Method, device and application of ethylene separation and recovery

Abstract

The present application relates to the technical field of ethylene production and recovery, and discloses a method and device for ethylene separation and recovery and application thereof.The method separates and separates the tail gas washing by gradually cooling the cracking gas, simultaneously increases the combination of high-pressure deethanization operation and C2 full fraction hydrogenation reaction, makes the component cutting more efficient and energy-saving, the reflux of the high-pressure deethanization tower makes the green oil removal more thorough, the ethylene in the tail gas is effectively controlled and reduced, the total refrigeration amount required can be reduced, the method does not need the traditional circulating ethylene to wash the green oil, avoids the back mixing of the circulating ethylene in the ethylene rectification process, reduces the fractionation load of the deethanization tower and the ethylene rectification tower, and the method can reduce the total refrigeration amount required, the ethylene recovery rate is high, the energy consumption is low, and the operation is stable and efficient.

Description

Technical Field

[0001] This invention relates to the field of ethylene production and recycling technology, specifically to a method, apparatus, and application for ethylene separation and recycling. Background Technology

[0002] Ethylene is one of the most important basic chemical raw materials in the world today. It is mainly produced by thermally cracking feedstocks such as ethane, propane, LPG, and naphtha to create a mixture of hydrocarbon gases containing ethylene. The hydrocarbon components are then separated and purified to obtain the ethylene product. Therefore, it is necessary to continuously develop advanced methods for ethylene production, recovery, separation, and purification to improve ethylene recovery rates and reduce energy consumption.

[0003] Traditional cryogenic separation units in ethylene plants mainly consist of a cold box and a demethanizer system. In the cold box, the cracked gas is cooled and condensed at low temperatures (down to -170°C), separating the heavier liquid phase and the lighter gas phase in a gas-liquid separator. The gas phase undergoes further condensation and separation multiple times, ultimately separating hydrogen, methane, and ethylene. The demethanizer system is crucial for cracked gas separation in the cryogenic process, separating the lighter components (methane and hydrogen) from C2 and higher-weight components. It must ensure both ethylene recovery and the purity of the C2 components. It is also one of the systems with the highest cooling load; pre-cooling of the feedstock for the demethanizer accounts for approximately 40% of the total cooling load, and the demethanizer itself accounts for approximately 12%, meaning the demethanizer system accounts for about 50% of the total cooling load. The demethanizer system is affected by many factors; if it operates poorly, it can result in high ethylene content in the tail gas, leading to ethylene loss.

[0004] A small amount of acetylene is produced during the steam cracking process to produce ethylene. To avoid the impact of acetylene on subsequent polymerization processes, a C2 selective hydrogenation process is commonly used to control the acetylene content below 1 ppmv and increase ethylene production. The rate-controlling steps for acetylene hydrogenation to ethylene are acetylene adsorption and the first step of acetylene hydrogenation; the rate-controlling step for hydrogenation to ethane is the second step of ethylene hydrogenation. High temperature and low hydrogen-to-acetylene ratio conditions are favorable for obtaining high ethylene selectivity. In the post-hydrogenation process, C2 product fraction hydrogenation is usually used. The gas phase separated after partial condensation in the deethaner top condenser is fed into the reactor. This gas phase contains a relatively high amount of acetylene (approximately 0.5-2.5%), and the high acetylene content increases the number of hydrogenation stages required.

[0005] Green oil is a byproduct of acetylene hydrogenation polymerization. The amount of green oil produced is related to the depth of hydrogenation. At higher reaction temperatures, the catalyst's ability to adsorb hydrogen weakens, allowing several acetylene molecules to adsorb simultaneously at the active sites, leading to polymerization and the formation of acetylene oligomers, thus significantly increasing the amount of green oil produced. Traditional green oil scrubbing towers use liquid ethylene from an ethylene distillation column to contact and countercurrently scrub the wet gas stream from the reactor. The ethylene liquid containing green oil at the bottom of the scrubbing tower is then returned to the de-ethanizer. While effective, this method typically uses 5-10% of the net stream for scrubbing, resulting in high operating costs.

[0006] In summary, current methods recover approximately 99.5% of the ethylene in the cracked gas during the cryogenic system separation process, while the remaining 0.5% of ethylene is carried by the hydrogen and methane mixture tail gas. The high acetylene concentration during C2 hydrogenation has several drawbacks. Firstly, it increases reactor temperature, significantly reduces catalyst selectivity, and results in a low conversion rate of acetylene to ethylene, leading to a large amount of green oil production. Secondly, it necessitates the use of recycled ethylene to wash the green oil, causing severe backmixing of the circulating materials within the ethylene product refining unit. This increases the system's propylene refrigeration load demand, resulting in high overall energy consumption for the entire unit. Summary of the Invention

[0007] The purpose of this invention is to overcome the problems in existing ethylene separation and recovery technologies, such as high acetylene concentration during C2 hydrogenation. High acetylene concentration during C2 hydrogenation leads to increased reactor temperature, significantly reduced catalyst selectivity, low acetylene-to-ethylene conversion, and high green oil production. Furthermore, washing green oil with recycled ethylene easily causes severe backmixing of circulating materials within the ethylene product refining unit, increasing the system's propylene refrigeration load and resulting in high overall energy consumption. This invention provides a method, apparatus, and application for ethylene separation and recovery. This method involves staged cooling and separation of cracked gas, followed by tail gas washing for gas-liquid separation. It also combines high-pressure deethanerization with C2 full-fraction hydrogenation, resulting in more efficient and energy-saving component separation. The reflux flow of the high-pressure deethanerization tower ensures more thorough green oil removal, effectively controlling and reducing ethylene in the tail gas. This also reduces the required total refrigeration capacity. Furthermore, this method eliminates the need for traditional recycled ethylene to wash green oil, avoiding backmixing of recycled ethylene during ethylene distillation and reducing the fractionation load of the deethanerization and ethylene distillation towers. This method achieves an ethylene recovery rate of ≥99.94%, an improvement of at least 0.26% compared to traditional ethylene separation processes. The total refrigeration load is reduced by at least 3% compared to conventional separation processes. In the long run, the ethylene recovery rate is higher, the operating energy consumption is lower, and the operating cost is more competitive.

[0008] To achieve the above objectives, a first aspect of the present invention provides a method for ethylene separation and recovery, wherein the method includes:

[0009] S1. The cracked gas is subjected to C4 separation to obtain gaseous material A and liquid material A, which are C4- and above heavy components. After cooling, gaseous material A, which is C4- and above heavy components, is subjected to C3 separation to obtain gaseous material B and liquid material B, which are C3- and above heavy components. After cooling, gaseous material B, which is C3- and above heavy components, is subjected to C2 recovery to obtain gaseous material C and liquid material C, which are C2- and above heavy components. After cooling, gaseous material C, which is C2- and above heavy components, is subjected to methane / hydrogen first separation to obtain gaseous material D, which is a mixture of C1 and hydrogen, and liquid material D. After cooling, gaseous material D, which is a mixture of C1 and hydrogen, is subjected to methane / hydrogen second separation to obtain gaseous crude hydrogen material E and liquid material E.

[0010] The liquid phase material D is divided into a first liquid phase material D and a second liquid phase material D. The first liquid phase material D is sent to the C2 recovery operation as a refrigerant, and the second liquid phase material D is sent to the gas-liquid separation operation as a washing liquid.

[0011] S2. Perform methane stripping on liquid phase material A and liquid phase material B to obtain C3 and above heavy component gas phase material F and C1 and below light component liquid phase material F; perform demethanizing operation on liquid phase material C and C3 and above heavy component gas phase material F to obtain C2 and above heavy component gas phase material G and liquid C2 material G; perform gas-liquid separation operation on C2 and above heavy component gas phase material G to obtain C2 and above heavy component gas phase material I and liquid phase material I, and reflux liquid phase material I for demethanizing operation; perform high-pressure deethane removal operation on liquid phase material F and below light component liquid phase material F to obtain mixed C2 gas phase material H and liquid material H; perform C2 full-fraction hydrogenation reaction on mixed C2 gas phase material H to remove alkynes to obtain material J, and perform distillation on material J to obtain ethylene and ethane.

[0012] A second aspect of the present invention provides an apparatus for ethylene separation and recovery, wherein the apparatus comprises: a C3 separation device 1, a stripping tower 2, a C2 recovery tower 3, a demethanizer 4, a tail gas scrubbing tower 5, a first methane / hydrogen separation device 6, a high-pressure deethaner 8, a hydrogenation reaction system 10, a distillation tower 12, a first feed separation tank 13, a second methane / hydrogen separation device 14, and a cold box 15;

[0013] The top outlet of the first feed separator 13 is connected to the inlet of the C3 separator 1, the top outlet of the C3 separator 1 is connected to the first inlet of the C2 recovery tower 3, the first outlet of the C2 recovery tower 3 is connected to the inlet of the first methane / hydrogen separator 6, and the top outlet of the first methane / hydrogen separator 6 is connected to the inlet of the second methane / hydrogen separator 14, for the purpose of separating the cracked gas in stages.

[0014] The bottom outlet of the methane / hydrogen first separation device 6 is provided with a first splitter and a second splitter. The first splitter is connected to the third inlet of the C2 recovery tower 3, and the second splitter is connected to the first inlet of the tail gas scrubbing tower 5.

[0015] The first outlet of the demethanizer 4 is connected to the second inlet of the tail gas scrubbing tower 5;

[0016] The bottom outlet of the first feed separator 13 is connected to the first inlet of the stripping tower 2, the bottom outlet of the C3 separator 1 is connected to the second inlet of the stripping tower 2, the first outlet of the stripping tower 2 is connected to the second inlet of the demethanizer 4, and the second outlet of the C2 recovery tower 3 is connected to the first inlet of the demethanizer 4.

[0017] The second outlet of the stripping tower 2 is connected to the first inlet of the high-pressure deethanizer 8, the first outlet of the high-pressure deethanizer 8 is connected to the inlet of the hydrogenation reaction system 10, and the outlet of the hydrogenation reaction system 10 is connected to the inlet of the distillation tower 12.

[0018] The third aspect of the present invention provides an application of the method described in the first aspect or the apparatus described in the second aspect in ethylene separation and recovery.

[0019] Through the above technical solutions, the ethylene separation and recovery method, apparatus, and application provided by this invention achieve the following beneficial effects: the cracked gas is cooled and separated in stages, and the tail gas is washed for gas-liquid separation. At the same time, the combination of high-pressure deethanerization and C2 full-fraction hydrogenation reaction is added, making component cutting more efficient and energy-saving. Meanwhile, the reflux flow of the high-pressure deethanerization tower makes the green oil removal more thorough, so that the ethylene in the tail gas is effectively controlled and reduced. At the same time, it can reduce the required total refrigeration capacity. Furthermore, this method does not require traditional recycled ethylene to wash the green oil, avoiding backmixing of recycled ethylene during ethylene distillation, reducing the fractionation load of the deethanerization tower and the ethylene distillation tower. This method can reduce the required total refrigeration capacity, has a high ethylene recovery rate, low energy consumption, and stable and efficient operation. Attached Figure Description

[0020] Figure 1 This is a schematic diagram of the ethylene separation process using the low-pressure method in Comparative Example 1.

[0021] Figure 2 This is a schematic diagram of an ethylene separation and recovery device provided in Embodiment 1 of the present invention;

[0022] Figure 3 This is a schematic diagram of an ethylene separation and recovery device provided in Embodiment 2 of the present invention;

[0023] Figure 4This is a schematic diagram of an ethylene separation and recovery device provided in Example 3;

[0024] Figure 5 This is a schematic diagram of an ethylene separation and recovery device provided in Example 4.

[0025] Explanation of reference numerals in the attached figures

[0026] 1. C3 separation equipment; 2. Stripping tower; 3. C2 recovery tower; 4. Methanogen removal tower; 5. Tail gas scrubbing tower; 6. Methane / hydrogen primary separation equipment; 7. Expansion-recompressor; 8. High-pressure ethane removal tower; 9. Low-pressure ethane removal tower; 10. Hydrogenation reaction system; 11. Green oil separator; 12. Distillation tower; 13. First feed separator; 14. Methane / hydrogen secondary separation equipment; 15. Cold box; 33. Propylene refrigerant; 34. Ethylene refrigerant;

[0027] 1-1. Demethanizer I; 1-2. Deethaner I; 1-3. Ethylene Distillation Column I; 1-4. Cold Box I; 1-5. Feed Separator I for the First Demethanizer; 1-6. Feed Separator I for the Second Demethanizer; 1-7. Feed Separator I for the Third Demethanizer; 1-10. Reflux Tank I for the Demethanizer; 1-11. Reflux Tank I for the Deethaner; 1-12. C2 Hydrogenation Reactor I; 1-13. Green Oil Washing Tank I; 1-33. Propylene Refrigerant; 1-34. Binary Refrigerant. Detailed Implementation

[0028] The endpoints and any values ​​of the ranges 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 endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0029] In this invention, the ethylene recovery rate (%) = content of polymer-grade ethylene product / (content of ethylene in the cracked gas entering the cold box - content of ethylene in the total feed to the high-pressure deethaner + content of polymer-grade ethylene product) * 100%.

[0030] The first aspect of this invention provides a method for ethylene separation and recovery, wherein the method includes:

[0031] S1. The cracked gas is subjected to C4 separation to obtain gaseous material A and liquid material A, which are C4- and above heavy components. After cooling, gaseous material A, which is C4- and above heavy components, is subjected to C3 separation to obtain gaseous material B and liquid material B, which are C3- and above heavy components. After cooling, gaseous material B, which is C3- and above heavy components, is subjected to C2 recovery to obtain gaseous material C and liquid material C, which are C2- and above heavy components. After cooling, gaseous material C, which is C2- and above heavy components, is subjected to methane / hydrogen first separation to obtain gaseous material D, which is a mixture of C1 and hydrogen, and liquid material D. After cooling, gaseous material D, which is a mixture of C1 and hydrogen, is subjected to methane / hydrogen second separation to obtain gaseous crude hydrogen material E and liquid material E.

[0032] The liquid phase material D is divided into a first liquid phase material D and a second liquid phase material D. The first liquid phase material D is sent to the C2 recovery operation as a refrigerant, and the second liquid phase material D is sent to the gas-liquid separation operation as a washing liquid.

[0033] S2. Perform methane stripping on liquid phase material A and liquid phase material B to obtain C3 and above heavy component gas phase material F and C1 and below light component liquid phase material F; perform demethanizing operation on liquid phase material C and C3 and above heavy component gas phase material F to obtain C2 and above heavy component gas phase material G and liquid C2 material G; perform gas-liquid separation operation on C2 and above heavy component gas phase material G to obtain C2 and above heavy component gas phase material I and liquid phase material I, and reflux liquid phase material I for demethanizing operation; perform high-pressure deethane removal operation on liquid phase material F and below light component liquid phase material F to obtain mixed C2 gas phase material H and liquid material H; perform C2 full-fraction hydrogenation reaction on mixed C2 gas phase material H to remove alkynes to obtain material J, and perform distillation on material J to obtain ethylene and ethane.

[0034] In this invention, by combining high-pressure deethanementation and C2 full-fraction hydrogenation, the mixed C2 gaseous material H obtained by high-pressure deethanementation does not contain C3 (especially a mixture of propyne and propadiene, i.e., MAPD) ​​and C4 (especially dienes), which makes the reaction less prone to runaway.

[0035] In this invention, all the mixed C2 gaseous material H obtained from the high-pressure deethanement operation is subjected to a C2 full-fraction hydrogenation reaction, which results in a lower acetylene concentration and indirectly controls the H2 / acetylene ratio, thereby controlling the reaction temperature rise. This leads to higher ethylene conversion and selectivity of the catalyst in the subsequent distillation process, thus improving the ethylene product yield.

[0036] According to the present invention, by way of example, the conditions for the hydrogenation reaction include: SOR-EOR: 65-120°C. Wherein, SOR refers to the operating conditions at the initial stage of the catalyst reaction in the hydrogenation reaction, and EOR refers to the operating conditions at the final stage of the catalyst reaction in the hydrogenation reaction. SOR-EOR: 65-120°C means that the catalyst reaction temperature at the initial stage of the hydrogenation reaction is not lower than 65°C, and the catalyst reaction temperature at the final stage of the hydrogenation reaction is not higher than 120°C.

[0037] In this invention, a condenser is provided during the C3 separation operation, and the refrigerant is preferably an ethylene refrigerant with a temperature of -50°C to -80°C.

[0038] In this invention, the gaseous material B, consisting of C3-lean and C3+ heavy components obtained from the C3 separation operation, contains virtually no C3. During the demethanization operation, pure C2 can be separated and directly used as a portion of the reflux in the high-pressure deethaner operation. This not only reduces the equipment load during high-pressure deethanerization but also reduces the condensation volume in the reflux condenser.

[0039] In this invention, the pyrolysis gas is preferably compressed, alkaline washed, and dried before cooling. The methods of compression, alkaline washing, and drying are all conventional in the art and will not be described in detail here.

[0040] According to the present invention, the method further includes removing green oil from the material J after condensation before distillation to obtain gaseous C2 material M and liquid C2 material M.

[0041] According to a preferred embodiment of the present invention, the gaseous C2 material M is distilled to obtain ethylene and ethane.

[0042] According to a preferred embodiment of the present invention, the liquid C2 material M is returned to the high-pressure deethanerization operation for reflux.

[0043] According to a preferred embodiment of the present invention, the liquid phase C2 material G is returned to the high-pressure deethanerization operation for reflux.

[0044] According to a preferred embodiment of the present invention, the heavy component is combined with the liquid phase C2 material G and then returned to the high-pressure deethanerization operation for reflux.

[0045] According to a preferred embodiment of the present invention, the liquid phase material H is subjected to a low-pressure deethane removal operation to obtain a gas phase material L and a liquid phase material L, and the gas phase L is condensed and refluxed to a high-pressure deethane removal operation.

[0046] In this invention, the above operations can maintain the high selectivity of the catalyst, improve the ethylene conversion rate during long-term operation, reduce the amount of green oil produced, eliminate the need for circulating ethylene to wash the green oil, and reduce operating energy consumption.

[0047] According to a preferred embodiment of the present invention, the gaseous crude hydrogen material E obtained from the second methane / hydrogen separation operation is condensed and the cold energy is recovered to obtain hydrogen.

[0048] In this invention, it is preferable to send the gaseous crude hydrogen material E into the cold box and the C2 recovery operation as a refrigerant to recover the cold energy and obtain hydrogen.

[0049] According to a preferred embodiment of the present invention, the liquid phase material E obtained from the second methane / hydrogen separation operation is cooled to obtain low-pressure methane.

[0050] According to a preferred embodiment of the present invention, the first liquid phase material D obtained from the first methane / hydrogen separation operation is fed into the C2 recovery operation as a refrigerant, and the cold energy is recovered to obtain medium-pressure methane.

[0051] According to a preferred embodiment of the present invention, the C2- and C2- or higher heavy component gaseous material I is compressed and the cold energy is recovered to obtain high-pressure methane.

[0052] In this invention, the method further includes a refrigeration system that provides cooling capacity, and the refrigeration system is also capable of recovering the cooling capacity generated in each step. The medium providing cooling capacity to the refrigeration system of this invention can be propylene and / or ethylene.

[0053] According to some embodiments of the present invention, the pyrolysis gas contains 15-25 vol% hydrogen, 16-28 vol% methane and 25-45 vol% ethylene.

[0054] In this invention, the components of the cracked gas include H2 and C1-C6 components. The contents of hydrogen, methane, and ethylene each meet the above-mentioned ranges. Other components (e.g., ethane, acetylene, C3 and above components) are allowed to fluctuate within a certain range during operation and have no significant impact on the results of this invention. They are not key control indicators, as long as the total content of H2 and C1-C6 components is 100%.

[0055] According to some preferred embodiments of the present invention, the pyrolysis gas is cooled to -30°C to -45°C before the C4 separation operation is performed.

[0056] According to some preferred embodiments of the present invention, before performing the C3 separation operation, the gaseous material A of the C4-depleted and C4-above heavy components is cooled to -60°C to -78°C.

[0057] According to some preferred embodiments of the present invention, before performing the C2 recovery operation, the gaseous material B of the C3-depleted and C3-above heavy components is cooled to -93°C to -101°C.

[0058] In this invention, the pyrolysis gas gradually separates the heavier hydrocarbon condensate from the lighter component gas at intermediate temperatures (e.g., -30°C to -45°C, -60°C to -78°C, and -93°C to -101°C), avoiding cooling this portion of liquid to the lowest temperature, which helps to reduce the total required refrigeration capacity and at the same time reduce the ethylene content in the exhaust gas.

[0059] According to some preferred embodiments of the present invention, the conditions for the C4 separation operation include: a column top temperature of -25°C to -40°C and a pressure of 2.9 MPaG to 4.1 MPaG.

[0060] The conditions for the C3 separation operation include: a column top temperature of -45°C to -60°C, a pressure of 2.8 MPaG to 4 MPaG, and a column bottom temperature of -30°C to -40°C.

[0061] According to some preferred embodiments of the present invention, the conditions for the C2 recovery operation include: a top temperature of -110°C to -135°C, a pressure of 2.8 MPaG to 4 MPaG, and a bottom temperature of -105°C to -130°C.

[0062] According to some preferred embodiments of the present invention, the operating conditions of the stripping tower include: a top temperature of -35°C to -60°C, a pressure of 2.8 MPaG to 4 MPaG, and a bottom temperature of 10°C to 30°C.

[0063] According to some preferred embodiments of the present invention, the conditions for the demethanizing operation include: a top temperature of -90°C to -99°C, a pressure of 2.7 MPaG to 3.8 MPaG, and a bottom temperature of -20°C to 10°C.

[0064] According to some preferred embodiments of the present invention, the conditions for the high-pressure ethane removal operation include: a top temperature of -25°C to -5°C, a pressure of 2 MPaG to 2.8 MPaG, and a bottom temperature of 50°C to 80°C.

[0065] According to some preferred embodiments of the present invention, the low-pressure ethane removal operating conditions include: a top temperature of -20°C to 10°C, a pressure of 1.6 MPaG to 2.4 MPaG, and a bottom temperature of 50°C to 70°C.

[0066] According to some preferred embodiments of the present invention, the distillation conditions include: a top temperature of -40°C to -10°C, a pressure of 1.5 MPaG to 2.6 MPaG, and a bottom temperature of -15°C to 20°C.

[0067] In this invention, the operating conditions of each operation meet the above-mentioned range, which can further improve the ethylene recovery rate and reduce energy consumption.

[0068] According to the present invention, preferably, the content of C4 and above heavy components in the gaseous material A, which is depleted of C4 and above heavy components, is controlled to be ≤500ppm.

[0069] According to the present invention, the content of C3 and above heavy components in the gaseous material B, which is depleted of C3 and above heavy components, is ≤100ppm.

[0070] According to the present invention, the content of C2 and above heavy components in the gaseous material C of the C-depleted C2 and above heavy components is ≤100ppm.

[0071] According to the present invention, the alkyne content in material J is ≤1ppm.

[0072] According to the present invention, the content of C3 and above heavy components in the gas phase F of the C3 and above heavy components is ≤100ppm, and the content of methane in the liquid phase obtained from the bottom of the stripping tower is ≤100ppmv.

[0073] According to the present invention, the content of C2 and above heavy components in the gaseous material I of C2 and above heavy components is ≤500ppm.

[0074] In this invention, the gas-liquid separation operation (i.e., the tail gas scrubbing tower 5 described below) serves to reduce the loss of C2 components with the gas phase. The recovered C2 components are refluxed with the methane liquid phase through the demethanizing tower 4. Therefore, the components in the liquid phase material I do not require strict control. Preferably, according to one embodiment of the present invention, the methane content in the liquid phase material I is ≤100ppm.

[0075] According to the present invention, the methane content in the liquid phase C2 material G is ≤100ppm.

[0076] According to the present invention, the content of heavy components of C3 and above in the gas phase material H is ≤500ppm, and the content of light components of C2 and below in the liquid phase material H is ≤300ppm.

[0077] In this invention, by controlling the content of each of the above components to meet the above range, the ethylene content in the exhaust gas can be further reduced and the ethylene recovery rate can be improved.

[0078] A second aspect of the present invention provides an apparatus for ethylene separation and recovery, wherein the apparatus comprises: a C3 separation device 1, a stripping tower 2, a C2 recovery tower 3, a demethanizer 4, a tail gas scrubbing tower 5, a first methane / hydrogen separation device 6, a high-pressure deethaner 8, a hydrogenation reaction system 10, a distillation tower 12, a first feed separation tank 13, a second methane / hydrogen separation device 14, and a cold box 15;

[0079] The top outlet of the first feed separator 13 is connected to the inlet of the C3 separator 1, the top outlet of the C3 separator 1 is connected to the first inlet of the C2 recovery tower 3, the first outlet of the C2 recovery tower 3 is connected to the inlet of the first methane / hydrogen separator 6, and the top outlet of the first methane / hydrogen separator 6 is connected to the inlet of the second methane / hydrogen separator 14, for the purpose of separating the cracked gas in stages.

[0080] The bottom outlet of the methane / hydrogen first separation device 6 is provided with a first splitter and a second splitter. The first splitter is connected to the third inlet of the C2 recovery tower 3, and the second splitter is connected to the first inlet of the tail gas scrubbing tower 5.

[0081] The first outlet of the demethanizer 4 is connected to the second inlet of the tail gas scrubbing tower 5;

[0082] The bottom outlet of the first feed separator 13 is connected to the first inlet of the stripping tower 2, the bottom outlet of the C3 separator 1 is connected to the second inlet of the stripping tower 2, the first outlet of the stripping tower 2 is connected to the second inlet of the demethanizer 4, and the second outlet of the C2 recovery tower 3 is connected to the first inlet of the demethanizer 4.

[0083] The second outlet of the stripping tower 2 is connected to the first inlet of the high-pressure deethanizer 8, the first outlet of the high-pressure deethanizer 8 is connected to the inlet of the hydrogenation reaction system 10, and the outlet of the hydrogenation reaction system 10 is connected to the inlet of the distillation tower 12.

[0084] In this invention, the C3 separation device 1 can be a C3 washing tower or a second feed separation tank connected in series with the first feed separation tank 13. The bottom materials of both the first feed separation tank 13 and the second feed separation tank 1 are used as feed for the high-pressure deethaner 8.

[0085] According to the present invention, the apparatus further includes a low-pressure deethaner 9 and a green oil separator 11 disposed between the hydrogenation reaction system 10 and the distillation column 12.

[0086] In this invention, the low-pressure deethaner is not necessary, but can be used selectively when the pressure of the deethaner is too high and the reactor temperature is too high.

[0087] In this invention, the heat source for the reboiler in the high-pressure deethanizer 8 and the reboiler in the low-pressure deethanizer 9 is preferably quench water from within the device. The bottom temperature of the high-pressure deethanizer 8 and the low-pressure deethanizer 9 is preferably ≤70°C.

[0088] In this invention, preferably, the heat source for the reboiler in the high-pressure deethanizer 8 and the reboiler in the low-pressure deethanizer 9 is low-pressure superheated steam. The refrigerant for the condenser in the low-pressure deethanizer 9 is preferably propylene refrigerant with a temperature range of -16°C to -40°C.

[0089] In this invention, a reactor effluent water cooler, a reactor feed heat exchanger, and a reactor effluent cooler are sequentially provided between the green oil separation tank 11 and the hydrogenation reaction system 10. The reactor effluent cooler preferably uses propylene refrigerant at -16°C to -40°C to further cool and condense the reactor effluent before sending it into the green oil separation tank 11. The mixed C2 obtained at the top of the green oil separation tank 11 does not contain acetylene or green oil and is sent into the distillation column 12.

[0090] In this invention, preferably, a dryer is provided between the green oil separation tank and the ethylene distillation column to remove trace amounts of moisture that may be generated during the hydrogenation reaction.

[0091] According to some preferred embodiments of the present invention, the second outlet of the demethanizer 4 is connected to the second inlet of the high-pressure deethanizer 8, the outlet of the green oil separator 11 is connected to the third inlet of the high-pressure deethanizer 8, the second outlet of the demethanizer 4 and the outlet of the green oil separator 11 are combined and connected to the third inlet of the high-pressure deethanizer 8, and the low-pressure deethanizer 9 is connected to the second outlet of the high-pressure deethanizer 8.

[0092] In this invention, the equipment is connected in the manner described above, which can maintain the high selectivity of the catalyst, improve the ethylene conversion rate during long-term operation, reduce the amount of green oil produced, eliminate the need for circulating ethylene to wash the green oil, and reduce operating energy consumption.

[0093] In this invention, it is important to understand that acetylene hydrogenation is a strongly exothermic reaction. All the gaseous phase from the top of the high-pressure deethaner 8 is fed into the hydrogenation reaction system 10. Due to dilution, the acetylene concentration in the gas is lower, resulting in a lower temperature rise in the catalyst bed of the hydrogenation reaction system 10. This maintains high selectivity (acetylene to ethylene) and better controls the thermal stability of the hydrogenation reaction system 10, reducing the risk of runaway.

[0094] In this invention, the number of stages in each reactor of the hydrogenation reaction system 10 can be one or two, depending on the concentration of acetylene in the top gaseous material of the high-pressure deethaner 8. This enables long-term, highly selective operation of the catalyst, resulting in high ethylene conversion, low green oil production, and eliminating the need for recycled ethylene as green oil washing liquid, thus reducing the operating energy consumption of the distillation column 12.

[0095] According to the present invention, the apparatus further includes: an expansion-recompressor 7 disposed between the first outlet of the tail gas scrubbing tower 5 and the inlet of the cold box 15; the second outlet of the tail gas scrubbing tower 5 being connected to the third inlet of the demethanizer 4; the top outlet of the low-pressure deethaner 9 being connected to the fourth inlet of the high-pressure deethaner 8; the top outlet of the methane / hydrogen second separation device 14 being connected to the second inlet of the C2 recovery tower 3; and the third and fourth outlets of the C2 recovery tower 3 being connected to the cold box 15.

[0096] In this invention, the tail gas scrubbing tower 5 is equipped with a packing section to separate the liquid phase material D obtained from the bottom of the methane / hydrogen first separation device 6 into a first liquid phase D and a second liquid phase material D. The second liquid phase material D enters the packing section above the tail gas scrubbing tower 5 as a scrubbing liquid to remove the C2- and C2- and higher heavy component gas phase material I obtained from the tail gas scrubbing tower 5, and further recover ethylene, so that the ethylene concentration in the C2- and C2- and higher heavy component gas phase material I obtained from the top of the tail gas scrubbing tower 5 is ≤500ppmv.

[0097] According to the present invention, both the stripping tower 2 and the demethanizing tower 4 are equipped with reboilers and condensers.

[0098] In this invention, the external heat at the bottom of the stripping tower 2 must ultimately be removed by the refrigeration unit. Therefore, by using propylene refrigerant to recover the cold energy in the tower bottom, it can offset part of the increase in the refrigeration unit load.

[0099] In this invention, the heat source for the reboiler of the stripping tower is preferably a propylene refrigerant at 35℃-80℃ for recovering cold energy; the cold source for the condenser of the stripping tower is preferably an ethylene refrigerant at -50℃-80℃. The content of the C3- and C3- and higher-weight gaseous material F obtained from the top of the stripping tower 2 is preferably less than or equal to 100 ppmv; the content of the C1- and lower-weight light components F obtained from the bottom of the stripping tower 2 is preferably ≤100 ppmv.

[0100] In this invention, the reboiler heat source in the demethanizer 4 is preferably a propylene refrigerant at 10℃-40℃ for cold recovery. The condenser cold source in the demethanizer 4 is preferably an ethylene refrigerant at -95℃ to -101℃. The ethylene concentration in the C2- and higher-weight heavy component gaseous material G obtained from partial condensation in the demethanizer 4 is ≤1000ppmv.

[0101] According to the present invention, the C2 recovery tower 3 is provided with an upper countercurrent condenser and a lower countercurrent condenser, wherein the upper countercurrent condenser is located above the lower countercurrent condenser.

[0102] In this invention, the top outlet of the second methane / hydrogen separation unit 14 returns to the cold box 15 sequentially through the cold box 15 and the upper countercurrent condenser of the C2 recovery tower 3. The bottom first splitter of the first methane / hydrogen separation unit 6 returns to the cold box 15 through the lower countercurrent condenser of the C2 recovery tower 3.

[0103] In this invention, the countercurrent condenser installed in the C2 recovery tower 3 can be a single stage or multiple stages, depending on the load requirements of the condenser.

[0104] According to the present invention, the C3 separation device 1 is provided with a condenser.

[0105] According to the present invention, reboilers are provided in the bottom of the high-pressure deethanizer 8, the bottom of the low-pressure deethanizer 9, and the distillation column 12.

[0106] According to the present invention, preferably, the hydrogenation reaction system 10 is provided with at least two reactors connected in parallel.

[0107] In this invention, the hydrogenation reaction system 10 is equipped with two parallel reactors, one in operation and one on standby, for switching operations during hydrogenation catalyst regeneration. Preferably, each reactor is equipped with at least two sections of catalyst beds connected in series with top inlet and bottom outlet, and an inter-section water cooler.

[0108] According to the present invention, the apparatus further includes a refrigeration system that provides cooling capacity to the cold box 15 and all condensers in the system, and recovers cooling capacity for the reboilers of the stripping column 2, the demethanizer column 4, and the distillation column 12. Preferably, the refrigeration system includes propylene refrigerant 33 and ethylene refrigerant 34.

[0109] In this invention, according to a specific embodiment, the top of the distillation column 12 is provided with a pasteurization section to remove light component impurities from the ethylene product. The final ethylene product (purity ≥99.95%) that meets the stringent purity requirements for polymer production is drawn from trays 8-10. The circulating ethane product at the bottom of the distillation column 12 is recovered and cracked in a gas cracking furnace.

[0110] The third aspect of the present invention provides an application of the method described in the first aspect or the apparatus described in the second aspect in ethylene separation and recovery.

[0111] The present invention will be described in detail below through embodiments.

[0112] In the following examples, ethylene recovery rate (%) = content of polymer-grade ethylene product / (content of ethylene in the cracked gas entering the cold box - content of ethylene in the total feed to the high-pressure deethaner + content of polymer-grade ethylene product) * 100%;

[0113] In the comparative example, ethylene recovery rate (%) = content of polymer-grade ethylene product / (content of ethylene in the cracked gas entering the cold box - content of ethylene in the total feed to the de-ethaner + content of polymer-grade ethylene product) * 100%;

[0114] In the following embodiments, the same reference numerals represent the same parts.

[0115] Example 1

[0116] Combination Figure 2 A method for separating and recovering ethylene is provided, the method comprising:

[0117] S1. The pyrolysis gas 20, after being compressed, alkali washed and dried, is sequentially separated through the first feed separator 13, C3 separator 1, C2 recovery tower 3, methane / hydrogen first separator 6 and methane / hydrogen second separator 14 connected in sequence.

[0118] The gaseous material A (containing 500 ppmv of C4 and above heavy components) obtained from the top of the first feed separator 13 (temperature -37℃, operating pressure 3.6 MPaG), the gaseous material B (containing 100 ppmv of C3 and above heavy components) obtained from the top of the C3 separator 1 (temperature -57℃, operating pressure 3.5 MPaG), the gaseous material C (containing 100 ppmv of C2 and above heavy components) obtained from the top of the C2 recovery tower 3 (temperature -124℃, operating pressure 3.4 MPaG), and the gaseous material D (containing a mixture of C1 and hydrogen) obtained from the top of the first methane / hydrogen separator 6 (temperature -140℃) are respectively sent to the cold box 15 for further cooling.

[0119] The bottom outlet of the methane / hydrogen first separation device 6 is provided with a first splitter and a second splitter. The first splitter is connected to the third inlet of the C2 recovery tower 3, and the second splitter is connected to the first inlet of the tail gas scrubbing tower (5). It is used to separate the liquid material D obtained at the bottom of the methane / hydrogen first separation device 6 into a first liquid material D and a second liquid material D. The first liquid material D is sent to the C2 recovery tower 3 as a refrigerant to recover the cold energy and obtain medium-pressure methane 16. The second liquid material D is sent to the tail gas scrubbing tower 5 as a scrubbing liquid.

[0120] The liquid material E obtained from the bottom of the methane / hydrogen second separation device 14 is passed through the cold box 15 to recover the cold energy and obtain low-pressure methane 18. The gaseous crude hydrogen material E obtained from the top of the methane / hydrogen second separation device 14 is passed through the cold box 15 and the upper countercurrent condenser in the C2 recovery tower 3 as a refrigerant and then sent back to the cold box to recover the cold energy and obtain hydrogen 17.

[0121] S2. The bottom outlet of the first feed separator 13 is connected to the first inlet of the stripping tower 2, the bottom outlet of the C3 separation device 1 is connected to the second inlet of the stripping tower 2, and the first outlet of the stripping tower 2 is connected to the second inlet of the demethanizer 4. This is used to send the liquid phase material A obtained from the bottom of the first feed separator 13 and the liquid phase material B obtained from the bottom of the C3 separation device 1 (bottom temperature -38℃) into the stripping tower 2 for stripping (pressure 3.4MPaG). The top of the stripping tower 2 (temperature -42℃) yields a gas phase material F of C3 and above heavy components (C3 and above heavy components content is 100ppmv), and the bottom of the stripping tower 2 (temperature 20℃) yields a liquid phase material F of C1 and below light components.

[0122] S3. The first outlet of the stripping tower 2 is connected to the second inlet of the demethanizing tower 4, the second outlet of the C2 recovery tower 3 is connected to the first inlet of the demethanizing tower 4, and the second outlet of the C2 recovery tower 3 is connected to the first inlet of the demethanizing tower 4, for feeding the gas phase material F of the lean C3 and C3 and above heavy components and the liquid phase material C obtained from the C2 recovery operation into the demethanizing tower 4 for demethanizing operation (tower top temperature -96℃, tower bottom temperature -6℃, operating pressure 3.1MPag); the second outlet of the tail gas scrubbing tower (5) is connected to the demethanizing tower (4) The third inlet of the demethanizer 4 is connected to the gaseous material G of lean C2 and C2 and above heavy components obtained from the top of the demethanizer 4. After partial condensation by the top condenser, the gaseous material G is sent to the bottom of the packing section of the tail gas scrubbing tower 5 for gas-liquid separation. The liquid material I obtained from the bottom of the tail gas scrubbing tower 5 is sent to the top of the demethanizer 4 for reflux. The gaseous material I of lean C2 and C2 and above heavy components (C2 and above heavy component content is 500 ppmv) obtained from the top of the tail gas scrubbing tower 5 is sent back to the cold box 15 through the expansion-recompressor 7 to recover the cold energy and obtain high-pressure methane 19.

[0123] S4. A green oil separator 11 is provided between the hydrogenation reaction system 10 and the distillation column 12. The second outlet of the demethanizer 4 is connected to the second inlet of the high-pressure deethanizer 8, and the outlet of the green oil separator 11 is connected to the third inlet of the high-pressure deethanizer 8. The second outlet of the demethanizer 4 and the outlet of the green oil separator 11 are combined and connected to the third inlet of the high-pressure deethanizer 8. This is used to combine the liquid C2 material G obtained from the bottom of the demethanizer 4 and the liquid C2 material M obtained from the bottom of the green oil separator 11 and send them to the top of the high-pressure deethanizer 8 for reflux. By using the expansion-recompressor 7 and an open refrigeration cycle using the JT effect, the demand for external refrigeration is minimized.

[0124] S5. The second outlet of the stripping tower 2 is connected to the first inlet of the high-pressure deethanizer 8, the first outlet of the high-pressure deethanizer 8 is connected to the inlet of the hydrogenation reaction system 10, and the outlet of the hydrogenation reaction system 10 is connected to the inlet of the distillation tower 12. This is used to feed the liquid phase material F containing lean C1 and lower light components into the high-pressure deethanizer 8 for high-pressure deethanization, performing non-clear separation of the C2 fraction. The liquid phase material H obtained at the bottom of the high-pressure deethanizer 8 (bottom temperature 70℃) is fed into the low-pressure deethanizer 9 for separation of C2 and C3 fractions. The mixed C2 gaseous material H obtained from the top of ethane tower 8 is sent to the hydrogenation reaction system 10 for C2 full-fraction hydrogenation reaction (operating conditions SOR-EOR: 65-120℃) to remove alkynes. After partial condensation in the condenser, it is sent to the green oil separator 11 to remove green oil. The gaseous C2 material M obtained from the top of the green oil separator 11 is sent to the distillation column 12 for purification of polymerization-grade ethylene products. The liquid C2 material M obtained from the bottom of the green oil separator 11 and the liquid C2 material G obtained from the bottom of the demethanizer 4 are combined and sent to the top of the high-pressure deethanizer 8 for reflux.

[0125] S6. The gaseous material L obtained from the top of the low-pressure ethane stripper 9 (temperature is 5°C, pressure is 2.2MPaG) is condensed by the top condenser and fed into the lower section of the high-pressure ethane stripper 8. The liquid material L obtained from the bottom of the low-pressure ethane stripper 9 (temperature is 65°C) is recovered as propylene product.

[0126] S7. The gaseous material obtained from the top of distillation column 12 (-30℃, 1.9MPag) is condensed by propylene refrigerant to obtain liquid phase, which is then returned to the top of distillation column 12 as reflux. Polymer-grade ethylene product (purity 99.95% v) is collected from the 10th plate of the upper section of distillation column 12. The reboiler of distillation column 12 uses propylene refrigerant as a heat source to recover cold energy, improve heat integration, and thus reduce the load on the propylene generator. The ethane (liquid phase) obtained from the bottom of distillation column 12 (temperature -5℃) is fed into the cracking furnace as raw material for cracking after cold energy recovery in the cold box.

[0127] The above method processes ethylene on a scale of 1.25 million tons / year, and the amount of cracked gas fed into the cold box 15 is 3.13 million tons / year, of which ethylene content is 1.2575 million tons / year.

[0128] After processing using the above method, the total feed rate to the high-pressure deethaner 8 is 2,428,400 tons / year, of which ethylene accounts for 1,256,700 tons / year and acetylene accounts for 18,600 tons / year; the total feed rate to the distillation column 12 is 1,479,700 tons / year, and the distillation yields 1,265,700 tons / year of polymer-grade ethylene product. The calculated ethylene recovery rate is 99.94%. The total power of the refrigeration unit is 53,180 kW.

[0129] Example 2

[0130] The method is the same as in Example 1, except that, Figure 3 As shown, a low-pressure ethane stripper 9 is not installed.

[0131] Example 3

[0132] The method is the same as in Example 1, except that, Figure 4 As shown, the second outlet of the demethanizer 4 is connected to the inter-section inlet of the high-pressure deethaner 8, rather than being merged with the outlet of the green oil separator 11.

[0133] Example 4

[0134] The method is the same as in Example 3, except that, Figure 5 As shown, a low-pressure ethane stripper 9 is not installed.

[0135] Comparative Example 1

[0136] like Figure 1 As shown, after preliminary compression, alkaline washing, and drying, the pyrolysis gas enters the cold box I1-4 and is condensed in stages. The condensate is fed to the first demethanizer feed separator I1-5 (-72℃, 3.6MPaG), the second demethanizer feed separator I1-6 (-95℃, 3.55MPaG), and the third demethanizer feed separator I1-7 (-136℃, 3.5MPaG). The liquid phase separated from these separators is used as feed and is sent to the demethanizer I1-1 in four streams. The gas phase separated from the top reflux tank I1-10 (-131℃, 0.6MPaG) of the demethanizer I1-1 is returned to the cold box I1-4 to recover the cooling capacity. The liquid phase separated from the bottom of the demethanizer I1-1 (-50℃, 0.65MPaG) is sent back to the cold box I1-4 to recover the cooling capacity. The gas phase (PaG) is fed into the deethanizer I1-2. The top gas phase (-19℃, 2.2MPaG) of the deethanizer I1-2 is partially condensed by the top condenser and sent to the deethanizer reflux tank I1-11 to separate the liquid phase for reflux. The gas phase separated from the deethanizer reflux tank I1-11 is sent to the C2 hydrogenation reactor I1-12 (SOR-EOR: 45-70℃) for reaction. The effluent from the C2 hydrogenation reactor I1-12 is washed in the green oil washing tank I1-13 with recycled ethylene drawn from the middle section of the ethylene tower. The top gas phase of the green oil washing tank I1-13 is dried and sent to the ethylene distillation tower I1-3. Polymer-grade ethylene product is obtained at the top of the ethylene distillation tower I1-3 (-36℃, 1.6MPaG).

[0137] Using the same method as in Example 1, the cracked gas fed into the cold boxes 1-4 is 3.13 million tons / year, containing 1.2575 million tons / year of ethylene; the total feed into the deethaner 1-2 is 2.4247 million tons / year, containing 1.2534 million tons / year of ethylene and 18,600 tons / year of acetylene; the total feed into the ethylene tower 1-3 is 1.4965 million tons / year, yielding 1.2624 million tons / year of polymer-grade ethylene product through distillation. The calculated ethylene recovery rate reaches 99.68%. The total power of the refrigeration unit is 55010 kW.

[0138] Comparing the results of Example 1 and Comparative Example 1, it can be seen that the ethylene recovery rate using the method of Example 1 reaches 99.94%, which is 0.26% higher than the traditional ethylene separation process in Comparative Example 1. This allows for the recovery of an additional 3300 tons / year of ethylene, increasing output value by nearly 31 million yuan. The total refrigeration load is reduced by 3% compared to the traditional separation process in Comparative Example 1, and the compressor power is reduced by 1830KW. In Example 2, the absence of a low-pressure deethaner 9 increases the separation load of the high-pressure deethaner 8, resulting in energy inefficiency. In Example 3, the influent from the demethanizer 4 enters the interstage of the high-pressure deethaner 8, causing some C2 component backmixing, which also increases the separation load of the high-pressure demethanizer, resulting in energy inefficiency. Compared to Example 3, Example 4 also lacks a low-pressure deethaner 9, causing some feed backmixing in the high-pressure deethaner 8, resulting in energy inefficiency and higher energy consumption than Examples 1-3. However, the total energy consumption of Examples 1-4 is lower than that of Comparative Example 1.

[0139] The preferred embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for separating and recovering ethylene, characterized in that, The method includes: S1. The cracked gas is subjected to C4 separation to obtain gaseous material A and liquid material A, which are C4- and above heavy components. After cooling, gaseous material A, which is C4- and above heavy components, is subjected to C3 separation to obtain gaseous material B and liquid material B, which are C3- and above heavy components. After cooling, gaseous material B, which is C3- and above heavy components, is subjected to C2 recovery to obtain gaseous material C and liquid material C, which are C2- and above heavy components. After cooling, gaseous material C, which is C2- and above heavy components, is subjected to methane / hydrogen first separation to obtain gaseous material D, which is a mixture of C1 and hydrogen, and liquid material D. After cooling, gaseous material D, which is a mixture of C1 and hydrogen, is subjected to methane / hydrogen second separation to obtain gaseous crude hydrogen material E and liquid material E. The liquid phase material D is divided into a first liquid phase material D and a second liquid phase material D. The first liquid phase material D is sent to the C2 recovery operation as a refrigerant, and the second liquid phase material D is sent to the gas-liquid separation operation as a washing liquid. S2. Liquid phase material A and liquid phase material B are subjected to methane stripping to obtain C3- and above heavy component gas phase material F and C1- and below light component liquid phase material F; liquid phase material C and C3- and above heavy component gas phase material F are subjected to demethanization to obtain C2- and above heavy component gas phase material G and liquid C2 material G; C2- and above heavy component gas phase material G is subjected to gas-liquid separation to obtain C2- and above heavy component gas phase material I and liquid material I, and liquid material I is refluxed for demethanization; C1- and below light component liquid phase material F is subjected to high-pressure deethanerization to obtain mixed C2 gas phase material H and liquid material H; mixed C2 gas phase material H is subjected to C2 full-fraction hydrogenation to remove alkynes to obtain material J, and material J is distilled to obtain ethylene and ethane.

2. The method according to claim 1, wherein, The method further includes removing green oil from material J after condensation before distillation to obtain gaseous C2 material M and liquid C2 material M. Preferably, the gaseous C2 material M is distilled to obtain ethylene and ethane; Preferably, the liquid C2 material M is returned to the high-pressure ethane removal operation for reflux use; Preferably, the liquid C2 material G is returned to the high-pressure ethane removal operation for reflux use; Preferably, the heavy components are combined with the liquid C2 material G and then returned to the high-pressure deethane removal operation for reflux. Preferably, the liquid phase material H is subjected to a low-pressure deethane removal operation to obtain a gas phase material L and a liquid phase material L, and the gas phase L is condensed and refluxed to a high-pressure deethane removal operation.

3. The method according to claim 1 or 2, wherein, The gaseous crude hydrogen material E obtained from the second methane / hydrogen separation operation is used as a refrigerant, and the cold energy is recovered to obtain hydrogen. Preferably, the liquid phase material E obtained from the second methane / hydrogen separation operation is cooled to obtain low-pressure methane; Preferably, the first liquid phase material D obtained from the first methane / hydrogen separation operation is fed into the C2 recovery operation as a refrigerant, and the cold energy is recovered to obtain medium-pressure methane; Preferably, the C2- and C2- or higher heavy component gaseous material I is compressed and the cold energy is recovered to obtain high-pressure methane.

4. The method according to claim 2, wherein, The cracked gas contains 15-25 vol% hydrogen, 16-28 vol% methane, and 25-45 vol% ethylene; Preferably, the pyrolysis gas is cooled to -30°C to -45°C before the C4 separation operation is performed; Preferably, before performing the C3 separation operation, the gaseous material A containing C4- and higher heavy components is cooled to -60°C to -78°C. Preferably, before performing the C2 recovery operation, the gaseous material B containing C3- and higher heavy components is cooled to -93°C to -101°C. Preferably, the conditions for the C4 separation operation include: a column top temperature of -25°C to -40°C and a pressure of 2.9 MPaG to 4.1 MPaG; The conditions for the C3 separation operation include: a column top temperature of -45°C to -60°C, a pressure of 2.8 MPaG to 4 MPaG, and a column bottom temperature of -30°C to -40°C. Preferably, the conditions for the C2 recovery operation include: a column top temperature of -110°C to -135°C, a pressure of 2.8 MPaG to 4 MPaG, and a column bottom temperature of -105°C to -130°C. Preferably, the conditions for the methane stripping operation include: a top temperature of -35°C to -60°C, a pressure of 2.8 MPaG to 4 MPaG, and a bottom temperature of 10°C to 30°C. Preferably, the conditions for the demethanization operation include: a top temperature of -90°C to -99°C, a pressure of 2.7 MPaG to 3.8 MPaG, and a bottom temperature of -20°C to 10°C. Preferably, the conditions for the high-pressure ethane removal operation include: a column top temperature of -25°C to -5°C, a pressure of 2 MPaG to 2.8 MPaG, and a column bottom temperature of 50°C to 80°C. Preferably, the conditions for the low-pressure ethane removal operation include: a column top temperature of -20°C to 10°C, a pressure of 1.6 MPaG to 2.4 MPaG, and a column bottom temperature of 50°C to 75°C. Preferably, the distillation conditions include: a top temperature of -40°C to -10°C, a pressure of 1.5 MPaG to 2.6 MPaG, and a bottom temperature of -15°C to 20°C.

5. The method according to claim 2, wherein, The content of C3 and above heavy components in the gaseous material B, which is low in C3 and above heavy components, is ≤100ppm. Preferably, the content of C2 and above heavy components in the gaseous material C of the C-phase material is ≤100ppm; Preferably, the alkyne content in material J is ≤1 ppm; Preferably, the content of C3 and above heavy components in the gas phase F, which is leaning towards C3 and above, is ≤100ppm, and the content of methane in the liquid phase material F, which is leaning towards C1 and below, is ≤100ppm. Preferably, the content of C2 and above heavy components in the gaseous material I is ≤500ppm; Preferably, the methane content in the liquid phase C2 material G is ≤100ppm; Preferably, the content of C3 and higher heavy components in the gaseous material H is ≤500ppm; Preferably, the content of light components of C2 and below in the liquid phase material H is ≤300ppm.

6. An apparatus for separating and recovering ethylene, characterized in that, The apparatus includes: a C3 separation unit (1), a stripping tower (2), a C2 recovery tower (3), a demethanizer (4), a tail gas scrubbing tower (5), a first methane / hydrogen separation unit (6), a high-pressure deethaner (8), a hydrogenation reaction system (10), a distillation tower (12), a first feed separator (13), a second methane / hydrogen separation unit (14), and a cold box (15); The top outlet of the first feed separator (13) is connected to the inlet of the C3 separator (1), the top outlet of the C3 separator (1) is connected to the first inlet of the C2 recovery tower (3), the first outlet of the C2 recovery tower (3) is connected to the inlet of the first methane / hydrogen separator (6), and the top outlet of the first methane / hydrogen separator (6) is connected to the inlet of the second methane / hydrogen separator (14), for the purpose of separating the cracked gas step by step; The bottom outlet of the methane / hydrogen first separation device (6) is provided with a first splitter and a second splitter. The first splitter is connected to the third inlet of the C2 recovery tower (3), and the second splitter is connected to the first inlet of the tail gas scrubbing tower (5). The first outlet of the demethanizer (4) is connected to the second inlet of the tail gas scrubbing tower (5); The bottom outlet of the first feed separator (13) is connected to the first inlet of the stripping tower (2), the bottom outlet of the C3 separator (1) is connected to the second inlet of the stripping tower (2), the first outlet of the stripping tower (2) is connected to the second inlet of the demethanizer (4), and the second outlet of the C2 recovery tower (3) is connected to the first inlet of the demethanizer (4). The second outlet of the stripping tower (2) is connected to the first inlet of the high-pressure deethaner (8), the first outlet of the high-pressure deethaner (8) is connected to the inlet of the hydrogenation reaction system (10), and the outlet of the hydrogenation reaction system (10) is connected to the inlet of the distillation tower (12).

7. The apparatus according to claim 6, wherein, The apparatus also includes a low-pressure deethaner (9) and a green oil separator (11) located between the hydrogenation reaction system (10) and the distillation column (12); Preferably, the second outlet of the demethanizer (4) is connected to the second inlet of the high-pressure deethanizer (8); Preferably, the outlet of the green oil separator (11) is connected to the third inlet of the high-pressure ethane stripper (8); Preferably, the second outlet of the demethanizer (4) and the outlet of the green oil separator (11) are combined and connected to the third inlet of the high-pressure deethaner (8); Preferably, the low-pressure deethaner (9) is connected to the second outlet of the high-pressure deethaner (8).

8. The apparatus according to claim 6 or 7, wherein, The device further includes an expansion-recompressor (7) installed between the first outlet of the exhaust gas scrubbing tower (5) and the inlet of the cold box (15); Preferably, the second outlet of the tail gas scrubbing tower (5) is connected to the third inlet of the demethanizing tower (4); Preferably, the top outlet of the low-pressure ethane stripper (9) is connected to the fourth inlet of the high-pressure ethane stripper (8); Preferably, the top outlet of the second methane / hydrogen separation device (14) is connected to the second inlet of the C2 recovery tower (3), and the third and fourth outlets of the C2 recovery tower (3) are connected to the cold box (15).

9. The apparatus according to claim 6 or 7, wherein, Both the stripping tower (2) and the demethanizing tower (4) are equipped with reboilers and condensers; Preferably, the C2 recovery tower (3) is provided with an upper countercurrent condenser and a lower countercurrent condenser, wherein the upper countercurrent condenser is located above the lower countercurrent condenser; Preferably, the C3 separation device (1) is equipped with a condenser; Preferably, reboilers are provided in the bottom of the high-pressure ethane stripper (8), the bottom of the low-pressure ethane stripper (9), and the distillation column (12); Preferably, the hydrogenation reaction system (10) is provided with at least two reactors connected in parallel; Preferably, the reactor is provided with at least two sections of catalyst beds connected in series, with the inlet at the top and the outlet at the bottom; Preferably, the device further includes a refrigeration system that provides cooling capacity to the cold box and the condenser.

10. The application of the method according to any one of claims 1-5 or the apparatus according to any one of claims 6-9 in ethylene separation and recovery.

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