A purification process of co, a production process and equipment of coal-to-ethanol

Abstract

This invention relates to the field of coal chemical technology and discloses a CO purification process, a coal-to-ethanol production process and equipment. The CO purification process includes: (1) passing a raw gas containing CO, H2 and acidic gases into a membrane pre-separation unit to separate CO-rich non-permeable gas and H2-rich permeable gas; (2) the CO-rich non-permeable gas enters a purification device to remove acidic gases; (3) separating the CO-rich non-permeable gas after removing acidic gases to obtain CO, wherein the separation method includes one or more of membrane separation, cryogenic separation, and pressure swing adsorption separation. By setting up a membrane pre-separation unit, this invention pre-concentrates CO, effectively stabilizing the gas volume and composition entering the subsequent purification unit. Even with changes in the type of raw coal and fluctuations in the CO content in the syngas, it can still achieve stable control of the purity and flow rate of the CO product, making it suitable for applications such as coal-to-ethanol where a stable CO supply is required.

Description

Technical Field

[0001] The present invention relates to the technical field of coal chemical industry, and particularly relates to a process for purifying CO, a production process and equipment for producing ethanol from coal. Background Art

[0002] Carbon monoxide, as an important basic chemical raw material, is widely used in the synthesis of methanol, acetic acid, ethanol and various carbonylation reaction processes. For different downstream uses and product index requirements, the existing industrial process routes for purifying carbon monoxide mainly include cryogenic separation, membrane separation and pressure swing adsorption, etc.

[0003] Among them, the cryogenic separation and pressure swing adsorption processes are applicable to the application scenarios with relatively high requirements for the purity of carbon monoxide products, and can achieve the stable production of high-purity carbon monoxide, but usually the process is relatively complex, and the equipment investment and operation energy consumption are relatively high; the membrane separation process has the characteristics of simple process, flexible start-stop, relatively low investment, etc., and is applicable to the occasions with relatively low requirements for the purity of carbon monoxide.

[0004] Taking the process of producing ethanol from syngas as an example, the raw material gas usually has relatively low requirements for the purity of carbon monoxide, and at the same time, a certain proportion of hydrogen components need to be retained to meet the reaction requirements of the carbonylation catalyst in a hydrogen-containing environment. Generally, the requirements for the raw material gas in the ethanol synthesis process are: the volume fraction of carbon monoxide is not less than about 94.5%, the volume fraction of hydrogen is not more than 5%, and the rest are a small amount of inert or weakly inert impurity components such as nitrogen, argon and methane. The content of the above impurities in the raw material syngas is limited and usually does not have a significant adverse impact on the ethanol synthesis reaction. Therefore, in such application scenarios, there is usually no need to carry out deep separation on them.

[0005] Based on the above characteristics, at present, the carbon monoxide purification unit配套建设 with the ethanol production device from syngas mostly adopts the membrane separation process to meet the product requirements. On the one hand, the membrane separation process can retain a certain proportion of hydrogen components on the non-permeate gas side while ensuring the purity of carbon monoxide, avoiding the complex operation of completely removing hydrogen by using the cryogenic separation or pressure swing adsorption process and then additionally supplementing pure hydrogen; on the other hand, only when the content of inert components such as nitrogen, argon or methane in the raw material gas significantly increases, will the cryogenic separation or pressure swing adsorption process be considered.

[0006] In coal gasification processes, regardless of whether dry pulverized coal entrained gasification technology or coal-water slurry entrained gasification technology is used, the resulting crude syngas typically contains hydrogen, carbon monoxide, carbon dioxide, hydrogen sulfide, water vapor, and small amounts of nitrogen, argon, and methane. After non-shift gas heat recovery, the water vapor in the crude syngas exists in a saturated state at room temperature. Subsequently, it passes through a low-temperature methanol washing unit to remove acidic components such as carbon dioxide and hydrogen sulfide, as well as trace amounts of water, ultimately yielding purified syngas with hydrogen and carbon monoxide as the main components, and containing small amounts of nitrogen, argon, and methane.

[0007] However, in actual industrial operation, the type of raw coal used in coal gasification units often varies. When the type of raw coal changes, the ratio of hydrogen to carbon monoxide in the output syngas usually fluctuates significantly, and the two often exhibit an inverse relationship, meaning that when the hydrogen content increases, the carbon monoxide content decreases accordingly. Especially when the carbon monoxide content decreases significantly, the original carbon monoxide purification unit, regardless of whether it uses membrane separation, pressure swing adsorption, or cryogenic separation processes, will be greatly affected, easily leading to the flow rate or purity of the carbon monoxide product failing to meet the stable operation requirements of downstream units.

[0008] Taking the scenario where the hydrogen content in syngas increases and the carbon monoxide content decreases as an example, if membrane separation technology is continued to be used to extract carbon monoxide, more membrane modules and increased permeate compressor capacity are typically required. If pressure swing adsorption (PSA) technology is used, the adsorption tower size needs to be increased, the adsorbent loading increased, and the number or capacity of carbon monoxide compressors increased. If cryogenic separation technology is used, liquid nitrogen often needs to be added, the heat exchanger size increased, and the number of stages in the carbon monoxide compressor unit increased to accommodate the lower carbon monoxide inlet partial pressure. Furthermore, to meet the extraction load requirements, the processing capacity of the shift converter, purification unit, and gas separation unit needs to be increased accordingly, allowing more non-shifted gas to enter the carbon monoxide purification unit, thereby further increasing system investment and operating energy consumption.

[0009] Although load fluctuations and coal type variations are typically considered during the plant design phase, the design load range of existing plants is generally limited to approximately 60% to 110% of the rated load. When the design margin is large, the unit energy consumption under normal operating conditions is high, resulting in poor economic efficiency. Conversely, when the design margin is small, it is difficult to effectively cope with fluctuations in syngas composition caused by changes in raw coal type, which can easily affect the stable supply of carbon monoxide products.

[0010] Therefore, existing carbon monoxide purification technologies generally suffer from problems such as difficulty in equipment adjustment, significant increase in investment and energy consumption, and slow response speed when dealing with fluctuations in the composition of raw syngas, especially significant changes in carbon monoxide content. It is difficult to achieve a stable and flexible gas supply to downstream plants that use carbon monoxide as raw material while ensuring economic efficiency. Summary of the Invention

[0011] This invention provides a CO purification process, a coal-to-ethanol production process and equipment, to solve the problem in the prior art that when the composition of syngas fluctuates due to changes in the type of raw coal, especially when the CO content decreases, the carbon monoxide purification unit is unable to achieve stable control of the purity and supply of carbon monoxide products without significantly increasing equipment investment and operating energy consumption.

[0012] In a first aspect, the present invention provides a CO purification process, comprising the following steps: (1) The feed gas containing CO, H2 and acidic gas is passed into the membrane pre-separation unit to separate CO-rich non-permeable gas and H2-rich permeable gas. (2) CO-rich non-permeable gas enters the purification device to remove acidic gases; (3) CO is obtained by separating the CO-rich non-permeable gas after removing acidic gases, wherein the separation method includes one or more of membrane separation process, cryogenic separation process, and pressure swing adsorption separation process.

[0013] In one embodiment of the present invention, in step (1), the H2-rich permeate pressure in the membrane pre-separation unit is 0.1 MPaG to 3 MPaG, and the CO-rich non-permeate pressure is 1.5 MPaG to 8.5 MPaG; for example, the H2-rich permeate pressure in the membrane pre-separation unit is 0.1 MPaG, 0.15 MPaG, 1 MPaG, 2 MPaG, and 3 MPaG; and the CO-rich non-permeate pressure is 1.5 MPaG, 2.5 MPaG, 3.5 MPaG, 4.5 MPaG, 5.5 MPaG, 6.5 MPaG, 7.5 MPaG, and 8.5 MPaG.

[0014] In one embodiment of the present invention, in step (2), the purification device includes a shift gas purification device and a non-shift gas purification device. The CO-rich non-permeable gas enters the non-shift gas purification device to remove acidic gas, and the H2-rich permeable gas enters the shift gas purification device to remove acidic gas.

[0015] In one embodiment of the present invention, in step (2), the H2-rich permeate gas is pressurized to the pressure of the shift gas pipeline and then combined with the shift gas and sent to the shift gas purification device to remove acidic gas.

[0016] In one embodiment of the present invention, in step (1), the raw material gas includes non-shift gas; And / or, in step (3), the separation method is selected from membrane separation process.

[0017] In a second aspect, the present invention provides a CO purification device, including a non-shift gas system, a membrane pre-separation unit, a purification device for removing acidic components from the gas, and a gas separation device. The outlet of the non-conversion gas system is connected to the inlet of the membrane pre-separation unit, the outlet of the membrane pre-separation unit is connected to the inlet of the purification device, the outlet of the purification device is connected to the inlet of the gas separation device, and the outlet of the gas separation device is used to output CO.

[0018] The purification device is used to remove acidic components such as carbon dioxide and hydrogen sulfide from non-shifted gas and shifted gas to obtain purified gas that meets the requirements of subsequent carbon monoxide purification, hydrogen purification, and methanol and ethanol synthesis. The purification device can be a low-temperature methanol washing device.

[0019] Thirdly, the present invention provides a coal-to-ethanol production process, comprising the following steps: S1 involves reacting raw coal with oxygen in a coal gasification reaction to form crude syngas; S2 The crude synthesis gas is fed into the shift gas system and the non-shift gas system respectively to obtain shift gas and non-shift gas; S3 Non-shift gas is introduced into the membrane pre-separation unit to separate CO-rich non-permeable gas and H2-rich permeable gas; After the S4 H2-rich permeate gas is mixed with the shift gas and the acid gas is removed, part of it is sent to the hydrogen extraction unit to purify H2, and the other part is sent to the methanol unit. The H2 purified by the pressure swing adsorption unit is sent to the ethanol unit, and the tail gas generated is sent to the methanol unit. After removing acidic gases from the CO-rich non-permeable gas, CO is purified and sent to the ethanol unit. The resulting tail gas is sent to the methanol unit, and the methanol produced in the methanol unit is sent to the ethanol unit. In the ethanol unit, methanol, CO, and H2 react to produce ethanol.

[0020] In one embodiment of the present invention, in step S4, the H2-rich permeate gas is pressurized and mixed with the shift gas before being sent into the purification device to remove acidic gases. In one embodiment of the present invention, in step S4, the purified H2 is pressurized and then sent to an ethanol unit.

[0021] In one embodiment of the present invention, in step S4, the CO-rich non-permeable gas is sent to a gas separation device for separation and purification after the acidic gas is removed to obtain CO. Preferably, the gas separation device includes one or more of a membrane separation unit, a cryogenic separation unit, and a pressure swing adsorption unit; preferably, it is a membrane separation unit.

[0022] Fourthly, the present invention provides a coal-to-ethanol production apparatus, comprising: The gasification unit has its inlet connected to the raw coal feed pipeline and the oxygen feed pipeline. The feed inlet of the shift gas system is connected to the discharge outlet of the gasification unit. The non-shift gas system has its inlet connected to the outlet of the gasification unit. A membrane pre-separation unit, the inlet of which is connected to the outlet of the non-conversion gas system; A shift gas purification device is used to remove acidic components from a gas, and its inlet is connected to the outlet of the shift gas system and the membrane pre-separation unit. A non-shift gas purification device is used to remove acidic components from a gas, and its inlet is connected to the outlet of the non-shift gas system and the membrane pre-separation unit. The hydrogen extraction unit has its inlet connected to the outlet of the shift gas purification device. A gas separation device, the inlet of which is connected to the outlet of a non-shift gas purification device; The methanol unit has its inlet connected to the outlet of the shift gas purification unit and the tail gas outlet of the gas separation unit. The ethanol unit has its inlet connected to the outlet of the hydrogen extraction unit, the gas separation unit, and the methanol unit.

[0023] The technical solution of this invention has the following advantages: This invention pre-separates the feed gas containing CO, H2, and acidic gases by installing a membrane pre-separation unit before the purification device. This pre-treats the feed gas and preferentially obtains CO-rich non-permeable gas, thus effectively stabilizing the CO partial pressure and flow rate entering subsequent purification and enhancement units even when the feed gas composition fluctuates, especially when the carbon monoxide content decreases and the hydrogen content increases. The CO-rich non-permeable gas, after removing acidic gases, is then purified for CO using one or more separation methods such as membrane separation, cryogenic separation, or pressure swing adsorption. This not only reduces the processing load and energy consumption of subsequent purification units but also avoids the equipment overload and operational instability problems caused by simply relying on scale-up of the device or increasing the processing gas volume. This achieves adjustable and stable supply of carbon monoxide product purity and flow rate, improving the process's adaptability to changes in feed gas composition and overall operational economy.

[0024] Additional aspects and advantages of the invention will be set forth in part in the description which follows, and in part will be obvious from the description, or may be learned by practice of the invention. Attached Figure Description

[0025] To more clearly illustrate the specific embodiments of the present invention or the technical solutions in the prior art, the drawings used in the description of the specific embodiments or the prior art will be briefly introduced below. Obviously, the drawings described below are some embodiments of the present invention. For those skilled in the art, other drawings can be obtained from these drawings without creative effort.

[0026] Figure 1 This is a schematic diagram of the assembly structure of the coal-to-ethanol production equipment of the present invention.

[0027] Explanation of reference numerals in the attached figures: 100. Gasification unit; 200. Air separation unit; 300. Shift conversion unit; 310. Non-shift gas system; 320. Shift gas system; 330. Membrane pre-separation unit; 400. Purification unit; 410. Non-shift gas purification system; 420. Shift gas purification system; 500. Sulfur recovery unit; 600. Gas separation unit; 610. Membrane separation unit; 620. Pressure swing adsorption unit; 630. Temperature swing adsorption unit; 700. Methanol unit; 800. Ethanol unit; 1. Raw coal; 2. Oxygen; 3. Crude syngas; 4. Crude syngas a; 5. Crude syngas b; 6A. Non-shift gas a; 6B. Non-shift gas b; 6C. Non-shift gas c; 6X. Non-permeate gas; 6Y. H2-rich permeate gas; 7. Shift gas; 7A. Shift gas a; 8. Purified non-shift gas; 9. Purified shift gas; 10. Hydrogen extraction shift gas; 11. Methanol Synthesis feedstock gas a; 11A, Methanol synthesis feedstock gas b; 12, Methanol synthesis feedstock gas c; 13, Methanol; 14, Purified CO; 15, Hydrogen; 16, CO-rich purge gas; 17, H2-rich purge gas; 18, Acid gas; 21, Membrane separation unit permeate gas; 22, Pressure swing adsorption tail gas; 23, Temperature swing adsorption purified gas; 24, Mixed gas; 25, Temperature swing adsorption tail gas; 26, Liquid wax product from methanol plant; 27, Mixed alcohol product from methanol plant; 28, By-product fuel gas from methanol plant; 29, By-product fuel gas; 30, Mixed fuel gas; 31, Ethanol product; 32, Ethyl acetate product; 33, By-product heavy components from ethanol plant; 34, Wastewater from ethanol plant; 40, Gasified acid gas; 41, Converted acid gas; 42, Boiler acid gas; 43, Flue gas; 44, Sulfur product; 45, CO2 gas a; 46, CO2 gas b. Detailed Implementation

[0028] The following embodiments are provided to better understand the present invention, but the following embodiments do not constitute a limitation on the content and scope of protection of the present invention. Any product that is the same as or similar to the present invention, derived by any person under the guidance of the present invention or by combining the features of the present invention with other prior art, falls within the scope of protection of the present invention.

[0029] Unless otherwise defined, all technical and scientific terms used herein have the same meaning as commonly understood by one of ordinary skill in the art to which this invention pertains; the terminology used herein is for the purpose of describing particular embodiments only and is not intended to limit the invention; the terms “comprising” and “having” and any variations thereof in the text of this invention are intended to cover non-exclusive inclusion.

[0030] In the description of the embodiments of the present invention, the technical terms "first", "second", etc. are used only to distinguish different objects and should not be construed as indicating or implying relative importance or implicitly indicating the number, specific order or primary and secondary relationship of the indicated technical features.

[0031] In this document, the term "embodiment" means that a particular feature, structure, or characteristic described in connection with an embodiment may be included in at least one embodiment of the invention. The appearance of this phrase in various places throughout the specification does not necessarily refer to the same embodiment, nor is it a separate or alternative embodiment mutually exclusive with other embodiments. It will be explicitly and implicitly understood by those skilled in the art that the embodiments described herein can be combined with other embodiments.

[0032] The "range" disclosed in this invention is defined by a lower limit and an upper limit. A given range is defined by selecting a lower limit and an upper limit, which define the boundaries of the specific range. This range can include or exclude endpoints and can be arbitrarily combined; that is, any lower limit can be combined with any upper limit to form a range. In this invention, unless otherwise stated, the numerical range "ab" represents a shortened representation of any combination of real numbers from a to b, where a and b are real numbers. For example, the numerical range "0-5" indicates that all real numbers between "0-5" have been listed herein; "0-5" is merely a shortened representation of these numerical combinations. Furthermore, when a parameter is described as an integer ≥ 2, it is equivalent to disclosing that the parameter can be, for example, an integer 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, etc.

[0033] In the description of the embodiments of this invention, the term "and / or" is merely a description of the relationship between related objects, indicating that three relationships can exist. For example, A and / or B can represent: A existing alone, A and B existing simultaneously, and B existing alone. Additionally, the character " / " in this document generally indicates that the preceding and following related objects have an "or" relationship.

[0034] In the description of the embodiments of the present invention, the term "at least one" refers to one or more (including two).

[0035] Unless otherwise specified, all experimental steps or conditions in the examples were performed according to conventional experimental procedures and conditions in the art. Reagents or instruments whose manufacturers are not specified are all commercially available products.

[0036] The present invention will now be described with reference to specific embodiments. It should be noted that these embodiments are merely descriptive and do not limit the present invention in any way.

[0037] Unless otherwise specified, all experimental steps or conditions in the examples were performed according to conventional experimental procedures and conditions in the art. Reagents or instruments whose manufacturers are not specified are all commercially available products.

[0038] This invention provides a CO purification process and equipment. The feed gas includes a non-shift gas containing CO, H2, and an acidic gas, wherein the acidic gas includes one or more of CO2 and H2S. The CO purification equipment includes a non-shift gas system 310, a membrane pre-separation unit 330, a purification device 400 for removing acidic components from the gas, and a gas separation device 600. The outlet of the non-shift gas system 310 is connected to the inlet of the membrane pre-separation unit 330, the outlet of the membrane pre-separation unit 330 is connected to the inlet of the purification device 400, the outlet of the purification device 400 is connected to the inlet of the gas separation device 600, and the outlet of the gas separation device 600 is used to output CO.

[0039] Specifically, the feed gas output from the non-shift gas system 310 first enters the membrane pre-separation unit 330 for pre-separation treatment. In the membrane pre-separation unit 330, the feed gas is separated into CO-rich non-permeable gas and H2-rich permeable gas. The pressure of the H2-rich permeable gas in the membrane pre-separation unit is 0.1 MPaG to 3 MPaG, and the pressure of the CO-rich non-permeable gas is 1.5 MPaG to 8.5 MPaG. Through the pre-separation action of the membrane pre-separation unit 330, the CO component in the feed gas is enriched on the non-permeable gas side, while the H2 component is enriched on the permeable gas side, thereby achieving preliminary separation of CO and H2 in the feed gas.

[0040] The CO-rich non-permeable gas output from the membrane pre-separation unit 330 enters the purification device 400 for acid gas removal treatment to remove acidic components such as CO2 and H2S, resulting in CO-rich gas after acid gas removal. The purification device is used to remove acidic components from the gas and can employ a low-temperature methanol washing device.

[0041] Furthermore, the purification device 400 includes a shift gas purification system 420 and a non-shift gas purification system 410. Specifically, the CO-rich non-permeable gas output from the membrane pre-separation unit enters the non-shift gas purification system 410 to remove acidic gases; the H2-rich permeable gas output from the membrane pre-separation unit enters the shift gas purification system 420 to remove acidic gases.

[0042] The H2-rich permeate gas is first pressurized by a compressor to the pressure of the shift gas pipeline, and then sent to the shift gas purification system 420 for acid gas removal. This setup allows the H2-rich gas obtained from the permeate side of the membrane pre-separation unit to be incorporated into the existing shift gas treatment process, reducing system modifications and facilitating the comprehensive utilization of subsequent hydrogen resources.

[0043] The CO-rich non-permeable gas, after being purified by the non-shift gas purification system 410 to remove acidic gases, is further fed into a gas separation device 600 for separation to obtain CO product. The separation method in the gas separation device 600 includes one or more of membrane separation, cryogenic separation, and pressure swing adsorption separation. In this embodiment, the preferred separation method is membrane separation. That is, the CO-rich non-permeable gas after deacidification enters the membrane separation unit for further purification. The non-permeable gas side of the membrane separation unit outputs the purified CO product, and the permeable gas side outputs H2-rich gas.

[0044] Through the above process, CO in the feed gas is first pre-concentrated by a membrane pre-separation unit, then its acidic components are removed by a purification device, and finally further separated and purified by a gas separation device to obtain a CO product that meets downstream requirements. Compared with processes that directly proceed to subsequent deacidification and CO purification without membrane pre-separation, this embodiment, by setting up a membrane pre-separation unit at the front end, can effectively increase the CO component content in the CO-rich gas entering the subsequent purification and refining units, reduce the processing load of the subsequent gas separation device, and facilitate the stable acquisition of CO products with the required flow rate and purity.

[0045] Example 1 like Figure 1 As shown, the present invention provides a coal-to-ethanol production equipment and process, as detailed below: Raw coal 1 and oxygen 2 from air separation unit 200 undergo a coal gasification reaction in gasification unit 100 to form crude syngas 3. The crude syngas 3 is divided into crude syngas a4 and crude syngas b5. An ammonia scrubbing tower is installed between gasification unit 100 and shift converter 300. Before entering shift converter 300, crude syngas a4 and crude syngas b5 undergo ammonia scrubbing treatment in the ammonia scrubbing tower. After ammonia scrubbing, crude syngas a4 is sent to the non-shift gas system 310 of shift converter 300 for heat recovery, and crude syngas b5 is sent to the shift gas system 320 of shift converter 300 for heat recovery. The gasified acid gas 40 output from gasification unit 100 and the shift acid gas 41 output from shift converter 300 are combined to form boiler acid gas 42, which is then supplied to the boiler.

[0046] The shift gas system 320 outputs shift gas 7, and the non-shift gas system 310 outputs non-shift gas a6A and non-shift gas b6B. Non-shift gas a6A is fed into the membrane pre-separation unit 330, where it is initially concentrated, resulting in non-permeate gas 6X on its non-permeate gas side. This non-permeate gas 6X is combined with non-shift gas b6B to form non-shift gas c6C, which is then fed into the non-shift gas purification system 410 of the purification device 400 for acid gas removal. The permeate gas side of the membrane pre-separation unit 330 yields H2-rich permeate gas 6Y, which is combined with shift gas 7 to form shift gas a7A. Shift gas a7A is then fed into the shift gas purification system 420 of the purification device 400 for acid gas removal.

[0047] The acidic gas 18 output by the purification device 400 mainly consists of H2S and CO2. This acidic gas 18 is transported to the sulfur recovery device 500 to produce sulfur product 44. The flue gas 43 generated by the sulfur recovery device 500 is then transported to the boiler. Simultaneously, the CO2 gas generated by the purification device 400 is split into CO2 gas a45 for hopper pressurization and CO2 gas b46 for pulverized coal conveying. CO2 gas a45 for hopper pressurization is transported to the gasification device 100 for hopper pressurization, and CO2 gas b46 for pulverized coal conveying is also transported to the gasification device 100 for pulverized coal conveying.

[0048] The non-shift gas purification system 410 outputs purified non-shift gas 8, which is then sent to the membrane separation unit 610 of the gas separation device 600 for CO purification. The non-permeate gas side of the membrane separation unit 610 outputs purified CO14, which is sent to the downstream ethanol unit 800 for a carbonylation reaction. The permeate gas side of the membrane separation unit 610 produces membrane separation unit permeate gas 21, which is sent to the methanol unit 700.

[0049] The shift gas purification system 420 outputs purified shift gas 9, which is divided into hydrogen extraction shift gas 10 and methanol synthesis feed gas a11. Methanol synthesis feed gas a11 is combined with permeate gas 21 from the membrane separation unit to form methanol synthesis feed gas b11A.

[0050] Hydrogen-reduced gas 10 is fed into pressure swing adsorption (PSA) unit 620 for hydrogen purification. The purified hydrogen 15 is then pressurized by a hydrogen compressor and sent to the ethanol unit 800 to provide fresh hydrogen for the hydrogenation stage of ethanol synthesis. The PSA tail gas 22 output from PSA unit 620 is pressurized by a compressor and sent to the feed gas pipeline of methanol unit 700.

[0051] Methanol synthesis feedstock gas b11A is mixed with mixed gas 24 to form methanol synthesis feedstock gas c12. Methanol synthesis feedstock gas c12 is fed into the inlet of the circulating compressor of methanol unit 700, and after pressurization, it enters the methanol synthesis reactor as methanol synthesis feedstock gas to produce methanol 13. Methanol 13 is then fed into ethanol unit 800 as feedstock. In addition to outputting methanol 13, methanol unit 700 also outputs methanol unit liquid wax product 26, methanol unit mixed alcohol product 27, and methanol unit by-product fuel gas 28.

[0052] The ethanol unit 800 includes a dimethyl ether section, a carbonylation section, a hydrogenation section, and an ethanol refining section. Purified CO 14, hydrogen 15, and methanol 13 participate in the reaction within the ethanol unit 800. The ethanol unit 800 outputs ethanol product 31, ethyl acetate product 32, heavy by-product components 33, and ethanol unit wastewater 34. Simultaneously, the ethanol unit 800 discharges CO-rich purge gas 16, H2-rich purge gas 17, and ethanol unit by-product fuel gas 29.

[0053] The CO-rich purge gas 16 and H2-rich purge gas 17 discharged from the ethanol unit 800 are sent to the temperature swing adsorption unit 630 for separation and recovery. After processing by the temperature swing adsorption unit 630, temperature swing adsorption purified gas 23 and temperature swing adsorption tail gas 25 are obtained. Among them, the temperature swing adsorption purified gas 23 and the pressure swing adsorption tail gas 22 are combined to form a mixed gas 24, which is then mixed with methanol synthesis feedstock gas b11A to form methanol synthesis feedstock gas c12; the temperature swing adsorption tail gas 25 is combined with the by-product fuel gas 29 of the ethanol unit and the by-product fuel gas 28 of the methanol unit output from the methanol unit 700 to form a mixed fuel gas 30, which is then output.

[0054] As a comparison, the present invention provides a reference operating condition. The only difference between the reference operating condition and the production equipment of Example 1 is that the membrane pre-separation unit 330 is not set. The other device configurations, connection relationships and process flow are the same as those of Example 1.

[0055] Specifically, in this baseline operating condition, the non-shifted gas output from the non-shifted gas system 310 no longer enters the membrane pre-separation unit 330 for pre-concentration treatment, but is directly sent to the non-shifted gas purification system 410 of the purification device 400 for purification. That is, in this baseline operating condition, there are no two streams of gas: non-permeate gas 6X and H2-rich permeate gas 6Y as in Example 1, and all the non-shifted gas output from the non-shifted gas system 310 enters the subsequent non-shifted gas purification system 410 as a single gas stream.

[0056] The specific production process is as follows: Raw coal 1 and oxygen 2 react in gasification unit 100 to generate crude syngas. The crude syngas enters non-shift gas system 310 for heat recovery to obtain non-shift gas d1. The non-shift gas d1 is directly sent to the subsequent non-shift gas purification system 410.

[0057] The non-shifted gas d1 contains 64.36 vol% CO, 26.67 vol% H2, and a corresponding saturated water content of 0.24 vol%. Due to its low water content, the CO volume fraction in the non-shifted gas d1 can be approximated as a dry basis composition.

[0058] The non-shifted gas d1 further enters the non-shifted gas purification system 410, where impurities are removed to obtain purified non-shifted gas d1. The volume fraction of CO in the purified non-shifted gas d1 is increased to 70.13 v, and the volume fraction of H2 is 29.21 v.

[0059] Under this reference operating condition, the purified non-shift gas d1 is fed into the membrane separation unit 610 of the gas separation device 600 for separation to obtain non-permeate gas, i.e., purified CO d1, and membrane separation unit permeate gas d1, wherein the membrane separation unit permeate gas d1 is hydrogen-rich gas.

[0060] Using the operating parameters of the membrane separation unit under the above-mentioned baseline conditions as the baseline conditions, the process operation under different feed gas composition conditions is compared and analyzed. The relevant parameters are detailed in Table 1.

[0061] Table 1: Purity and Flow Parameters under Baseline Operating Conditions

[0062] Comparative Example 1 This comparative example is based on the analysis of a coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the feed composition of the membrane separation unit 610 in this comparative example has changed, that is, the CO volume fraction in the purified non-shift gas d1 is reduced by about 3v compared with the baseline operating condition.

[0063] Under these conditions, the membrane separation unit 610 in the gas separation device 600 is still required to produce the same purified CO product flow rate and purity as under the baseline operating conditions. Based on this, the composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 2.

[0064] Table 2: Purity and Flow Rate Parameters

[0065] Under this operating condition, to maintain the same purified carbon monoxide product flow rate and purity as the baseline condition, the flow rate of the purified non-shift gas d1 entering the membrane separation unit 610 is significantly increased. As shown in Table 2, the gas flow rate of the purified non-shift gas d1 reaches 97600.00 Nm³. 3 / h, significantly higher than the 70802 Nm under the baseline operating condition. 3 The throughput is approximately 138% of the baseline operating rate. This throughput significantly exceeds the design load range of the upstream related equipment, making it difficult for the system to maintain stable operation within the baseline operating range.

[0066] Simultaneously, under this operating condition, the flow rate of the membrane separation unit permeate gas d1 generated by the membrane separation unit 610 also increases significantly. As shown in Table 2, the gas flow rate of the membrane separation unit permeate gas d1 is 58743.39 Nm³. 3 / h, while the gas flow rate of permeate gas d1 in the membrane separation unit under the baseline operating conditions is 31165 Nm. 3 The former is approximately 1.88 times that of the latter, nearly twice. Since the permeate gas d1 from the membrane separation unit needs to be pressurized by the permeate gas compressor before being sent to the downstream system, under essentially the same inlet and outlet pressure conditions, the compressor's processing load will increase significantly with the increase in permeate gas volume, resulting in a substantial increase in compressor power demand, reaching approximately 8000 kW, about twice that of the baseline operating condition. Therefore, without the membrane pre-separation unit 330, when the CO volume fraction in the feed to the membrane separation unit 610 decreases by approximately 3v%, if the CO product flow rate and purity corresponding to the baseline operating condition are still required, the load on the membrane separation unit 610 and its associated compression system will increase significantly, leading to a marked increase in system energy consumption and equipment burden, making stable operation difficult.

[0067] Comparative Example 2 This comparative example is based on the coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the CO volume fraction in the purified non-shift gas d1 in this comparative example is reduced by approximately 3v%.

[0068] Under these conditions, the purified non-shift gas d1 is separated by the membrane separation unit 610 in the gas separation device 600, and the flow rate of the membrane separation unit permeate gas d1 generated by the membrane separation unit 610 is limited to not exceeding the maximum processing capacity of the permeate gas compressor. Under these constraints, the composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 3.

[0069] Table 3: Purity and Flow Rate Parameters

[0070] The maximum processing capacity of the membrane separation permeate compressor is approximately 110% of the baseline operating condition, corresponding to a maximum throughput of approximately 35,000 Nm³ / h. When the membrane separation permeate compressor reaches this maximum processing load, the permeate flow rate d1 produced by the membrane separation unit 610 is close to the system's allowable upper limit. At this point, even if the number of membrane modules is further increased, the hydrogen permeation rate cannot be significantly increased. Under this constraint, the purity of the purified CO d1 output from the membrane separation unit 610 will be limited, with a CO purity of approximately 93% v.

[0071] Although the purity of the aforementioned carbon monoxide product still meets the basic requirements for use in the ethanol synthesis process, its lower purity than the baseline operating level will increase the amount of impurity gases entering the carbonylation reaction loop of the ethanol unit, leading to an increase in the amount of purge gas emitted from the carbonylation reaction loop. This increased purge gas emission not only causes the loss of effective components but also adversely affects the economic efficiency and stability of the system operation.

[0072] Comparative Example 3 This comparative example is based on the coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the CO volume fraction in the purified non-shift gas d1 in this comparative example is reduced by approximately 3v%.

[0073] Under these conditions, the purified non-shifted gas d1 is separated by the membrane separation unit 610 in the gas separation device 600, and the amount of purified non-shifted gas d1 entering the membrane separation unit 610 is maintained to be basically consistent with the baseline operating conditions. Based on this, with the goal of maximizing the purity of the purified CO d1 product, the operating parameters of the membrane separation unit 610 are adjusted. The composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 4.

[0074] Table 4: Purity and Flow Rate Parameters

[0075] As shown in Table 4, under the condition that the inlet flow rate of the purified non-shift gas d1 entering the membrane separation unit 610 remains constant at approximately 70800 Nm³ / h, the CO purity in the purified CO d1 product can be increased to approximately 96 v% by adjusting the operating parameters of the membrane separation unit 610. However, under this condition, the flow rate of the membrane separation unit permeate gas d1 generated by the membrane separation unit 610 increases to 41466.70 Nm³ / h, which is significantly higher than the maximum processing capacity of the membrane separation permeate gas compressor of 35000 Nm³ / h, exceeding its maximum allowable load.

[0076] This indicates that, without the membrane pre-separation unit 330, when the CO volume fraction in the purified non-shift gas d1 decreases by approximately 3v%, even if the intake gas flow rate into the membrane separation unit 610 is maintained at approximately the same as the baseline operating condition, and the operating parameters of the membrane separation unit 610 are optimized with the goal of improving the purity of the purified CO d1 product, the flow rate of the permeate gas d1 in the membrane separation unit will still exceed the maximum processing capacity of the matching compressor, making it difficult to achieve stable operation under the constraints of the existing equipment capacity.

[0077] Comparative Example 4 This comparative example is based on the coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the CO volume fraction in the purified non-shift gas d1 in this comparative example is reduced by approximately 3v%.

[0078] Under these conditions, the purified non-shift gas d1 is separated by the membrane separation unit 610 in the gas separation device 600, and the flow rate of the purified CO d1 product is maintained at a level basically consistent with the baseline operating conditions. The operating parameters of the membrane separation unit 610 are adjusted with the goal of maximizing the purity of the purified CO d1 product. The composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 5.

[0079] Table 5: Purity and Flow Rate Parameters

[0080] As shown in Table 5, while maintaining the purified CO d1 product flow rate at approximately 38,000 Nm³ / h, the CO purity in the purified CO d1 product can be increased to approximately 96% by adjusting the operating parameters of the membrane separation unit 610. However, under this condition, the permeate gas d1 flow rate generated by the membrane separation unit 610 is 54,239.15 Nm³ / h, significantly higher than the maximum processing capacity of the membrane separation permeate gas compressor (35,000 Nm³ / h), exceeding it by approximately 19,239.15 Nm³ / h, which is about 55% of the maximum processing capacity. Under these conditions, the membrane separation unit 610 can no longer further increase the hydrogen permeate flow rate and the purity of the purified CO d1 product by further increasing the number of membrane modules.

[0081] Meanwhile, to maintain the flow rate of the purified CO d1 product, the flow rate of the purified non-shift gas d1 entering the membrane separation unit 610 increases to 92,500.00 Nm³ / h, significantly higher than the baseline operating condition of 70,802 Nm³ / h, approximately 131% of the baseline operating condition. This processing capacity exceeds the design load range of the upstream related equipment, causing the system to be unable to operate stably within the baseline operating condition range.

[0082] This indicates that, without the membrane pre-separation unit 330, when the CO volume fraction in the purified non-shift gas d1 decreases by approximately 3v%, if a higher product purity is pursued while maintaining the CO d1 product flow rate essentially unchanged, it will not only cause the permeate gas d1 flow rate of the membrane separation unit to significantly exceed the maximum processing capacity of the matching compressor, but also significantly increase the amount of raw material gas entering the membrane separation unit 610, making it difficult to meet the capacity constraints of the existing equipment, and the system will also be unable to operate stably.

[0083] Comparative Example 5 This comparative example is based on the coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the CO volume fraction in the purified non-shift gas d1 in this comparative example is reduced by approximately 8% v%.

[0084] Under these conditions, the purified non-shifted gas d1 is separated by the membrane separation unit 610 in the gas separation device 600. The throughput of purified non-shifted gas d1 entering the membrane separation unit 610 is maintained at a level consistent with the baseline operating conditions. The operating parameters of the membrane separation unit 610 are adjusted with the goal of maximizing the purity of the purified CO d1 product. The composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 6.

[0085] Table 6: Purity and Flow Parameters

[0086] As shown in Table 6, when the inlet flow rate of the purified non-shift gas d1 entering the membrane separation unit 610 is maintained at approximately 70,000 Nm³ / h, the CO purity in the purified CO d1 product can be increased to approximately 92% by adjusting the operating parameters of the membrane separation unit 610. However, under this condition, the flow rate of the membrane separation unit permeate gas d1 generated by the membrane separation unit 610 is 37,986.41 Nm³ / h, which exceeds the maximum processing capacity of the membrane separation permeate gas compressor of 35,000 Nm³ / h. Therefore, the membrane separation unit 610 can no longer further increase the hydrogen permeate flow rate and the purity of the purified CO d1 product by further increasing the number of membrane modules.

[0087] Meanwhile, under the aforementioned operating conditions, the purified CO d1 flow rate output by the membrane separation unit 610 was 32013.59 Nm³ / h, significantly lower than the baseline operating condition of 39637 Nm³ / h; the CO purity in the purified CO d1 was approximately 92 vol%, also significantly lower than the baseline operating condition of 98.338%. This indicates that, without the membrane pre-separation unit 330, when the CO volume fraction in the purified non-shift gas d1 decreases by approximately 8 vol%, even if the feed rate of the membrane separation unit 610 is kept essentially unchanged, adjusting the operating parameters with the goal of improving the purity of the purified CO d1 product still cannot simultaneously meet the permeate compressor's processing capacity constraints and the CO product flow rate and purity requirements corresponding to the baseline operating condition, making it difficult for the system to achieve stable operation.

[0088] Comparative Example 6 This comparative example is based on the coal-to-ethanol production system without the membrane pre-separation unit 330 described in the baseline operating condition. Compared with the baseline operating condition, the CO volume fraction in the purified non-shift gas d1 in this comparative example is reduced by approximately 8% v%.

[0089] Under these conditions, the purified non-shift gas d1 is separated by the membrane separation unit 610 in the gas separation device 600, and the flow rate of the membrane separation unit permeate gas d1 generated by the membrane separation unit 610 is limited to the maximum processing capacity of the membrane separation permeate gas compressor, which is approximately 35,000 Nm³ / h. Under this constraint, with the goal of maximizing the purity and flow rate of the purified CO d1 product, the operating parameters of the membrane separation unit 610 are adjusted. The composition and flow rate parameters of the inlet and outlet streams of the membrane separation unit 610 under the corresponding operating conditions are shown in Table 7.

[0090] Table 7: Purity and Flow Rate Parameters

[0091] As shown in Table 7, when the flow rate of the permeate gas d1 in the membrane separation unit is limited by the maximum processing capacity of the membrane separation permeate gas compressor, the CO purity in the purified CO d1 product is approximately 92% v%, and the product flow rate is approximately 30,000 Nm³ / h. Both are significantly lower than the product purity and flow rate corresponding to the baseline operating conditions. This indicates that, without the membrane pre-separation unit 330, when the CO volume fraction in the purified non-shift gas d1 decreases by approximately 8% v%, even if the flow rate of the permeate gas d1 in the membrane separation unit is controlled within the processing capacity range of the matching compressor, it is still difficult to simultaneously ensure that the flow rate and purity of the purified CO d1 product meet the requirements of the baseline operating conditions.

[0092] Comparing Comparative Examples 1 to 6, without the membrane pre-separation unit 330, when the CO component in the syngas decreases and the H2 component increases, the CO component in the purified non-shift gas will decrease, making the CO purification unit in the gas separator 600 a critical link limiting the stable operation of the system. At this point, whether by increasing the throughput or separation load of the membrane separation unit 610, or optimizing operating parameters under compressor capacity constraints, it is difficult to simultaneously meet the CO product flow rate and purity requirements corresponding to the baseline operating conditions. This can easily lead to problems such as excessive permeate gas d1 flow rate in the membrane separation unit, overloaded compressor operation, excessive upstream system throughput, and decreased downstream ethanol synthesis capacity and increased purge gas volume. Therefore, relying solely on the existing non-shift gas treatment and CO purification pathways is insufficient to effectively address the system instability caused by fluctuations in feedstock composition.

[0093] To address the issue of decreased CO and increased H2 content in the syngas, this invention does not solve the problem by increasing the processing capacity of each unit. Simply increasing the design margin of each unit would not only waste equipment investment but also lead to increased system energy consumption. Therefore, this invention incorporates a membrane pre-separation unit 330 at the outlet of the shift converter 300 to pre-concentrate the CO component in the non-shifted gas. Instead of using pressure swing adsorption or cryogenic separation for CO pre-concentration, membrane separation is employed to concentrate the CO component in the non-shifted gas to the composition corresponding to or higher than the raw material's equilibrium baseline operating conditions.

[0094] Under baseline operating conditions, the CO content in the non-shifted gas of the membrane separation unit 610 in the gas separation unit 600 is approximately 70 v%, corresponding to a CO content of approximately 64 v% at the inlet of the non-shifted gas purification system 410 in the purification unit 400. Therefore, by pre-concentrating the non-shifted syngas at the outlet of the shift unit 300 to increase its CO content to approximately 64 v%, the requirements for stable operation of the membrane separation unit 610 in the downstream gas separation unit 600 can be met, and the required CO flow rate and purity can be guaranteed for supplying the downstream ethanol unit 800.

[0095] Based on the above design concept, the material balance of a specific pre-concentration scheme will be explained below.

[0096] Example 2 This embodiment analyzes the coal-to-ethanol production equipment and process described in Embodiment 1. Compared to the baseline operating condition, the CO volume fraction of the non-shift gas d1 in the baseline operating condition is 64.36 v. In this embodiment, the CO volume fraction of the non-shift gas a6A fed into the membrane pre-separation unit 330 is reduced by approximately 3 v. The remaining device configuration, connection relationships, and process flow are the same as in Embodiment 1.

[0097] In this embodiment, non-shift gas a6A is fed into the membrane pre-separation unit 330 for CO component pre-concentration. By adjusting the operating parameters of the membrane pre-separation unit 330, the CO component content in the non-permeable gas 6X output from the non-permeable gas side of the membrane pre-separation unit 330 reaches approximately 64% v%. The non-permeable gas 6X and non-shift gas b6B are combined to form non-shift gas c6C, which is then sent to the non-shift gas purification system 410 of the purification device 400 for subsequent processing. The permeable gas side of the membrane pre-separation unit 330 outputs H2-rich permeable gas 6Y. The H2-rich permeable gas 6Y is combined with shift gas 7 to form shift gas a7A, which is then sent to the shift gas purification system 420 of the purification device 400 for subsequent processing. In this embodiment, the pressure of the H2-rich permeable gas in the membrane pre-separation unit 330 is 0.15 MPaG, and the pressure of the CO-rich non-permeable gas is 3.5 MPaG. Under the corresponding operating conditions, the composition and flow rate parameters of the inlet and outlet streams of the membrane pre-separation unit 330 are shown in Table 2.1.

[0098] Table 2.1: Material Parameter Table

[0099] When the CO volume fraction in the non-shift gas a6A fed into the membrane pre-separation unit 330 decreases by approximately 3% compared to the baseline operating conditions, after pre-concentrating the CO component in the non-shift gas a6A through the membrane pre-separation unit 330, the CO volume fraction in the non-permeable gas 6X output from the non-permeable gas side of the membrane pre-separation unit 330 can be increased to 64.434%. This allows the CO volume fraction in the purified non-shift gas 8 obtained after subsequent treatment by the purification device 400 to recover to 70.217%, with a gas flow rate of 70856.13 Nm³. 3 / h. It can be seen that the non-shifted gas fed into the purification device 400 and the gas separation device 600 is basically equivalent to the original reference operating conditions in terms of gas volume and composition purity, thus meeting the requirements for stable operation of the subsequent gas separation device 600.

[0100] Simultaneously, H2-rich permeate gas 6Y is obtained from the permeate gas side of the membrane pre-separation unit 330. This H2-rich permeate gas 6Y, after pressurization, is combined with the shift gas 7 and enters the shift gas purification system 420 of the purification device 400. As shown in Table 2.1, the gas flow rate of the H2-rich permeate gas 6Y is 8015.06 Nm³. 3 Since the gas volume is relatively limited, after it is incorporated into the gas exchange side, the overall processing load of the purification device 400 is still within the design allowable range and will not cause the system to operate under significant overload.

[0101] Furthermore, by employing the aforementioned CO pre-concentration process, only a membrane pre-separation unit 330 needs to be added to the outlet of the shift converter 300, along with a limited number of membrane modules. This allows for effective pre-concentration of the CO component in the non-shift gas, maintaining the downstream purification unit 400 and gas separation unit 600 in near-baseline operating conditions. Compared to addressing raw material composition fluctuations by significantly increasing the design margins of upstream and downstream units, this approach results in significantly lower equipment investment and operating energy consumption, thus demonstrating better overall economic efficiency.

[0102] Example 3 Example 3 analyzes the coal-to-ethanol production equipment and process described in Example 1. Compared to the baseline operating conditions, when the CO volume fraction in the non-shift syngas fed into the membrane pre-separation unit 330 decreases by approximately 5v%, the CO content in the non-permeable gas output from the non-permeable side of the membrane pre-separation unit 330 is maintained at the same level as the original baseline operating conditions, approximately 64v%, by adjusting the operating parameters of the membrane pre-separation unit 330. In this example, the pressure of the H2-rich permeable gas in the membrane pre-separation unit 330 is 0.15 MPaG, and the pressure of the CO-rich non-permeable gas is 3.5 MPaG. Under these conditions, the composition and flow rate parameters of the inlet and outlet streams of the newly added membrane pre-separation unit 330 are shown in Table 2.2.

[0103] Table 2.2: Material Parameter Table

[0104] After the carbon monoxide (CO) in the non-shift gas is pre-concentrated by the membrane pre-separation unit 330, the non-shift gas sent to the purification unit 400 and the gas separation unit 600 can be restored to a level that is basically equivalent to the original reference operating conditions in terms of the amount of gas processed and the purity of composition. This ensures that the subsequent units can operate stably under conditions close to the reference operating conditions.

[0105] Meanwhile, the hydrogen-rich gas obtained from the permeate side of the membrane pre-separation unit 330 is pressurized by a compressor and then sent to the main inlet pipe of the purification unit 400. Since the additional gas volume in this part is limited, the overall operating load of the purification unit 400 is still within the design allowable range and will not cause significant overload to the gas-side treatment system.

[0106] By employing the aforementioned CO pre-concentration process, only a limited number of membrane modules are required to effectively regulate the CO composition in the non-shift gas. Compared to methods that rely on significantly increasing the design margin of upstream and downstream systems to accommodate fluctuations in feedstock composition, this approach results in significantly lower additional equipment investment and operating energy consumption, leading to better overall economic efficiency.

[0107] The material balance results from the newly added membrane pre-separation unit 330 show that simply installing the membrane pre-separation unit 330 at the outlet of the conversion unit 300 can effectively address system operation problems caused by significant fluctuations in raw material composition. If the membrane pre-separation unit 330 is not installed upstream, even if cryogenic separation or pressure swing adsorption (PSA) processes are used downstream for carbon monoxide purification, significant fluctuations in product purity and flow rate will still occur. Specifically, cryogenic separation is prone to insufficient cooling capacity and excessively low carbon monoxide pressure at the cold box outlet, leading to the inability of the original carbon monoxide compressor to operate properly. PSA, on the other hand, is prone to insufficient carbon monoxide production, and simply increasing the processing gas volume to compensate will adversely affect the operational stability of the adsorption tower.

[0108] This invention provides a membrane pre-separation unit 330 at the outlet of the conversion unit 300 to pre-concentrate the carbon monoxide component in the non-conversion syngas, restoring it to the composition level corresponding to the original design coal type operating conditions. This ensures that the operating loads of the downstream purification unit 400, gas separation unit 600, and methanol unit 700 do not exceed the baseline operating conditions.

[0109] Based on this, even if the ethanol unit 800 requires capacity expansion and modification in the future, the changes required for the entire plant will be relatively limited. Specifically, while keeping the gasification unit 100 unchanged, the production capacity of the ethanol unit 800 can be further increased by purchasing some methanol, further increasing the carbon monoxide concentration in the non-shift gas at the outlet of the membrane pre-separation unit 330, and correspondingly increasing the hydrogen production of the pressure swing adsorption unit 620, provided that the ethanol unit 800 itself has sufficient capacity expansion space.

[0110] Furthermore, the membrane pre-separation unit 330 described in this invention can be used in series with the separation unit for carbon monoxide extraction in the gas separation device 600, thereby ensuring that the output carbon monoxide product flow rate and purity meet the process requirements of the downstream syngas-to-ethanol unit. The carbon monoxide purification unit in the gas separation device 600 can employ membrane separation technology, cryogenic separation technology, or pressure swing adsorption technology. In the scenario of technical transformation of existing ethanol plants, this invention requires fewer new equipment. The membrane pre-separation unit 330 has a compact footprint and low investment cost, but can significantly improve the stability of carbon monoxide product supply and greatly reduce the problem of excessively high partial pressure of alcohols, esters, and ethers in carbon monoxide-rich purge gas due to high hydrogen content. At the same time, it minimizes the amount of carbon monoxide-rich purge gas, thereby avoiding adverse effects on downstream unit capacity due to fluctuations in carbon monoxide gas purity.

[0111] Obviously, the above embodiments are merely illustrative examples for clear explanation and are not intended to limit the implementation. Those skilled in the art will recognize that other variations or modifications can be made based on the above description. It is neither necessary nor possible to exhaustively list all possible implementations here. However, obvious variations or modifications derived therefrom are still within the scope of protection of this invention.

Claims

1. A CO purification process, characterized in that, Includes the following steps: (1) The feed gas containing CO, H2 and acidic gas is passed into the membrane pre-separation unit to separate CO-rich non-permeable gas and H2-rich permeable gas. (2) CO-rich non-permeable gas enters the purification device to remove acidic gases; (3) CO is obtained by separating the CO-rich non-permeable gas after removing acidic gases, wherein the separation method includes one or more of membrane separation process, cryogenic separation process, and pressure swing adsorption separation process.

2. The CO purification process according to claim 1, characterized in that, In step (1), the pressure of the H2-rich permeate gas in the membrane pre-separation unit is 0.1 MPaG to 3 MPaG, and the pressure of the CO-rich non-permeate gas is 1.5 MPaG to 8.5 MPaG.

3. The CO purification process according to claim 1 or 2, characterized in that, In step (2), the purification device includes a shift gas purification device and a non-shift gas purification device. The CO-rich non-permeable gas enters the non-shift gas purification device to remove acidic gas, and the H2-rich permeable gas enters the shift gas purification device to remove acidic gas.

4. The CO purification process according to claim 3, characterized in that, In step (2), the H2-rich permeate gas is pressurized to the pressure of the shift gas pipeline and then combined with the shift gas and sent to the shift gas purification device to remove acidic gas.

5. The CO purification process according to any one of claims 1-4, characterized in that, In step (1), the raw material gas includes non-shift gas; And / or, in step (3), the separation method is selected from membrane separation process.

6. A CO purification apparatus, characterized in that, This includes a non-shift gas system, a membrane pre-separation unit, a purification device for removing acidic components from the gas, and a gas separation device. The outlet of the non-conversion gas system is connected to the inlet of the membrane pre-separation unit, the outlet of the membrane pre-separation unit is connected to the inlet of the purification device, the outlet of the purification device is connected to the inlet of the gas separation device, and the outlet of the gas separation device is used to output CO.

7. A coal-to-ethanol production process, characterized in that, Includes the following steps: S1 involves reacting raw coal with oxygen in a coal gasification reaction to form crude syngas; S2 The crude synthesis gas is fed into the shift gas system and the non-shift gas system respectively to obtain shift gas and non-shift gas; S3 Non-shift gas is introduced into the membrane pre-separation unit to separate CO-rich non-permeable gas and H2-rich permeable gas; After the S4 H2-rich permeate gas is mixed with the shift gas and the acid gas is removed, part of it is sent to the hydrogen extraction unit to purify H2, and the other part is sent to the methanol unit. The H2 purified by the pressure swing adsorption unit is sent to the ethanol unit, and the tail gas generated is sent to the methanol unit. After removing acidic gases from the CO-rich non-permeable gas, CO is purified and sent to the ethanol unit. The resulting tail gas is sent to the methanol unit, and the methanol produced in the methanol unit is sent to the ethanol unit. In the ethanol unit, methanol, CO, and H2 react to produce ethanol.

8. The coal-to-ethanol production process according to claim 7, characterized in that, In step S4, the H2-rich permeate gas is pressurized and mixed with the shift gas before being sent to the purification device to remove acidic gases. And / or, in step S4, the purified H2 is pressurized and sent to the ethanol unit.

9. The coal-to-ethanol production process according to claim 7 or 8, characterized in that, In step S4, the CO-rich non-permeable gas, after the acidic gas is removed, is sent to a gas separation device for separation and purification to obtain CO. Preferably, the gas separation device includes one or more of a membrane separation unit, a cryogenic separation unit, and a pressure swing adsorption unit; preferably, it is a membrane separation unit.

10. A coal-to-ethanol production equipment, characterized in that, include: The gasification unit has its inlet connected to the raw coal feed pipeline and the oxygen feed pipeline. The feed inlet of the shift gas system is connected to the discharge outlet of the gasification unit. The non-shift gas system has its inlet connected to the outlet of the gasification unit. A membrane pre-separation unit, the inlet of which is connected to the outlet of the non-conversion gas system; A shift gas purification device is used to remove acidic components from a gas, and its inlet is connected to the outlet of the shift gas system and the membrane pre-separation unit. A non-shift gas purification device is used to remove acidic components from a gas, and its inlet is connected to the outlet of the non-shift gas system and the membrane pre-separation unit. The hydrogen extraction unit has its inlet connected to the outlet of the shift gas purification device. A gas separation device, the inlet of which is connected to the outlet of a non-shift gas purification device; The methanol unit has its inlet connected to the outlet of the shift gas purification unit and the tail gas outlet of the gas separation unit. The ethanol unit has its inlet connected to the outlet of the hydrogen extraction unit, the gas separation unit, and the methanol unit.

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