CN110156553A - A separation process for direct conversion of methane to produce ethylene products - Google Patents

Abstract

The invention relates to the field of methane conversion technology and petrochemical industry, in particular to a separation process for directly converting methane to produce ethylene products. A separation process for the direct conversion of methane to ethylene products. The product S0 of the direct conversion of methane to ethylene is dehydrated, decarburized and dried to obtain rich in methane, ethylene, ethane and C2+ hydrocarbons (C2+ is hydrocarbons with a carbon number greater than 2) Class, mainly propylene and propane) mixture S1, the mixture S1 is input to the input port of the lower end of the methane fractionation tower, and the solvent S2 is input to the upper end input port of the methane fractionation tower. The invention solves the product separation problem of methane conversion to ethylene.

Description

A separation process for direct conversion of methane to produce ethylene products

technical field

The invention relates to the field of methane conversion technology and petrochemical industry, in particular to a separation process for direct conversion of methane to produce ethylene products.

Background technique

Ethylene is not only an important bulk petrochemical product, but also a bulk chemical raw material. At present, the routes for producing ethylene include petroleum-based routes and coal-based routes, among which the petroleum-based routes are mainly hydrocarbon cracking methods, that is, ethylene is produced by tubular thermal cracking with ethane and propane as cracking raw materials. However, as the demand for olefins continues to increase, only ethylene and propane as cracking raw materials are far from meeting the market demand for olefins, and the cracking raw materials have begun to develop towards heavy ones. In recent years, the coal-based ethylene route MTO technology developed by the Dahua Institute of the Chinese Academy of Sciences and the Shanghai Petrochemical Engineering Company has become a new type of ethylene production method. MTO technology, that is, carbon-containing resources (coal, natural gas or biomass) are converted to olefins through methanol, and through this method, coal gasification, gas purification, methanol synthesis, methanol refining, methanol to olefins, olefin separation and other steps are used to obtain ethylene and propylene products . This method provides an alternative method for the production of ethylene and propylene under the background of my country's energy structure of "lack of oil, less gas, and relatively rich coal". However, this method has the disadvantages of long synthesis route and excessive investment in equipment.

Judging from the existing published technologies, there are few separation processes for OCM products, but the low-carbon hydrocarbon components obtained from the OCM product components and the petrochemical route are very similar, the only difference is that the product composition of the two is quite different . Therefore, reviewing the patented technology for the separation of low-carbon hydrocarbons in the petrochemical route has reference significance for the development of separation processes for OCM products.

At present, the patented technologies for the separation of methane and ethylene mainly include cryogenic distillation and solvent absorption-desorption. The cryogenic distillation method proposed in the early years has been widely used because of its high technical maturity, but its disadvantage is that the required cryogenic temperature is below -140°C, and the resulting compression refrigeration cycle operation costs are extremely high. Material selection and processing requirements are high, and equipment investment is large. For this reason, McCue et al. (1990) proposed two-stage cryogenic separation, and changed the cryogenic to two stages, that is, first cryogenic to -38-17°C, and then cryogenic to <-38°C, which greatly reduced the overall cryogenic temperature. demand. However, even so, high operating costs and large investment in equipment are still the main problems of cryogenic distillation.

On the contrary, the solvent extraction (absorption)-desorption method can increase the cryogenic temperature from -165°C to -40°C, greatly reducing the operating cost, for example: Mehra (1988, 1988.05.10) proposed based on Solvent extraction-desorption separation process, using polyglycol dimethyl ether (boiling point 250 ° C), N-methylpyrrolidone (boiling point 203 ° C), N, N-dimethylformamide (boiling point 153 ° C), propylene carbonate Ester (boiling point 242°C), sulfolane (boiling point 285°C), ethylene triacetate (/), C8-C10 aromatics (boiling point 144°C), etc. are used as extraction (absorption) agents to reduce operating costs and increase ethylene recovery . However, this process also has the following problems: the extraction agent used has a higher boiling point, which makes the required temperature at the lower end of the solvent fractionation column higher, which is difficult to achieve with conventional steam heating methods. Other unconventional heating methods (such as: electricity) It is bound to cause an increase in operating costs.

Bauer (1996, 1996.05.28) also proposed a process for recovering C2+ from cracked gas. This process also uses the solvent absorption-desorption method, using alkanes with a molar mass of 50-75 g/mol (such as n-pentane, iso Pentane or its mixture) as solvent, achieved better results. However, this process has the following problems: Although n-pentane, isopentane or their mixtures have good absorption capacity for C2+ hydrocarbons, it is difficult to avoid a certain amount of C5 hydrocarbon solvents in the vapor phase recovery at the top of the absorption tower. , resulting in solvent loss; at the same time, the boiling point difference between C5 hydrocarbons and C3 hydrocarbons is small, and solvent recovery is difficult;

Minkkinen (2002, 2002.03.19) proposed a solution absorption-regeneration process for separating ethane and ethylene. This process uses a C3 fraction hydrocarbon scrubber-solvent absorption tower. The selected solvent is from toluene, pentane, hexane, and Choose from toluene-benzene mixtures. Although this process has many advantages, the use of toluene and toluene-benzene mixture still has the problems of high temperature in the solvent fractionation tower and low solvent absorption capacity, making it difficult to use conventional heating methods, resulting in increased operating costs; while using pentane , hexane solvent, there is a relatively high solvent loss; in addition, the process is relatively long, suitable for feed compositions with light component (CH4-) content that is equivalent or even slightly higher than C2+ content, and light C3+ hydrocarbon solvent circulation requires additional For C3+ hydrocarbons, the loss during the cycle cannot be ignored.

Li Lixin (2008, 2008.10.15) proposed a separation method for light hydrocarbon products in MTO/MTP reaction products. The phase is sent to the absorption steam tower; the top product of the absorption steam tower is sent to the oil absorption tower after being cooled, and the oil absorption tower uses ethane as a lean absorbent to evaporate the ethylene in the feed, and the light gas is discharged from the top of the tower, and the tower The kettle product is returned to the absorption and distillation tower as rich absorbent. Although this process can obtain ethylene products with a polymerization purity greater than 99.95% and a recovery rate greater than 99.6%, the vapor phase of the oil absorption tower overhead steam contains relatively high content (～30～40%) of ethane, resulting in a large amount of ethane The loss of ethane solvent affects the economics of the process.

Jiang Jiagui (2014, 2014.01.23) proposed a medium-cooled solvent washing separation method for low-carbon hydrocarbons. This method uses a mixture of ethane and C4 hydrocarbons, or a mixture of ethane and C4+ hydrocarbons, or a mixture of ethane and C5+ hydrocarbons. mixture, or a mixture of propane and any of the above mixtures as absorbent. However, from the perspective of engineering experience, although this method has the advantages of reducing investment and energy consumption, it is difficult to ensure that the light components produced in the vapor phase at the top of the demethanizer do not contain ethylene and ethane, which results in the loss of C2 hydrocarbons , which reduces the recovery rate of C2 hydrocarbons, especially ethylene, and the economy is damaged.

In the utility model CN205024119U, Li Xingang (2015, 2015.06.10) proposed a three-tower device for recovering ethylene and ethane by oil absorption and dry gas. This device includes a compressor, an absorption tower, a desorption tower and a rectification tower. C5 is used as the absorption agent. Although this device has the advantages of low investment and low operating costs, the solvent used in the patent is similar to that reported by Bauer, and the specific components of C5 are not disclosed, the creativity is not obvious, and the technical disclosure is insufficient; meanwhile, its absorption tower The fuel tail gas produced by the top still contains a considerable amount of ethylene and ethane, resulting in the loss of the target product.

Contents of the invention

The technical problem to be solved by the present invention is: how to fully separate the product of direct conversion of methane to ethylene, so as to effectively utilize the product.

The technical solution adopted in the present invention is: a separation process for the direct conversion of methane to ethylene products, the product S0 of direct conversion of methane to ethylene is dehydrated, decarburized and dried to obtain methane, ethylene, ethane and C2+ hydrocarbons (C2+ is a hydrocarbon with a carbon number greater than 2, mainly propylene and propane) mixture S1, the mixture S1 is input to the input port at the lower end of the methane fractionation tower, and the solvent S2 is input to the upper end input port of the methane fractionation tower, and the mixture S1 and solvent S2 are mixed in methane Fully contact and conduct mass transfer and heat transfer in the fractionation tower, the methane gas containing only a trace of ethylene, that is, rich methane S3, is obtained at the output port of the upper end of the methane fractionation tower, and the mixture S4 of C2+ and solvent is obtained at the output port of the lower end of the methane fractionation tower; The C2+ and the solvent mixture S4 enter the input port of the solvent fractionation tower for fractionation, the C2-3 hydrocarbon mixture S5 is obtained at the output port at the upper end of the solvent fractionation tower, and the solvent-enriched liquid S6 containing a trace amount of C2+ is obtained at the output port at the lower end of the solvent fractionation tower. The rich solvent liquid S6 is input to the input port of the upper end of the methane fractionation tower for recycling as solvent S2; the C2-3 hydrocarbon mixture S5 enters the input port of the ethylene fractionation tower for fractionation, and the ethylene product S7 is obtained at the output port of the upper end of the ethylene fractionation tower, A mixture S8 of ethane and propylene is obtained from the outlet at the lower end of the fractionation tower.

As a preferred method, the mixture S8 of ethane and propylene enters the input port of the ethane fractionation tower, obtains ethane S9 at the output port at the upper end of the ethane fractionation tower, and obtains the propylene product S10 at the lower end output port of the ethane fractionation tower; The alkane S9 is input into the ethane dehydrogenation reactor of the reaction section, which is used to increase the production of ethylene, and the solvent-rich liquid S6 containing a trace amount of C2+ is obtained at the output port at the lower end of the solvent fractionation tower, and the solvent-rich liquid S6 is input to the upper end input port of the ethylene absorption tower, The methane gas containing only a small amount of ethylene, that is, rich methane S3, is obtained at the output port of the upper end of the methane fractionation tower and input to the lower end input port of the ethylene absorption tower. The outlet outputs methane gas S13 that hardly contains ethylene, and the outlet of the lower part of the ethylene absorption tower obtains the mixture S11 of methane, ethylene and solvent; the mixture S11 is decompressed and enters the input port of the separator for flash evaporation, and the methane-rich vapor phase is output from the upper output port of the separator S12 undergoes deflagration combustion (the amount of S12 is small, the pressure is also low, and the recovery cost is high, and engineering generally does this), and the solvent S2 is output from the output port at the lower end of the separator (a small amount of impurities are dissolved, but it does not affect the use as a solvent. ), the solvent S2 enters the input port at the upper end of the methane fractionation tower.

Operating temperature at the top of the methane fractionation tower: -40~-20°C

Operating temperature of methane fractionation tower still: 90-140°C

Operating pressure of methane fractionation tower: 20~30bar

Operating temperature at the top of the solvent fractionation tower: -40～-20°C

Solvent Fractionation Tower Tank Operation Temperature: 175～210℃

Solvent fractionation tower top operating pressure: 12 ~ 20bar

Operating temperature at the top of the ethylene fractionation tower: -30 to -10°C

Operation temperature of ethylene fractionation tower bottom: 0～20℃

Operating pressure at the top of the ethylene fractionation tower: 20~30bar

Operating temperature at the top of the ethane fractionation tower: -20~0°C

Operating temperature of ethane fractionation tower still: 20～60℃

Operating pressure at the top of the ethane fractionation tower: 15~25bar

Operating temperature at the top of the ethylene absorption tower: -30～-10°C

Operating temperature of ethylene absorption tower kettle: -30～-10℃

Operating pressure at the top of the ethylene absorption tower: 40~60bar

Separator operating temperature: -30～-10°C

Separator operating pressure: 20~30bar

The mass ratio of the solvent S2 entering the methane fractionation tower to the mixture S1 is 6:1-10:1, and the optimal ratio of the ethylene recovery rate and the operating cost of the solvent fractionation tower is 6.5:1-9.5:1.

The solvent S2 is a mixture of pentane and o-xylene, and the mass fraction can usually be 1:4 to 4:1, and the optimum value is 1:1 to 4:1.

Before S6 is introduced into the ethylene absorption tower, it should usually be cooled to: -40～-20°C

The mass ratio of methane/other products in the feed S1 to be treated ranges from 0.4:1 to 1.4:1, and the optimum range is 0.5:1 to 1:1.

The invention solves the product separation problem of methane conversion to ethylene. The recovery rate of ethylene in the process proposed by the invention is over 99%. At the same time, compared with the traditional process, the minimum operating temperature of the system is increased by nearly 100°C to -40°C (even -20℃), greatly reducing equipment investment and operating costs. Simultaneously, the technology that the present invention proposes overcomes the following deficiencies of the absorption (or extraction)-desorption technology of patent report in the past:

When the boiling point of the solvent is high, the absorption capacity is insufficient, and the steam used for solvent regeneration is of high grade, resulting in high operating costs;

When the boiling point of the solvent is low, it is easy to be taken out by methane during the absorption process, resulting in too much solvent loss; the mass ratio of methane/main product in cracked gas or methanol-to-olefin products (the sum of the mass flow rate of methane and ethylene, ethane and propylene) The ratio) is very low, usually far less than 1, and for the OCM reaction product, its ratio is usually higher, and the traditional process may not be suitable.

Description of drawings

Fig. 1 is a product separation process block diagram of the present invention;

Fig. 2 is the product separation process block diagram of alternative scheme 1 of the present invention;

Fig. 3 is the product separation process block diagram of alternative scheme 2 of the present invention;

Fig. 4 is a block diagram of the product separation process of the simplified scheme of the present invention.

Detailed ways

The present invention will be further described below in conjunction with the accompanying drawings.

Example 1

The product S0 of direct conversion of methane to ethylene is dehydrated, decarburized and dried to obtain a mixture S1 rich in methane, ethylene, ethane and C2+ hydrocarbons (C2+ is hydrocarbons with a carbon number greater than 2, mainly propylene and propane). As shown in accompanying drawing 1, mixture S1 enters the input port of the lower end of the methane fractionation tower, the solvent is imported into the upper end input port of the methane fractionation tower, and the bottom liquid stream S4 is output from the output port of the lower end of the methane fractionation tower, and the upper end output port of the methane fractionation tower obtains rich Methane vapor phase S3.

Conditions for this step:

Operating temperature at the top of the methane fractionation tower: -20°C

Operating temperature of methane fractionation tower bottom: 116°C

Operating pressure of methane fractionation tower: 25bar

From the output port at the lower end of the methane fractionation tower, the bottom liquid stream S4 enters the solvent fractionation tower 3, the C2+ mixture S5 is obtained at the output port at the upper end of the tower, and the relatively pure lean liquid S6 obtained at the output port at the lower end is cooled and circulated, and then sent to into the upper end input port of the methane fractionation tower.

Conditions for this step:

Solvent fractionation tower top operating temperature: -25°C

Solvent Fractionation Tower Tank Operating Temperature: 188°C

Solvent fractionation tower top operating pressure: 17bar

(3) The stream S5 obtained from the upper outlet of the solvent fractionation tower enters the ethylene fractionator, the ethylene product S7 is obtained at the upper outlet of the ethylene fractionator, and ethane and its C2+ mixture S8 are obtained at the lower outlet.

Conditions for this step:

Operating temperature at the top of the ethylene fractionation tower: -21°C

Operating temperature of the ethylene fractionation tower tank: 11°C

Operating pressure at the top of the ethylene fractionator: 25bar

(4) The stream S8 obtained at the output port at the lower end of the solvent fractionation tower enters the ethane fractionation tower, and at the output port at the upper end of the tower, ethane S9 is circulated and enters the ethylene reactor from ethane dehydrogenation in the synthesis section for increasing the ethylene output, The propylene product S10 is obtained at the output port at its lower end.

Conditions for this step:

Operating temperature at the top of the ethane fractionation tower: -7°C

Operating temperature of ethane fractionation tower tank: 48°C

Operating pressure at the top of the ethane fractionation tower: 20bar

(5) In the methane-rich vapor phase S3 obtained at the outlet of the upper end of the methane fractionation tower, there is still a certain amount of C2 hydrocarbons (especially ethylene), and the methane vapor phase S3 enters the lower end of the ethylene absorption tower and comes from the outlet at the bottom of the solvent fractionation tower The rich solvent liquid S6 enters the input port of the upper end of the ethylene absorption tower, and the gas-liquid two-phase of S3 and S6 fully contacts for mass transfer and heat transfer, and the methane gas S13 that hardly contains C2 hydrocarbons is obtained at the output port of the upper end of the ethylene absorption tower, and is absorbed in the ethylene absorption tower. The bottom production liquid S11 containing a small amount of C2 hydrocarbons, methane and a large amount of solvent is obtained at the lower end of the tower.

Conditions for this step:

Operating temperature at the top of the ethylene absorption tower: -30°C

Operating temperature of ethylene absorption tower kettle: -30°C

Operating pressure at the top of the ethylene absorption tower: 50bar

(6) The tower bottom production liquid S11 that contains a small amount of C2 hydrocarbons, methane and a large amount of solvent obtained by the output port at the lower end of the ethylene absorption tower enters the separator under reduced pressure, and obtains rich methane and a small amount of C2 hydrocarbon gas at the output port at the upper end of the separator. Phase S12, the vapor phase S12 is injected into the input port at the lower end of the ethylene absorption tower 6 or directly discharged and burned, and the solvent S2 (containing a small amount of methane, C2 hydrocarbons and a large amount of solvent) is obtained at the output port at the lower end of the separator, and the solvent S2 liquid is injected into the methane The input port at the upper end of the fractionation column.

Conditions for this step:

Separator 7 operating temperature: -30°C

Separator 7 operating pressure: 25bar

The mass flow ratio of the circulating solvent S2 to the feed S1 is 6.3:1-6.4:1.

The recycled solvent S2 used is mainly a mixture of n-pentane and o-xylene, which is composed of 50/50 n-pentane and o-xylene.

Before S6 enters the ethylene absorption tower, it should usually be cooled to a temperature of -30°C

Using the S1 stream shown in Table 1 as the feed, using the separation process disclosed in this patent, the product results obtained through simulation calculation are listed in Table 1. It can be seen that the S1 stream has a relatively high methane/other product ratio (the ratio of the mass flow of methane to the sum of the mass flow of ethylene, ethane and propylene reaches 0.683:1), and this stream enters the separation process and is placed at the upper end of the methane fractionation tower. The output port obtains stream S3, which contains all methane, hydrogen and carbon monoxide in the feed, and contains almost no ethylene, and obtains stream S4 at the outlet of the lower end of the methane fractionation column, which contains almost all ethylene in the feed, and has achieved a high Ethylene recovery rate; the stream S4 is further fed into the solvent fractionation tower, and almost all of the ethylene, ethane and propylene mixture S5 is obtained at the output port at the upper end, and the lean liquid S6 containing relatively pure solvent is obtained at the output port at the lower end, which is recycled into the ethylene absorption tower top.

Example 2

The difference from Comparative Document 1 is that:

Operating temperature at the top of the methane fractionation tower: -40°C

Operating temperature of methane fractionation tower bottom: 112°C

Operating pressure of methane fractionation tower: 25bar

Solvent fractionation tower top operating temperature: -26°C

Solvent Fractionation Tower Tank Operating Temperature: 187°C

Solvent fractionation tower top operating pressure: 17bar

Operating temperature at the top of the ethylene fractionation tower: -21°C

Operating temperature of the ethylene fractionation tower tank: 11°C

Operating pressure at the top of the ethylene fractionator: 25bar

Operating temperature at the top of the ethane fractionation tower: -7°C

Operating temperature of ethane fractionation tower tank: 48°C

Operating pressure at the top of the ethane fractionation tower: 20bar

Operating temperature at the top of the ethylene absorption tower: -15°C

Operating temperature of ethylene absorption tower kettle: -16°C

Operating pressure at the top of the ethylene absorption tower: 50bar

Separator 7 operating temperature: -10°C

Separator 7 operating pressure: 20bar

The mass flow ratio of the circulating solvent S2 to the feed S1 is 6.1:1-6.2:1.

The recycled solvent S3 used is mainly a mixture of n-pentane and o-xylene, which is composed of 50/50 n-pentane and o-xylene.

The S6 should usually be cooled to a temperature of -20°C before entering the ethylene absorption tower

Using the S1 stream shown in Table 2 as the feed, using the separation process disclosed in this patent, the product results obtained through simulation calculations are listed in Table 2. It can be seen that the S1 stream has a relatively high methane/other product ratio (the ratio of the mass flow rate of methane to the sum of the mass flow rates of ethylene, ethane and propylene reaches 0.721:1), and this stream enters the separation process and is placed at the upper end of the methane fractionation tower. The output port obtains stream S3, which contains all methane, hydrogen and carbon monoxide in the feed, and contains almost no ethylene, and obtains stream S4 at the output port at the lower end of the methane fractionation column, which contains almost all ethylene in the feed, and has achieved a high Ethylene recovery rate; stream S4 further enters the solvent fractionation tower, and obtains almost all ethylene, ethane and propylene mixture S5 at the output port at the upper end, and obtains lean liquid S6 with relatively pure solvent at the output port at the lower end, and circulates into the upper end of the ethylene absorption tower input port.

Example 3

The difference from Example 1 is:

Operating temperature at the top of the methane fractionation tower: -40°C

Operating temperature of methane fractionation tower bottom: 123°C

Operating pressure of methane fractionation tower: 30bar

Solvent fractionation tower top operating temperature: -38°C

Solvent Fractionation Tower Tank Operating Temperature: 206°C

Solvent fractionation tower top operating pressure: 12bar

Operating temperature at the top of the ethylene fractionation tower: -13°C

Operating temperature of ethylene fractionation tower bottom: 20°C

Operating pressure at the top of the ethylene fractionator: 30bar

Operating temperature at the top of the ethane fractionation tower: -18°C

Operating temperature of ethane fractionation column tank: 36°C

Operating pressure at the top of the ethane fractionation tower: 15bar

Operating temperature at the top of the ethylene absorption tower: -14°C

Operating temperature of ethylene absorption tower kettle: -17°C

Operating pressure at the top of the ethylene absorption tower: 50bar

Separator 7 operating temperature: -17°C

Separator 7 operating pressure: 30bar

The mass flow ratio of the circulating solvent S2 to the feed S1 is 6.1:1-6.2:1.

The recycled solvent S2 used is mainly a mixture of n-pentane and o-xylene, and is composed of 20/80 wt% n-pentane and o-xylene.

The S6 should usually be cooled to a temperature of -25°C before entering the ethylene absorption tower

Using the S1 stream shown in Table 3 as the feed, using the separation process disclosed in this patent, the product results obtained through simulation calculations are listed in Table 3. It can be seen that the S1 stream has a relatively high ratio of methane/main product (the ratio of methane mass flow to the sum of ethylene, ethane and propylene mass flow is up to 0.721:1), and this stream enters the separation process at the upper end of the methane fractionation tower The outlet obtains stream S3, which contains all methane, hydrogen and carbon monoxide in the feed, and hardly contains ethylene, and obtains stream S4 at the outlet of the lower end of the methane fractionation column 2, which contains almost all ethylene in the feed, and has achieved a high ethylene recovery rate; the stream S4 is further fed to the solvent fractionation tower, and almost all of the ethylene, ethane and propylene mixture S5 is obtained at the output port at the upper end, and the lean liquid S6 with relatively pure solvent is obtained at the output port at the lower end, which is recycled into the ethylene absorption tower top.

Example 4

The difference from Example 1 is:

Operating temperature at the top of the methane fractionation tower: -40°C

Operating temperature of methane fractionation tower bottom: 138°C

Operating pressure of methane fractionation tower: 30bar

Solvent fractionation tower top operating temperature: -38°C

Solvent Fractionation Tower Tank Operating Temperature: 210°C

Solvent fractionation tower top operating pressure: 12bar

Operating temperature at the top of the ethylene fractionation tower: -13°C

Operating temperature of ethylene fractionation tower bottom: 20°C

Operating pressure at the top of the ethylene fractionator: 30bar

Operating temperature at the top of the ethane fractionation tower: -18°C

Operating temperature of ethane fractionation tower tank: 35°C

Operating pressure at the top of the ethane fractionation tower: 15bar

Operating temperature at the top of the ethylene absorption tower: -18°C

Operating temperature of ethylene absorption tower kettle: -17°C

Ethylene absorption tower top operating pressure: 50bar

Separator operating temperature: -17°C

Separator operating pressure: 30bar

The mass flow ratio of the circulating solvent S2 to the feed S1 is 6.1:1-6.2:1.

The solvent S2 used is mainly a mixture of n-pentane and o-xylene, with a composition of 20/80 wt% n-pentane/o-xylene.

The S6 should usually be cooled to a temperature of -20°C before entering the methane fractionation tower

Using the S1 stream shown in Table 4 as the feed, using the separation process disclosed in this patent, the product results obtained through simulation calculations are listed in Table 4. It can be seen that the S1 stream has a relatively high ratio of methane/main product (the ratio of methane mass flow to the sum of ethylene, ethane and propylene mass flow is as high as 1:1), and this stream enters the separation process, and at the upper end of the methane fractionation tower The output port obtains stream S3, which contains all methane, hydrogen and carbon monoxide in the feed, and contains almost no ethylene, and obtains stream S4 at the outlet of the lower end of the methane fractionation column, which contains almost all ethylene in the feed, and has achieved a high Ethylene recovery rate; the stream S4 is further fed into the solvent fractionation tower, and almost all the ethylene, ethane and propylene mixture S5 is obtained at the output port at the upper end, and the lean liquid S6 with relatively pure solvent is obtained at the output port at the lower end, which is recycled into the ethylene absorption tower upper end.

Example 5

The difference from Example 1 is:

Operating temperature at the top of the methane fractionation tower: -40°C

Operating temperature of methane fractionation tower kettle: 92°C

Operating pressure of methane fractionation tower: 30bar

Solvent fractionation tower top operating temperature: -20°C

Solvent Fractionation Tower Tank Operating Temperature: 175°C

Solvent fractionation tower top operating pressure: 20bar

Operating temperature at the top of the ethylene fractionation tower: -29°C

Operating temperature of ethylene fractionation tower bottom: 1°C

Operating pressure at the top of the ethylene fractionator: 20bar

Operating temperature at the top of the ethane fractionation tower: 2°C

Operating temperature of the ethane fractionation tower tank: 59°C

Operating pressure at the top of the ethane fractionation tower: 25bar

Operating temperature at the top of the ethylene absorption tower: -18°C

Operating temperature of ethylene absorption tower kettle: -16°C

Operating pressure at the top of the ethylene absorption tower: 40bar

Separator operating temperature: -16°C

Separator operating pressure: 30bar

The mass flow ratio of the circulating solvent S2 to the feed S1 is 7.95:1-8.052:1.

The solvent S2 used is mainly a mixture of n-pentane and o-xylene, with a composition of 80/20 wt% n-pentane/o-xylene.

The S6 should usually be cooled to a temperature of -40°C before entering the methane fractionation tower

Using the S1 stream shown in Table 5 as the feed, using the separation process disclosed in this patent, the product results obtained through simulation calculations are listed in Table 5. It can be seen that the ratio of methane/main product (the ratio of methane mass flow to the sum of ethylene, ethane and propylene mass flow reaches 0.5:1), this stream enters this separation process, and stream S3 is obtained at the output port at the upper end of the methane fractionation tower, It contains all methane, hydrogen and carbon monoxide in the feed, and contains almost no ethylene, and obtains stream S4 at the outlet of the lower end of the methane fractionation tower, which contains almost all ethylene in the feed, and achieves a high ethylene recovery rate; stream S4 Further feed the solvent fractionation tower 3, obtain almost all ethylene, ethane and propylene mixture S5 at the output port at its upper end, obtain lean liquid S6 with relatively pure solvent at the output port at its lower end, and circulate into the upper end of the ethylene absorption tower.

The invention solves the product separation problem of methane conversion to ethylene. The recovery rate of ethylene in the process proposed by the invention is over 99%. At the same time, compared with the traditional process, the minimum operating temperature of the system is increased by nearly 100°C to -40°C (even -20℃), greatly reducing equipment investment costs and operating costs.

Claims (10)

1. A separation process for methane direct conversion to ethylene product, characterized in that: the product S0 of methane direct conversion to ethylene is dehydrated, decarburized and dried to obtain rich in methane, ethylene, ethane and C2+ hydrocarbons (C2+ series Hydrocarbons with a carbon number greater than 2, mainly propylene and propane) mixture S1, the mixture S1 is input to the input port at the lower end of the methane fractionation tower, the solvent S2 is input to the upper end input port of the methane fractionation tower, and the mixture S1 and solvent S2 are in the methane fractionation tower Fully contact and conduct mass transfer and heat transfer, obtain methane gas containing only a trace of ethylene at the output port of the methane fractionation tower, that is, methane-enriched S3, and obtain a mixture of C2+ and solvent S4 at the output port of the lower end of the methane fractionation tower; the C2+ and The solvent mixture S4 enters the input port of the solvent fractionation tower for fractionation, the C2-3 hydrocarbon mixture S5 is obtained at the output port at the upper end of the solvent fractionation tower, and the rich solvent liquid S6 containing a trace amount of C2+ is obtained at the output port at the lower end of the solvent fractionation tower. S6 is input to the input port of the upper end of the methane fractionation tower as the solvent S2 for recycling; the C2-3 hydrocarbon mixture S5 enters the input port of the ethylene fractionation tower for fractionation, and the ethylene product S7 is obtained at the output port of the upper end of the ethylene fractionation tower, and the S7 is obtained at the lower end of the ethylene fractionation tower The output port obtains the mixture S8 of ethane and propylene. 2. the separation process of a kind of methane direct conversion system ethylene product according to claim 1, is characterized in that: the mixture S8 of ethane, propylene enters the input port of ethane fractionation tower, obtains ethane at the output port of the upper end of ethane fractionation tower Alkane S9, the propylene product S10 is obtained at the outlet at the lower end of the ethane fractionation tower; the ethane S9 is input into the ethane dehydrogenation reactor of the reaction section for increasing the production of ethylene, and at the outlet at the lower end of the solvent fractionation tower, propylene product S10 is obtained at the outlet at the lower end of the solvent fractionator. The rich solvent liquid S6, the rich solvent liquid S6 is input to the input port of the upper end of the ethylene absorption tower, and the methane gas containing only a trace of ethylene is obtained at the output port of the upper end of the methane fractionation tower, that is, the rich methane S3 is input to the lower end input port of the ethylene absorption tower, and the rich solvent liquid S6 and methane-enriched S3 are in full contact for mass transfer and heat transfer. The output port at the upper end of the ethylene absorption tower outputs methane gas S13 that hardly contains ethylene, and the output port at the lower part of the ethylene absorption tower obtains a mixture of methane, ethylene and solvent S11; the mixture S11 is decompressed It enters the input port of the separator for flashing, and outputs the methane-rich vapor phase S12 from the output port of the upper end of the separator for deflagration combustion (the amount of S12 is small, the pressure is also low, and the recovery cost is high, and engineering generally does this). The output port at the lower end outputs the solvent S2 (a small amount of impurities are dissolved, but it does not affect its use as a solvent), and the solvent S2 enters the input port at the upper end of the methane fractionation tower. 3. the separation process of a kind of methane direct conversion system ethylene product according to claim 2, is characterized in that: Operating temperature of methane fractionation tower bottom: 90-140 ℃ Operating pressure of methane fractionation tower: 20~30 bar Solvent fractionation tower top operating temperature: -40~-20 ℃ Operating temperature of solvent fractionation tower kettle: 175~210 ℃ Solvent fractionation tower top operating pressure: 12~20 bar Operating temperature at the top of the ethylene fractionation tower: -30~-10 ℃ Operating temperature of ethylene fractionation tower bottom: 0~20 ℃ Operating pressure at the top of the ethylene fractionator: 20~30 bar Operating temperature at the top of the ethane fractionation tower: -20~0 ℃ Operating temperature of ethane fractionation tower bottom: 20~60 ℃ Operating pressure at the top of the ethane fractionation tower: 15~25 bar Operating temperature at the top of the ethylene absorption tower: -30~-10 ℃ Operating temperature of ethylene absorption tower tank: -30~-10 ℃ Operating pressure at the top of the ethylene absorption tower: 40~60 bar Separator operating temperature: -30~-10 ℃ Separator operating pressure: 20~30 bar. 4. A separation process for direct conversion of methane to ethylene products according to claim 2, characterized in that: the mass ratio of the solvent S2 entering the methane fractionation tower to the mixture S1 is 6:1~10:1. 5. A separation process for direct conversion of methane to ethylene products according to claim 4, characterized in that: the mass ratio of the solvent S2 entering the methane fractionation tower to the mixture S1 is 6.5:1~9.5:1. 6. A separation process for direct conversion of methane to ethylene products according to claim 2, characterized in that: the solvent S2 is a mixture of pentane and o-xylene, with a mass ratio of 1:4~4:1. 7. A separation process for direct conversion of methane to ethylene products according to claim 6, characterized in that: the solvent S2 is a mixture of pentane and o-xylene, with a mass ratio of 1:1 to 4:1. 8. A separation process for direct conversion of methane to ethylene products according to claim 2, characterized in that: S6 should usually be cooled to -40~-20°C before being introduced into the ethylene absorption tower. 9. A separation process for direct conversion of methane to ethylene products according to claim 2, characterized in that the mass ratio of methane to other products in the mixture S1 ranges from 0.4:1 to 1.4:1. 10. A separation process for direct conversion of methane to ethylene products according to claim 2, characterized in that the mass ratio of methane to other products in the mixture S1 ranges from 0.5:1 to 1:1.