of the tower is 0-100 ℃, the temperature of the bottom of the tower is 0-200 ℃, and the operating pressure is 0-4 MPa;

preferably, the number of theoretical plates of the regeneration tower is 5-100, the temperature of the top of the tower is 0-100 ℃, the temperature of the bottom of the tower is 0-200 ℃, and the operating pressure is 0-3 MPa.

9. An apparatus for carrying out the method of any one of claims 1-8, wherein the apparatus comprises:

10. The apparatus of claim 9, further comprising a decarbonization unit coupled in fluid communication upstream of the first heat exchanger;

preferably, the apparatus further comprises an absorbent recovery column connected in fluid communication to the absorption desorber and the extractive separation column;

preferably, the multistage methanation reactor comprises a first-stage methanation reactor, a second-stage methanation reactor and a fourth-stage methanation reactor;

preferably, the second heat exchanger comprises heat exchangers respectively connected in fluid communication downstream of the first-stage methanation reactor, the second-stage methanation reactor and the fourth-stage methanation reactor.