CN107602325B - Separation method of methanol-to-olefin reaction gas - Google Patents
Separation method of methanol-to-olefin reaction gas Download PDFInfo
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Abstract
The invention relates to the field of methanol to olefin, and discloses a method for separating reaction gas from methanol to olefin, which comprises the following steps: (1) introducing the methanol-to-olefin reaction gas into a compression unit for compression treatment; (2) cooling the gas phase obtained by compression treatment; (3) in the presence of an absorbent, performing demethanization treatment on a gas-liquid two-phase obtained after cooling in a demethanizer; (4) carrying out depropanization treatment on the tower bottom product obtained by the demethanization treatment in a depropanization tower; (5) drying the tower top product of the depropanization treatment and then carrying out deethanization treatment in a deethanization tower. The method of the invention removes light components such as methane in methanol to olefin (MTO/DMTO) reaction gas by using gasoline component as an absorbent, can obviously improve the purity and recovery rate of ethylene and propylene, and can obviously reduce the energy consumption of the system.
Description
Technical Field
The invention relates to the field of methanol-to-olefin, in particular to a separation method of methanol-to-olefin reaction gas.
Background
The energy structure of China is mainly coal which is used as basic energy and important chemical raw materials of China for a long time. The modern coal chemical industry is based on clean coal technology and is an important direction for efficient clean conversion and utilization of coal. In 2014, the annual ethylene yield of China has increased to 1700 million tons. The great demand of national economy for ethylene and propylene mainly depends on petroleum, so that the annual average consumption of petroleum in China is increased to more than 5 hundred million tons. The modern coal chemical industry represented by the technology of preparing olefin from coal undoubtedly provides important support for national economy and social development, and is also a requirement for guaranteeing national energy safety. The breakthrough of the technology for preparing low-carbon olefin from methanol and the industrialization thereof mark that China is in the leading level in the world in the field.
The complete set of methanol-to-olefin technology consists of reaction technology and separation technology. The reaction technology takes catalyst development and reactor development as the core; the separation technology takes reaction products as raw materials, and produces polymerization-grade ethylene and propylene products through the processes of impurity removal, compression, separation and the like, and the core of the separation technology is the development and design of impurity removal and separation processes.
The separation process to obtain polymer grade ethylene and propylene from methanol to olefins (MTO/DMTO) product gas is generally divided into four units: the device comprises a compression unit, an impurity removal unit, a cold separation unit and a hot separation unit. Wherein the compression unit has the main task of increasing the (MTO/DMTO) product gas to the pressure required for the subsequent process. The impurity removing unit has the main tasks of removing unreacted raw materials such as methanol, intermediate reaction products such as dimethyl ether (DME) and acid gases such as carbon dioxide and NO from (MTO/DMTO) product gasxAnd the like, and partial impurities can be removed from the compressor section by water washing, alkali washing and other operations according to different required operation conditions. The cold separation unit is a separation unit in a temperature range below 0 ℃, such as demethanization, deethanization, ethylene rectification and the like. The thermal separation process is a separation unit at a temperature range of above 0 ℃, and comprises the processes of depropanization, propylene rectification and the like.
The removal of the light component methane from the impurities is key to obtaining high purity polymer grade ethylene. The most common methods at present are cryogenic separation and absorption separation. Patent application CN200810201217.3 discloses a method for separating light hydrocarbon products from MTO/MTP reaction products, which adopts ethane as an absorbent to remove light components in the products through two-stage absorption processes of an absorption tower and a demethanizer, the cooling temperature is about-65 ℃, and finally 90-99.99% of ethylene can be obtained through rectification. The absorption temperature required by using ethane as an absorbent is lower, an ethylene refrigeration compressor at the temperature of-100 ℃ is required in engineering application, and even a methane compressor at the temperature of-135 ℃ is required to provide cold energy, so that the investment of equipment is increased, and the energy consumption is increased.
Patent application CN201110279313.1 discloses a process flow for separating low-carbon hydrocarbons and methanol-to-olefin (M-OS/MTO) gas, which adopts the bottom effluents of a front demethanizer, a depropanizer and a debutanizer as absorbents to absorb three components of carbon and carbon in product gas and separate light fractions such as methane from the product gas. The process requires demethanization at temperatures around-40 ℃ and in this process is front-end deethanization.
Patent application CN201310216950.3 discloses a system and a method for recovering ethylene from MTO/DMTO product gas, wherein a cold energy recoverer is additionally arranged behind a demethanizer, so that the system energy consumption is reduced, the ethylene recovery rate is more than 99.7%, but a cold box is required, and the energy consumption is higher.
Patent application WO2009/015507 discloses a method for preparing polymer-grade low-carbon olefin by separating methanol cracking gas, wherein methane is removed by an absorption separation method, but the lowest process temperature is about-91 ℃, and the energy consumption is high.
The demethanization process of the methods disclosed in the above patent documents is performed at 0 ℃ or lower, and is all pre-deethanized or propane, so that methane and other light hydrocarbons such as ethane, ethylene and propylene are easily formed into hydrocarbon hydrates with water, such hydrates easily cause freezing and blocking of valves and pipelines at 0 ℃ or lower, and the energy consumption of the above methods is also large.
Disclosure of Invention
The invention aims to overcome the defects that the method for separating the methanol-to-olefin reaction gas in the prior art has high energy consumption, is easy to cause pipeline blockage and the like, and provides the method for separating the methanol-to-olefin reaction gas.
In order to achieve the above object, the present invention provides a method for separating a methanol-to-olefin reaction gas, comprising:
(1) introducing the methanol-to-olefin reaction gas into a compression unit for compression treatment;
(2) cooling the gas phase obtained by compression treatment;
(3) in the presence of an absorbent, performing demethanization treatment on a gas-liquid two-phase obtained after cooling in a demethanizer;
(4) carrying out depropanization treatment on the tower bottom product obtained by the demethanization treatment in a depropanization tower;
(5) drying the tower top product subjected to the depropanization treatment, and then performing deethanization treatment in a deethanizer;
wherein the compression treatment is performed so that the obtained gas phase pressure is 3.2-6.0MPa, membrane separation treatment is performed in the compression treatment process, the cooling temperature in step (2) is 30-60 ℃, the absorbent is gasoline, and the demethanization treatment conditions comprise: the number of theoretical plates of the demethanizer is 20-60, the operating pressure is 3.0-5.8MPa, the temperature at the top of the tower is 30-90 ℃, and the temperature at the bottom of the tower is 50-130 ℃.
The methanol-to-olefin (MTO/DMTO) reaction gas separation method has the following advantages:
(1) in the method, the gasoline component is used as an absorbent to remove light components such as methane in the methanol-to-olefin (MTO/DMTO) reaction gas, so that the raw material of the absorbent is easy to obtain and the cost is low;
(2) methane and light fractions in methanol-to-olefin (MTO/DMTO) product gas are removed in an operating temperature range of more than 30 ℃, so that water and hydrocarbons are effectively inhibited from forming hydrates;
(3) in the method, the drying process does not need to be placed at the outlet of the compression unit, and only the top gas of the deethanizer needs to be dried, so that the liquid-phase condensate at the outlet of the compression unit is prevented from being dried, and the consumption of nitrogen and energy consumption in public works are greatly reduced;
(4) in the method, the membrane separation unit is arranged between the compressor sections, so that hydrogen and part of methane in the feed gas can be removed, and the energy consumption of the compressor is reduced;
(5) in the method, the minimum temperature in the whole process is not lower than-40 ℃, an ethylene refrigerating unit is not needed, the material can be common low-temperature carbon steel, the energy consumption is low, and the investment is saved;
(6) the method can obtain polymer-grade ethylene and propylene products with high purity, and the recovery rate of ethylene is more than 95 percent and the recovery rate of propylene is more than 95 percent.
Additional features and advantages of the invention will be set forth in the detailed description which follows.
Drawings
FIG. 1 is a flow chart of a compression process according to an embodiment of the present invention.
FIG. 2 is a flow chart of a methanol to olefin reaction gas separation process according to an embodiment of the present invention.
Detailed Description
The following describes in detail specific embodiments of the present invention. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.
The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.
The pressure in the present invention means absolute pressure.
The reaction gas of the methanol-to-olefin (MTO/DMTO) reactor is provided with a large amount of saturated water, and the impurity removal process usually adopts alkaline washing, water washing and other operations, so that the material in the separation process contains a large amount of saturated water. The operation temperature of the existing demethanization process is below 0 ℃, and the operation pressure is more than 2.0 MPa. Under the condition, methane and other light hydrocarbons such as ethane, ethylene, propylene and the like are easy to form hydrocarbon hydrates with water, and the hydrates are easy to cause the freezing and blocking phenomena of valves and pipelines at the temperature of below 0 ℃. Therefore, methanol to olefin (MTO/DMTO) reaction gas generally requires a drying process after passing through a compression unit before entering a subsequent separation unit. Drying the feed gas at the outlet end of the compression unit is chosen because conventional methanol to olefins product gas separation processes typically begin with a deethanization process. Ethane and lighter components are first separated from the feed gas, in which case gaseous components such as methane, carbon monoxide, nitrogen, etc. need to be removed from the relatively light ethylene and ethane in the product gas in a subsequent separation process, which is done in a cold separation process below 0 ℃, so that saturated water has to be removed at the outlet end of the compression unit.
Another problem with the prior deethanization of the dried and compressed product gas is that it is not conducive to the optimization of the heat exchange network. Because the operation pressure of the deethanizer is about 2.5MPa, the operation temperature of the tower top is about-20 ℃, and the operation temperature of the tower bottom is about 65 ℃. The gas at the top of the deethanizer enters a cold separation unit to start the deethanization and demethanization process, and the liquid at the bottom of the deethanizer enters a hot separation unit to start the depropanization process. The pressure of the cold separation process of deethylenization and demethanization is generally about 2.3MPa, and the operating temperature is-35 ℃ or below. In order to avoid the deep cooling process, the tower bottom liquid of the depropanizer is required to be used as an absorbent. The operating pressure of the depropanizing tower is about 1.5MPa, the operating temperature of the tower top is about 45 ℃, and the operating temperature of the tower kettle is about 90 ℃. Therefore, the system needs to cool the bottom liquid of the depropanizer at about 90 ℃ to a temperature range below-35 ℃. And the temperature difference between the material flow needing to be heated in the thermal separation unit and the tower bottom liquid of the depropanizing tower is small, so that the thermal recovery is not facilitated, and a large amount of cold energy needs to be input.
Therefore, in order to overcome the above-mentioned drawbacks of the prior art due to prior deethanization, the present invention provides a method for separating a methanol-to-olefin reaction gas, comprising:
(1) introducing the methanol-to-olefin reaction gas into a compression unit for compression treatment;
(2) cooling the gas phase obtained by compression treatment;
(3) in the presence of an absorbent, performing demethanization treatment on a gas-liquid two-phase obtained after cooling in a demethanizer;
(4) carrying out depropanization treatment on the tower bottom product obtained by the demethanization treatment in a depropanization tower;
(5) drying the tower top product subjected to the depropanization treatment, and then performing deethanization treatment in a deethanizer;
wherein the compression treatment is performed so that the obtained gas phase pressure is 3.2-6.0MPa, membrane separation treatment is performed in the compression treatment process, the cooling temperature in step (2) is 30-60 ℃, the absorbent is gasoline, and the demethanization treatment conditions comprise: the number of theoretical plates of the demethanizer is 20-60, the operating pressure is 3.0-5.8MPa, the temperature at the top of the tower is 30-90 ℃, and the temperature at the bottom of the tower is 50-130 ℃.
In the present invention, the methanol-to-olefin reaction gas may contain hydrogen, nitrogen oxides, carbon monoxide, carbon dioxide, dimethyl ether, water, methane, methanol, ethane, ethylene, acetylene, propane, propylene, propadiene, n-butane, isobutane, 1-butene, 2-butene, 1, 3-butadiene, cis-2-butene, trans-2-butene, n-pentane, isopentane, 1-pentene, n-hexane, isohexane, hexene, benzene, toluene, p-xylene, m-xylene, o-xylene, and 1,3, 5-trimethylbenzene. Among them, the content of ethylene may be 25 to 95 mol%, the content of propylene may be 5 to 49 mol%, the content of methane may be 0 to 10 mol%, the content of hydrogen may be 0 to 10 mol%, and the total content of other components may be 0 to 30 mol%.
According to the method of the present invention, preferably, in step (1), the compression unit includes 3-5 compressor sections, and the pressure of the methanol-to-olefin reaction gas after compression treatment by the compression unit is gradually increased to 3.2-6.0 MPa. Wherein, the temperature of the outlet of each section of the compressor section is generally not more than 140 ℃ according to the process requirements, the temperature of the inlet of each section is about 40 ℃, and if the temperature is overhigh, the compressor section is generally cooled by circulating water at the temperature of 25-35 ℃.
According to the method provided by the invention, in addition to the membrane separation treatment, alkali washing and water washing are preferably carried out in the compression treatment process, wherein the sequence of water washing and alkali washing is not particularly required, when the raw material gas contains more acidic gases, alkali washing can be carried out first, then water washing is carried out, and when the raw material gas contains more methanol, dimethyl ether and the like, alkali washing can be carried out first, and then alkali washing is carried out. The order of the alkali washing and the water washing as an integrated process and the membrane separation is not particularly limited, and it is preferable to perform the alkali washing and the water washing first and then the membrane separation treatment, so that the life of the membrane used in the membrane separation can be improved.
According to the process of the present invention, preferably, the alkali washing and the water washing are carried out at an operating pressure of 0.3 to 3.0 MPa. Typically, the caustic wash and water wash steps may be carried out at the outlet of the compressor section of stages 1-3.
In the present invention, the water washing is performed to remove organic oxides such as methanol, dimethyl ether, etc. from the raw material gas. Generally, desalted water is sprayed into the tower from the top of the water washing tower to fully contact with the raw material gas. Depending on the operating pressure, the water washing process may absorb a part of the hydrocarbon components in the feed gas, so that most of the methanol and the DME can be recovered by the prior art, and then the hydrocarbon components can be recovered by stripping and the like, and then returned to the compression unit. Wherein, the water washing process is a conventional process in the field.
In the present invention, the alkali washing is performed to remove acidic gases such as carbon dioxide from the raw material gas. Generally, alkali liquor is sprayed into the tower from the top of the alkaline tower to fully contact with the raw material gas. The alkaline washing tower is provided with an alkaline washing section and a water washing section, and gas at the top of the alkaline washing tower enters the compression unit. Among them, the alkali washing process is a process conventional in the art.
According to the method of the present invention, preferably, the membrane separation conditions include: the operating pressure is 1.5-4.0MPa, so that hydrogen and methane can be separated from the raw material gas more effectively. The membrane separation unit separates hydrogen and methane from the raw material gas by utilizing different permeability of the membrane to methane, hydrogen and other components, and stores the hydrogen and methane in the buffer tank, so that the hydrogen and methane can be prepared from the gas stored in the buffer tank by adopting a pressure swing adsorption hydrogen production technology. The pressure swing adsorption hydrogen production technology is well known in the art and will not be described in detail herein. The purity of the hydrogen obtained by separation in the invention is 99-99.999%. Pressure swing adsorption hydrogen production is abbreviated as PSA hydrogen production.
In the present invention, since methanol to olefin (MTO/DMTO) reaction gas contains a large amount of saturated water, an operation such as alkaline washing and water washing is required between compressor stages, and a water-phase condensate or a three-phase effluent containing a gas phase, a water phase and an oil phase may be generated as the pressure increases and the cooling temperature decreases. Wherein the gas phase component is a hydrocarbon material containing a small amount of saturated water, the water phase is water containing a very small amount of hydrocarbon material, and the oil phase is a hydrocarbon material containing heavier components of C2-C5. The water phase generated by each compressor section can be returned to the inlet of the previous compressor section, and finally sent to a stripping tower to recover hydrocarbon substances, so that the loss of raw materials is avoided. The oil phase generated after cooling the gas in the compression process can enter a depropanizer after being pressurized, and the oil phase generated after cooling the gas at the outlet of the compressor can enter a demethanizer.
According to the method of the present invention, preferably, in the step (2), the cooling is performed to a temperature of 30 to 60 ℃. The cooled raw material gas is gas-liquid two-phase and can be directly sent into the demethanizer without being dried. Wherein, the used refrigerant can be general circulating water or cooling water with the temperature of about 7 ℃, and the cold water can be provided by a lithium bromide absorption refrigerator. The lithium bromide refrigerator adopts an absorption refrigeration process, uses waste heat steam of a factory as a heat source, and has the advantage of low energy consumption.
According to the method of the invention, preferably, in the step (3), the absorbent is gasoline, mainly C5-C8 hydrocarbons and derivatives thereof and a small amount of heavy components above C8. The gasoline may be any gasoline known in the art, for example, the properties of which may include: the initial boiling point is 25-55 ℃, the dry point is 120-250 ℃, the sulfur content is less than 50ppm, and the sulfur-containing material can be at least one of light naphtha, heavy naphtha, catalytic reforming gasoline, catalytic cracking gasoline, fluidized catalytic cracking FCC gasoline, coal-to-liquids oil and biogasoline, but is not limited to the above. Preferably, the gasoline contains C6-C7 components, C6-C7 components account for 10-60% of the total weight of the gasoline, and the sulfur content of the gasoline is less than 50ppm, so that the absorption effect can be improved.
The amount of the absorbent used in the present invention is not particularly limited, and for example, the molar ratio of the amount of the absorbent used to the gas-liquid two-phase obtained after cooling in step (2) may be 1:10 to 10:1, preferably 1:5 to 5:1, and more preferably 1:4 to 4: 1.
The process according to the invention wherein the absorption temperature of the absorbent is lower the better the absorption, i.e. the lower the demethanizer overhead C2 and heavier components are lost, but the higher the energy consumption, preferably the absorbent temperature is in the range of 30-60℃, most preferably 30-50℃.
According to the method of the present invention, the demethanizer can be divided into an absorption section and a stripping section. The absorption section is used for absorbing C2 and heavier components in a gas phase and reducing the content of C2 components in light components such as overhead methane, preferably, the absorption section of the demethanizer can be added with an intercooler to ensure the absorption effect, more preferably, the temperature of the intercooler is 30-60 ℃, and most preferably 30-50 ℃. The stripping section is used for controlling the methane content in the tower bottom liquid, and the lower the methane content is, the more beneficial the purity of the finally prepared ethylene product is ensured. However, as the methane content in the bottom of the demethanizer decreases, the temperature of the bottom increases, and the loss of C2 and heavier components at the top of the tower increases, and the methane content in the bottom liquid of the tower is preferably controlled to be 1-10000 ppm.
According to the method of the present invention, in step (4), the depropanizing treatment conditions may be depropanizing conditions conventional in the art, and preferably, the depropanizing treatment conditions include: the theoretical plate number of the depropanizing tower is 15-50, the operation pressure is 1.5-3.5MPa, the tower top temperature is 10-50 ℃, and the tower bottom temperature is 100-160 ℃, so that the recovery rate of the ethylene and the propylene obtained by final separation can be further improved.
According to the method, in the step (4), the top of the depropanizing tower mainly contains C2 and C3 components, the bottom of the depropanizing tower mainly contains C4 and heavier components and also contains a small amount of C3 components, the gas phase at the top of the depropanizing tower needs to be cooled to reflux due to the entrainment of a certain amount of C4 and heavier components, and the temperature at the top of the depropanizing tower can be controlled within the range of 10-50 ℃, preferably within the range of 15-35 ℃. The content of the C3 component is mainly controlled in the tower bottom of the depropanizing tower, the lower the content of the C3 component in the tower bottom is, the higher the recovery rate of the C2 and C3 components of the system is, and the higher the temperature of the tower bottom is; the higher the temperature of the tower kettle is, the higher the entrainment of the C4 at the top of the tower is, and the higher the energy consumption is required. Therefore, the molar content of the column bottom C3 component can be controlled within the range of 1-25%, preferably within the range of 5-20%.
According to the method provided by the invention, in the step (4), preferably, the tower bottom liquid obtained by the depropanization treatment is recycled as the absorbent for the demethanization treatment, and because a certain temperature difference exists between the depropanization tower bottom liquid and the demethanization tower and between the depropanization tower bottom liquid and the deethanization tower, more preferably, the tower bottom liquid obtained by the depropanization treatment is firstly used as a heat source for the demethanization treatment and the deethanization treatment before being recycled as the absorbent, and is finally cooled to 30-60 ℃ step by the circulating water and the chilled water, preferably to 30-50 ℃, so that the energy consumption can be reduced, and the cost can be saved. In addition, a small part of the absorbent is discharged with the top gas phase of the demethanizer, therefore, in the method of the present invention, it is preferable to mix the bottom liquid obtained from the depropanization treatment with a part of fresh absorbent and then introduce the mixture into the demethanizer to ensure the absorbent flow rate of the absorber in the system, wherein the ratio of the bottom liquid obtained from the depropanization treatment recycled as the absorbent to the fresh absorbent is not particularly limited, and for example, the weight ratio of the two may be 1: 0-0.1.
According to the method of the present invention, in step (5), the drying is a drying process conventional in the art, aiming at removing saturated water from the depropanizer overhead gas. The drying method of the molecular sieve drying system can be adopted, and the specific method can comprise the following steps: the dryer needs to be alternately operated from start to finish, and when the dryer is alternately operated, one of the previously operated dryers needs to be subjected to a regeneration operation. During the regeneration operation, the nitrogen is required to be heated by high-pressure steam for regenerating the drying agent. The drying process gas pressure produces only a small pressure drop, typically not exceeding 200 kPa.
According to the method of the invention, in the step (5), the depropanized overhead product is dried, cooled to-10 ℃ to 30 ℃ and then subjected to deethanization treatment. The cooling can be carried out by cold energy generated in the ethylene rectification and recovery process so as to further reduce the energy consumption.
According to the method of the present invention, preferably, the condition of the deethanizing treatment may be a condition of the deethanizing treatment conventional in the art, and preferably, the condition of the deethanizing treatment includes: the number of theoretical plates of the deethanizer is 25-120, the operating pressure is 1.0-3.0MPa, the temperature at the top of the tower is-30 ℃ to-10 ℃, and the temperature at the bottom of the tower is 50-80 ℃, so that the recovery rate of ethylene and propylene finally obtained by separation can be further improved.
According to the method of the present invention, the deethanizer can clearly separate the C2 and C3 components, and by controlling the deethanizer conditions within the above preferred ranges, the propylene mole content in the deethanizer overhead product can be made not higher than 0.5%, preferably not higher than 0.1%.
According to the method provided by the invention, preferably, the tower top product subjected to deethanization in the step (5) is subjected to hydrogenation and ethylene rectification in sequence, ethylene is obtained from an ethylene rectification tower by a conventional method or an intermediate extraction method according to process requirements, a certain amount of methane possibly exists in the material sent into the ethylene rectification tower by the deethanization tower, and the intermediate extraction method is preferably adopted, so that a polymerization-grade ethylene product can be obtained, when the intermediate extraction method is adopted, the ethylene rectification tower bottom product is ethane, ethylene is extracted from the side line of the tower bottom, the 2 nd to 8 th tower plates can be selected from the top to the bottom at a common extraction position, and the ethylene rectification tower top product is preferably returned to a compressor section with similar pressure. When conventional processes are employed, the overhead product is ethylene and the bottoms product is ethane. Among these, hydrogenation is for converting acetylene to ethylene, and this technique can be provided by a known technique.
According to the method provided by the invention, preferably, the overhead product of the deethanization treatment in the step (5) is hydrogenated, cooled to-35 ℃ to-10 ℃ and then introduced into an ethylene rectifying tower for ethylene rectification. Wherein the cooling can be performed by cold energy generated in the ethylene rectification recovery process, so as to further reduce energy consumption.
The conditions for ethylene rectification according to the process of the present invention may be those conventional in the art, and may include, for example: the theoretical plate number of the ethylene rectifying tower is 100-150, the operation pressure is 0.8-2.6MPa, the tower top temperature is-39 ℃ to-15 ℃, and the tower bottom temperature is-20 ℃ to 0 ℃.
According to the method provided by the invention, preferably, the tower bottom product subjected to deethanization treatment in the step (5) is subjected to propylene rectification, propylene is obtained from a propylene rectification tower through an intermediate extraction method, the tower bottom product of the propylene rectification tower is propane, and the tower top product of the propylene rectification tower is returned to a compressor section with similar pressure. The specific method of the intermediate extraction method employed herein may be the same as the intermediate extraction method for obtaining ethylene, except that the extraction position is selected from the 2 nd to 15 th inter-tray plates counted from the top.
The conditions for the rectification of propylene according to the process of the present invention may be those conventional in the art, and may include, for example: the theoretical plate number of the propylene rectifying tower is 110-400, the operation pressure is 0.8-2.6MPa, the tower top temperature is 10-55 ℃, and the tower kettle temperature is 35-60 ℃.
The recovery rate of the ethylene separated by the method is more than 95 percent, and the recovery rate of the propylene is more than 95 percent.
In the following examples and comparative examples, methanol to olefin reaction gas was composed of ethylene, propylene, methane, hydrogen and other components (nitrogen, nitrogen oxides, carbon monoxide, carbon dioxide, dimethyl ether, water, methanol, ethane, acetylene, propane, propadiene, n-butane, isobutane, 1-butene, 2-butene, 1, 3-butadiene, cis-2-butene, trans-2-butene, n-pentane, isopentane, 1-pentene, n-hexane, isohexane, hexene, benzene, toluene, p-xylene, m-xylene, o-xylene and 1,3, 5-trimethylbenzene), among them, the content of methane was 4.1 mol%, the content of hydrogen was 1.8 mol%, the content of ethylene was 40.7 mol%, the content of propylene was 37.5 mol%, the content of water was 5.7 mol%, and the content of other components was 10.2 mol%.
Example 1
This example is intended to illustrate the process for separating a methanol to olefin reaction gas according to the present invention.
(1) As shown in fig. 1, a methanol-to-olefin reaction gas raw material is introduced into a compression unit for compression treatment, the compression unit consists of 3 sections of compressor sections, the outlet temperature of each section of the compressor section is not more than 140 ℃, the inlet temperature of each section is 40 ℃, 30 ℃ circulating water is adopted for cooling, the pressure of the methanol-to-olefin reaction gas after compression treatment by the compression unit is increased to 3.9MPa step by step, wherein after the outlet of the 1 st section of the compressor section is discharged from a tank (the pressure is 0.55MPa), alkali washing treatment and water washing treatment are sequentially carried out by an alkali washing tower and a water washing tower which are arranged, alkali liquor enters the tower from the tower top of the alkali washing tower in the alkali washing treatment process and is fully contacted with the raw material gas, desalted water is adopted in the water washing treatment process and is sprayed into the tower from the tower top of the water washing tower to be fully contacted with the raw material gas, and a product after. In addition, a membrane separation unit is arranged at the outlet of the 2 nd compressor section, and the membrane separation conditions are as follows: the operation pressure is 2.2MPa, the hydrogen and the methane are separated from the raw material gas and are stored in a buffer tank, and the removal rate of the methane is 68 percent. And preparing hydrogen and methane (fuel gas) from the gas stored in the buffer tank by adopting a PSA (pressure swing adsorption) hydrogen production technology, wherein the purity of the prepared hydrogen is 99.999%. Gas phase, water phase and oil phase products are generated in the compression treatment process, water phase condensate generated by each section of compressor section returns to the inlet of the previous section of compressor section, and finally the condensate is sent to a stripping tower to recover light hydrocarbon substances so as to avoid raw material loss, the oil phase generated after the gas at the outlet of the 3 rd section of compressor is cooled enters a demethanizer, and no oil phase is generated between the compressor sections;
(2) cooling the gas phase obtained by compression treatment to 30 ℃, wherein the cooled raw material gas is a gas-liquid two-phase gas and can be directly sent to a demethanizer without being dried, and the used refrigerant is cooling water at about 7 ℃;
(3) as shown in fig. 2, absorbent gasoline (initial boiling point 42 ℃, dry point 187 ℃, C6-C7 component in the gasoline 42.7% of the total weight, sulfur content less than 50ppm) is cooled to 30 ℃, and then introduced into a demethanizer, the gas-liquid two-phase stream obtained after cooling is introduced into the demethanizer for demethanization, the molar ratio of the absorbent usage to the stream entering the demethanizer after cooling the compressor outlet gas is 1:1.91, the absorption section of the demethanizer is provided with an intercooler, the temperature of the intercooler is 30 ℃, the theoretical plate number of the demethanizer is controlled to be 50, the operating pressure is 3.7MPa, the overhead temperature is 34.9 ℃, and the bottom temperature of the tower is controlled to be 63.4 ℃, so that the methane content in the tower bottom liquid is controlled to be 4786 ppm;
(4) introducing tower bottom liquid obtained by the demethanization into a depropanization tower for depropanization treatment; the number of theoretical plates of the depropanizing tower is 40, the operating pressure is 2.6MPa, the tower top temperature is 31.3 ℃, the tower bottom temperature is 139.6 ℃, wherein the molar content of the C3 component in the tower bottom liquid is 1.53 percent;
(5) the tower bottom liquid obtained by the depropanization treatment is used as a heat source for the demethanization treatment and the deethanization treatment, and is cooled to 30 ℃ by circulating water, a part of the tower bottom liquid (light hydrocarbon) is sent out of a battery compartment and sent to a debutanizer, and the rest of the tower bottom liquid is used as a circulating solvent and mixed with a part of fresh absorbent and then sent into the demethanizer (the weight ratio of the circulating solvent to the fresh absorbent is 1:0.08), so that the molar ratio of the consumption of the absorbent to the flow of the cooled compressor outlet gas entering the demethanizer is 1: 1.91. Drying the tower top product subjected to depropanization treatment by adopting a molecular sieve drying system, and then cooling to 6.8 ℃, wherein the cooling is carried out by cold energy generated in the ethylene rectification recovery process, and the specific operation method of the molecular sieve drying system comprises the following steps: the dryer needs to be alternately operated from start to finish, and when the dryer is alternately operated, one of the previously operated dryers needs to be subjected to a regeneration operation. During the regeneration operation, the nitrogen is required to be heated by high-pressure steam for regenerating the drying agent. Introducing the cooled tower top product subjected to depropanization into a deethanizer for deethanization, wherein the number of theoretical plates of the deethanizer is 80, the operating pressure is 2.3MPa, the tower top temperature is-22.8 ℃, and the tower bottom temperature is 53.6 ℃, so that the molar content of propylene in the tower top product of the deethanizer is 0.005%;
(6) after hydrogenation, cooling the tower top product subjected to deethanization in the step (5) to-28.3 ℃, introducing the cooled tower top product into an ethylene rectifying tower for ethylene rectification, wherein the cooling can be carried out by cold energy generated in the ethylene rectification recovery process, the number of theoretical plates of the ethylene rectifying tower is 100, the operating pressure is 1.8MPa, the tower top temperature is-33.4 ℃, the tower bottom temperature is-9.2 ℃, the product of the ethylene rectifying tower is ethane, and the product of tower top gas is ethylene;
(7) and (3) performing propylene rectification on the tower bottom product subjected to deethanization treatment in the step (5), obtaining propylene from a propylene rectification tower by adopting an intermediate extraction method, wherein the tower bottom product of the propylene rectification tower is propane, the tower top product of the propylene rectification tower is returned to a compressor section with similar pressure, the propylene extraction position is selected from 2 nd to 15 th tower plates counted from top to bottom, the number of theoretical plates of the propylene rectification tower is 170, the operating pressure is 1.95MPa, the tower top temperature is 18.9 ℃, and the tower bottom temperature is 53.7 ℃.
Example 2
This example is intended to illustrate the process for separating a methanol to olefin reaction gas according to the present invention.
(1) Introducing a methanol-to-olefin reaction gas raw material into a compression unit for compression treatment, wherein the compression unit consists of 3 sections of compressor sections, the outlet temperature of each section of the compressor section is not more than 140 ℃, the inlet temperature of each section is 40 ℃, 30 ℃ circulating water is adopted for cooling, the pressure of the methanol-to-olefin reaction gas after compression treatment by the compression unit is gradually increased to 3.9MPa, wherein after the methanol-to-olefin reaction gas is discharged out of a tank at the outlet of the 1 st compressor section (the pressure is 0.65MPa), alkali washing treatment and water washing treatment are sequentially carried out by an alkali washing tower and a water washing tower, alkali liquor is sprayed into the tower from the top of the alkali washing tower in the alkali washing treatment process and fully contacts with the raw material gas, desalted water is sprayed into the tower from the top of the water washing tower in the water washing treatment process and fully contacts with the raw material gas, and the product after water washing is sent. In addition, a membrane separation unit is arranged at the outlet of the 3 rd compressor section, and the membrane separation conditions are as follows: the operation pressure is 2.0MPa, the hydrogen and the methane are separated from the raw material gas and are stored in a buffer tank, and the removal rate of the methane is 75 percent. And preparing hydrogen and methane (fuel gas) from the gas stored in the buffer tank by adopting a PSA hydrogen production technology, wherein the purity of the prepared hydrogen is 99.9%. Gas phase, water phase and oil phase products are generated in the compression treatment process, water phase condensate generated by each section of compressor section returns to the inlet of the previous section of compressor section, and finally the condensate is sent to a stripping tower to recover light hydrocarbon substances so as to avoid raw material loss, the oil phase generated after the gas at the outlet of the 3 rd section of compressor is cooled enters a demethanizer, and no oil phase is generated between the compressor sections;
(2) cooling the gas phase obtained by compression treatment to 35 ℃, wherein the cooled raw material gas is a gas-liquid two-phase gas and can be directly sent to a demethanizer without being dried, and the used refrigerant is cooling water at about 7 ℃;
(3) as shown in fig. 2, absorbent gasoline (initial boiling point is 32 ℃, dry point is 192 ℃, C6-C7 components in the gasoline account for 51.2% of the total weight, and sulfur content is less than 50ppm) is cooled to 35 ℃, and then introduced into a demethanizer, the gas-liquid two-phase stream obtained after cooling is introduced into the demethanizer for demethanization treatment, the molar ratio of the usage of the absorbent to the stream entering the demethanizer after cooling the gas at the outlet of the compressor is 1:1.71, an intermediate cooler is arranged at the absorption section of the demethanizer, the temperature of the intermediate cooler is 35 ℃, the number of theoretical plates of the demethanizer is controlled to be 35, the operating pressure is 3.6MPa, the temperature at the top of the tower is 42.1 ℃, and the temperature at the bottom of the tower is 65.1 ℃, so that the methane content in the liquid of the tower is controlled to be 4600 ppm;
(4) introducing tower bottom liquid obtained by the demethanization into a depropanization tower for depropanization treatment; the number of theoretical plates of the depropanizing tower is 35, the operating pressure is 2.7MPa, the tower top temperature is 31.3 ℃, the tower bottom temperature is 139.4 ℃, wherein the molar content of the C3 component in the tower bottom liquid is 1.54 percent;
(5) taking tower bottom liquid obtained by the depropanization treatment as a heat source for the demethanization treatment and the deethanization treatment, cooling to 35 ℃ by circulating water, sending a part of the tower bottom liquid (light hydrocarbon) out of a battery compartment to a debutanizer, mixing the rest circulating solvent and a part of fresh absorbent, and introducing the mixture into the demethanizer (the weight ratio of the circulating solvent to the fresh absorbent is 1:0.07), and ensuring that the molar ratio of the consumption of the absorbent to the flow of the cooled compressor outlet gas entering the demethanizer is 1: 1.71. Drying the tower top product subjected to depropanization treatment by adopting a molecular sieve drying system, and then cooling to 7.3 ℃, wherein the cooling is carried out by cold energy generated in the ethylene rectification recovery process, and the specific operation method of the molecular sieve drying system comprises the following steps: the dryer needs to be alternately operated from start to finish, and when the dryer is alternately operated, one of the previously operated dryers needs to be subjected to a regeneration operation. During the regeneration operation, the nitrogen is required to be heated by high-pressure steam for regenerating the drying agent. Introducing the cooled tower top product subjected to depropanization into a deethanizer for deethanization, wherein the number of theoretical plates of the deethanizer is 70, the operating pressure is 2.4MPa, the tower top temperature is-21 ℃, and the tower bottom temperature is 55.9 ℃, so that the molar content of propylene in the tower top product of the deethanizer is 0.005%;
(6) after hydrogenation, cooling the tower top product subjected to deethanization in the step (5) to-27.8 ℃, introducing the cooled tower top product into an ethylene rectifying tower for ethylene rectification, wherein the cooling can be carried out by cold energy generated in the ethylene rectification recovery process, the number of theoretical plates of the ethylene rectifying tower is 110, the operating pressure is 1.7MPa, the tower top temperature is-34.6 ℃, the tower bottom temperature is-13.5 ℃, the product of the ethylene rectifying tower is ethane, and the product of tower top gas is ethylene;
(7) and (3) performing propylene rectification on the tower bottom product subjected to deethanization treatment in the step (5), obtaining propylene from a propylene rectification tower by adopting an intermediate extraction method, wherein the tower bottom product of the propylene rectification tower is propane, the tower top product of the propylene rectification tower is returned to a compressor section with similar pressure, the propylene extraction position is selected from 2 nd to 15 th tower plates counted from top to bottom, the number of theoretical plates of the propylene rectification tower is 180, the operating pressure is 1.85MPa, the tower top temperature is 19.2 ℃, and the tower bottom temperature is 51.2 ℃.
Example 3
This example is intended to illustrate the process for separating a methanol to olefin reaction gas according to the present invention.
(1) Introducing a methanol-to-olefin reaction gas raw material into a compression unit for compression treatment, wherein the compression unit consists of 4 sections of compressor sections, the outlet temperature of each section of the compressor section is not more than 140 ℃, the inlet temperature of each section is 40 ℃, 30 ℃ circulating water is adopted for cooling, the pressure of the methanol-to-olefin reaction gas after compression treatment by the compression unit is gradually increased to 3.9MPa, wherein after the methanol-to-olefin reaction gas is discharged out of a tank at the outlet of the 2 nd compressor section (the pressure is 1.2MPa), alkali washing treatment and water washing treatment are sequentially carried out by an alkali washing tower and a water washing tower which are arranged, alkali liquor is sprayed into the tower from the top of the alkali washing tower in the alkali washing treatment process and fully contacts with the raw material gas, desalted water is sprayed into the tower from the top of the water washing tower in the water washing treatment process and fully contacts with the raw material gas, and the product after water. In addition, a membrane separation unit is arranged at the outlet of the 3 rd compressor section, and the membrane separation conditions are as follows: the operation pressure is 2.6MPa, the hydrogen and the methane are separated from the raw material gas and are stored in a buffer tank, and the removal rate of the methane is 70.7 percent. And preparing hydrogen and methane (fuel gas) from the gas stored in the buffer tank by adopting a PSA hydrogen production technology, wherein the purity of the prepared hydrogen is 99.9%. Gas phase, water phase and oil phase products are generated in the compression treatment process, water phase condensate generated by each section of compressor section returns to the inlet of the previous section of compressor section, and finally the condensate is sent to a stripping tower to recover light hydrocarbon substances so as to avoid raw material loss, the oil phase generated after the gas at the outlet of the 3 rd section of compressor is cooled is pressurized and then enters a depropanizing tower, and the oil phase generated after the gas at the outlet of the 4 th section of compressor is cooled enters the depropanizing tower;
(2) cooling the gas phase obtained by compression treatment to 40 ℃, wherein the cooled raw material gas is a gas-liquid two-phase gas and can be directly sent to a demethanizer without being dried, and the used refrigerant is cooling water at about 7 ℃;
(3) as shown in fig. 2, absorbent gasoline (initial boiling point 51 ℃, dry point 243 ℃, C6-C7 component in the gasoline 42.1% of the total weight, sulfur content less than 50ppm) is cooled to 40 ℃, and then introduced into a demethanizer, the gas-liquid two-phase stream obtained after cooling is introduced into the demethanizer for demethanization, the molar ratio of the absorbent usage to the stream entering the demethanizer after cooling the compressor outlet gas is 1:1.39, the absorption section of the demethanizer is provided with an intercooler, the temperature of the intercooler is 40 ℃, the theoretical plate number of the demethanizer is controlled to be 45, the operating pressure is 3.7MPa, the overhead temperature is 45.6 ℃, and the bottom temperature is 71.6 ℃, so that the methane content in the bottom liquid is controlled to be 4250 ppm;
(4) introducing tower bottom liquid obtained by the demethanization into a depropanization tower for depropanization treatment; the number of theoretical plates of the depropanizing tower is 35, the operating pressure is 2.6MPa, the tower top temperature is 31.3 ℃, the tower bottom temperature is 139.6 ℃, wherein the molar content of the C3 component in the tower bottom liquid is 1.58 percent;
(5) the tower bottom liquid obtained by the depropanization treatment is used as a heat source for the demethanization treatment and the deethanization treatment, and is cooled to 40 ℃ by circulating water, a part of the tower bottom liquid (light hydrocarbon) is sent out of a battery compartment and sent to a debutanizer, and the rest of the tower bottom liquid is used as a circulating solvent and mixed with a part of fresh absorbent and then sent into the demethanizer (the weight ratio of the circulating solvent to the fresh absorbent is 1:0.06), so that the molar ratio of the consumption of the absorbent to the flow of the cooled compressor outlet gas entering the demethanizer is 1: 1.39. Drying the tower top product subjected to depropanization treatment by adopting a molecular sieve drying system, and then cooling to 5.6 ℃, wherein the cooling is carried out by cold energy generated in the ethylene rectification recovery process, and the specific operation method of the molecular sieve drying system comprises the following steps: the dryer needs to be alternately operated from start to finish, and when the dryer is alternately operated, one of the previously operated dryers needs to be subjected to a regeneration operation. During the regeneration operation, the nitrogen is required to be heated by high-pressure steam for regenerating the drying agent. Introducing the cooled depropanized tower top product into a deethanizer for deethanizing, wherein the number of theoretical plates of the deethanizer is 60, the operating pressure is 2.3MPa, the tower top temperature is-22.9 ℃, and the tower bottom temperature is 53.6 ℃, so that the molar content of propylene in the deethanized tower top product is 0.003%;
(6) hydrogenating the tower top product subjected to deethanization in the step (5), cooling to-28.2 ℃, introducing into an ethylene rectifying tower for ethylene rectification, wherein the cooling can be carried out by cold energy generated in the ethylene rectification recovery process, the theoretical plate number of the ethylene rectifying tower is 110, the operating pressure is 1.9MPa, the tower top temperature is-31.5 ℃, the tower bottom temperature is-9.2 ℃, the ethylene rectifying tower bottom product is ethane, obtaining ethylene from the ethylene rectifying tower by adopting an intermediate extraction method, returning the ethylene rectifying tower top product to a compressor section with similar pressure, and selecting an ethylene extraction position to count the 2-8 tower plates from top to bottom;
(7) and (3) performing propylene rectification on the tower bottom product subjected to deethanization treatment in the step (5), obtaining propylene from a propylene rectification tower by adopting an intermediate extraction method, wherein the tower bottom product of the propylene rectification tower is propane, the tower top product of the propylene rectification tower is returned to a compressor section with similar pressure, the propylene extraction position is selected from 2 nd to 15 th tower plates counted from top to bottom, the number of theoretical plates of the propylene rectification tower is 260, the operating pressure is 1.95MPa, the tower top temperature is 18.9 ℃, and the tower bottom temperature is 53.7 ℃.
Example 4
This example is intended to illustrate the process for separating a methanol to olefin reaction gas according to the present invention.
The methanol to olefin reaction gas was separated as in example 1, except that the compressor exit gas, the absorbent and the cooler in the demethanizer were all cooled to 60 ℃ and the molar ratio of the absorbent usage to the stream entering the demethanizer after cooling the compressor exit gas was 1.04: 1.
Example 5
This example is intended to illustrate the process for separating a methanol to olefin reaction gas according to the present invention.
The methanol-to-olefin reaction gas was separated according to the method of example 1, except that the membrane separation was directly placed after the water scrubber, i.e. after the alkali wash and the water scrubber, the membrane separation treatment was directly performed without the compression treatment, and the membrane separation conditions were as follows: the operating pressure was 1.0 MPa.
Comparative example 1
The methanol to olefin reaction gas was separated as in example 1, except that the number of theoretical plates of the demethanizer was controlled to 18, the operating pressure was 3.3MPa, the overhead temperature was 68.1 ℃ and the bottom temperature was 59 ℃.
TABLE 1
Note: the steam is saturated steam of 0.8 MPa.
As can be seen from the data in Table 1, the recovery rate of ethylene and the recovery rate of propylene obtained by separation by the method are respectively over 99 percent and over 99 percent, and the method has low energy consumption and high operation flexibility and can adjust energy consumption process parameters according to product income.
The method utilizes the gasoline component as an absorbent to remove light components such as methane in the methanol-to-olefin (MTO/DMTO) reaction gas, and removes methane and light fractions in the methanol-to-olefin (MTO/DMTO) product gas in an operating temperature range of more than 30 ℃, thereby effectively inhibiting water and hydrocarbons from forming hydrates. In addition, the drying process of the invention does not need to be placed at the outlet of the compression unit, and only the top gas of the deethanizer needs to be dried, so that the liquid-phase condensate at the outlet of the compression unit is prevented from being dried, the treatment capacity in the drying process is greatly reduced, and the consumption and energy consumption of the nitrogen in public works are greatly reduced. The method converts the demethanization process into the thermal separation process, is favorable for optimizing a heat exchange network, and does not need to input a refrigerant below minus 40 ℃. Has the characteristics of low investment, high material recovery rate and the like.
The preferred embodiments of the present invention have been described in detail, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.
It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.
In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the disclosure of the present invention as long as it does not depart from the spirit of the present invention.
Claims (16)
1. A method for separating methanol-to-olefin reaction gas is characterized by comprising the following steps:
(1) introducing the methanol-to-olefin reaction gas into a compression unit for compression treatment;
(2) cooling the gas phase obtained by compression treatment;
(3) in the presence of an absorbent, performing demethanization treatment on a gas-liquid two-phase obtained after cooling in a demethanizer;
(4) carrying out depropanization treatment on the tower bottom product obtained by the demethanization treatment in a depropanization tower;
(5) drying the tower top product subjected to the depropanization treatment, and then performing deethanization treatment in a deethanizer;
wherein, in the step (1), the compression treatment is carried out to obtain a gas phase pressure of 3.2-6.0MPa, and a membrane separation treatment is carried out in the compression treatment, wherein the membrane separation condition comprises that: the operating pressure is 1.5-4.0 MPa;
wherein, in the step (2), the cooling is carried out to the temperature of 30-60 ℃;
wherein, in the step (3), the absorbent is gasoline, the gasoline contains C6-C7 components, the C6-C7 components account for 10-60% of the total weight of the gasoline, the sulfur content of the gasoline is less than 50ppm, the initial boiling point of the gasoline is 25-55 ℃, the dry point of the gasoline is 120-250 ℃, the temperature of the absorbent is 30-60 ℃, and the condition of the demethanization treatment comprises the following steps: the number of theoretical plates of the demethanizer is 20-60, the operating pressure is 3.0-5.8MPa, the temperature at the top of the tower is 30-90 ℃, and the temperature at the bottom of the tower is 50-130 ℃;
wherein, the tower bottom liquid obtained by the depropanization treatment is recycled as the absorbent for the demethanization treatment.
2. The process according to claim 1, wherein the temperature of the absorbent is 30-50 ℃.
3. The method according to claim 1, wherein in step (2), the cooling is carried out to a temperature of 30-50 ℃.
4. The method of claim 1, wherein the compression unit comprises a 3-5-stage compressor section, and the pressure of the methanol-to-olefin reaction gas is increased to 3.2-6.0MPa step by step after the compression treatment of the compression unit.
5. A method according to claim 1 or 4, wherein a caustic wash and a water wash are performed during the compression treatment.
6. The method according to claim 5, wherein the alkaline and aqueous washes are performed at an operating pressure of 0.3-3.0 MPa.
7. The method of claim 1, wherein the method further comprises: and performing pressure swing adsorption on the gas obtained by the membrane separation treatment to prepare hydrogen so as to obtain the hydrogen.
8. The method of claim 1, wherein the depropanizing conditions comprise: the number of theoretical plates of the depropanizing tower is 15-50, the operating pressure is 1.5-3.5MPa, the tower top temperature is 10-50 ℃, and the tower bottom temperature is 100-160 ℃.
9. The method of claim 1, wherein the conditions of the deethanization process comprise: the number of theoretical plates of the deethanizer is 25-120, the operating pressure is 1.0-3.0MPa, the temperature at the top of the tower is-30 ℃ to-10 ℃, and the temperature at the bottom of the tower is 50-80 ℃.
10. The method of claim 1, wherein the method further comprises: in the step (5), the tower top product after the depropanization treatment is dried, cooled to the temperature of minus 10 ℃ to 30 ℃ and then subjected to deethanization treatment.
11. The method of claim 1 or 10, wherein the method further comprises: and (3) sequentially carrying out hydrogenation and ethylene rectification on the tower top product subjected to deethanization in the step (5), obtaining ethylene from an ethylene rectification tower, returning the tower top product of the ethylene rectification tower to a compressor section with similar pressure, wherein the tower bottom product of the ethylene rectification tower is ethane.
12. The method of claim 11, wherein the conditions for ethylene rectification comprise: the theoretical plate number of the ethylene rectifying tower is 100-150, the operation pressure is 0.8-2.6MPa, the tower top temperature is-39 ℃ to-15 ℃, and the tower bottom temperature is-20 ℃ to 0 ℃.
13. The method of claim 1 or 10, wherein the method further comprises: and (5) performing propylene rectification on the tower bottom product subjected to deethanization treatment in the step (5), and obtaining propylene from a propylene rectification tower, wherein the tower bottom product of the propylene rectification tower is propane.
14. The process of claim 13, wherein the conditions for the rectification of propylene comprise: the theoretical plate number of the propylene rectifying tower is 110-400, the operation pressure is 0.8-2.6MPa, the tower top temperature is 10-55 ℃, and the tower kettle temperature is 35-60 ℃.
15. The method of claim 1, wherein the method further comprises: and an oil phase generated after cooling the gas at the outlet of the compressor enters a demethanizer.
16. The method of claim 1, wherein the bottoms from the depropanizing treatment is used as a heat source for the demethanizing and deethanizing treatments before being recycled as the absorbent.
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