Method for recovering c2 and c3 components from refinery dry gas, and apparatus and use thereof

Abstract

This invention belongs to the field of chemical engineering, specifically disclosing a method, apparatus, and application for recovering C2 and C3 components from refinery dry gas. The method includes: compressing, cooling, absorbing, high-pressure decarbonizing, low-pressure decarbonizing, and desorption steps of the refinery dry gas. It also includes reabsorption and redesorption steps. The method provided by this invention is based on shallow-cooled oil absorption technology. While simplifying the process, reducing cooling consumption, and lowering equipment investment, it increases the ethane + ethylene content in C2 enriched gas and the propane + propylene content in C3 enriched gas, respectively, while achieving a high recovery rate. This allows for more rational and efficient utilization of refinery dry gas resources and has broad industrial application prospects.

Description

Technical Field

[0001] This invention belongs to the field of chemical engineering, specifically relating to a method, apparatus and application for recovering C2 and C3 components from refinery dry gas. Background Technology

[0002] Refinery dry gas is an important resource in the petrochemical industry, mainly derived from primary and secondary crude oil processing, such as atmospheric and vacuum distillation, catalytic cracking (FCC), catalytic cracking (DCC), hydrocracking, and delayed coking. It contains a large amount of light hydrocarbons, including hydrogen, methane, ethane, ethylene, propane, and propylene. Based on its composition, dry gas containing a significant amount of unsaturated components such as ethylene, propylene, and butene is called unsaturated dry gas, such as FCC dry gas and DCC dry gas; while dry gas containing little or no unsaturated components is called saturated dry gas, such as atmospheric and vacuum distillation dry gas, hydrocracking dry gas, and reforming PSA tail gas. Currently, much refinery dry gas is still burned as industrial and domestic fuel gas, resulting in a serious waste of resources. If C2 enriched gas, rich in ethylene and ethane, can be separated from refinery dry gas and used for ethylene production, a large amount of light oil used as feedstock for ethylene cracking can be saved. This is an effective way to improve the comprehensive utilization rate of resources, expand the sources of ethylene feedstock, and reduce the cost of ethylene production.

[0003] As a feedstock for ethylene plants, the higher the content of C2 components (ethylene and ethane) in the C2 refined gas, the higher the ethylene yield and the lower the energy consumption per unit of ethylene product. Therefore, the purity of the C2 components in the C2 refined gas product should be maximized, while the content of hydrogen, methane, and C3 and higher components should be minimized. Furthermore, the C3 components such as propylene and propane in the dry gas are also valuable resources. For example, dilute propylene can be purified to produce polymer-grade propylene, and propane can be used to produce propylene and ethylene through propane dehydrogenation or steam cracking. Therefore, efficient utilization of refinery dry gas requires maximizing the recovery rate of both C2 and C3 components.

[0004] Currently, the most commonly used methods for dry gas recovery in refineries in China are pressure swing adsorption (PSA) and oil absorption, each with its own characteristics. PSA utilizes the different adsorption selectivity of the adsorbent for each component in a gas mixture and the characteristic that the adsorbent's adsorption capacity varies with adsorption pressure to separate different components. This method can operate at room temperature and has low energy consumption, but it requires large equipment and has relatively low C2 recovery rates and product purity. Oil absorption primarily utilizes the different solubilities of the absorbent for each component in a gas mixture to achieve separation. Generally, the absorbent first absorbs C2 and heavier components, separating non-condensable gases such as methane and hydrogen. Then, distillation is used to separate the individual components from the absorbent. This method is characterized by small scale, strong adaptability, and low investment costs, making it one of the most competitive dry gas recovery technologies currently available for refineries.

[0005] CN101759518A discloses a method for recovering refinery catalytic dry gas using shallow-cooled oil absorption technology. The absorption temperature is increased to shallow cooling (5-15℃), thus eliminating the need for a propylene chiller, drying of the feedstock, and decarbonization. The process is simple and achieves a high recovery rate of C2 components. A re-absorption tower is also included to recover the absorbent entrained in the absorption tail gas. However, the enriched gas product obtained by this method is a mixture mainly composed of C2 and C3 components, with a low C2 component content.

[0006] US10052581B1 discloses a method for recovering ethylene cracking feedstock from FCC dry gas. This method involves first subjecting the FCC dry gas to amine washing for desulfurization and decarbonization, followed by hydrogenation for alkyne removal. After these refining processes, the FCC dry gas enters a pressure swing adsorption (PSA) tower to remove light components such as methane and hydrogen. Following a series of separation processes, the final product is an ethylene cracking feedstock primarily composed of ethane and ethylene. However, this method is complex because the PSA tower may operate below 0°C. To prevent freezing, the dry gas needs to undergo cooling, phase separation, and dehydration drying before entering the PSA tower, resulting in a relatively large investment in the equipment.

[0007] CN101063048A discloses a method for separating FCC dry gas using a medium-cooled oil absorption process. This process comprises steps such as compression, removal of acidic gases, drying and purification, absorption, desorption, cold energy recovery, and coarse separation. It can concentrate ethane and ethylene from the dry gas and has advantages such as low absorbent cost and low loss. However, this process requires cooling the dry gas to -30°C to -40°C, making it a medium-cooled separation process that requires propylene refrigerant. This results in a large cooling load, a complex process, and high investment and energy consumption.

[0008] CN109748771A discloses a method for recovering C2 components from refinery dry gas. This method operates at a shallow cooling temperature (5-15°C), does not require a propylene refrigeration system or a drying system, and produces a high C2 component content in the separated enriched gas. The process is simple and energy-efficient. However, due to the high C3 component content in the circulating absorbent, a significant amount of C3 components is lost in the tail gas at the top of the absorber, failing to effectively recover the C3 components from the dry gas.

[0009] CN109749780A discloses a method for recovering refinery dry gas, which also employs shallow cooling operating conditions and improves the purity of C2 components in the C2 enriched gas by adding an ethane tower. The enriched gas obtained using this method has a high C2 component content, and the content of C3 and higher components can be controlled below 1.5% (mol), while also exhibiting high C2 and C3 recovery rates. However, this method requires an additional enriched gas compressor to increase the feed pressure of the ethane tower due to the high ethane tower pressure, resulting in significant equipment investment. Furthermore, the added rotating equipment may lead to operational instability risks.

[0010] In summary, currently used industrial dry gas recovery methods generally separate enriched gas, mainly composed of a mixture of C2 and C3 components, from the dry gas. The concentration of C2 in the enriched gas is relatively low. To reduce the C3 content in the enriched gas, existing processes generally employ two methods. One method is to achieve the separation of C2 and C3 components by lowering the operating temperature, but this requires the addition of propylene refrigeration and decarbonization / drying systems, increasing cooling energy consumption and equipment investment, and making the process more complex. The other method is to maintain operation at a slightly cool temperature, which eliminates the need for propylene refrigeration and drying systems, but suffers from problems such as low C3 recovery rates or the need to add an enriched gas compressor. Summary of the Invention

[0011] The purpose of this invention is to address the problems of high investment, large cooling energy consumption, and low C3 component recovery rate in existing refinery dry gas recovery processes when separating and concentrating ethane, ethylene, propane, and propylene components. This invention provides a method, apparatus, and application for recovering C2 and C3 components from refinery dry gas. This method can increase the C2 component (ethane + ethylene) content in C2 concentrated gas to over 90% (mol) and the C3 component (propane + propylene) content in C3 concentrated gas to over 80% (mol). It achieves high recovery rates for both C2 and C3 components, eliminates the need for drying and dehydration facilities and propylene refrigeration compressors, and also eliminates the need for concentrated gas compressors, thus reducing equipment investment and energy consumption.

[0012] To achieve the above objectives, a first aspect of the present invention provides a method for recovering C2 and C3 components from refinery dry gas, characterized in that the method comprises:

[0013] (1) Compression and cooling: The dry gas from the refinery is compressed and cooled in sequence to obtain cooled compressed gas;

[0014] (2) Absorption: The cooled compressed gas is introduced into the middle of the absorption tower and makes a first countercurrent contact with the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower receives the first absorption tail gas and the bottom of the absorption tower receives the first absorption rich liquid.

[0015] (3) High-pressure decarbonization II: The first rich absorption liquid is passed into the middle of the high-pressure decarbonization II tower for the first separation. The top of the high-pressure decarbonization II tower is obtained as concentrated C2 gas, and the bottom of the high-pressure decarbonization II tower is obtained as the first liquid phase.

[0016] (4) Low-pressure decarbonization II: The first liquid phase is introduced into the upper part of the low-pressure decarbonization II tower for the second separation. The top of the low-pressure decarbonization II tower yields liquid phase C2 and C3 components, and the bottom of the low-pressure decarbonization II tower yields the second liquid phase. The liquid phase C2 and C3 components are returned to step (3) for recycling decarbonization.

[0017] (5) Desorption: The second liquid phase is introduced into the middle of the desorption tower for desorption. The top of the desorption tower is C3 enriched gas, and the bottom of the desorption tower is the third liquid phase. Part of the third liquid phase is returned to step (2) as a circulating absorbent.

[0018] A second aspect of the present invention provides an apparatus for recovering C2 and C3 components from dry gas, characterized in that the apparatus comprises: a compressor, a cooler, an absorption tower, a high-pressure decarbonization tower, a low-pressure decarbonization tower, and a desorption tower;

[0019] The compressor and the cooler are connected;

[0020] The middle part of the absorption tower is connected to the cooler, which is used to make a first countercurrent contact between the cooled compressed gas from the cooler and the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower discharges the first absorption tail gas, and the bottom of the absorption tower discharges the first absorption rich liquid.

[0021] The middle section of the high-pressure decarbonization tower is connected to the bottom of the absorption tower, and is used to perform a first separation on the first absorption rich liquid. The top of the high-pressure decarbonization tower discharges concentrated C2 gas, and the bottom of the high-pressure decarbonization tower discharges the first liquid phase.

[0022] The upper part of the low-pressure decarbonization tower is connected to the lower part of the high-pressure decarbonization tower, and is used to perform a second separation on the first liquid phase. The top of the low-pressure decarbonization tower discharges liquid phase C2 and C3 components, the bottom of the low-pressure decarbonization tower discharges the second liquid phase, and the liquid phase C2 and C3 components are returned to the high-pressure decarbonization tower for recycling and decarbonization.

[0023] The middle part of the desorption tower is connected to the bottom of the low-pressure decarbonization tower, and is used to desorb the second liquid phase. The top of the desorption tower discharges the C3 enriched gas, and the bottom of the desorption tower obtains the third liquid phase. Part of the third liquid phase is returned to the top of the absorption tower as a circulating absorbent.

[0024] A third aspect of the present invention provides an application of the above-described method and / or apparatus in refinery dry gas recovery.

[0025] Through the above technical solutions, the method, apparatus, and application for recovering C2 and C3 components from refinery dry gas provided by the present invention achieve the following beneficial effects:

[0026] (a) This invention can extract C2 and C3 components from various saturated or unsaturated refinery dry gases, with the C2 enriched gas containing no less than 90 mol% of C2 components (ethane + ethylene) and the C3 enriched gas containing no less than 80 mol% of C3 components (propane + propylene). Compared with the traditional dry gas recovery technology that mixes C2 and C3 components together, the resources of ethane, ethylene, propane and propylene in the dry gas are utilized more rationally and effectively.

[0027] (b) In this invention, a high-pressure and low-pressure decarbonization two-tower design is adopted, which can control the tower bottom temperature within 130°C while ensuring the separation accuracy requirements of C2 and C3 components. This avoids problems such as polymerization and coking of unsaturated hydrocarbons in refinery dry gas at high temperatures, and is suitable for the treatment of various saturated and unsaturated dry gases.

[0028] (c) The minimum operating temperature of the equipment is generally 5-20℃. It does not require propylene refrigeration, decarbonization and drying systems, and requires less investment, simple operation and low cooling load.

[0029] (d) In this invention, the separation of C2 and C3 components does not require the addition of a concentrated gas compressor, saving equipment investment and improving the stability of the device operation;

[0030] (d) In this invention, the recovery rate of C2 component (ethane + ethylene) is not less than 95% and the recovery rate of C3 component (propane + propylene) is not less than 95%, which is high and has good separation effect. Attached Figure Description

[0031] The above and other objects, features and advantages of the present invention will become more apparent from the more detailed description of exemplary embodiments of the invention in conjunction with the accompanying drawings, wherein the same reference numerals generally represent the same parts.

[0032] Figure 1 This is a schematic diagram of the process for extracting concentrated C2 and C3 components from refinery dry gas according to Embodiment 1 of the present invention.

[0033] Figure 2 This is a schematic diagram of the process for extracting concentrated C2 and C3 components from refinery dry gas according to Embodiment 2 of the present invention.

[0034] Explanation of reference numerals in the attached figures

[0035] 1. Refinery dry gas; 2. Compressor; 3. Cooler; 4. Absorber; 5. High-pressure decarbonization tower II; 6. Low-pressure decarbonization tower II; 7. Desorption tower; 8. Supplementary absorbent; 9. Light hydrocarbon extraction; 10. First absorption tail gas; 11. C2 enriched gas; 12. C3 enriched gas; 13. Circulating absorbent; 14. Reabsorption tower; 15. Redesorption tower; 16. Second absorption tail gas; 17. Third absorption tail gas; 18. Reabsorbent; 19. C4 components. Detailed Implementation

[0036] The endpoints of the ranges and any values ​​disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0037] In this invention, unless otherwise stated, the pressure is gauge pressure.

[0038] The first aspect of this invention provides a method for recovering C2 and C3 components from refinery dry gas, characterized in that the method comprises:

[0039] (1) Compression and cooling: The dry gas from the refinery is compressed and cooled in sequence to obtain cooled compressed gas;

[0040] (2) Absorption: The cooled compressed gas is introduced into the middle of the absorption tower and makes a first countercurrent contact with the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower receives the first absorption tail gas and the bottom of the absorption tower receives the first absorption rich liquid.

[0041] (3) High-pressure decarbonization II: The first rich absorption liquid is passed into the middle of the high-pressure decarbonization II tower for the first separation. The top of the high-pressure decarbonization II tower is obtained as concentrated C2 gas, and the bottom of the high-pressure decarbonization II tower is obtained as the first liquid phase.

[0042] (4) Low-pressure decarbonization II: The first liquid phase is introduced into the upper part of the low-pressure decarbonization II tower for the second separation. The top of the low-pressure decarbonization II tower yields liquid phase C2 and C3 components, and the bottom of the low-pressure decarbonization II tower yields the second liquid phase. The liquid phase C2 and C3 components are returned to step (3) for recycling decarbonization.

[0043] (5) Desorption: The second liquid phase is introduced into the middle of the desorption tower for desorption. The top of the desorption tower is C3 enriched gas, and the bottom of the desorption tower is the third liquid phase. Part of the third liquid phase is returned to step (2) as a circulating absorbent.

[0044] This invention employs two decarbonization towers, one high-pressure and one low-pressure, to avoid coking of unsaturated hydrocarbons due to excessively high tower bottom temperatures. It also achieves higher separation accuracy for C2 and C3 gases, making it suitable for the treatment and recovery of various saturated and unsaturated dry gases.

[0045] Furthermore, the method does not require decarbonization and drying, nor does it require compression and pressurization of the enriched gas. The liquid phase light component is returned to step (3) for recycling decarbonization, and the third liquid phase is returned to step (2) as a recycling absorbent, thereby maximizing resource utilization, saving costs, and reducing investment.

[0046] In this invention, there is no particular limitation on the refinery dry gas. It can be saturated refinery dry gas, such as atmospheric and vacuum distillation dry gas, hydrogenation dry gas, and reforming PSA tail gas; or it can be unsaturated refinery dry gas, such as catalytic cracking (FCC) dry gas and catalytic pyrolysis (DCC) dry gas.

[0047] According to the present invention, there is no particular limitation on the number of stages of dry gas compression treatment in refineries, but 2-3 stages are preferred.

[0048] According to the present invention, the compression increases the pressure of the refinery dry gas to 2-5 MPaG, preferably 3-4 MPaG.

[0049] Furthermore, the compression ensures that the temperature of the dry gas does not exceed 120°C;

[0050] According to the present invention, the cooling process results in the temperature of the compressed cold air being 5-20°C, preferably 10-15°C.

[0051] According to the present invention, the temperature of the absorbent is 5-20°C, preferably 10-15°C.

[0052] In this invention, there are no particular limitations on the absorbent, such as a refinery mixed C4 component, a post-etherified C4 component, or liquefied petroleum gas, preferably a C4 component mainly composed of butane and butene.

[0053] In this invention, there is no particular limitation on the amount of absorbent used; it can be determined by those skilled in the art based on their needs or common knowledge of the prior art.

[0054] According to the present invention, the number of theoretical plates in the absorption tower is 20-50, preferably 30-40.

[0055] According to the present invention, the operating pressure of the absorption tower is 2-5 MPaG, preferably 3-4 MPaG.

[0056] According to the present invention, the temperature at the top of the absorption tower is 10-30°C.

[0057] According to the present invention, the bottom temperature of the absorption tower is 100-130°C.

[0058] In this invention, the bottom temperature of the absorption tower is controlled at 100-130℃, which ensures that the light components such as methane and hydrogen in the bottom of the absorption tower are reduced to below 1 mol%.

[0059] According to the present invention, the theoretical number of plates in the high-pressure decarbonization tower is 20-50, preferably 30-40.

[0060] According to the present invention, the operating pressure of the high-pressure decarbonization tower is 2-4 MPaG, preferably 2.5-3.5 MPaG.

[0061] According to the present invention, the top temperature of the high-pressure decarbonization tower is 10℃-40℃, preferably 15-20℃.

[0062] According to the present invention, the bottom temperature of the high-pressure decarbonization tower is 100℃-130℃.

[0063] In this invention, by adjusting the pressure and top temperature of the high-pressure decarbonization tower within the aforementioned range, the content of C3 and above components in the top gas is controlled, thereby ensuring that the content of C2 components in the C2 enrichment gas is not less than 90 mol.

[0064] According to the present invention, the theoretical number of plates in the low-pressure decarbonization tower is 20-50, preferably 25-35.

[0065] According to the present invention, the operating pressure of the low-pressure decarbonization tower is 1-3 MPaG, preferably 1-2 MPaG.

[0066] According to the present invention, the top temperature of the low-pressure decarbonization tower is 10℃-50℃, preferably 15-40℃.

[0067] According to the present invention, the bottom temperature of the low-pressure decarbonization tower is 100℃-130℃.

[0068] In this invention, the operating pressure of the low-pressure tower is lower than that of the high-pressure tower.

[0069] In this invention, by adjusting the pressure and bottom temperature of the low-pressure decarbonization tower within the above-mentioned range, the content of C2 in the bottom liquid phase is controlled, thereby ensuring that the recovery rate of C2 components (ethane + ethylene) in the C2 enrichment gas is not less than 95%.

[0070] In this invention, a portion of the third liquid phase is returned to step (2) as a circulating absorbent, and a small portion is extracted as light hydrocarbons. Simultaneously, a small amount of fresh absorbent can be added as needed to the absorption tower to replace the heavy components in the system and maintain the balance of the C4 components within the system.

[0071] According to the present invention, the number of theoretical plates in the desorption tower is 20-50, preferably 30-40.

[0072] According to the present invention, the operating pressure of the desorption tower is 1-4 MPaG, preferably 1-2 MPaG.

[0073] According to the present invention, the top temperature of the desorption tower is 10℃-50℃, preferably 15-40℃.

[0074] According to the present invention, the bottom temperature of the desorption tower is 100℃-130℃.

[0075] In this invention, the separation accuracy of C3 and C4 components in the desorption tower is controlled within the above-mentioned range by adjusting the pressure and temperature of the desorption tower, thereby controlling the recovery rate of C3 component (propane + propylene) to be not less than 95 mol%, and the content of C3 component (propane + propylene) in the C3 concentrate at the top of the tower to be not less than 90 mol%.

[0076] In this invention, in step (2), due to phase equilibrium, the first absorption tail gas at the top of the absorption tower will contain a portion of C4 absorbent. A reabsorption tower and a redesorption tower can be set up to recover this portion of C4 component.

[0077] According to the present invention, the method further includes step (6) reabsorption: the first absorption tail gas is introduced into the bottom of the reabsorption tower and makes a second countercurrent contact with the reabsorbent from the top of the reabsorption tower, the top of the reabsorption tower receives the second absorption tail gas, and the bottom of the reabsorption tower receives the second absorption rich liquid.

[0078] According to the present invention, the method further includes step (7) re-desorption: the second absorption rich liquid is introduced into the middle of the re-desorption tower, and after analysis, it is condensed and separated. The third absorption tail gas is obtained at the top of the re-desorption tower, the C4 component is obtained at the top of the re-desorption tower, and the fourth liquid phase is obtained at the bottom of the re-desorption tower.

[0079] In this invention, the C4 component is returned to step (2) as an absorbent for recycling, and the fourth liquid phase is returned to step (6) as a reabsorbent.

[0080] In this invention, the absorbent is a mixture of the supplementary absorbent, the circulating absorbent, and the C4 component obtained at the top of the re-desorption tower.

[0081] In this invention, there is no particular limitation on the reabsorbent, which may be, for example, stabilized gasoline, heavy naphtha, aromatic raffinate and refinery C7 components. Preferably, the reabsorbent is stabilized gasoline or heavy naphtha.

[0082] In this invention, there is no particular limitation on the amount of the reabsorbent, which can be determined by those skilled in the art based on their needs or common knowledge of the prior art.

[0083] According to the present invention, the theoretical number of plates in the reabsorption tower is 15-25, preferably 18-20.

[0084] According to the present invention, the operating pressure of the reabsorption tower is 2-4 MPaG, preferably 2.8-3.8 MPaG.

[0085] According to the present invention, the top temperature of the reabsorption tower is 10-30°C.

[0086] According to the present invention, the bottom temperature of the reabsorption tower is 20-60°C.

[0087] According to the present invention, the theoretical plate number of the re-desorption tower is 25-35, preferably 28-30;

[0088] The operating pressure of the reabsorption tower is 0.5-1.5 MPaG, preferably 0.6-1.0 MPaG;

[0089] The top temperature of the re-desorption tower is 40-80℃, preferably 50-60℃;

[0090] The bottom temperature of the re-desorption tower is 130-220℃, preferably 140-160℃.

[0091] A second aspect of the present invention provides an apparatus for recovering C2 and C3 components from refinery dry gas, characterized in that the apparatus comprises: a compressor, a cooler, an absorption tower, a high-pressure decarbonization tower, a low-pressure decarbonization tower, and a desorption tower;

[0092] The compressor and the cooler are connected;

[0093] The middle part of the absorption tower is connected to the cooler, which is used to make a first countercurrent contact between the cooled compressed gas from the cooler and the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower discharges the first absorption tail gas, and the bottom of the absorption tower discharges the first absorption rich liquid.

[0094] The middle part of the high-pressure decarbonization tower is connected to the bottom of the absorption tower, and is used to perform the first separation of the first absorption rich liquid. The top of the high-pressure decarbonization tower discharges the C2 enriched gas, and the bottom of the high-pressure decarbonization tower discharges the first liquid phase.

[0095] The upper part of the low-pressure decarbonization tower is connected to the lower part of the high-pressure decarbonization tower, and is used to perform a second separation on the first liquid phase. The top of the low-pressure decarbonization tower discharges liquid phase C2 and C3 components, the bottom of the low-pressure decarbonization tower discharges the second liquid phase, and the liquid phase C2 and C3 components are returned to the high-pressure decarbonization tower for recycling and decarbonization.

[0096] The middle part of the desorption tower is connected to the bottom of the low-pressure decarbonization tower, and is used to desorb the second liquid phase. The top of the desorption tower discharges the C3 enriched gas, and the bottom of the desorption tower obtains the third liquid phase. Part of the third liquid phase is returned to the top of the absorption tower as a circulating absorbent.

[0097] In this invention, the separation of C2 and C3 components does not require a propylene refrigeration system or decarbonization and drying processes, nor does it require an additional enrichment gas compressor, thus saving equipment investment, reducing energy consumption, and improving the stability of the device operation.

[0098] According to the present invention, the apparatus further includes a reabsorption tower and a redesorption tower;

[0099] The lower part of the reabsorption tower is connected to the top of the absorption tower, and is used to make the first absorption tail gas and the reabsorbent from the upper part of the reabsorption tower in a second countercurrent contact. The second absorption tail gas is discharged from the top of the reabsorption tower, and the second absorption rich liquid is discharged from the bottom of the reabsorption tower.

[0100] The middle part of the reabsorption tower is connected to the bottom of the reabsorption tower, which is used to desorb the second rich absorption liquid and then condense and separate it. The third absorption tail gas is obtained at the top of the reabsorption tower, the C4 component is discharged from the upper part of the reabsorption tower, and the fourth liquid phase is discharged from the bottom of the reabsorption tower.

[0101] In this invention, the C4 component is returned to the top of the absorption tower as an absorbent, and the fourth liquid phase is returned to the top of the reabsorption tower as a reabsorbent.

[0102] In this invention, the top of the high-pressure decarbonization tower, the low-pressure decarbonization tower, the desorption tower, and the re-desorption tower also includes a condensation reflux tank for reducing the temperature at the top of the tower.

[0103] A third aspect of the present invention provides an application of the above-described method or apparatus in the recovery of dry gas in a refinery.

[0104] Combination Figure 1 The method and apparatus described in this invention will be further explained below.

[0105] The refinery dry gas 1 is compressed in 2-3 stages by compressor 2 to increase the pressure to 2-5 MPaG. The temperature of the compressed dry gas does not exceed 120℃. The compressed dry gas is then cooled to 5-20℃ by cooler 3 with 7℃ low temperature water before being sent to the middle of absorber 4.

[0106] The compressed and cooled gas comes into first countercurrent contact with the absorbent from the top of the absorption tower 4, which has been cooled to 5-20°C by water at 7°C. The first absorption tail gas 10 obtained from the top of the absorption tower 4 is sent to the fuel gas pipeline or hydrogen recovery system, or for other uses. The first absorption rich liquid obtained from the bottom of the absorption tower 4 is sent to the middle of the high-pressure decarbonization tower 5.

[0107] The first rich absorbent solution is fed into the middle of the high-pressure decarbonization tower 5 for the first separation. The top of the high-pressure decarbonization tower 5 yields concentrated C2 gas 11, and the bottom of the high-pressure decarbonization tower 5 yields the first liquid phase, which is then fed into the low-pressure decarbonization tower 5.

[0108] The first liquid phase is subjected to a second separation in the low-pressure decarbonization tower 5. The liquid phase C2 and C3 obtained from the top of the low-pressure decarbonization tower are returned to the recirculation decarbonization tower in step (3), and the second liquid phase obtained from the bottom of the low-pressure decarbonization tower is sent to the desorption tower.

[0109] The second liquid phase is analyzed in the desorption tower 7. C3 enriched gas 12 is obtained at the top of the desorption tower 7, and the third liquid phase is obtained at the bottom of the desorption tower 7. Among them, part of the third liquid phase is returned to step (2) as a circulating absorbent 13, and a small part of the third liquid phase is extracted as light hydrocarbon 9.

[0110] Combination Figure 2 The method and apparatus described in this invention will be further preferred in the following description.

[0111] As mentioned above, the first absorption tail gas 10 is sent to a fuel gas pipeline network, a hydrogen recovery system, or for other uses. The first absorption tail gas 10 will contain some C4 absorbent, which can be recovered by setting up a reabsorption tower 14 and a redesorption tower 15.

[0112] The first absorption tail gas 10 is introduced into the bottom of the reabsorption tower and comes into a second countercurrent contact with the reabsorbent 18 from the top of the reabsorption tower 14. The second absorption tail gas 16 obtained from the top of the reabsorption tower 14 is sent to the fuel gas pipeline or hydrogen recovery system or for other uses. The second absorption rich liquid obtained from the bottom of the reabsorption tower 14 is sent to the middle of the redesorption tower 15.

[0113] After the second rich absorbent is re-desorbed in the re-desorption tower 15, it is condensed and separated. The third absorbent tail gas 17 obtained from the top of the re-desorption tower 15 is sent to the fuel gas pipeline network, and the fourth liquid phase is obtained from the bottom of the re-desorption tower 15.

[0114] The top of the re-desorption tower 15 yields C4 component 19, which is returned to step (2) as an absorbent.

[0115] The fourth liquid phase is returned to step (6) as reabsorbent 18.

[0116] The present invention will be described in detail below through embodiments. In the following embodiments, the dry gas involved is a mixture of catalytic dry gas, reformed PSA tail gas and disproportionated fuel gas from a certain refinery. The specific dry gas composition is shown in Table 1.

[0117] Table 1

[0118]

[0119]

[0120] Example 1

[0121] Adopting such Figure 1 The schematic diagram shown illustrates the process of extracting concentrated C2 and C3 components from refinery dry gas.

[0122] The device includes a compressor 2, a cooler 3, an absorption tower 4, a high-pressure decarbonization tower 5, a low-pressure decarbonization tower 6, and a desorption tower 7.

[0123] The compressor 2 and the cooler 3 are connected.

[0124] The refinery dry gas 1 is compressed to 4 MPaG by the compressor 2 in three stages. The temperature of the compressed dry gas is 95.3℃. The compressed dry gas is then cooled to 15℃ by the cooler 3 with 7℃ low temperature water before being sent to the middle of the absorption tower 4.

[0125] The circulating absorbent 13 from the bottom of desorption tower 7 is cooled to 15°C with 7°C low-temperature water and then injected from the top of the absorption tower by a pump, making first countercurrent contact with the compressed and cooled gas. The absorption tower 4 has a theoretical plate number of 39, an operating pressure of 3.85 MPaG, a top temperature of 19.3°C, and a bottom temperature of 110.8°C. The top of the absorption tower 4 receives the first absorption tail gas 10 (mainly unabsorbed methane, hydrogen, and a small amount of entrained absorbent) which is discharged to the fuel gas pipeline. The bottom of the absorption tower 4 receives the first absorption rich liquid, which is sent to the middle of the high-pressure decarbonization tower 5.

[0126] The first rich absorbent liquid from the bottom of the absorption tower 4 enters the high-pressure decarbonization tower 5 for the first separation by pressure difference. The high-pressure decarbonization tower 5 has a theoretical plate number of 38, an operating pressure of 3.4 MPaG, a top temperature of 16.5℃, and a bottom temperature of 120.5℃. The top of the high-pressure decarbonization tower 5 yields C2 enriched gas 11, and the bottom of the high-pressure decarbonization tower 5 yields the first liquid phase, which is sent to the middle of the low-pressure decarbonization tower 6.

[0127] In this embodiment, the reflux condenser temperature at the top of the high-pressure decarbonization tower is controlled at 11.3℃, which ensures that the C2 component (ethane + ethylene) content in the concentrated C2 gas is not less than 90 mol%.

[0128] In the high-pressure decarbonization tower 5, the first liquid phase still contains some C2 components. It enters the low-pressure decarbonization tower 6 via pressure difference for further separation of C2 and C3 components. The low-pressure decarbonization tower 6 has a theoretical plate number of 28, an operating pressure of 1.65 MPaG, a top temperature of 43.8℃, and a bottom temperature of 111.3℃. The C2 and C3 components separated at the top of the low-pressure decarbonization tower 6 are returned to the high-pressure decarbonization tower 5 for cyclic decarbonization, while the second liquid phase enters the middle section of the desorption tower 7.

[0129] By adjusting the temperature of the bottom of column 6 of the low-pressure decarbonization tower to 117℃, the content of C2 components (ethane + ethylene) in the second liquid phase collected from the bottom of the tower can be controlled to within 0.5 mol%, ensuring that the recovery rate of C2 components (ethane + ethylene) is not less than 95%.

[0130] The second liquid phase is desorbed in desorption tower 7, which has a theoretical plate number of 38, an operating pressure of 1.3 MPaG, a top temperature of 47.1°C, and a bottom temperature of 111.8°C. The top of the desorption tower 7 yields C3 enriched gas 13, and the bottom of the desorption tower 7 yields the third liquid phase. A portion of the third liquid phase is returned to step (2) as a circulating absorbent 13, and a small portion of the third liquid phase is extracted as light hydrocarbons 9.

[0131] In this embodiment, the reflux temperature at the top of the desorption tower 7 is controlled at 40°C, which ensures that the content of C3 components (propane + propylene) in the C3 enrichment gas is not less than 80 mol%. The outlet temperature of the reboiler at the bottom of the desorption tower 7 is controlled at 116°C, which can control the content of C3 components in the liquid phase collected from the bottom of the tower to within 4 mol%, and can ensure that the recovery rate of C3 components (propane + propylene) is not less than 95%.

[0132] A small portion of light hydrocarbons is extracted from the third liquid phase collected from the bottom of desorption tower 7 to balance the heavy components in the system. At the same time, a small amount of fresh absorbent 8 is added, i.e., supplementary absorbent, and the amount of supplementary absorbent added is 7333.7 kg / h.

[0133] In this embodiment, the main components and flow rates of the inlet and outlet devices are shown in Table 2. The C2 enrichment gas contains 92.60 mol% of C2 components (ethylene + ethane), and the C3 enrichment gas contains 82.62 mol% of C3 components (propane + propylene). The recovery rate of C2 components (ethane + ethylene) is 96.1%, and the recovery rate of C3 components (propane + propylene) is 95.8%.

[0134] Table 2

[0135]

[0136] Example 2

[0137] Use Figure 2The schematic diagram shown illustrates the process of extracting concentrated ethane, ethylene, propane, and propylene components from refinery dry gas.

[0138] The device includes a compressor 2, a cooler 3, an absorption tower 4, a high-pressure decarbonization tower 5, a low-pressure decarbonization tower 6, a desorption tower 7, a reabsorption tower 14, and a re-desorption tower 15.

[0139] The compressor 2 and the cooler 3 are connected.

[0140] The refinery dry gas 1 is compressed to 4 MPaG by the compressor 2 in three stages. The temperature of the compressed dry gas is 95.3℃. The compressed dry gas is then cooled to 15℃ by the cooler 3 with 7℃ low temperature water before being sent to the middle of the absorption tower 4.

[0141] The circulating absorbent 13 from the bottom of desorption tower 7 is cooled to 15°C with 7°C low-temperature water and then injected from the top of the absorption tower by a pump, making first countercurrent contact with the compressed and cooled gas. The absorption tower 4 has a theoretical plate number of 39, an operating pressure of 3.85 MPaG, a top temperature of 19.4°C, and a bottom temperature of 111°C. The top of the absorption tower 4 receives the first absorption tail gas 10 (mainly unabsorbed methane, hydrogen, and a small amount of entrained absorbent) which is discharged to the fuel gas pipeline. The bottom of the absorption tower 4 receives the first absorption rich liquid, which is sent to the middle of the high-pressure decarbonization tower 5.

[0142] The first rich absorbent liquid from the bottom of the absorption tower 4 enters the high-pressure decarbonization tower 5 for the first separation by means of pressure difference. The high-pressure decarbonization tower 5 has a theoretical plate number of 38, an operating pressure of 3.4 MPaG, a top temperature of 16.5℃, and a bottom temperature of 120.4℃. The top of the high-pressure decarbonization tower 5 yields C2 enriched gas 11, and the bottom of the high-pressure decarbonization tower 5 yields the first liquid phase, which is sent to the middle of the low-pressure decarbonization tower 6.

[0143] In this embodiment, the top condenser reflux temperature of the high-pressure decarbonization tower 5 is controlled at 11.3℃, which ensures that the content of C2 components (ethane + ethylene) in the C2 enrichment gas 11 is not less than 90 mol.

[0144] In the high-pressure decarbonization tower 5, the first liquid phase still contains some C2 components, which enter the low-pressure decarbonization tower 6 for further separation of C2 and C3 components based on the pressure difference. The high-pressure decarbonization tower 5 has a theoretical plate number of 28, an operating pressure of 1.65 MPaG, a tower top temperature of 42.4℃, and a tower bottom temperature of 112℃. The C2 and some C3 components separated at the top of the low-pressure decarbonization tower 6 are returned to the high-pressure decarbonization tower 5 for cyclic decarbonization, while the second liquid phase enters the middle of the desorption tower 7.

[0145] By adjusting the temperature of the bottom of column 6 of the low-pressure decarbonization tower to 118.2℃, the content of C2 components (ethane + ethylene) in the first liquid phase collected from the bottom of the tower can be controlled within 0.5 mol%, so as to ensure that the recovery rate of C2 components (ethane + ethylene) is not less than 95%.

[0146] The second liquid phase is desorbed in desorption tower 7, which has a theoretical plate number of 38, an operating pressure of 1.3 MPaG, a top temperature of 47°C, and a bottom temperature of 112.9°C. The top of the desorption tower 7 yields C3 enriched gas 13, and the bottom of the desorption tower 7 yields the third liquid phase. A portion of the third liquid phase is returned to step (2) as a circulating absorbent 13, and a small portion of the third liquid phase is extracted as light hydrocarbons 9.

[0147] In this embodiment, the reflux temperature at the top of the desorption tower 7 is controlled at 40°C, which ensures that the content of C3 components (propane + propylene) in the C3 enrichment gas is not less than 80 mol%. The outlet temperature of the reboiler at the bottom of the desorption tower 7 is controlled at 117.6°C, which can control the content of C3 components in the liquid phase collected from the bottom of the tower to within 5 mol%, so as to ensure that the recovery rate of C3 components (propane + propylene) is not less than 95%.

[0148] A small portion of light hydrocarbons is extracted from the third liquid phase collected from the bottom of desorption tower 7 to balance the heavy components in the system. At the same time, a small amount of fresh absorbent 8 is added, i.e., supplementary absorbent, and the amount of supplementary absorbent added is 3133.6 kg / h.

[0149] The first absorption tail gas 10 carries a portion of C4 absorbent, which is introduced into the bottom of the reabsorption tower 14. There, it undergoes a second countercurrent contact with the reabsorbent 18 from the top of the reabsorption tower 14, which has been cooled to 15°C by water at 7°C. The reabsorption tower has a theoretical plate count of 20, an operating pressure of 3.75 MPaG, a top temperature of 23°C, and a bottom temperature of 27.4°C. The second absorption tail gas is discharged from the top of the reabsorption tower and fed into the fuel gas pipeline. The second rich absorbent solution discharged from the bottom of the reabsorption tower is sent to the middle section of the re-desorption tower 15.

[0150] The second rich absorbent solution enters the middle of the re-desorption tower 15 under pressure difference for re-desorption, and then undergoes condensation and separation. The re-desorption tower 15 has a theoretical plate number of 26, an operating pressure of 0.67 MPaG, a top temperature of 57.7℃, and a bottom temperature of 146.1℃. The top of the re-desorption tower 15 produces a third absorbent tail gas 17 which is discharged into the fuel gas pipeline. The upper part of the re-desorption tower 15 produces C4 component 19, and the bottom of the re-desorption tower 15 produces a fourth liquid phase. The C4 component 19 is returned to step (2) as an absorbent, and the fourth liquid phase is returned to step (6) as a reabsorbent 18.

[0151] In this embodiment, the main components and flow rates of the inlet and outlet devices are shown in Table 3. The C2 enrichment gas contains 92.61 mol% of C2 components (ethylene + ethane), and the C3 enrichment gas contains 82.61 mol% of C3 components (propylene + propane). The recovery rate of C2 components (ethylene + ethane) is 95.8%, and the recovery rate of C3 components (propylene + propane) is 95.7%.

[0152] Table 3

[0153]

[0154] Comparing Tables 2 and 3, it can be seen that the C2 and C3 enriched gas products in Examples 1-5 have higher content and recovery rates. The C2 enriched gas contains more than 93.01 mol% of the C2 component (ethane + ethylene), and the C3 enriched gas contains more than 82.62 mol% of the C3 component (propane + propylene). The recovery rate of the C2 component (ethane + ethylene) is more than 96.1%, and the recovery rate of the C3 component (propane + propylene) is more than 95.8%. However, because Example 2 added a reabsorption-desorption process to recover the C4 absorbent in the tail gas, the supplementary flow rate of n-butane was reduced by 4200.1 kg / h.

[0155] This invention eliminates the need for propylene refrigerant and additional enrichment compressors. While reducing cooling energy consumption and saving investment, it separates C2 and C3 enriched gases from refinery dry gas. The purity of the C2 component (ethane + ethylene) in the C2 enriched gas is not less than 90%, and the purity of the C3 component (propane + propylene) in the C3 enriched gas is not less than 80%. The recovery rates of both C2 (ethane + ethylene) and C3 (propane + propylene) components are not less than 95%, achieving a more rational and efficient utilization of refinery dry gas and possessing broad industrial application prospects.

[0156] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. A method for recovering C2 and C3 components from refinery dry gas, characterized in that, The method includes: (1) Compression and cooling: The dry gas from the refinery is compressed and cooled in sequence to obtain cooled compressed gas; (2) Absorption: The cooled compressed gas is introduced into the middle of the absorption tower and makes a first countercurrent contact with the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower receives the first absorption tail gas and the bottom of the absorption tower receives the first absorption rich liquid. (3) High-pressure decarbonization II: The first rich absorption liquid is passed into the middle of the high-pressure decarbonization II tower for the first separation. The top of the high-pressure decarbonization II tower is obtained as concentrated C2 gas, and the bottom of the high-pressure decarbonization II tower is obtained as the first liquid phase. (4) Low-pressure decarbonization II: The first liquid phase is introduced into the upper part of the low-pressure decarbonization II tower for the second separation. The top of the low-pressure decarbonization II tower yields liquid phase C2 and C3 components, and the bottom of the low-pressure decarbonization II tower yields the second liquid phase. The liquid phase C2 and C3 components are returned to step (3) for recycling decarbonization II. (5) Desorption: The second liquid phase is introduced into the middle of the desorption tower for desorption. The top of the desorption tower is C3 enriched gas, and the bottom of the desorption tower is the third liquid phase. Part of the third liquid phase is returned to step (2) as a circulating absorbent. The operating pressure of the high-pressure decarbonization tower is 3.4-4 MPaG; The top temperature of the high-pressure decarbonization tower is 16.5℃-40℃; The bottom temperature of the high-pressure decarbonization tower is 120.4℃-130℃; The operating pressure of the low-pressure decarbonization tower is 1.65-3 MPaG; The top temperature of the low-pressure decarbonization tower is 42.4℃-50℃; The bottom temperature of the low-pressure decarbonization tower is 111.3℃-130℃.

2. The method according to claim 1, wherein, The compression pressurizes the refinery dry gas to 2-5 MPaG; the compression ensures that the temperature of the dry gas does not exceed 120°C.

3. The method according to claim 2, wherein, The compression pressurizes the refinery dry gas to 3-4 MPaG.

4. The method according to claim 1, wherein, The cooling process keeps the temperature of the compressed cold air at 5-20°C.

5. The method according to claim 4, wherein, The cooling process brings the temperature of the compressed cold air to 10-15°C.

6. The method according to claim 1, wherein, The temperature of the absorbent is 5-20℃.

7. The method according to claim 6, wherein, The temperature of the absorbent is 10-15℃.

8. The method according to any one of claims 1-7, wherein, The theoretical plate number of the absorption tower is 20-50; The operating pressure of the absorption tower is 2-5 MPaG; The temperature at the top of the absorption tower is 10-30℃; The bottom temperature of the absorption tower is 100-130℃.

9. The method according to claim 8, wherein, The theoretical plate number of the absorption tower is 30-40; The operating pressure of the absorption tower is 3-4 MPaG.

10. The method according to any one of claims 1-7, wherein, The theoretical number of plates in the high-pressure decarbonization tower is 20-50.

11. The method according to claim 10, wherein, The theoretical number of plates in the high-pressure decarbonization tower is 30-40.

12. The method according to any one of claims 1-7, wherein, The theoretical number of plates in the low-pressure decarbonization tower is 20-50.

13. The method according to claim 12, wherein, The theoretical number of plates in the low-pressure decarbonization tower is 25-35.

14. The method according to any one of claims 1-7, wherein, The theoretical plate number of the desorption tower is 20-50; The operating pressure of the desorption tower is 1-4 MPaG; The temperature at the top of the desorption tower is 10℃-50℃; The bottom temperature of the desorption tower is 100℃-130℃.

15. The method according to claim 14, wherein, The theoretical plate number of the desorption tower is 30-40; The operating pressure of the desorption tower is 1-2 MPaG; The temperature at the top of the desorption tower is 15-40℃.

16. The method according to any one of claims 1-7, wherein, The method further includes: (6) Reabsorption: The first absorption tail gas is introduced into the bottom of the reabsorption tower and comes into a second countercurrent contact with the reabsorbent from the top of the reabsorption tower. The top of the reabsorption tower receives the second absorption tail gas, and the bottom of the reabsorption tower receives the second absorption rich liquid. (7) Re-desorption: The second absorption rich liquid is introduced into the middle of the re-desorption tower, and after analysis, it is condensed and separated. The third absorption tail gas is obtained at the top of the re-desorption tower, the C4 component is obtained at the top of the re-desorption tower, and the fourth liquid phase is obtained at the bottom of the re-desorption tower. The C4 component is returned to step (2) as an absorbent; The fourth liquid phase is returned to step (6) as a reabsorbent.

17. The method according to claim 16, wherein, The theoretical plate number of the reabsorption tower is 15-25; The operating pressure of the reabsorption tower is 2-4 MPaG; The top temperature of the reabsorption tower is 10-30℃; The bottom temperature of the reabsorption tower is 20-60℃.

18. The method according to claim 17, wherein, The theoretical number of plates in the reabsorption tower is 18-20; The operating pressure of the reabsorption tower is 2.8-3.8 MPaG.

19. The method according to claim 16 or 17, wherein, The theoretical plate number of the re-desorption tower is 25-35; The operating pressure of the re-desorption tower is 0.5-1.5 MPaG; The top temperature of the re-desorption tower is 40-80℃; The bottom temperature of the re-desorption tower is 130-220℃.

20. The method according to claim 19, wherein, The theoretical plate number of the re-desorption tower is 28-30; The operating pressure of the re-desorption tower is 0.6-1 MPaG; The top temperature of the re-desorption tower is 50-60℃; The bottom temperature of the re-desorption tower is 140-160℃.

21. An apparatus for recovering C2 and C3 components from refinery dry gas, characterized in that, The device includes: a compressor, a cooler, an absorption tower, a high-pressure decarbonization tower, a low-pressure decarbonization tower, and a desorption tower; The compressor and the cooler are connected; The middle part of the absorption tower is connected to the cooler, which is used to make a first countercurrent contact between the cooled compressed gas from the cooler and the absorbent from the top of the absorption tower to absorb C2 and heavier components. The top of the absorption tower discharges the first absorption tail gas, and the bottom of the absorption tower discharges the first absorption rich liquid. The middle part of the high-pressure decarbonization tower is connected to the bottom of the absorption tower, and is used to perform the first separation of the first absorption rich liquid. The top of the high-pressure decarbonization tower discharges the C2 enriched gas, and the bottom of the high-pressure decarbonization tower discharges the first liquid phase. The upper part of the low-pressure decarbonization tower is connected to the lower part of the high-pressure decarbonization tower, and is used to perform a second separation on the first liquid phase. The top of the low-pressure decarbonization tower discharges the second absorption tail gas, the bottom of the low-pressure decarbonization tower discharges the second liquid phase, and the second absorption tail gas is returned to the high-pressure decarbonization tower for recycling and decarbonization. The middle part of the desorption tower is connected to the bottom of the low-pressure decarbonization tower, and is used to desorb the second liquid phase. The top of the desorption tower discharges the C3 enriched gas, and the bottom of the desorption tower obtains the third liquid phase. Part of the third liquid phase is returned to the top of the absorption tower as a circulating absorbent.

22. The apparatus according to claim 21, wherein, The apparatus further includes a reabsorption tower and a redesorption tower; The lower part of the reabsorption tower is connected to the top of the absorption tower, and is used to make the first absorption tail gas and the reabsorbent from the upper part of the reabsorption tower in a second countercurrent contact. The second absorption tail gas is discharged from the top of the reabsorption tower, and the second absorption rich liquid is discharged from the bottom of the reabsorption tower. The middle part of the reabsorption tower is connected to the bottom of the reabsorption tower, which is used to condense and separate the second rich absorption liquid after analysis. The third absorption tail gas is obtained at the top of the reabsorption tower, the C4 component is discharged from the upper part of the reabsorption tower, and the fourth liquid phase is discharged from the bottom of the reabsorption tower. The C4 component is returned to the top of the absorption tower as an absorbent; The fourth liquid phase is returned to the top of the reabsorption tower as a reabsorbent.

23. The application of the method according to any one of claims 1-20 and / or the apparatus according to claim 21 or 22 in refinery dry gas recovery.

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