A method, reactor and process system for preparing methanol into olefins in a fluidized bed

Abstract

The present invention relates to a method, a reactor and a process system for producing olefins from methanol in a fluidized bed, belonging to the field of chemical engineering technology. After being preheated, the methanol raw material enters a fluidized bed coupling reactor and contacts and reacts with microsphere dual catalysts respectively containing molecular sieves with CHA and MFI structures. The catalyst deactivated by carbon deposition is continuously circulated to the regenerator and completely regenerated by passing air; the reaction products are separated in the subsequent separation process to obtain ethylene and propylene, and hydrogen and fuel gas are by-produced; ethane, propane, C4 and above components are returned to the coupling reactor for millisecond catalytic cracking and heat coupling, and the catalyst is selectively pre-carbonized; or C4 is separated, butane and other by-products are returned to the coupling reactor, and butene and part of ethylene enter a flexible switching fixed bed reactor and contact with WO3 / SiO2 catalyst to generate propylene. The present invention adopts complete regeneration of the dual catalyst, flexible utilization of by-products, selective pre-carbonization and heat coupling, improving the conversion performance and olefin yield, etc.

Description

Technical Field

[0001] The present invention relates to a technology for preparing low-carbon olefins from methanol, and in particular to a method, a reactor and a process system for preparing olefins from methanol in a fluidized bed, and belongs to the technical field of chemical industry. Technical Background

[0002] Ethylene and propylene are the basic chemical products with the largest production and consumption in the world. They are also basic organic raw materials with a wide range of uses. The annual output of olefins can measure a country's level of chemical development.

[0003] The production of ethylene and other products mainly comes from petroleum processing, such as steam cracking of petroleum fractions. However, the process of converting methanol produced from coal into ethylene and propylene catalytically is also a very interesting technical route. Its products have been proven to be fully suitable for the production of polyolefins and other products, and have important strategic significance and application prospects. After the industrialization of methanol to olefins (MTO) technology, it has quickly formed an industrial scale. At present, methanol to olefins accounts for a significant proportion of my country's ethylene and propylene production capacity, and further technological research and development and innovative inventions have also received high attention.

[0004] In 1984, Mobil Corporation of the United States reported that it had improved the methanol-to-gasoline process and used it to prepare low-carbon olefins (MTO) from methanol (Chang, CD et al., J. Catal., SB, 289 (1984)). Since then, many related research institutions in the United States, Germany, Japan and China have carried out research and development of methanol-to-olefins, such as the early research and development work of ExxonMobil Corporation in the United States reported in U.S. Patents USP3931349 and USP3998899.

[0005] Many aluminosilicates or molecular sieves have been tried as catalysts for the methanol to olefins reaction. For example, large-pore molecular sieves such as Y zeolite and mordenite can catalyze the conversion of methanol into hydrocarbons, but C5 + The selectivity of hydrocarbons is high, while the selectivity of low-carbon olefins is low; while small-pore molecular sieves such as erionite, T zeolite, and chabazite can improve the selectivity of olefins, they are deactivated very quickly due to coking. Therefore, more research is focused on molecular sieves such as SAPO-34 and ZSM-5, such as the contents described in the prior art such as EPA541915 and DE4009459.

[0006] In recent years, methanol to low-carbon olefin technology has received more and more attention. According to the different product distribution, methanol to low-carbon olefin technology is mainly divided into MTO technology and MTP technology. MTO technology generally uses a fluidized bed reactor and SAPO-34 molecular sieve catalyst to produce low-carbon olefins mainly composed of ethylene and propylene in one step, and its product P / E ratio is very low; the technical characteristics are described in USP5817906 and USP6710218.

[0007] The MTP process for converting methanol into propylene mainly uses ZSM-5 molecular sieve catalyst and adiabatic fixed bed reactor, and its notable features are high propylene selectivity and P / E ratio, as disclosed in W02006136433, EP448000, WO192190, and DE1020050159232.

[0008] In the USP4499327 patent, a detailed study was conducted on the application of silicoaluminophosphate molecular sieve catalysts in the methanol to olefins process. It is believed that SAPO-34 molecular sieve is the preferred catalyst for the methanol to olefins (MTO) process. It has high selectivity for low-carbon olefins and high activity. It can make the reaction time of methanol to low-carbon olefins reach less than 10 seconds, and even reach the reaction time range of the riser. Therefore, the research on reactors and process devices has received great attention and has been reported a lot.

[0009] For example, United Oil Company of America disclosed in USP6166282 a methanol to olefins unit that combines a dense bed with a fast fluidized bed reactor. The product gas and the catalyst are quickly separated to prevent the occurrence of secondary reactions. The catalyst inventory is greatly reduced compared to a traditional bubbling fluidized bed reactor, but the yield of low-carbon olefins is low.

[0010] CN1723262 discloses a multi-stage riser reaction device with a central catalyst loop for the process of converting oxides into light olefins, but it also has the problem of low light olefin yield; CN1166478 discloses a dense phase bed circulating fluidized reaction device for preparing olefins from methanol, and the cracking reaction is carried out in an upward dense phase bed circulating fluidized reaction device. The device is easy to remove the reaction heat and make the process easy to control.

[0011] There are many reports on the invention contents that focus on improving the reactor structure in the prior art. For example, CN205024117 discloses a methanol to olefin equipment with a dense phase fluidized bed reactor with a specific structure transverse grid at the bottom. By breaking bubbles to enhance gas-solid mixing and shortening the upper free space of the reaction equipment, the gas reaction residence time can be shortened to 1 to 2 seconds, thereby improving the product selectivity and olefin yield; CN104437274 also discloses a dense phase fluidized bed reactor technology for methanol to olefin, using a specific gas pre-lift pipe to extend into the bed of the dense phase fluidized bed reactor, through radial ventilation to the bed, combined with the special baffle design in the reactor, greatly improve the selectivity of low-carbon olefins in the product of the single-pass reaction process, and the residence time is controlled within 0.5 to 1.5 seconds. Different from the reactor technology that generally adopts the upward movement of materials and the contact of catalysts, CN1156416C uses a gas-solid parallel downward fluidized bed ultra-short contact reactor to suppress the occurrence of side reactions by timely stopping the reaction.

[0012] In order to cope with complex reaction-regeneration conditions, the technology of multiple reaction zones and regeneration zones has also been widely reported. In CN115106028, a tower plate-type component is used to separate the reactor into several reaction zones or purge zones along the reactor axis, so as to realize the hierarchical utilization of catalysts, and the raw material processing capacity is distributed by setting the volume of the separated zones, thereby improving the catalyst utilization rate and product selectivity; CN102464535 discloses a technology for preparing olefins from methanol by providing a reactor with baffles to form three reaction zones; CN103664439 also adopts three fluidized bed reaction zones; and CN103739427 also combines the down-going bed reaction zone with the fast fluidized bed reaction zone and two riser reaction zones to solve the problem of low yield of low-carbon olefins.

[0013] In addition to the use of multi-reaction zone technology, multi-reactor and regenerator technology is also used. CN104098428 uses a two-stage fluidized bed reactor to improve the propylene yield; CN107043317 uses a technology of two reactors sharing one regenerator; CN104672040 uses multiple reactors plus multiple regenerators; CN 203725133U uses multiple regenerators and multiple regeneration areas to deal with different degrees of carbon deposition and deactivated catalysts; CN102276390 uses multiple reactors and reaction areas to adjust the product ratio; CN104672045 uses a dense phase fluidized bed and a riser reactor, as well as multiple regeneration areas; CN104667836 reduces the construction cost by lowering the total elevation of the device through the equal-height layout of two reactors and one regenerator; CN105130727 uses a three-stage series fluidized bed reactor to improve the production capacity of a unit catalyst; CN110117214 uses a three-riser reactor and a fully regenerated catalyst and pre-carbon deposition treatment to convert C4 + hydrocarbons and improve selectivity and yield.

[0014] In summary, there are many technical reports involving multiple reactors in the prior art, which, while meeting the purpose of solving the reaction and improving the olefin yield, also increase the complexity of the device and operation and the investment cost.

[0015] There are also many reports on the technology of dividing the catalyst into two or more strands. For example, CN101357874 divides the catalyst into different flow directions for treatment to solve the problems of low yield of light olefins, difficult to control reactor temperature, uneven mixing of catalysts in the reaction zone, etc.; CN102464524 divides the catalyst into a riser and a fluidized bed reaction zone respectively after diversion; CN102274752 divides the regenerated catalyst into two parts for treatment to solve the problem of high temperature of the regenerated catalyst entering the reactor; CN102295505 divides the catalyst to be regenerated into two strands for separate treatment.

[0016] In order to further improve the yield of low-carbon olefins, how to make full use of the by-products in the reaction process has received great attention. For example, in CN101177374, methanol or dimethyl ether conversion reaction, ethylene and methanol alkylation reaction, and C4 and above heavy components are subjected to catalytic cracking reaction to increase the propylene yield; CN112299939 also adds a cracking reactor for converting C4 + Hydrocarbons increase the overall ethylene and propylene yields; CN205024117U discloses a process unit comprising a methanol to olefins reactor, a regenerator and a cracking reactor.

[0017] In order to make more reasonable use of the heat in the reaction, the prior art reports the technology of coupling two processes / systems or two raw materials. For example, CN113845401 and CN2165136096U couple the ethane oxidative dehydrogenation (ODHE) process system with the methanol to olefins (MTO) process system; CN103804110 couples the organic oxygen-containing compound to produce low-carbon olefins with the C5 + The coupling of hydrocarbon catalytic cracking effectively utilizes the high temperature of the regenerated catalyst, shortens the reaction induction period and increases the olefin yield; CN113800993 and CN109651038 couple the dehydrogenation of low-carbon alkanes with methanol to olefins, using the regeneration agent to provide heat for the endothermic reaction of the dehydrogenation of low-carbon alkanes. The cooled catalyst is more suitable for the MTO reaction temperature, thereby increasing the olefin yield and space utilization.

[0018] CN102875285 discloses a method for converting methanol into low-carbon olefins by combining methanol conversion and catalytic cracking of light naphtha; CN115124399 couples the methanol production of olefins with the high-temperature cracking reaction of olefins to achieve the production of all C4 +High temperature cracking of olefins improves the cracking conversion rate and the yield of low-carbon olefins; CN1102190542 and CN102190538 use a plurality of reactors including a riser reactor and a downer reactor for reaction coupling in the coupling method of olefin production from methanol and catalytic cracking of hydrocarbons above C4; CN112876333 also includes a semi-regenerated catalyst for two-stream treatment in the coupling olefin cracking of methanol to olefins; CN101402538 also includes the reaction of olefin production from methanol and the recycling of hydrocarbons above C4; by returning C4 and above components to the main reactor for cracking reaction after separation, although the olefin yield is improved, there is often a problem of limited C4 processing capacity, especially low conversion rate, involving C4 + Component cracking is more often done by separately setting up an auxiliary reactor / regeneration cracking device, such as CN103539597, CN103539607, etc.

[0019] CN111423302 adds a pre-reactor for converting C4 and / or C5 obtained from the separated product + Hydrocarbons, whose regenerated catalyst is completely regenerated and pre-carbonized in the pre-reactor; CN102190548 also includes catalyst pre-carbonization and partial carbonization, and the methanol feedstock and the catalyst are contacted and reacted in a fast fluidized bed. Due to the competition for the reaction raw materials, the improvement of olefin selectivity is limited, and the fast fluidization also aggravates the catalyst wear and the increase in catalyst consumption; CN101333141 also uses a fast bed reaction device, which also brings the problem of increased catalyst consumption.

[0020] CN100586910C, CN102190539, CN102190539 and other prior arts also disclose a combined process including methanol to olefins, catalytic cracking reaction of heavy components above C4, and anti-disproportionation reaction of ethylene and butene to obtain propylene with high selectivity; CN101177374 discloses a method for preparing propylene with high selectivity including methanol to olefins, catalytic cracking of heavy components above C4, and alkylation reaction of ethylene and methanol. However, there are few reports on the technology of preparing MTO including ethylene with high selectivity through combined process technology.

[0021] In addition, in the method for producing olefins from methanol in CN103664450, the methanol feedstock enters a fluidized bed reactor and contacts with a catalyst including a SAPO-34 molecular sieve and a ZSM-5 molecular sieve fine powder auxiliary having an average particle size of 5 to 35 microns to improve the conversion rate of butene-2, which involves the application of two molecular sieve active components in the MTO reaction and olefin cracking reaction process, but there are relatively few reports on this type of technology, and it is mainly used to increase propylene production.

[0022] In summary, in order to further improve the yield of low-carbon olefins in the process of methanol to olefins, especially the yield of ethylene, improve the catalyst conversion performance under process conditions, regulate the reaction induction period and catalyst pre-carbon deposition, how to more efficiently convert and utilize by-products, how to improve cracking activity, heat extraction efficiency and thermal coupling in the reaction, and face the increasingly complex processes and equipment, there is still a need for continuous improvement of methanol to olefins (MTO) technology, including catalysts and molecular sieves, conversion methods, reactors and process systems. Summary of the invention

[0023] At present, there are two main technical routes for methanol to olefins (MTO), namely MTO route and MTP route, and the catalysts used are mainly SAPO-34 molecular sieve and ZSM-5 molecular sieve. ZSM-5 molecular sieve is a widely used molecular sieve, belonging to MFI type molecular sieve, with three-dimensional through ten-membered ring folded channels (0.51nm×0.55nm) and straight channels (0.53nm×0.56nm), which are very conducive to the diffusion of raw material and product hydrocarbon molecules, but the distribution of reaction products tends to produce more propylene.

[0024] SAPO-34 molecular sieve belongs to the CHA type molecular sieve, which has an ellipsoidal cage and an octagonal ring pore structure. The cage size is 1.1nm×0.65nm, while the octagonal ring pore diameter is only 0.38nm. The unique three-dimensional pore octagonal ring structure, suitable acidic center and good hydrothermal stability make it very suitable for MTO reaction.

[0025] SSZ-13 molecular sieve also belongs to CHA type molecular sieve, with an ellipsoidal cage (0.73nm×1.2nm) of eight-membered ring structure and a three-dimensional cross-channel structure. The channel size is 0.38nm×0.38nm, and its specific surface area can reach 700m 2 / g, good hydrothermal stability, a large number of surface proton acidic centers and exchangeable cations are its characteristics, which make it have some differences in catalytic performance compared with SAPO-34 molecular sieve.

[0026] The two catalysts composed of active components with different structures have different reaction mechanisms and product distributions due to the different pore sizes and morphologies. The products of SAPO-34 and SSZ-13 are mixed olefins of ethylene, propylene and butene, with a high selectivity for ethylene. The reaction products of ZSM-5 are mixed olefins mainly containing high contents of propylene, butene and pentene, especially with a high yield of propylene.

[0027] In addition, the catalysts composed of these two molecular sieves with different structural types have very different coking and deactivation rates. The ZSM-5 molecular sieve type catalyst cokes relatively slowly, while the SAPO-34 and SSZ-13 molecular sieve type catalysts coke and deactivate quickly. The coking loss of methanol is large, which increases the energy consumption of catalyst regeneration. However, the latter two molecular sieves also have differences in acidic characteristics, which also brings differences in adsorption, catalytic performance and coking. This not only provides room for selection when applied in the process, and the convenience of using the different properties of molecular sieves for complementary use, but also brings problems of coordination and coordination of process conditions and reactors and regenerators.

[0028] As far as the prior art has known, there is no particularly satisfactory method for converting methanol into individual olefins, such as ethylene or propylene (C Hang, Methanol Conversion to Light Olefins, Catal. Rev.-Sci. Eng., 1984, 26 (3&4), 323-345). In order to improve the selectivity and yield of light olefins in the MTO process, other modifying elements have been introduced into the active components of the catalyst in the prior art (M. Stocker, Methanol-to-hydrocarbons: catalytic materials and their behavior, Microp. Mesop. Mater., 1999, 29, 3-48); for example, JM Thomas et al. used nickel to modify SAPO-34 molecular sieve and obtained about 95% ethylene selectivity at 250° C. (Chem. Mater., 1991, 3 (4), 669-672). However, metal modification on silicon-phosphorus-aluminum molecular sieves with low acid center density is also likely to have an adverse effect on the acid center density and catalytic activity, which requires further research and improvement.

[0029] Since SAPO-34 molecular sieve is the key to the realization of MTO technology, the methanol to olefins reaction catalyzed by SAPO-34 has an obvious induction period, that is, in the initial stage of the reaction, the conversion rate and olefin selectivity increase with the increase of catalyst carbon deposition, but the catalyst will quickly deactivate after carbon deposition. The MTO reaction must adopt a circulating fluidized bed process to regenerate the deactivated catalyst by burning carbon, and restore the catalyst's methanol conversion activity through complete regeneration or partial regeneration. In the reaction-regeneration cycle, it is inevitable to involve whether the catalyst is completely regenerated, which has an important impact on the reaction conversion process. The regeneration degree of the regenerated catalyst, that is, whether the deactivated catalyst is completely regenerated or incompletely regenerated with a small amount of carbon deposition, will affect whether the catalyst is effectively utilized throughout the process. In practice, it is found that since the product distribution is actually controlled by the catalyst shape selectivity, although carbon deposition on the catalyst will lead to deactivation, it is beneficial to improve the selectivity of small molecular olefins ethylene and propylene, but too high carbon deposition will reduce the selectivity of low-carbon olefins, block the pores, and reduce activity.

[0030] Therefore, controlling the amount of carbon deposition on catalysts in industrial reactors or performing controlled pre-carbon deposition treatment, especially pre-carbon deposition on regenerated catalysts after complete regeneration, is a very important reaction technology that affects the conversion results and product yield and distribution in the process. This requires exploring methods, devices and systems that meet the requirements of the reaction conditions when preparing molecular sieves and catalysts, improving process methods and reaction equipment.

[0031] Since it is generally believed in the art that when the amount of carbon deposits of the catalyst in the MTO process reaches a certain control level, it is able to effectively improve the yield of light olefins and improve the selectivity of product distribution, it is generally believed in the prior art that the uniform mixing of carbon deposits inside the reaction zone is very important. In addition to the adjustment of process conditions such as reaction temperature, by adjusting and controlling the reaction time when carbon deposits are generated, modifying the acid strength and acid density distribution of the acidic active centers that generate carbon deposits, and the type and ratio of molecular sieves and carriers, it is a relatively direct and effective method to influence, change and control the amount and distribution of carbon deposits, and improve the yield of light olefins and product distribution characteristics. However, in the specific practice process, it is inevitable to encounter the problem that the amount and distribution of carbon deposits are difficult to adjust and control, and it is easy to lead to defects such as poor results, poor product distribution and low yield of light olefins.

[0032] In the reaction pathway of methanol conversion to produce low-carbon olefins, it can basically be regarded as consisting of the following two key steps, namely, methanol dehydration to produce olefins, and olefin conversion to aromatics + alkanes. These two reaction steps constitute a series reaction of methanol to produce low-carbon olefins, and the target product olefin is essentially an intermediate product. In order to reduce the production of by-product alkanes and increase the yield of product olefins, it is necessary to quickly separate the catalyst from the gas product and stop the reaction in time to reduce the occurrence of secondary reactions.

[0033] The second thing that needs attention is that the reaction of methanol conversion to produce hydrocarbon mixture is a highly exothermic reaction. For example, at 400°C, the reaction heat released for every kg of methanol converted is about 1.74MJ. If no heat removal measures are taken for such a large exothermic heat, it is easy to cause excessive local adiabatic temperature rise in the reaction area, which can reach 250°C. Therefore, certain measures must be taken to utilize and control the exothermic heat and temperature rise of the reaction process. In actual operation, due to the problem of reaction heat, a large amount of water is often added to the raw materials as a diluent or heat carrier, but a large amount of water vaporizes and condenses during the process, which will increase the energy consumption of the process and reduce the heat utilization efficiency and the space-time yield of the target product. The use of thermal coupling is undoubtedly an efficient measure.

[0034] The third issue that needs attention is that when methanol is converted into ethylene and propylene through the MTO reaction process, a large amount of mixed C4 and C5 will inevitably be produced as by-products. + The carbon-based yield of the components accounts for about 13% to 15% in total. The usual utilization method is to directly sell them as by-products or separate and purify them for utilization. The added value of these by-products is low, and their reuse will be directly related to the economy of the entire methanol to olefins process.

[0035] The aforementioned SAPO-34, SSZ-13 and other molecular sieve catalysts after carbon deposition are beneficial to the selectivity of methanol to olefins reaction. The SAPO-34 and SSZ-13 molecular sieve regeneration agents in a high-activity state can also carry out some hydrocarbon cracking reactions. On the other hand, the catalyst entering the reactor at a high temperature under fully regenerated conditions will also cause side reactions, reducing the yield of the target product. This is a contradictory problem that needs to be solved. + The cracking reaction of hydrocarbons in contact can increase the production of ethylene and propylene, and can also remove the excess heat they carry and inhibit the side reactions in the dilute phase section of the MTO reactor, thus achieving the effect of killing two birds with one stone.

[0036] Therefore, the high temperature regenerator from the regenerator is + The hydrocarbons are pre-carbonized in the cracking reactor before entering the methanol reactor. During the regeneration catalyst transportation process, C4 or C4 + Logistics has also become a common technical means to increase the yield of ethylene and propylene and reduce the unit consumption of methanol. Increasing the production of ethylene and propylene products by cracking olefin cracking technology in the process system has strong technical and economic benefits, especially good process rationality. However, due to the limited cracking activity and heat carried by the regenerated catalyst, its effect will be inhibited.

[0037] In order to solve the various problems and shortcomings existing in the above-mentioned prior art, the present invention further improves the yield of low-carbon olefins in the process of preparing olefins from methanol, adjusts and controls the catalyst pre-carbon deposition that affects the reaction induction period, makes the reaction activity and selectivity higher and the conversion of by-products more efficient, and fully couples and utilizes the heat during the conversion of different reaction processes, as well as improves and simplifies the process and the device as well as the operation process. In fact, it is to provide improved technologies including methods, catalysts and molecular sieves, reactors and process systems.

[0038] Therefore, the present invention provides a method for converting methanol to olefins in a fluidized bed, which is characterized by comprising the following specific steps: (1) preheating the raw methanol to 120-300° C. and feeding it into the dense phase section at the bottom of the fluidized bed coupled reactor through a methanol feed distributor, and heating the reactor at a temperature of 400-500° C., a pressure of 0.1-1 MPa, and a mass space velocity of 0.5-4 hours. -1 , methanol / water weight ratio 1: (0-5), methanol / catalyst weight ratio 1: (4-10) process conditions;

[0039] (2) contacting with a dual catalyst system consisting of a microsphere catalyst of a ZSM-5 molecular sieve modified with chromium and phosphorus on a high-pore volume γ-Al2O3 carrier and a microsphere catalyst of a SSZ-13 and SAPO-34 molecular sieve on a low-pore volume SiO2 carrier in a weight ratio of 1: (10-100) to carry out dehydration reaction conversion; and completely regenerating the carbon-deactivated catalyst by passing air at 550-720° C. in a regenerator through continuous circulation, so that the carbon content of the regenerated catalyst is reduced to <0.1wt%;

[0040] (3) After the products enter the subsequent separation process, the separated ethane, propane, butane or C4 components, C5 and above components are mixed and injected into and returned to the dilute phase section of the fluidized bed coupled reactor through a nozzle in a vertical direction to the outlet end of the regenerated catalyst, and undergo a millisecond-level catalytic cracking reaction and a partial dehydrogenation reaction with a contact time of 0.1 to 1 second with the regenerated high-temperature catalyst, and the dual catalyst system is selectively pre-carbonized non-uniformly, and thermal coupling is achieved between the reactions in different regions of the reactor;

[0041] (4) according to the step (3), the C4 component is returned to the fluidized bed coupling reactor, or the C4 component is separated into butane and butene, butane is returned to the fluidized bed reactor, butene and part of the ethylene are sent to the flexible fixed bed reactor, and heated at a temperature of 150 to 450 ° C, a pressure of 0.1 to 1.5 MPa, and a temperature of 0.5 to 1.5 hours. -1At a mass space velocity, the catalyst is contacted with a WO3 / SiO2 strip or pellet catalyst therein for reaction, and after the catalyst is deactivated, it is regenerated by passing air at a temperature of 350-620°C; the reaction product is mixed with the fluidized bed product and enters the subsequent separation process; (5) the mixed product gas of the above reaction is compressed, carbon dioxide is removed, and it is dehydrated and dried, and then separated by a separation process comprising C3-C4 separation, C2-C3 separation, ethylene distillation, propylene distillation and PSA pressure swing adsorption, to obtain product ethylene and propylene within the range of ethylene / propylene yield ratio of 0.8-1.2, and obtain by-product hydrogen and fuel gas.

[0042] The present invention provides a method for converting methanol to olefins in a fluidized bed, characterized in that the C4 components are separated, and the separated butenes and part of the ethylene enter a flexible switching fixed bed reactor in a reaction state for reaction and conversion, and the product ethylene / propylene ratio is reduced; the separated butane is mixed with other hydrocarbon by-products and then returned to the fluidized bed coupling reactor for cracking. Alternatively, the C4 components are directly mixed with other hydrocarbon by-products without separation and then returned to the fluidized bed coupling reactor, and all the ethylene is used as a product, and the product ethylene / propylene ratio is increased; the flexible switching fixed bed reactor is in a state of stopping feeding or regenerating.

[0043] The method for converting methanol to olefins provided by the present invention is characterized in that the microsphere catalyst of the ZSM-5 molecular sieve modified by chromium and phosphorus loaded on the high pore volume γ-Al2O3 carrier is composed of 25wt% to 50wt% of the ZSM-5 molecular sieve with a silicon-aluminum molar ratio of 10 to 15, 0.5wt% to 2wt% of phosphorus oxide, 0.5wt% to 2wt% of chromium oxide, and the balance of the high pore volume γ-Al2O3 carrier with a pore volume of 0.6 to 1.2 ml / g; the average microsphere radius of the catalyst is 60 to 100 microns, and the pore volume is 0.25 to 0.5 ml / g.

[0044] The present invention provides a method for producing olefins from methanol in a fluidized bed, characterized in that the microsphere catalyst loaded with SSZ-13 and SAPO-34 molecular sieves on the low pore volume SiO2 carrier is composed of 25wt% to 50wt% of SSZ-13 and SAPO-34 molecular sieves in a weight ratio of 1:(5 to 10) and a balance of a low pore volume silica carrier with a pore volume of 0.1 to 0.3 ml / g; the average microsphere radius of the catalyst is 60 to 100 microns, and the pore volume is 0.1 to 0.25 ml / g.

[0045] The present invention provides a method for producing olefins from methanol in a fluidized bed, characterized in that the WO3 / SiO2 strip or pellet catalyst is composed of a carrier containing 5wt% to 20wt% tungsten oxide and the remainder of silicon dioxide, and is prepared by impregnation or kneading to load active components, and extrusion or rolling forming; the pore capacity of the catalyst is 0.1 to 0.5 ml / g.

[0046] The present invention provides a method for producing olefins from methanol in a fluidized bed, characterized in that the amount of carbon deposited by the non-uniform selective pre-carbon deposition of the catalyst accounts for 5% to 30% of the carbon content of the catalyst to be regenerated after deactivation, and the weight ratio of the carbon deposited amount per unit weight of the ZSM-5 molecular sieve catalyst to the carbon deposited amount per unit weight of the SSZ-13 and SAPO-34 molecular sieve catalysts is 1:(0.1 to 0.6).

[0047] The present invention provides a method for producing olefins from methanol in a fluidized bed, characterized in that the strong endothermic reaction of the millisecond-level catalytic cracking reaction of by-product hydrocarbons and the dehydrogenation reaction of some alkanes in the fluidized bed coupling reactor are thermally coupled with the strong exothermic reaction of the methanol dehydration reaction; and the coking reaction and the non-uniform selective pre-carbon deposition process of the dual catalysts are reaction-coupled.

[0048] At present, MTO reactors include bubbling fluidized bed, turbulent fluidized bed, fast fluidized bed and other fluidized types, but still have problems such as low selectivity for low-carbon olefins. For example, in traditional riser reactors, raw materials and catalysts enter from the bottom of the reactor, and the logistics direction is upward. The axial mixing of the catalyst and gas components is very serious, which will inevitably cause a large number of secondary reactions, thereby reducing the production of olefins in the product. The particle residence time of the commonly used cyclone separator is about 0.5 seconds, which cannot fully meet the needs of rapid gas-solid separation with high selectivity for olefins, and there is also a problem of insufficient heat utilization.

[0049] When the two-stage cyclone separator used in the existing fluidized bed reactor technology is used for gas-solid separation, the interface of the fluidized bed catalyst is not very clear. The reaction process relies solely on gravity sedimentation to reduce the catalyst content of the gas at the inlet of the cyclone separator. The distance between the reaction zone and the inlet of the gas-solid separator is generally far. After leaving the ideal fluidized bed reaction, the reaction product gas will still stay in the catalyst environment for a long time. The actual reaction time is uncertain, making it difficult to actually terminate and control the reaction, and will also lead to an increase in side reactions. It is necessary to solve the existing problems in a targeted manner, that is, through coupling reaction functions, suppressing secondary reactions, controlling reaction time, improving the methanol processing capacity and low-carbon olefin selectivity of the fluidized bed reactor, and thermal coupling.

[0050] Therefore, the present invention also provides a fluidized bed methanol to olefins reactor, characterized in that the reactor is a fluidized bed coupled reactor with multi-layer stepped heat extraction, comprising a sealed shell with a larger upper part and a smaller lower part and a reaction vertical pipe, wherein the top of the shell is provided with 2 to 4 cyclone separators, an output gas outlet and a vent, and the bottom of the cyclone separator is provided with a vertical particle output pipe; the lower part of the shell is provided with a raw material inlet distribution pipe and a heat extraction coil located above the raw material distribution pipe; the nozzle of the by-product hydrocarbon return pipeline is perpendicular to the outlet of the regenerated catalyst pipeline; 3 to 8 layers of grid plates and 1 to 5 layers of steam coils are horizontally fixedly installed in the reaction vertical pipe at the lower part of the dense phase section of the reactor; the catalyst pipeline to be regenerated is connected to the fluidized bed regenerator through a control valve and a conveying pipeline, and the regenerated catalyst pipeline of the fluidized bed regenerator is also connected to the fluidized bed reactor through a control valve and a conveying pipe.

[0051] The present invention also provides a fluidized bed methanol to olefin process system, characterized in that it includes a reaction raw material methanol, a fluidized bed microsphere dual catalyst, a fixed bed catalyst, an alkali solution, process water and steam, fuel gas, nitrogen, compressed air; a fluidized bed coupling reactor, a fluidized bed regenerator, a flexible switching fixed bed reactor; a gas-solid separator, a washing tower, a depropanizer, a debutanizer, a demethanizer, a deethanizer, an ethylene distillation tower, a propylene distillation tower, a pressure swing adsorption device, an adsorption separator, a dryer; a heating furnace, a heat exchanger, a cooler; a material circulation pump, a gas compressor; the raw material, process water and the returned hydrocarbon by-product are respectively distributed by the reactor methanol feed after heat exchange or heating. The nozzle of the return pipeline of the reactor and hydrocarbon by-product enters the lower dense phase section of the fluidized bed coupled reactor and contacts with the regenerated dual catalyst system to carry out reaction conversion, selective pre-carbon deposition and thermal coupling; the temperature is controlled by steam stripping of the steam coil process and heat extraction of the heating pipe; hot air and fuel gas enter the regenerator through the feed distribution pipe to burn and completely regenerate the recycled catalyst to be regenerated; the reaction products are subjected to multi-stage gas-solid separation, multi-stage compression, carbon dioxide removal, drying, multi-stage distillation tower separation and pressure swing adsorption separation to obtain product propylene and ethylene, and by-product hydrogen and fuel gas; the by-product hydrocarbons are circulated back to the fluidized bed coupled reactor and enter the flexible switching fixed bed reactor for reaction conversion.

[0052] The molecular sieve and catalyst molding and manufacturing processes involved in the present invention; the start and stop of the fluidized bed reactor, the fluidization, circulation and addition of the catalyst in the fluidized bed reactor; the switching of raw materials, reactors and pipelines are conventional operations in the field, and are well known and used daily by ordinary technicians in the field. The raw materials and chemical reagents involved in the present invention can be easily obtained by commercial procurement.

[0053] It is well known to those skilled in the art that the conversion method, molecular sieve and catalyst, reactor and process system constitute the content, system and characteristics of the present invention, and are the most important factors affecting the catalytic conversion of alcohols and hydrocarbons; although innovative improvements based on the existing technology are also needed, they are different from the existing technology in terms of technical routes and principles, methods, key technical means, reactor core structure and process system composition.

[0054] Since the influence of the above factors on each other is nonlinear and faces many uncertainties, it is difficult to obtain direct inspiration from existing technologies, and it is also difficult to obtain the expected technical results through simple permutation and combination experiments based on existing technologies. Systematic research is needed to obtain valuable positive effects.

[0055] The present invention provides a method, reactor and process system for converting methanol into olefins in a fluidized bed, which have high-efficiency conversion performance and can make the conversion activity distribution in the reaction process, the coupled utilization of the processes and heat of different reactions, the selective and controllable pre-carbon deposition of the catalyst and the control of the change of bed temperature more balanced and reasonable.

[0056] Through the improvements and innovations made by the present invention, the drastic temperature rise, temperature change and imbalance of conversion activity in various regions of the reactor caused by the strong exothermic reaction of methanol conversion to olefins are slowed down, local overheating is reduced and eliminated; side reactions are suppressed, the catalyst after complete regeneration is controllably and selectively pre-carbonized, and the distribution of coking is adjusted and controlled, reducing the problems of rapid deactivation of active components in the catalyst dehydration reaction and insufficient cracking conversion activity; the yield of ethylene and propylene is improved, the hydrocarbon by-products are flexibly converted and the proportion of low-carbon olefins, especially ethylene, in the product distribution is guaranteed; the reaction-regeneration process of fine-tuning the product distribution and the utilization of hydrocarbon by-products is organically combined to make the conversion process more flexible. In addition, the process requirements and the complexity of the reaction equipment and process system are simplified; the integration of reaction and heat coupling utilization is improved; the production process is also made easier to operate and control, saving investment costs and improving the complexity and cost of operations and operation processes. BRIEF DESCRIPTION OF THE DRAWINGS

[0057] Figure 1 The present invention is a schematic flow diagram of a method and process system for producing propylene from methanol in a fluidized bed.

[0058] Figure 1 In: 1- raw material methanol; 2- fluidized bed methanol to olefins coupling reactor; 3- fixed bed catalytic reactor; 4- compression device; 5- carbon dioxide removal device; 6- dehydration and drying tower; 7- depropanizer; 8- demethanizer; 9- deethanizer; 10- debutanizer; 11- PSA pressure swing adsorption device; 12- ethylene distillation tower; 13, propylene distillation tower; 14- by-product fuel gas; 15- ethylene product; 16- by-product hydrogen; 17- propylene product.

[0059] Figure 2 The present invention is a schematic diagram of the structure of a fluidized bed coupled reactor for producing propylene from methanol in a fluidized bed.

[0060] Figure 2 In the figure: 1-raw material methanol feed distributor; 2-reactor (shell, dilute phase section); 18-by-product hydrocarbon return pipeline feed nozzle (vertical to the regenerated catalyst outlet); 19-regenerated catalyst delivery pipe; 20-horizontal delivery pipeline for catalyst to be regenerated; 21-lifting nitrogen and process air; 22-water vapor coil; 23-reactor temperature control and heat extraction coil; 24-first and second stage gas-solid cyclone separators; 25-export pipelines to the third and fourth stage gas-solid cyclone separators outside the reactor; 26-venting pipeline; 27-lower reaction vertical pipe and internal horizontal fixed grid plate. DETAILED DESCRIPTION

[0061] The following is a fluidized bed methanol to olefins method provided by the present invention, and the specific implementation methods and use effects of the reactor and reaction system are described in conjunction with the accompanying drawings and embodiments.

[0062] like Figure 2 As shown, the raw methanol enters the lower dense phase section reaction zone of the fluidized bed coupled reactor 2 and the lower reaction vertical pipe 27 through the feed pipeline and the feed distributor 1, contacts with the molecular sieve-containing dual catalyst system, reacts to generate a product flow containing ethylene and propylene, and enters the gas-solid rapid separation equipment 24. The deactivated catalyst enters the fluidized bed regenerator from the catalyst delivery pipeline 20 for regeneration.

[0063] After the regeneration is completed, the catalyst enters the fluidized bed coupled reactor from the regenerated catalyst delivery pipeline 19, reacts with the hydrocarbon by-products from the by-product hydrocarbon return pipeline and the feed nozzle 18, and causes the regenerated agent to undergo controllable selective pre-carbon deposition. The product gas merges with the product gas from the lower reaction area and enters the separation device 24.

[0064] according to Figure 1 The combined product gas enters the latter separation process after gas-solid separation 24, enters the washing tower 5 after multi-stage compression 4 to remove CO2, and enters the depropanizer 7 after dehydration and drying 6 to separate C4 hydrocarbons; then passes through the demethanizer 8 to separate hydrogen and methane, and enters the deethanizer 9 to separate C2 and C3 hydrocarbons.

[0065] The C2 hydrocarbons enter the ethylene distillation tower 12 for separation to obtain the product ethylene 15; the C3 hydrocarbons enter the propylene distillation tower 13 for separation to obtain the product propylene 17; the hydrogen and methane are separated by adsorption through the pressure swing adsorption device 11 to obtain the by-product fuel gas 14 and hydrogen 16; the separated ethane and propane are collected as hydrocarbon by-products and enter the hydrocarbon by-product return pipeline and nozzle 18 to return to the fluidized bed reactor 2, see Figure 1 and 2The process shown in .

[0066] according to Figure 1 In the process flow shown, the C4 hydrocarbons separated from the depropanizer 7 are not separated in one implementation, but are directly collected with ethane and propane and returned to the fluidized bed reactor 2. At this time, the fixed bed reactor 3 is in a state of stopping feeding, or can enter a catalyst regeneration state. In another implementation, the C4 hydrocarbons enter the debutanizer 10 to separate butane and butene, butane is collected with ethane and propane and returned to the reactor 2, butene and part of the ethylene from the ethylene rectification tower 12 enter the fixed bed reactor 3 together, contact and react with the catalyst therein, and the reaction product gas is collected with the product gas of the reactor 2 and enters the subsequent separation process together.

[0067] In the fluidized bed coupled reactor 2, see Figure 2 , including a sealed shell 2 with a larger upper part and a smaller lower part and a reaction vertical pipe 27, the reaction vertical pipe contains a multi-layer stepped steam coil 22 for steam stripping and dehydrocarbon removal, and a multi-layer grid plate fixedly installed laterally; a cyclone separator 24, an output gas outlet 25 and a vent 26 are located at the top of the reactor, and a vertical particle output pipe is provided at the bottom of the cyclone separator to complete the gas-solid separation of the catalyst and the product gas.

[0068] The lower part of the reactor shell is provided with an inlet distribution pipe 1 for raw material methanol, and a heat extraction coil 23 is located above the raw material distribution pipe for heat extraction and temperature control; the nozzle 18 of the byproduct hydrocarbon return pipeline is perpendicular to the outlet 19 of the regeneration agent pipeline, forming a millimeter-scale catalytic cracking reaction area and a controllable selective pre-carbon deposition of the catalyst. Figure 2 .

[0069] The deactivated spent catalyst passes through the control valve at the bottom of the reaction vertical pipe and the spent catalyst delivery pipeline 20, and is lifted by nitrogen and process air 21 to enter the connected fluidized bed regenerator for regeneration; the regenerator structure is similar to Figure 2 Similarly, the regenerated agent after char regeneration is circulated back to the fluidized bed coupling reactor 2 through the control valve and the conveying pipe 19.

[0070] The fluidized bed dual catalyst is composed of a microsphere catalyst of a ZSM-5 molecular sieve modified with chromium and phosphorus loaded on a high-pore volume γ-Al2O3 carrier, and a microsphere catalyst of a SSZ-13 and SAPO-34 molecular sieve loaded on a low-pore volume SiO2 carrier. The molecular sieve is mixed with a carrier binder and slurried and then spray-dried and molded. The average radius of the catalyst microsphere is 60 to 100 microns, preferably 65 to 85 microns.

[0071] The fixed bed WO3 / SiO2 catalyst is prepared by kneading or impregnating the metal active component and the carrier binder, and forming the catalyst into a pellet by rolling, or forming the catalyst into a strip by extrusion.

[0072] The present invention is further described below by way of examples, but is not limited to the contents of the examples. In the examples, the composition analysis of the raw materials and the reaction products is performed using an Agilent 6890N gas chromatograph; the molecular sieve structure analysis is performed using an X-ray diffractometer; the catalyst composition analysis is performed using an X-ray fluorescence analyzer; the catalyst carbon deposition is measured using an infrared carbon analyzer, and the particle size distribution of the particles is measured using a laser particle size analyzer.

[0073] For other analytical tests, please refer to the relevant analytical methods in (National Standard for Test Methods for Petroleum and Petroleum Products, published by China Standards Press in 1989) and (Petrochemical Analysis Methods (RIPP Test Methods), published by Science Press in 1990).

[0074] Example 1

[0075] This example is used to illustrate the preparation of the phosphorus-containing MFI structure ZSM-5 molecular sieve required in the present invention.

[0076] Referring to the preparation steps and contents of the embodiments in the inventor's authorized patents CN101428232B, CN107010639B, and CN111268689B, and feeding according to the calculated composition, a phosphorus-containing, low silicon-aluminum ratio ZSM-5 molecular sieve with high cracking activity and activity stability was prepared.

[0077] The prepared product has a molecular sieve framework with a silicon-aluminum molar ratio of 13.1, contains 3.1 wt% of phosphorus oxide in its chemical composition, has an average molecular sieve crystal particle size of 2 microns, and has a light oil micro-anti-cracking activity of 52%. Its micropore capacity of 0.5 nanometer micropores is 0.164 ml / g, and its pore capacity distribution of 4 nanometer mesopores is 0.1 ml / g, thus having a multi-level pore structure that is easy for hydrocarbon molecules to diffuse.

[0078] Example 2

[0079] This example is used to illustrate the preparation of the SSZ-13 molecular sieve with CHA structure required in the present invention.

[0080] With reference to the preparation steps and contents of the embodiment in the patent CN108190907B authorized by the inventor, and by adding materials according to the calculated composition, a SSZ-13 molecular sieve with high thermal stability and strong acid density was prepared.

[0081] The average particle size of the prepared molecular sieve crystals is 4 microns and the specific surface area is 735 m 2 / g, collapse temperature 1105℃, NH3-TPD strong acid center density 1.5mmol / g, total acid content 1.85mmol / g, with high thermal stability and strong acid density distribution.

[0082] Example 3

[0083] This example is used to illustrate the preparation of the SAPO-34 molecular sieve with CHA structure required in the present invention.

[0084] With reference to the preparation steps and contents of Examples 32 to 38 in the prior art USP4440871, which has exceeded the patent protection period, the SAPO-34 molecular sieve with low acid density was obtained by adding materials and preparing according to the calculated composition.

[0085] The prepared product molecular sieve has a phosphorus-aluminum molar ratio of 0.8 and a specific surface area of ​​550 m 2 / g, collapse temperature 1021℃, NH3-TPD strong acid center density 0.2mmol / g, with a lower strong acid density distribution.

[0086] Example 4

[0087] This example is used to illustrate the preparation of the catalyst containing ZSM-5 molecular sieve required in the present invention.

[0088] Referring to the preparation steps and contents of the embodiment in the inventor's patent application CN114988447A, and adding materials according to the calculated composition, the γ-Al2O3 precursor aluminum hydroxide colloid before calcination was prepared, and it was mixed and slurried with the phosphorus-containing MFI-structured ZSM-5 molecular sieve obtained in Example 1 according to the calculated composition ratio of 40wt% of the molecular sieve.

[0089] Referring to the preparation steps of the embodiments in the inventor's authorized patents CN101204666B and CN109092353B, a microsphere catalyst was prepared by a spray drying molding method, and after being calcined at 550°C for 2 hours, it was impregnated with a calculated amount of chromium nitrate solution for modification and drying, and then activated by calcination at 450°C / 1 hour.

[0090] The prepared microsphere catalyst of chromium and phosphorus modified ZSM-5 molecular sieve on the high pore volume γ-Al2O3 carrier contains 1.2wt% phosphorus oxide, 1.5wt% chromium oxide and a specific surface area of ​​205 m 2 / g, pore capacity 0.35 ml / g, and average microsphere radius of the catalyst 82 microns.

[0091] Example 5

[0092] This example is used to illustrate the preparation of the catalyst containing SSZ-13 molecular sieve and SAPO-34 molecular sieve required in the present invention.

[0093] Referring to the preparation steps of Example 4, and adding materials according to the calculated composition, silica sol (industrial product, SiO220wt%) and the SSZ-13 and SAPO-34 molecular sieves obtained in Example 2 and Example 3 are mixed and beaten, and then spray-dried to form, washed with 3.5wt% ammonium phosphate solution and dried, and activated by calcination at 450℃ / 1 hour.

[0094] The prepared microsphere catalyst with SSZ-13 and SAPO-34 molecular sieves loaded on low pore volume SiO2 carrier has a specific surface area of ​​253 m 2 / g, pore capacity 0.14 ml / g, and average microsphere radius of the catalyst 81 microns.

[0095] Example 6

[0096] This example illustrates the preparation of the WO3 / SiO2 catalyst required for the flexible switching fixed bed in the present invention.

[0097] White carbon black powder (industrial product, SiO2>98wt%, particle size <200 mesh) and silica sol (same as above) were kneaded in a weight ratio of 1:2 and extruded into strips. After calcining at 450℃ for 1 hour, the strips were impregnated with calculated amount of ammonium metatungstate, dried and calcined at 500℃ / 2 hours for activation.

[0098] The prepared WO3 / SiO2 catalyst with tungsten oxide loaded on a low-pore volume SiO2 carrier contains 8wt% tungsten oxide in its chemical composition and has a specific surface area of ​​253 m 2 / g, pore capacity 0.32 ml / g, and the strip catalyst size is φ1.5mm×15mm.

[0099] Example 7

[0100] This embodiment is used to illustrate the implementation of the fluidized bed methanol to olefins method, reactor, process system and catalyst of the present invention when the fixed bed is flexibly switched in the process system in a reaction operation state.

[0101] Reference Figure 2 In a medium-sized fluidized bed simulation test device, a methanol-containing raw material is introduced into the lower dense phase section of a fluidized bed coupling reactor 2 through a methanol raw material inlet raw material distributor 1, and then, in the presence of steam, it is contacted with the dual catalyst system prepared in Example 4 and Example 5 to carry out a methanol to olefin reaction conversion.

[0102] After entering the stable operation stage, the catalyst is in a controllable selective pre-carbon deposition reaction. The weight ratio of Example 4 catalyst: Example 5 catalyst is 1:20; the double dense phase bed temperature is 475°C, the pressure is 0.13MPa, and the air velocity is 1.5 hours.-1 The weight ratio of water to methanol is 1.5: 1. The obtained product gas and the catalyst to be produced enter the subsequent separation process through the gas-solid separation 24 and the product gas outlet 25.

[0103] The catalyst to be regenerated enters the lower reaction vertical pipe 27, and the hydrocarbons carried by it are removed by steam stripping 22. At this time, the carbon content of the catalyst to be regenerated is 6.1wt%. After the flow rate is adjusted by the regulating valve, it enters the horizontal conveying pipe 20 for the catalyst to be regenerated by air through nitrogen and air 21. The structure of the fluidized bed regenerator is similar to that of the regenerator. Figure 2 Similarly, the bed temperature of the dense phase section of the regenerator is 690°C, the pressure is 0.14MPa, and the residence time is about 1 hour to obtain a fully regenerated regenerated catalyst. At this time, the carbon content of the regenerated catalyst is less than 0.1wt%, and the flue gas and regeneration agent delivery pipe 19 are removed by steam stripping and circulated back to the fluidized bed coupling reactor 2. The flue gas generated by the regeneration reaction enters the subsequent treatment facility and is discharged after purification.

[0104] Reference Figure 1 The product gas entering the latter separation process enters the washing tower 5 after multi-stage compression 4 to remove carbon dioxide by alkali washing and water washing, and enters the depropanizer 7 to separate C4 hydrocarbons after dehydration and drying 6; then enters the deethanizer 9 to separate C2 hydrocarbons and C3 hydrocarbons after hydrogen and methane are separated by the demethanizer 8. The C2 hydrocarbons enter the ethylene distillation tower 12 to separate and obtain product ethylene 15; the C3 hydrocarbons enter the propylene distillation tower 13 to obtain product propylene 17; hydrogen and methane are separated by adsorption through the pressure swing adsorption device 11 to obtain by-product fuel gas 14 and by-product hydrogen 16.

[0105] The separated ethane and propane are collected as hydrocarbon by-products and enter the hydrocarbon return pipeline and nozzle 18 to return to the fluidized bed reactor 2; the depropanizer 7 separates the C4 hydrocarbons, and in this embodiment, enters the debutanizer 10 to separate butane and butene.

[0106] After butane is collected with ethane and propane, it returns to the fluidized bed coupled reactor 2 through the byproduct hydrocarbon return pipeline and the nozzle 18 perpendicular to the regeneration agent pipeline outlet 19, and contacts with the high-temperature regeneration agent of >650°C from the regenerator at the regeneration agent outlet, and performs a rapid catalytic cracking reaction with a contact time of about 0.3 to 0.9 seconds, and performs controllable selective pre-carbon deposition on the dual catalyst system. Through sampling and analysis, the average carbon deposition amount of the catalyst in Example 4 is 2.1wt%, and the average carbon deposition amount of the catalyst in Example 5 is 1.2wt%.

[0107] The butene separated from the debutanizer 10 and 10 wt% of ethylene from the ethylene distillation tower 12 enter the flexible switch fixed bed reactor 3 together, and react with the catalyst of Example 6 filled therein. The fixed bed reaction temperature is 160°C, the pressure is 0.6 MPa, and the space velocity is 1.5 hours.-1 The product gas is combined with the product gas from reactor 2 and enters the subsequent separation process together.

[0108] The specific reaction conversion results are shown in Table 1. It can be seen from the results in Table 1 that by using the method, reactor and process system of the present invention, in the case of adding a flexibly switchable fixed bed reactor to a set of fluidized bed reaction-regeneration devices, there is no need to set up more and more complex reaction facilities to achieve the complete conversion process of methanol raw materials and hydrocarbon by-products. Compared with the multiple fluidized bed reactors and regenerators commonly used in the current existing industrial application technology, the equipment investment and operating costs can be greatly reduced.

[0109] Table 1. Product distribution and reaction results after stable operation of Example 7 (based on carbon basis):

[0110]

[0111] It can be seen from Table 1 that, calculated on a carbon basis, Example 7 has a high ethylene selectivity and a high carbon basis yield of 85.5 wt% for ethylene + propylene under a high ethylene / propylene distribution ratio.

[0112] Example 8

[0113] This embodiment illustrates the implementation of the fluidized bed methanol to olefins method, reactor, process system including catalyst of the present invention when the fixed bed is flexibly switched to be in a regeneration or feed-stopping state in the process system.

[0114] Reference Figure 2 The medium-sized fluidized bed simulation test device and catalyst, methanol feed and pre-carbon deposition method are the same as those in Example 7, but the weight ratio of the catalyst in Example 4 to the catalyst in Example 5 is 1:10; in the fluidized bed coupling reactor, the bed temperature of the dense phase section is 475°C, the pressure is 0.14 MPa, and the space velocity is 1.0 hour -1 , product gas and catalyst to be generated are obtained, and the gas-solid separation is carried out in the same manner as in Example 7, and the product gas enters the subsequent separation treatment process.

[0115] Through sampling and analysis, the carbon content of the catalyst to be regenerated is 6.4wt%. The catalyst to be regenerated is circulated into the regenerator for regeneration with air in the same manner as in Example 7. The bed temperature of the dense phase section of the regenerator is 695°C and the pressure is 0.15MPa. A fully regenerated regenerated catalyst with a carbon content of less than 0.1wt% is obtained. The regenerated agent is returned to the fluidized bed coupled reactor 2, and the flue gas is treated in the same manner as in Example 7 and then discharged.

[0116] Reference Figure 1, the product gas entering the rear separation process is separated in the same manner as in Example 7 to obtain product ethylene and product propylene, as well as byproduct hydrogen and byproduct fuel gas. The difference from Example 7 is that the carbon four hydrocarbon components separated by the depropanizer 7 in this embodiment do not enter the debutanizer 10 to separate butane and butene, but are directly collected with the separated ethane and propane as hydrocarbon byproducts, and then returned to the fluidized bed coupling reactor 2, and contacted with a high-temperature regeneration agent of >650°C at the regeneration agent outlet to perform a rapid catalytic cracking reaction with a reaction contact time of about 0.3 to 0.9 seconds, and selective pre-carbon deposition is performed on the dual catalyst system. Through sampling and analysis, the average carbon deposition amount of the catalyst in Example 4 is 2.8wt%, and the average carbon deposition amount of the catalyst in Example 5 is 1.3wt%. After the product gas and the product gas of the methanol dehydration reaction are collected, they are separated by gas-solid separation 24 and enter the rear separation process.

[0117] In this embodiment, all the ethylene separated from the ethylene distillation tower 12 enters the ethylene product pipeline as the ethylene product output system; the fixed bed reactor 3 is flexibly switched to be in a feed stop state or in a regeneration state, and the regeneration condition is regeneration with air at 620°C.

[0118] The specific conversion reaction results are shown in Table 2. It can be seen from the results in Table 2 that, by using the method, reaction and process system of the present invention, in the case of Example 8, even when only the fluidized bed reaction-regeneration device is in operation, no more reaction facilities are required to achieve the complete conversion process of the methanol feedstock and the hydrocarbon by-products, which can significantly reduce the equipment investment and operating costs compared to the multiple fluidized bed reactors and regenerators commonly used in the current existing industrial application technology.

[0119] Table 2. Product distribution and reaction results after stable operation of Example 8 (based on carbon basis):

[0120]

[0121] It can also be seen from Table 2 that, calculated on a carbon basis, Example 8 achieves a high carbon basis yield of 85.3% for ethylene + propylene under conditions of high ethylene selectivity and ethylene / propylene product distribution.

[0122] It should be noted that the above embodiments are only used to illustrate the technical solution of the present invention and are not intended to limit it. Although the present invention has been described in detail with reference to the preferred embodiments, those skilled in the art should understand that the technical solution of the present invention may be modified or replaced by equivalents without departing from the spirit and scope of the technical solution of the present invention.

Claims

1. A method for converting methanol to olefins in a fluidized bed, characterized by comprising the following steps: (1) The raw material methanol is preheated to 120-300°C and enters the dense phase section at the bottom of the fluidized bed coupling reactor through a methanol feed distributor. The temperature is 400-500°C, the pressure is 0.1-1 MPa, and the mass space velocity is 0.5-4 hours. -1 , methanol / water weight ratio 1: (0-5), methanol / catalyst weight ratio 1: (4-10) process conditions; (2) contacting with a dual catalyst system consisting of a microsphere catalyst of a ZSM-5 molecular sieve modified with chromium and phosphorus on a high-pore volume γ-Al2O3 carrier and a microsphere catalyst of a SSZ-13 and SAPO-34 molecular sieve on a low-pore volume SiO2 carrier in a weight ratio of 1: (10-100) to carry out dehydration reaction conversion; and completely regenerating the carbon-deactivated catalyst by passing air at 550-720° C. in a regenerator through continuous circulation, so that the carbon content of the regenerated catalyst is reduced to <0.1wt%; (3) After the products enter the subsequent separation process, the separated ethane, propane, butane or C4 components, C5 and above components are mixed and injected into and returned to the dilute phase section of the fluidized bed coupled reactor through a nozzle in a vertical direction to the outlet end of the regenerated catalyst, and undergo a millisecond-level catalytic cracking reaction and a partial dehydrogenation reaction with a contact time of 0.1 to 1 second with the regenerated high-temperature catalyst, and the dual catalyst system is selectively pre-carbonized non-uniformly, and thermal coupling is achieved between the reactions in different regions of the reactor; (4) according to the step (3), the C4 component is returned to the fluidized bed coupling reactor, or the C4 component is separated into butane and butene, butane is returned to the fluidized bed reactor, butene and part of the ethylene are sent to the flexible fixed bed reactor, and heated at a temperature of 150 to 450 ° C, a pressure of 0.1 to 1.5 MPa, and a temperature of 0.5 to 1.5 hours. -1 At the mass space velocity, it contacts and reacts with the WO3 / SiO2 strip or pellet catalyst. After the catalyst is deactivated, it is regenerated by passing air at a temperature of 350-620°C. The reaction product is mixed with the fluidized bed product and enters the subsequent separation process. (5) The mixed product gas of the above reaction is compressed, carbon dioxide is removed, and it is dehydrated and dried, and then subjected to a separation process comprising C3-C4 separation, C2-C3 separation, ethylene distillation, propylene distillation, and PSA pressure swing adsorption to obtain product ethylene and propylene in an ethylene / propylene yield ratio in the range of 0.8 to 1.2, and obtain by-products of hydrogen and fuel gas; The microsphere catalyst of the ZSM-5 molecular sieve modified by chromium and phosphorus loaded on the high-pore volume γ-Al2O3 carrier is composed of 25wt% to 50wt% of the ZSM-5 molecular sieve with a silicon-aluminum molar ratio of 10 to 15, 0.5wt% to 2wt% of phosphorus oxide, 0.5wt% to 2wt% of chromium oxide, and the balance of the high-pore volume γ-Al2O3 carrier with a pore volume of 0.6 to 1.2 ml / g; the average microsphere radius of the catalyst is 60 to 100 microns, and the pore volume is 0.25 to 0.5 ml / g; The microsphere catalyst loaded with SSZ-13 and SAPO-34 molecular sieves on the low pore volume SiO2 carrier is composed of 25wt% to 50wt% of SSZ-13 and SAPO-34 molecular sieves, the weight ratio of which is 1:(5-10), and the balance of low pore volume silica carrier with a pore volume of 0.1-0.3 ml / g; the average microsphere radius of the catalyst is 60-100 microns, and the pore volume is 0.1-0.25 ml / g; The WO3 / SiO2 strip or pellet catalyst is composed of 5wt% to 20wt% tungsten oxide and the remainder of silicon dioxide carrier, and is prepared by impregnation or kneading to load active components, and extrusion or rolling ball molding; the pore capacity of the catalyst is 0.1 to 0.5 ml / g; The method for producing methanol to olefins in a fluidized bed adopts a fluidized bed methanol to olefin reactor. The fluidized bed methanol to olefin reactor is a fluidized bed coupled reactor with multi-layer stepped heat extraction, comprising a sealed shell with a larger upper part and a smaller lower part and a reaction vertical pipe. The top of the shell is provided with 2 to 4 cyclone separators, an output gas outlet and a venting port, and the bottom of the cyclone separator is provided with a vertical particle output pipe; the lower part of the shell is provided with a raw material inlet distribution pipe, and the heat extraction and temperature control coil is located above the raw material distribution pipe; the nozzle of the by-product hydrocarbon return pipeline is perpendicular to the outlet of the regenerated catalyst pipeline; 3 to 8 layers of grid plates and 1 to 5 layers of steam coils are fixedly installed horizontally in the reaction vertical pipe at the lower part of the dense phase section of the reactor; the catalyst pipeline to be regenerated is connected to the fluidized bed regenerator through a control valve and a conveying pipeline, and the regenerated catalyst pipeline of the fluidized bed regenerator is also connected to the fluidized bed reactor through a control valve and a conveying pipe.

2. The method for producing methanol to olefins in a fluidized bed according to claim 1, characterized in that: The carbon four components are separated, and the separated butene and part of the ethylene enter the flexible switching fixed bed reactor in a reaction state for reaction and conversion, and the product ethylene / propylene ratio is reduced; the separated butane is mixed with other hydrocarbon by-products and then returned to the fluidized bed coupling reactor for cracking.

3. The method for producing methanol to olefins in a fluidized bed according to claim 1, characterized in that: The C4 component is directly mixed with other hydrocarbon by-products without separation and then returned to the fluidized bed coupling reactor, all ethylene is used as a product, and the product ethylene / propylene ratio is increased; the fixed bed reactor is flexibly switched to be in a feeding stop or regeneration state.

4. The method for producing methanol to olefins in a fluidized bed according to claim 1, characterized in that: The carbon deposit amount of the catalyst non-uniform selective pre-carbon deposition accounts for 5% to 30% of the carbon content of the catalyst to be regenerated after deactivation, and the weight ratio of the carbon deposit amount per unit weight of the ZSM-5 molecular sieve catalyst to the carbon deposit amount per unit weight of the SSZ-13 and SAPO-34 molecular sieve catalyst is 1:(0.1 to 0.6).

5. The method for producing methanol to olefins in a fluidized bed according to claim 1, characterized in that: The millisecond-level catalytic cracking reaction of byproduct hydrocarbons and the strong endothermic dehydrogenation reaction of some alkanes in the fluidized bed coupling reactor are thermally coupled with the strong exothermic dehydration reaction of methanol; the coking reaction and the dual catalyst non-uniform selective pre-carbon deposition process are reaction coupled.

6. The method for producing methanol to olefins in a fluidized bed according to claim 1, characterized in that: The following process system is adopted, which includes reaction raw material methanol, fluidized bed microsphere dual catalyst, fixed bed catalyst, alkali solution, process water and steam, fuel gas, nitrogen, compressed air; fluidized bed coupling reactor, fluidized bed regenerator, flexible switching fixed bed reactor; gas-solid separator, washing tower, depropanizer, debutanizer, demethanizer, deethanizer, ethylene distillation tower, propylene distillation tower, pressure swing adsorption device, adsorption separator, dryer; heating furnace, heat exchanger, cooler; material circulation pump, gas compressor; the raw material, process water and returned hydrocarbon by-product are respectively fed by the reactor methanol feed distributor and hydrocarbon by-product distributor after heat exchange or heating. The nozzle of the return pipeline enters the lower dense phase section of the fluidized bed coupled reactor and contacts with the regenerated dual catalyst system to carry out reaction conversion, selective pre-carbon deposition and thermal coupling; the temperature is controlled by steam stripping and heat extraction pipes of the steam coil process; hot air and fuel gas enter the regenerator through the feed distribution pipe to burn and completely regenerate the recycled catalyst to be regenerated; the reaction products are subjected to multi-stage gas-solid separation, multi-stage compression, carbon dioxide removal, drying, multi-stage distillation tower separation and pressure swing adsorption separation to obtain product propylene and ethylene, and by-product hydrogen and fuel gas; the by-product hydrocarbons are recycled back to the fluidized bed coupled reactor and enter the flexible switching fixed bed reactor for reaction conversion.

Images