WO2021089049A1 - 一种提高含氧化合物转化制低碳烯烃选择性的方法及其装置 - Google Patents

一种提高含氧化合物转化制低碳烯烃选择性的方法及其装置 Download PDF

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WO2021089049A1
WO2021089049A1 PCT/CN2020/127548 CN2020127548W WO2021089049A1 WO 2021089049 A1 WO2021089049 A1 WO 2021089049A1 CN 2020127548 W CN2020127548 W CN 2020127548W WO 2021089049 A1 WO2021089049 A1 WO 2021089049A1
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hydrocarbon
pooling
regenerant
reaction
catalyst
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PCT/CN2020/127548
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English (en)
French (fr)
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李群柱
李瑞昀
李莉
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洛阳维达石化工程有限公司
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Priority to EP20884737.6A priority Critical patent/EP4056547A4/en
Priority to US17/775,519 priority patent/US20220289643A1/en
Publication of WO2021089049A1 publication Critical patent/WO2021089049A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • C07C1/24—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms by elimination of water
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/1836—Heating and cooling the reactor
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
    • B01J8/26—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/82—Phosphates
    • C07C2529/84—Aluminophosphates containing other elements, e.g. metals, boron
    • C07C2529/85—Silicoaluminophosphates (SAPO compounds)

Definitions

  • the invention relates to a method and a device for improving the selectivity of oxygen-containing compound conversion to low-carbon olefins.
  • Low-carbon olefins defined here as ethylene and propylene
  • ethylene and propylene are two important basic organic chemical raw materials.
  • the consumption of chemicals produced with ethylene and propylene as basic raw materials is increasing, and the demand for chemicals produced from ethylene and propylene remains high.
  • ethylene and propylene are mainly produced through petroleum routes.
  • the cost of producing ethylene and propylene from petroleum resources continues to increase.
  • Oxygenated compounds include methanol, ethanol, dimethyl ether, methyl ethyl ether and so on.
  • the raw materials include coal, natural gas, biomass, etc.
  • methanol can be made from coal or natural gas, and the technology is very mature.
  • Oxygen-containing organic compounds represented by methanol or dimethyl ether are mainly produced from coal-based or natural gas-based synthesis gas.
  • the process for producing low-carbon olefins based on ethylene and propylene using oxygenates represented by methanol as raw materials currently mainly includes MTO technology from UOP/Hyro in the United States, DMTO technology from Dalian Institute of Chemical Technology, Chinese Academy of Sciences, and MTP technology from Lurgi, Germany. .
  • the process of preparing low-carbon olefins from methanol is characterized by rapid reaction, strong exothermicity, and relatively low solvent alcohol.
  • the reaction and regeneration are carried out in a continuous reaction-regeneration dense fluidized bed reactor.
  • the high-temperature oil and gas rich in low-carbon olefins such as ethylene and propylene needs to be quenched and washed with water.
  • the target products of the MTO process unit are ethylene and propylene, and the by-products are ethane, propane, components above C5 and fuel gas (dry gas).
  • MTO has become a hotspot and focus of research by industry insiders. People have conducted extensive research and exploration in terms of processing flow, catalyst, process conditions, and equipment structure, and achieved satisfactory results, but there are not many literature reports on how to improve the selectivity of low-carbon olefins.
  • a certain amount of coke deposited on the SAPO-34 catalyst can greatly increase the yield of low-carbon olefins in the reaction product (and there is an optimal coke deposition range, and the selectivity of low-carbon olefins is the highest). Therefore, it is necessary to properly control the amount of coke deposited on the catalyst entering the conversion reactor to achieve the purpose of improving the selectivity of low-carbon olefins.
  • the catalyst with the best coke deposit the more uniform the coke distribution on the catalyst bed, the higher the selectivity of low-carbon olefins in the product. Therefore, how to achieve uniform control of the catalyst coke distribution in the MTO reactor (zone) is one of the keys to improving the selectivity of low-carbon olefins.
  • SAPO-34 is considered to be the preferred catalyst for the MTO process.
  • SAPO-34 catalyst has high selectivity of low-carbon olefins and high activity, which can make the reaction time of methanol to low-carbon olefins reach less than 10 seconds, and even reach the reaction time range required by riser reaction.
  • the US20060025646 patent relates to a method for controlling the amount of coke deposited on the catalyst in the reaction zone of the MTO reactor.
  • Part of the coke deposit catalyst i.e., spent catalyst, referred to as spent catalyst
  • spent catalyst is sent to the regeneration zone for coke burning, and the other part of the coke deposit catalyst is combined with
  • the regenerated catalyst (referred to as the regenerant) is mixed and returned to the reaction zone to continue the reaction.
  • US 6162282 discloses a method for converting methanol into low-carbon olefins.
  • a fast fluidized bed reactor is used. After the reaction in the dense phase reaction zone with a lower gas velocity is completed, the reaction gas and the entrained catalyst rise to the fast zone together. Most of the entrained catalyst is initially separated. Due to the rapid separation of the reaction product and the catalyst, the occurrence of secondary reactions is effectively prevented.
  • the carbon-based yields of low-carbon olefins in this method are generally around 77%, and there is also the problem of low yields of low-carbon olefins.
  • the DMTO technology of the Dalian Institute of Physics, Chinese Academy of Sciences uses a turbulent bed reactor, which has a relatively low operating gas velocity, usually 0.6 to 1.0 m/s.
  • the yield of low-carbon olefin carbon base in this method is generally 78-80%.
  • the MTO process follows the hydrocarbon pool mechanism, and the active species in the hydrocarbon pool may be olefin species, aromatic hydrocarbon species or both at the same time.
  • the shape-selection effect is enhanced, the reaction activity is significantly improved, and the molecular sieve exhibits autocatalytic properties.
  • the active species in the hydrocarbon pool with catalytic effect are not stable, and further reaction with olefins will cause fused cyclization and cause the catalyst to coke and deactivate.
  • the induction period of the catalytic methanol conversion reaction usually takes a few minutes to complete, while the induction period of the active species forming the "hydrocarbon pool" for the shape-selective catalytic methanol conversion reaction to low-carbon olefins requires tens of minutes or even hundreds of minutes. Tens of times, so the catalyst bed in the circulating fluidized bed reactor (zone) has the problem of uneven distribution of active species in the "hydrocarbon pool".
  • hydrocarbons from methanol is a very complex reaction process, involving tens of thousands of reactions and intermediate products, and there may be hundreds of reaction pathways; between olefin products, between aromatic products, olefins and aromatics, and others.
  • the types and quantity (or content) of the active species in the "hydrocarbon pool” are changed: the reaction conditions change, and the "hydrocarbon pool” The types and numbers of active species will change.
  • hydrocarbon and hydrocarbon pool active species generated during the conversion reaction of (olefin) hydrocarbons with carbon four or more at high temperature eg 530 ⁇ 600°C
  • MTO reaction conditions eg: 470 ⁇ 480°C
  • the active species of the “hydrocarbon pool” that catalyzes the reaction of methanol conversion to hydrocarbons and the “hydrocarbon pool” that catalyzes the reaction of methanol conversion to low-carbon olefins are not the same or different Exactly the same.
  • the active species of the "hydrocarbon pool" in the circulating fluidized bed reactor used in industrial applications is dynamic and changing (the type and quantity of the active species change with the reaction conditions (mainly the reaction temperature)), so
  • the catalyst bed not only has the problem of coke distribution, but also has the problem of uneven distribution of active species in the "hydrocarbon pool", which will inevitably affect the catalytic activity and selectivity of the methanol-to-low-carbon olefin reaction. In fact, this is the main reason for the low yield of low-carbon olefins in the MTO unit. But for a long time, the above-mentioned problems have not been discovered by people, so far no one has conducted special research and reports.
  • the high-temperature regenerated catalyst directly enters the conversion reactor, and there is a temperature difference of hundreds of degrees (usually 150-300°C) with the catalyst in the conversion reactor, which will cause local overheating of the catalyst bed (the high-temperature regenerant itself and its surroundings). In turn, there are many side reactions during the conversion reaction, a large amount of coke is produced, and the selectivity of low-carbon olefins is poor.
  • the purpose of the present invention is to provide a pre-hydrocarbon pooling facility (or reaction space) on the premise of ensuring a good regeneration effect to perform "pre-hydrocarbon pooling" treatment on the regenerated catalyst, in order to form a shape-selective catalytic oxygen-containing compound system.
  • the reaction of the active species in the "hydrocarbon pool" required for the reaction of low-carbon olefins provides sufficient reaction time and reaction space, so that the regenerant is formed before entering the reactor and meets the requirements of the conversion reaction conditions, and has good reaction activity for the production of low-carbon olefins.
  • Cooling technology breaks the heat balance of the reaction regeneration system, reduces the temperature at which the regenerated catalyst enters the conversion reactor by setting a regenerative catalyst cooler, eliminates local overheating in the conversion reactor, optimizes the temperature distribution of the conversion reactor, and improves the yield of low-carbon olefins. rate.
  • the technical problem to be solved by the present invention is to set up a pre-hydrocarbon pooling facility (or reaction space) to perform "pre-hydrocarbon pooling" treatment on the regenerated catalyst, so as to meet the requirements of the conversion reaction conditions and have a good low-carbon olefin reaction Active and selective "hydrocarbon pool” active species, improve the "hydrocarbon pool” active species distribution and coke distribution of the catalyst in the conversion reactor, so as to increase the activity and selectivity of the regenerant oxygenates to produce low-carbon olefins, thereby shortening Or eliminate the "induction period" of the conversion of oxygenated compounds to low-carbon olefins.
  • the present invention provides a method for improving the selectivity of oxygen-containing compound conversion to low-carbon olefins.
  • the regenerated catalyst (referred to as the regenerant) from the regenerator and stripped enters the pre-hydrocarbon pooling facility, where it comes into contact with the activation medium, and the pre-treatment occurs. Chemical reactions such as hydrocarbon pooling, forming "hydrocarbon pool” active species (and a certain amount of coke).
  • the pre-hydrocarbon pooling regenerator (referred to as the pre-hydrocarbon pooling catalyst or pre-hydrocarbon pooling regenerator) enters the conversion reactor for recycling.
  • the main operating conditions of the pre-hydrocarbon pooling facility are as follows: the reaction temperature is 300-600°C (preferably 360-560°C, more preferably 400-530°C), and the contact time is less than 300 minutes (preferably 0.001-200 minutes) , More preferably 10 to 150 minutes).
  • the pre-hydrocarbon pooling facility includes a regenerant inlet (including a catalyst distributor), (after pre-hydrocarbon pooling) a regenerant outlet, an activation medium inlet (including a distributor), an activation medium outlet, or/and Fluidized medium inlet (including distributor), etc.
  • the said pre-hydrocarbon pooling facility adopts various reactors used in industry, including any one, two or more of fluidized bed reactor, moving bed and fixed bed reactor, or Their combination.
  • the fluidized bed reactors include bubbling bed reactors, turbulent bed reactors, fast bed reactors or riser reactors, etc.
  • the riser reactors can be of various equal diameters or variable diameters used in industry. Diameter riser reactor.
  • the pre-hydrocarbon pooling facility is operated by a low-speed dense-phase fluidized bed, and its superficial gas velocity (the ratio of the flow rate of the fluidizing medium to the cross section of the empty tower of the equipment) is less than 0.5m/s ( It is preferably 0.0001 to 0.3 m/s, more preferably 0.001 to 0.2 m/s).
  • the activation medium entering the pre-hydrocarbon pooling facility may be any one, two or more of oxygen-containing compound raw materials, reaction products, various hydrocarbons or other oxygen-containing compounds, or a mixture thereof.
  • the oxygen-containing compound raw materials are any one, two or more of methanol, ethanol, dimethyl ether, methyl ethyl ether, etc., or a mixture thereof.
  • the reaction product may be a reaction gas product without separation or one or more stages of separators (including cyclones, cyclones, etc.) to remove part or all of the catalyst, or after heat exchange, cooling, and water washing.
  • the latter reaction gas is either the reaction gas boosted by the reaction gas compressor from the downstream olefin product separation device (unit), or the stripping gas (containing steam and reaction products, etc.) from the top of the sewage stripping tower, or Any one, two or more of them, or a mixture of them.
  • the various hydrocarbons can be any of the products (including ethylene, propylene, ethane, propane, mixed C4, C5 and above fractions, fuel gas, etc.) from the downstream olefin product separation device (unit) , Two or more, or a mixture of them; the various hydrocarbons can also be any one, two or more of various olefins, aromatics or alkanes, or a mixture of them.
  • the other oxygen-containing compounds are any one, two or more of any organic oxygen-containing compounds (including various alcohols, ethers, esters, aldehydes, ketones, etc.), or a mixture thereof .
  • the catalyst is any catalyst used in industry, including SAPO-34, ZSM-5 molecular sieve catalyst and the like.
  • the pre-hydrocarbon pooling facility can be installed outside the conversion reactor or inside the conversion reactor; it can be connected to the conversion reactor as a whole, or it can be connected to it through a conveying pipe.
  • a regenerator cooler can be installed before the regenerant enters the pre-hydrocarbon pooling facility to reduce the temperature of the regenerator to 200-630°C (preferably 300-600°C, more preferably 360- 560°C).
  • the regenerator cooler can be arranged before the regenerator stripper to strengthen the regenerator Mixing to eliminate the radial temperature difference caused by non-uniform heat transfer and non-uniform flow, so as to ensure that the temperature of the regenerant is balanced, so as to meet the requirements of pre-hydrocarbon pooling reaction temperature control and downstream conversion reaction temperature control, and improve the accuracy of temperature control Sex and flexibility.
  • a catalyst mixing buffer space needs to be arranged downstream of the regenerator cooler to eliminate the radial temperature difference caused by non-uniform heat transfer and non-uniform flow. The temperature reaches equilibrium.
  • the regenerant stripper or/and the regenerant mixing buffer space adopts low-speed dense phase fluidized bed operation, and its superficial gas velocity (the ratio of the flow rate of the fluidizing medium to the cross section of the empty tower of the equipment) is less than 0.5m/s (Preferably 0.0001 to 0.3 m/s, more preferably 0.001 to 0.2 m/s).
  • the stripping or fluidization medium of the regenerant stripper or the catalyst mixing buffer space can be steam or other fluids or a mixture thereof (preferably steam).
  • the above-mentioned pre-hydrocarbon pooling reaction or/and the reaction temperature of the conversion reactor are mainly adjusted by adjusting the circulating amount of the regenerator (that is, adjusting the control element of the cold regenerator delivery pipe, such as a slide valve, etc.), or/and mainly by adjusting the cold regenerator Temperature control; and/or adjust the heat extraction load of the heat extractor of the pre-hydrocarbon pooling reactor to keep it at an optimal value, thereby optimizing the temperature distribution of the conversion reactor, and further improving the conversion reaction Selectivity to improve the yield of low-carbon olefins and other targeted products.
  • the temperature of the cold regenerator is controlled by adjusting the flow rate of the fluidizing medium entering the regenerator cooler and/or the flow rate of the cold catalyst returning to the regenerator and/or other parameters.
  • an internal heat extractor or/and an external heat extractor can also be arranged inside or outside the pre-hydrocarbon pooling reactor to maintain the pre-hydrocarbon pooling reaction system The calorie balance.
  • the reaction temperature of the pre-hydrocarbon pooling reactor is mainly controlled by adjusting the return amount of the pre-hydrocarbon pooling regenerator, or/and the flow rate of the heating medium or/and the flow rate of the fluidizing medium or/and other parameters.
  • the fluidization medium may be air, steam, or other fluids or a mixture thereof
  • the heat extraction medium may be water, steam, air or other fluids, etc., or a mixture thereof.
  • the catalyst cooler is a mature industrial equipment.
  • the method and device of the present invention can adopt various structural forms (such as upflow type, downflow type, etc.) used in industry (including fluidized catalytic cracking unit, MTO device, etc.), and the catalyst conveying channel
  • Various specific connection structures such as internal circulation pipes, Y-shaped, U-shaped external conveying (circulation) pipes, etc.
  • degassing (balance) pipes can also be used, with or without degassing (balance) pipes, and those of ordinary skill in the art have specific structures and connections.
  • the selection and control of the type and operating parameters are very clear, and do not constitute a limitation to any specific implementation of the concept of the present invention.
  • the conversion reactor may be various reactors used in industry, including any one, two or more of fluidized bed reactors, moving bed reactors, and fixed bed reactors, or a combination thereof.
  • the conversion reactor can be various fluidized bed (including bubbling bed, turbulent bed, fast bed, etc.) reactors or riser reactors used in industry (including fluidized catalytic cracking units, MTO units, etc.) Any one, two or more of them, or a combination thereof, the riser reactor can be various riser reactors of equal diameter or variable diameter used in industry.
  • the fluidized bed (including riser) reaction regeneration device is a mature industrial process, and it uses various reactors, regenerators, internal or external catalyst coolers (or heat collectors, including up-flow, down-flow, and back-mixing). External heat extractors, etc.), steam (gas) strippers, catalyst distributors, steam (gas) gas distributors, etc., all of which can be used in the present invention.
  • Those of ordinary skill in the art have specific structures, combinations, operations, and control processes.
  • the conversion reaction conditions, the separation of reaction products and the regeneration of the catalyst are all carried out according to conventional methods.
  • the spent catalyst is regenerated by coking under normal regeneration conditions in the regenerator, and the regeneration temperature is usually controlled at 550-800°C (preferably 600-730°C, more preferably 650-710°C), the conversion reaction temperature is usually 400-560°C (preferably 420-520°C, more preferably 450-500°C).
  • the method and device of the present invention can adopt various reaction regeneration types used in industry (including fluidized catalytic cracking unit, MTO device, etc.).
  • reaction regeneration types used in industry (including fluidized catalytic cracking unit, MTO device, etc.).
  • MTO device fluidized catalytic cracking unit
  • Those of ordinary skill in the art are very clear about its specific structure, combination type, operation and control process. , Does not constitute a limitation to any specific implementation of the concept of the present invention.
  • the method of the present invention for improving the selectivity of the conversion of oxygenated compounds to low-carbon olefins uses a pre-hydrocarbon pooling facility to perform "pre-hydrocarbon pooling" treatment on the regenerated catalyst, so that the regenerated catalyst is formed before entering the conversion reactor.
  • Hydrocarbon pool” active species and coke deposits to improve the distribution of active species and coke deposits on the catalyst in the conversion reactor, thereby shortening or eliminating the “induction period” of the conversion reaction, and increasing the regeneration agent oxygenated compounds to produce low-carbon olefins The catalytic activity and selectivity of the reaction.
  • regenerant cooling technology breaks the heat balance of the reaction regeneration system, reduces the temperature at which the regenerant enters the conversion reactor, eliminates local overheating in the conversion reactor caused by the excessive temperature of the regenerator, and makes the bed of the conversion reactor
  • the temperature distribution of the layer is more uniform, which greatly promotes ideal reactions such as the conversion of oxygenated compounds to low-carbon olefins, and inhibits non-ideal reactions such as thermal polymerization of low-carbon olefins, thereby increasing the reaction selectivity, further improving the yield of low-carbon olefins, and reducing
  • the coking rate of the catalyst that is, the carbon difference between the regenerating agent and the spent agent).
  • the use of low-temperature regenerator circulation reduces the hydrothermal deactivation of the regenerator during the transportation process (before the conversion reactor), improves the activity of the regenerator, and reduces the consumption of the catalyst.
  • the catalyst mixing buffer space can be set to strengthen the mixing of the catalyst, so that the temperature of the regenerant can be balanced, uniform and stable, and the accuracy and flexibility of the control of the pre-hydrocarbon pooling reaction temperature and the downstream conversion reaction temperature can be improved.
  • the adjustment of the operating conditions such as the reaction temperature of oxygenate conversion to olefins, the catalyst circulation amount, etc. is relatively independent and more flexible, and can be flexibly adjusted according to market conditions to achieve different product distributions.
  • the low-temperature regenerant can be used as a cold shock agent, directly enters the rapid separation facility including the inlet or the outlet to realize the rapid termination of the conversion reaction, thereby inhibiting non-ideal reactions such as thermal polymerization of low-carbon olefins, and further improving the yield of low-carbon olefins , which reduces the coking rate of the catalyst (that is, the carbon difference between the regenerating agent and the spent agent).
  • Figures 1 to 4 are typical schematic diagrams of the method and device for improving the selectivity of oxygen-containing compound conversion to low-carbon olefins according to the present invention.
  • Figure 1 is a schematic block diagram of the method and device for improving the selectivity of oxygenate conversion to low-carbon olefins according to the present invention (the pre-hydrocarbon pooling facility is separately installed outside the conversion reactor).
  • the method and device for improving the selectivity of oxygenate conversion to low-carbon olefins of the present invention include (conversion) reactor 1, rapid separation facility 2, spent agent stripper 3, regenerator 4 , Regenerant stripper 6, Pre-hydrocarbon pooling facility 8.
  • the oxygen-containing compound raw material 11 enters the (conversion) reactor 1 after preheating and increasing temperature, and contacts with the catalyst to carry out the oxygen-containing compound conversion to olefin reaction.
  • the reaction product and the entrained part of the deactivated catalyst go up to the gas-solid rapid separation facility 2 in the conversion reactor 1 to quickly separate the catalyst and the reaction product, and then sequentially enter the first and second separators (such as cyclones, etc.) for further removal
  • the reaction gas after heat exchange, cooling, and water washing enters the downstream product separation system for further separation, so as to obtain the desired target products (such as ethylene and propylene). Etc.), and by-products (ethane, propane, mixed C4, and C5 and above fractions and fuel gas).
  • the main operating conditions of the conversion reactor 1 are as follows: the reaction temperature is 400-560°C (preferably 420-520°C, more preferably 450-500°C), the reaction pressure is 0.11-0.4MPa, and the weight ratio of (recycling) regenerant to raw material (abbreviated as The ratio of agent to alcohol) is 0.1 to 0.5 (preferably 0.12 to 0.3, more preferably 0.15 to 0.25).
  • the spent catalyst (abbreviated as spent catalyst) from the reaction zone 22 of the conversion reactor 1 is stripped by the spent catalyst stripper 3, and then enters the regenerator 4 for coking regeneration, so that the activity of the catalyst in the conversion reactor 1 is satisfied. Response requirements.
  • the regeneration temperature is usually controlled at 550-800°C (preferably 600-730°C, more preferably 650-710°C).
  • regenerant from the regenerator 4 passes through the regenerant stripper 6 and then enters the pre-hydrocarbon pooling facility 8.
  • chemical reactions such as contact with the activation medium 12 and pre-hydrocarbon pooling occur to form "hydrocarbon pool” active species and a certain amount of coke, so as to shorten or eliminate the "induction period” of the conversion of oxygenated compounds to low-carbon olefins.
  • the "pre-hydrocarbon pooling" regenerant after leaving the pre-hydrocarbon pooling facility 8 enters the conversion reactor 1 for recycling.
  • the main operating conditions of the pre-hydrocarbon pooling facility are as follows: the reaction temperature is 300-600°C (preferably 360-560°C, more preferably 400-530°C), and the contact time is less than 300 minutes (preferably 0.001-200 minutes) , More preferably 10 to 150 minutes).
  • an internal heat extractor or/and an external heat extractor can also be provided inside and/or outside to maintain the pre-hydrocarbon The heat balance of the pooled reaction system.
  • the reaction temperature of the pre-hydrocarbon pooling facility 8 is mainly determined by adjusting the return amount of the pre-hydrocarbon pooling regenerant, or/and the flow rate of the fluidization medium entering the heat extractor and/or the heat extraction medium flow rate and/or other Parameters are controlled.
  • Figure 2 is a schematic diagram of a process flow diagram of the method and device for improving the selectivity of oxygen-containing compound conversion to low-carbon olefins according to the present invention.
  • the method and device for improving the selectivity of oxygenated compound conversion to low-carbon olefins of the present invention include a conversion reactor 1, a rapid separation facility 2 (including a cyclone distributor 24), and a spent agent stripping device.
  • Reactor 3 regenerator 4, regenerant cooler 5, regenerant stripper 6, pre-hydrocarbon pooling facility 8.
  • a partition plate 25 is provided to isolate the reaction zone 22 and the sedimentation zone 9 (including the dilute phase zone 9A and the dense phase zone 9B).
  • the entrained catalyst separated by the rapid separation facility 2 and the first and second cyclone separators 20 and 21 enters the dense phase zone 9B of the sedimentation zone 9.
  • the oxygen-containing compound raw material 11 is preheated and heated up, and then enters the (conversion) reactor 1 reaction zone 22 through the distributor 46, and contacts the catalyst to carry out the oxygen-containing compound conversion to olefin reaction.
  • the reaction product and the entrained catalyst go up to the gas-solid rapid separation facility 2 in the conversion reactor 1 to quickly separate the catalyst and the reaction product, and then enter the first and second cyclone separators 20, 21 to further remove the spent catalyst, and then pass After the three-stage cyclone separator further removes the fine catalyst powder, the reaction gas after heat exchange, cooling, and water washing successively enters the downstream product separation system for further separation, so as to obtain the desired target products (such as ethylene and propylene, etc.), and By-products (ethane, propane, mixed C4, and C5 and above fractions and fuel gas).
  • desired target products such as ethylene and propylene, etc.
  • By-products ethane, propane, mixed C4, and C5 and above fractions and fuel gas
  • the main operating conditions of the conversion reactor 1 are as follows: the apparent linear velocity of the reaction zone 22 of the conversion reactor 1 is less than 1.2 m/s (preferably 0.5 to 1.0 m/s), and the reaction temperature is 400 to 560°C (preferably 420 to 520°C, It is more preferably 450-500°C), the reaction pressure is 0.11-0.4 MPa, and the weight ratio of the (recycling) regenerant to the raw material (abbreviated as agent alcohol ratio) is 0.1-0.5 (preferably 0.12-0.3, more preferably 0.15-0.25).
  • an internal heat extractor or/and an external heat extractor (not shown in the figure) is provided to remove excess heat in the conversion reaction process.
  • the spent agent from the reaction zone 22 of the conversion reactor 1 is stripped by the spent agent stripper 3, and then enters the regenerator 4 through the conveying pipe 37 [including the control valve 38 and the catalyst distributor (not shown in the figure)].
  • the coke-burning regeneration is performed so that the activity of the catalyst in the reaction zone 22 of the conversion reactor 1 meets the reaction requirements.
  • the regeneration temperature is controlled at 600-730°C (preferably 650-710°C).
  • the regenerant from the regenerator 4 enters the regenerator cooler 5 and is cooled to 200-630°C (preferably 300-550°C, more preferably 360-500°C, more preferably 420-480°C) and passes through the regenerator stripper 6 (in order to After entering the pre-hydrocarbon pooling facility 8 to make the temperature of the regenerant reach equilibrium), it will enter the pre-hydrocarbon pooling facility 8 through the regenerant delivery pipe 35 [including the control valve 36 and the catalyst distributor (not shown in the figure)].
  • chemical reactions such as contact with the activation medium 12 and pre-hydrocarbon pooling occur to form the "hydrocarbon pool" active species and a certain amount of coke, so as to shorten or eliminate the "induction period" of the reaction.
  • the activation medium 12 passes through the regenerant bed from bottom to top, and the reaction gas after the pre-hydrocarbon pooling enters the dilute phase zone 9A in the upper part of the sedimentation zone 9 through the pipeline 10.
  • the main operating conditions of the pre-hydrocarbon pooling facility are as follows: the reaction temperature is 300-600°C (preferably 360-560°C, more preferably 400-530°C), and the contact time is less than 300 minutes (preferably 0.001-200 minutes) , More preferably 10 to 150 minutes).
  • the pre-hydrocarbon pooling facility preferably operates in a low-speed dense-phase fluidized bed, and its superficial gas velocity is less than 0.5m/s (preferably 0.0001-0.3m/s, more preferably 0.001-0.2m/s).
  • the "pre-hydrocarbon pooling" regenerant after leaving the pre-hydrocarbon pooling facility 8 passes through the regenerant delivery pipe 33 [including the control valve 34 and the catalyst distributor (not shown in the figure)] and enters the conversion reactor 1 settlement zone 9 After mixing with the entrained catalyst in the lower dense phase zone 9B (or the transition zone 28 or the delivery pipe 26), it passes through the catalyst circulation pipe 31 (including the control valve 32) and then enters the reaction zone 22 of the conversion reactor 1 for recycling.
  • the "pre-hydrocarbon pooling" regenerant can directly enter the reaction zone 22 of the conversion reactor 1 through the regenerant delivery pipe 33 [including the control valve 34 and the catalyst distributor (not shown in the figure)] for recycling.
  • the stripping medium 13, 14, 16, the fluidizing medium 15, 19, and the lifting medium 17 may be steam or other fluids (preferably steam).
  • the lifting medium 18 may be air or other gas (preferably air).
  • the heat extraction medium 50 may be water, steam or other fluids (preferably water).
  • an internal heat extractor or/and an external heat extractor can also be provided inside and/or outside to maintain the pre-hydrocarbon The heat balance of the pooled reaction system.
  • the reaction temperature of the pre-hydrocarbon pooling facility 8 is mainly determined by adjusting the return amount of the pre-hydrocarbon pooling regenerant, or/and the flow rate of the fluidization medium entering the heat extractor and/or the heat extraction medium flow rate and/or other Parameters are controlled.
  • the regenerant stripper 6 when the regenerant stripper 6 is arranged upstream of the regenerant cooler 5, the A catalyst mixing buffer space is arranged downstream of the regenerant cooler 5 to strengthen the mixing of the regenerant, and the temperature of the regenerant is balanced before entering the pre-hydrocarbon pooling facility 8, so as to satisfy the "pre-hydrocarbon pooling" reaction and the downstream conversion reactor 1 Requirement for optimal control of reaction temperature.
  • the regenerant stripper 6 (or the catalyst mixing buffer space) can be connected by a pipe.
  • the regenerant stripper 6 (or catalyst mixing buffer space) can also be composed of an integrated structure (equal diameter or variable diameter) with the regenerator cooler 5 (as shown in FIG. 2).
  • the regenerant stripper 6 (or catalyst mixing buffer space) is operated by a low-speed dense-phase fluidized bed, and its superficial gas velocity is less than 0.5m/s (preferably 0.0001 ⁇ 0.3m/s, more preferably 0.001 ⁇ 0.2m/s) .
  • Figure 3 is a schematic diagram of the catalyst pre-hydrocarbon pooling method and equipment structure of the present invention (countercurrent contact).
  • the catalyst pre-hydrocarbon pooling (or pre-activation) equipment of the present invention includes a regenerant inlet 101 (including a catalyst distributor 41), a regenerant outlet 102, and an activation medium inlet 103 (including a distributor 42) , The activation medium outlet 104, or/and the fluidization medium inlet 105 (including the distributor 43), etc.
  • the regenerated catalyst 30 from the regenerator enters the upper part of the pre-hydrocarbon pooling reactor through the regenerant delivery pipe 33 [including the control valve and the catalyst distributor, not shown in the figure] successively through the regenerant inlet 101 and the catalyst distributor 41.
  • the flow is in countercurrent contact with the activation medium 12, and chemical reactions such as pre-hydrocarbon pooling occur to form a "hydrocarbon pool" of active species and a certain amount of coke to shorten or eliminate the "induction period" of the reaction.
  • the activation medium 12 sequentially enters the bottom of the pre-hydrocarbon pooling reactor through the activation medium inlet 103 and the distributor 42 and passes through the regenerant bed from bottom to top.
  • the reaction gas 10 after pre-hydrocarbon pooling passes through the activation medium outlet 104 and enters the downstream oxygen-containing mixture conversion reactor (sedimentation zone).
  • a one-stage or two-stage cyclone separator can also be arranged in the pre-hydrocarbon pooling reactor, after removing the catalyst entrained by the reaction gas after the pre-hydrocarbon pooling, the reaction gas enters the three-stage cyclone separator Entrance.
  • the "pre-hydrocarbon pooling" regenerator 40 After leaving the pre-hydrocarbon pooling reactor, the "pre-hydrocarbon pooling" regenerator 40 passes through the "pre-hydrocarbon pooling" regenerator outlet 102 and through the regenerant delivery pipe 35 [including control valve and catalyst distributor, not shown in the figure Out], enter the conversion reactor for recycling.
  • an internal heat extractor or/and an external heat extractor (not shown in the figure) can be installed inside or/and outside the pre-hydrocarbon pooling reactor.
  • the reaction temperature of the pre-hydrocarbon pooling reactor can be adjusted by adjusting the return amount of the pre-hydrocarbon pooling regenerator, or/and the heat extraction medium flow rate or/and the fluidization medium flow rate entering the heat extractor or/and other Parameters are controlled.
  • the fluidizing medium 19 may be steam or other fluid (preferably steam).
  • the activation medium 12 is preferably a reaction gas.
  • the heat extraction medium can be water, steam or other fluids (preferably water).
  • the catalyst distributor can be any industrially used catalyst distributor, and the vapor (gas) gas distributor can be any industrially used gas distributor (including distributor plates, distributor pipes, etc.).
  • the main operating conditions of the pre-hydrocarbon pooling reactor are as follows: the reaction temperature is 300-600°C (preferably 360-560°C, more preferably 400-530°C), and the contact time is less than 300 minutes (preferably 0.001-200 minutes, more preferably 10 to 150 minutes).
  • the pre-hydrocarbon pooling reactor adopts low-speed dense phase fluidized bed operation, and its superficial gas velocity is less than 0.5m/s (preferably 0.0001-0.3m/s, more preferably 0.001-0.2m/s).
  • Fig. 4 is a schematic diagram of a method for pre-hydrocarbon pooling (or pre-activation) of a catalyst of the present invention and its equipment (cocurrent contact).
  • the catalyst pre-hydrocarbon pooling (or pre-activation) equipment of the present invention includes a regenerant inlet 101 (including a catalyst distributor 41), a regenerant outlet 102, and an activation medium inlet 103 (including a distributor 42) , The activation medium outlet 104, or/and the fluidization medium inlet 105 (including the distributor 43), etc.
  • the regenerated catalyst 30 from the regenerator enters the bottom of the pre-hydrocarbon pooling reactor through the regenerant delivery pipe 33 [including the control valve and the catalyst distributor, not shown in the figure] successively through the regenerant inlet 101 and the catalyst distributor 41; 12 enters the bottom of the pre-hydrocarbon pooling reactor through the activation medium inlet 103 and the distributor 42 in turn, and the two co-currently flow through the regenerator bed from bottom to top, and chemical reactions such as pre-hydrocarbon pooling occur to form "hydrocarbon pool” active species And a certain amount of carbon deposits to shorten or eliminate the "induction period" of the reaction.
  • reaction gas after pre-hydrocarbon pooling enters the downstream oxygen-containing mixture conversion reactor (sedimentation zone) through the activation medium outlet 104.
  • the "pre-hydrocarbon pooling" regenerator 40 After leaving the pre-hydrocarbon pooling reactor, the "pre-hydrocarbon pooling" regenerator 40 passes through the "pre-hydrocarbon pooling" regenerator outlet 102 and through the regenerant delivery pipe 35 [including control valve and catalyst distributor, not shown in the figure Out], enter the conversion reactor for recycling.
  • an internal heat extractor or/and an external heat extractor (not shown in the figure) can be installed inside or/and outside the pre-hydrocarbon pooling reactor.
  • the reaction temperature of the pre-hydrocarbon pooling reactor can be adjusted by adjusting the return amount of the pre-hydrocarbon pooling regenerator, or/and the heat extraction medium flow rate or/and the fluidization medium flow rate entering the heat extractor or/and other Parameters are controlled.
  • the fluidizing medium 19 may be steam or other fluid (preferably steam).
  • the activation medium 12 is preferably a reaction gas.
  • the heat extraction medium can be water, steam or other fluids (preferably water).
  • the catalyst distributor can be any industrially used catalyst distributor, and the vapor (gas) gas distributor can be any industrially used gas distributor (including distributor plates, distributor pipes, etc.).
  • the main operating conditions of the pre-hydrocarbon pooling reactor are as follows: the reaction temperature is 300-600°C (preferably 360-560°C, more preferably 400-530°C), and the contact time is less than 300 minutes (preferably 0.001-200 minutes, more preferably 10 to 150 minutes).
  • the pre-hydrocarbon pooling reactor adopts low-speed dense phase fluidized bed operation, and its superficial gas velocity is less than 0.5m/s (preferably 0.0001-0.3m/s, more preferably 0.001-0.2m/s).
  • the oxygenate raw material is methanol
  • the process flow shown in Figure 2 is adopted
  • the catalyst is SAPO-34
  • the methanol conversion reactor and regenerator are under normal operating conditions
  • the reaction gas is the activation medium
  • the pre-hydrocarbon pooling facility The main operating conditions are as follows: the apparent linear velocity is 0.1-0.2m/s, the pre-hydrocarbon pooling reaction temperature is -460°C, and the contact time is 60-80 minutes.
  • the active species of the hydrocarbon pool on the regenerated catalyst after regeneration are basically restored, and a reasonable level of coke deposition is formed.
  • the oxygenate raw material is methanol
  • the process flow shown in Figure 2 is adopted
  • the catalyst is SAPO-34
  • the methanol conversion reactor and regenerator are both under normal operating conditions
  • the reaction gas is the activation medium
  • the pre-hydrocarbon pooling facility The main operating conditions are as follows: the apparent linear velocity is 0.1-0.2m/s, the pre-hydrocarbon pooling reaction temperature is 460°C, and the contact time is 30-40 minutes.

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Abstract

一种提高含氧化合物转化制低碳烯烃选择性的方法，来自再生器的再生催化剂进入预烃池化设施，在此与活化介质接触、发生预烃池化反应，形成"烃池"活性物种，离开预烃池化设施的预烃池化再生催化剂，进入转化反应器循环使用。通过设置预烃池化设施，对再生催化剂进行"预烃池化"处理，使再生催化剂在进入转化反应器前形成"烃池"活性物种和积炭，改善转化反应器内催化剂的"烃池"活性物种分布和积炭分布，从而缩短或消除转化反应的"诱导期"，提高再生剂含氧化合物制低碳烯烃反应的催化活性和选择性。

Description

一种提高含氧化合物转化制低碳烯烃选择性的方法及其装置 技术领域
本发明涉及一种提高含氧化合物转化制低碳烯烃选择性的方法及其装置。
背景技术
低碳烯烃，这里定义为乙烯和丙烯，是两种重要的基本有机化工原料。随着经济的持续快速发展，以乙烯、丙烯为基本原料生产的化学品消费日益增加，乙烯、丙烯生产的化学品需求量居高不下。乙烯、丙烯传统上主要是通过石油路线制得，但由于石油资源有限的供应量及较高的价格，由石油资源生产乙烯、丙烯的成本不断增加。随着石油资源的短缺，采用非石油原料生产甲醇，再将甲醇催化转化生产乙烯、丙烯等轻质烯烃具有重要的意义。
近年来，人们开始大力发展替代能源转化技术，如含氧化合物转化制烯烃(OTO)的工艺，含氧化合物包括甲醇、乙醇、二甲醚、甲乙醚等。目前，有许多技术可用来生产含氧化合物，其原料包括煤、天然气、生物质等，例如甲醇就可以由煤或天然气制得，工艺已十分成熟。由于含氧化合物来源的广泛性，再加上其制取低碳烯烃工艺的经济性，所以由含氧化合物转化制烯烃(OTO)的工艺，特别是由甲醇转化制烯烃(MTO)的工艺受到越来越多的重视。
以甲醇或二甲醚为代表的含氧有机化合物，主要由煤基或天然气基的合成气生产。用以甲醇为代表的含氧化合物为原料生产以乙烯和丙烯为主的低碳烯烃工艺目前主要有美国UOP/Hyro公司的MTO技术、中国科学院大连化物所的DMTO技术和德国Lurgi公司的MTP技术。
以甲醇制取低碳烯烃工艺(简称MTO)的反应特点是快速反应、强放热、且剂醇比较低，在连续的反应-再生的密相流化床反应器中进行反应和再生。反应生成的富含乙烯和丙烯等低碳烯烃的高温油气，需要进行急冷和水洗，除去其中催化剂和降温后，送往下部烯烃分离系统进行分离。MTO工艺装置的目的产品是乙烯和丙烯，副产品为乙烷、丙烷、C5以上组分和燃料气(干气)，其中C4＝的碳基收率为10％左右。
近年来，MTO已成为业内人士研究的热点和重点。人们从加工流程、催化剂、工艺条件以及设备结构等方面进行了广泛的研究和探索，取得令人满意的成果，但有关如何提高低碳烯烃选择性的文献报道不多。
对于MTO技术而言，SAPO-34催化剂上一定量的积炭，可大大提高反应产物中低碳 烯烃的产率(且存在最佳积炭范围，低碳烯烃的选择性最高)。因此，要对进入转化反应器的催化剂积炭量进行适当的控制，进而达到提高低碳烯烃选择性的目的。此外，对于具有最佳积炭量的催化剂而言，催化剂床层的积炭分布越均匀，产物中低碳烯烃的选择性就越高。因此，如何实现MTO反应器(区)内催化剂积炭分布的均匀控制是提高低碳烯烃选择性的关键之一。
US4499327专利中对磷酸硅铝分子筛催化剂应用于甲醇制烯烃工艺进行了详细的研究，认为SAPO-34是MTO工艺的首选催化剂。SAPO-34催化剂具有很高的低碳烯烃选择性，而且活性较高，可使甲醇转化为低碳烯烃的反应时间达到小于10秒的程度，甚至达到提升管反应要求的反应时间范围内。
US20060025646专利中涉及一种控制MTO反应器反应区中催化剂积炭量的方法，是将积炭催化剂(即待生催化剂，简称待生剂)一部分送入再生区烧炭，另一部分积炭催化剂与再生催化剂(简称再生剂)混合后返回到反应区继续反应。但是，该方法中进入提升管反应器内的两股催化剂之间的炭差很大，致使反应器内催化剂积炭分布很不均匀；而反应器内含有较多炭的催化剂以及含有很少炭的催化剂都对低碳烯烃的选择性不利，致使低碳烯烃选择性变差、目的产物(低碳烯烃)收率降低。
US 6166282中公布了一种甲醇转化为低碳烯烃的方法，采用快速流化床反应器，在气速较低的密相反应区反应完成后，反应气体与其夹带的催化剂一起上升到快分区，初步分离出大部分夹带催化剂。由于反应产物与催化剂的快速分离，有效地防止了二次反应的发生。经模拟计算，与传统的鼓泡流化床反应器相比，该快速流化床反应器内径及催化剂所需藏量均大大减少。该方法的低碳烯烃碳基收率一般均在77％左右，也存在低碳烯烃收率较低的问题。
中国科学院大连化物所的DMTO技术采用湍流床反应器，其操作气速较低，通常为0.6～1.0m/s。该方法的低碳烯烃碳基收率一般均在78～80％。
在上述方法中，进入转化反应器内的再生剂与转化反应器内催化剂之间的碳差很大(而含有较多碳的催化剂以及含有很少碳的催化剂都对低碳烯烃的选择性不利)，都存在低碳烯烃选择性差、低碳烯烃收率低的问题。
大量研究表明，MTO过程遵循烃池机理，烃池活性物种可能是烯烃物种、芳烃物种或二者同时发挥作用。随着分子筛中烃池活性物种的增多，择形作用增强，反应活性显著提高，表现出自催化特性。然而，具有催化作用的烃池活性物种并不稳定，其与烯烃等进一步反应将发生稠环化而导致催化剂结焦失活。
现有的再生器中，通常采用高温(550～800℃)烧焦再生。研究表明，高温再生后的再生剂虽然仍带有“碳”，但能够催化甲醇制低碳烯烃反应的“烃池”活性物种经过高温再生后已经不存在了。
催化甲醇转化反应的诱导期通常需要几分钟即可完成，而形成择形催化甲醇转化制低碳烯烃反应的“烃池”活性物种的诱导期则需要几十分钟甚至上百分钟，两者相差几十倍，因此循环流化床反应器(区)中催化剂床层就存在“烃池”活性物种分布不均匀的问题。
然而，甲醇制烃类是一个十分复杂的反应过程，涉及的反应及中间产物有上万种之多，反应途径可能有上百种；烯烃产物之间、芳烃产物之间、烯烃与芳烃和其它烃类之间都存在平衡反应，它们之间的转化反应受到热力学平衡的限制，同时又受到动力学的制约。因此，“烃池”活性物种的形成与种类就受到热力学平衡和动力学的制约，“烃池”活性物种的种类和数量(或含量)都是变化的：反应条件发生变化，“烃池”活性物种的种类和数量都会发生变化。例如：高温(例如：530～600℃)下碳四以上的(烯)烃类的转化反应过程中生成的“炭”和“烃池”活性物种在MTO反应条件(如：470～480℃)下也不一定具有催化制低碳烯烃反应的活性，催化甲醇转化制烃类反应的“烃池”活性物种和催化甲醇转化制低碳烯烃反应的“烃池”活性物种并不相同或并不完全相同。
综上所述，工业应用的循环流化床反应器中“烃池”活性物种是动态的、变化的(其种类和数量等是随着反应条件(主要是反应温度)在变化的)，因此催化剂床层不仅存在积炭分布的问题，同时还存在“烃池”活性物种分布不均匀的问题，这样就势必影响甲醇制低碳烯烃反应的催化活性和选择性。事实上，这是MTO装置低碳烯烃收率低的主要原因。但是长期以来，上述问题一直没有被人们所发现，至今没有人进行专门的研究与报道。
因此，如何实现MTO反应器(区)内“烃池”活性物种(特别是能在转化反应条件下提高催化含氧化合物制低碳烯烃反应的活性和选择性的“烃池”活性物种)的均匀分布是提高低碳烯烃选择性的关键之一。
同时，高温再生催化剂直接进入转化反应器，与转化反应器内的催化剂存在数百度的温度差(通常为150～300℃)，将引起催化剂床层局部过热(高温再生剂本身及其周围)，进而导致在转化反应过程中副反应多，生焦量大，低碳烯烃的选择性差。
本发明的目的是在保证良好的再生效果的前提下，通过设置预烃池化设施(或反应空间)，对再生催化剂剂进行“预烃池化”处理，为形成择形催化含氧化合物制低碳烯烃反应所要求的“烃池”活性物种的反应提供足够的反应时间和反应空间，使再生剂在进入反应 器前形成符合转化反应条件要求、并具有良好的制低碳烯烃反应活性和选择性的“烃池”活性物种，以提高再生剂含氧化合物制低碳烯烃的活性和选择性，改善转化反应器内催化剂的“烃池”活性物种分布和积炭分布，同时采用再生剂冷却技术，打破反应再生系统的热量平衡，通过设置再生催化剂冷却器降低再生催化剂进入转化反应器的温度，消除转化反应器内的局部过热，优化转化反应器的温度分布，进而提高低碳烯烃收率。
发明内容
本发明要解决的技术问题是通过设置预烃池化设施(或反应空间)，对再生催化剂剂进行“预烃池化”处理，形成符合转化反应条件要求、并具有良好的制低碳烯烃反应活性和选择性的“烃池”活性物种，改善转化反应器内催化剂的“烃池”活性物种分布和积炭分布，以提高再生剂含氧化合物制低碳烯烃的活性和选择性，从而缩短或消除含氧化合物转化制低碳烯烃反应的“诱导期”。同时降低再生催化剂进入转化反应器的温度，消除转化反应器内的局部过热，优化转化反应器(反应区)内的反应温度分布，进一步提高再生剂含氧化合物制低碳烯烃(即乙烯和丙烯)的活性和选择性，进而提高低碳烯烃收率。
本发明提供一种提高含氧化合物转化制低碳烯烃选择性的方法，来自再生器、汽提后的再生催化剂(简称再生剂)进入预烃池化设施，在此与活化介质接触、发生预烃池化等化学反应，形成“烃池”活性物种(和一定的积炭)。离开预烃池化设施预烃池化后的再生剂(简称预烃池化催化剂或预烃池化再生剂)，进入转化反应器循环使用。
所述的预烃池化设施(反应器)的主要操作条件如下：反应温度300～600℃(优选360～560℃，更优选400～530℃)、接触时间小于300分钟(优选0.001～200分钟，更优选10～150分钟)。
所述的预烃池化设施(反应器)包括再生剂进口(包括催化剂分配器)、(预烃池化后)再生剂出口、活化介质进口(包括分布器)、活化介质出口，或/和流化介质入口(包括分布器)等。
所述的预烃池化设施(反应器)采用工业上使用的各种反应器，包括流化床反应器、移动床、固定床反应器中的任意一种、两种或多种，或是它们的组合。所述的流化床反应器包括鼓泡床反应器、湍流床反应器、快速床反应器或提升管反应器等，所述的提升管反应器可以是工业上使用的各种等直径或变直径的提升管反应器。优选地，所述的预烃池化设施(反应器)采用低速密相流化床操作，其表观气速(流化介质流量与设备的空塔横截面之比)小于0.5m/s(优选0.0001～0.3m/s，更优选0.001～0.2m/s)。
所述的进入预烃池化设施的活化介质可以是含氧化合物原料、反应产物、各种烃类 或其它含氧化合物中的任意一种、两种或多种，或者是它们的混合物。
所述的含氧化合物原料是甲醇、乙醇、二甲醚、甲乙醚等的任意一种、两种或多种，或者是它们的混合物。
所述的反应产物可以是未经分离或经一级或多级分离器(包括旋风分离器、旋流分离器等)除去部分或全部催化剂的反应气体产物，或者是经过换热、冷却、水洗后的反应气，或者是来自下游烯烃产物分离装置(单元)反应气压缩机升压后的反应气，或者是来自污水汽提塔顶的汽提气(含有蒸汽和反应产物等)，或者是它们中的任意一种、两种或多种，或者是它们的混合物。
所述的各种烃类可以是来自下游烯烃产物分离装置(单元)的产品(包括乙烯、丙烯、乙烷、丙烷、混合碳四、碳五以上的馏分和燃料气等)中的任意一种、两种或多种，或者是它们的混合物；所述的各种烃类还可以是各种烯烃、芳烃或烷烃中的任意一种、两种或多种，或者是它们的混合物。
所述的其它含氧化合物是任何有机含氧化合物(包括各种醇类、醚类、酯类、醛类、酮类等)中的任意一种、两种或多种，或者是它们的混合物。
所述的催化剂是工业上使用的任何催化剂，包括SAPO-34、ZSM-5分子筛催化剂等。
所述的预烃池化设施可以设置于转化反应器外部，也可以设置于转化反应器内部；可以与转化反应器连为一体，也可以通过输送管与其相连。本领域普通技术人员对其具体结构、连接型式、操作条件和控制过程非常清楚,不构成对本发明构思的任何具体实施方式的限制。
为实现预烃池化反应所需的反应温度，可以在再生剂进入预烃池化设施前设置再生剂冷却器将再生剂温度降低到200～630℃(优选300～600℃，更优选360～560℃)。
进一步地，为实现预烃池化反应温度或/和下游含氧化合物转化反应器的反应温度的优化控制，所述的再生剂冷却器可设置在再生剂汽提器前，以强化再生剂的混合，消除非均匀传热和非均匀流动而产生的径向温差，以保证使再生剂温度达到均衡，从而满足预烃池化反应温度控制和下游转化反应温度控制的要求，提高温度控制的精准性和灵活性。再生剂冷却器设置再生剂汽提器后时，还需在所述的再生剂冷却器下游设置催化剂混合缓冲空间，以消除非均匀传热和非均匀流动而产生的径向温差，使再生剂温度达到均衡。
所述的再生剂汽提器或/和再生剂混合缓冲空间采用低速密相流化床操作，其表观气速(流化介质流量与设备的空塔横截面之比)小于0.5m/s(优选0.0001～0.3m/s，更优选 0.001～0.2m/s)。所述的再生剂汽提器或催化剂混合缓冲空间的汽提或流化介质可以是蒸汽或其他流体或者是它们的混合物(优选蒸汽)。
本领域普通技术人员对其具体结构、连接型式、操作和控制过程非常清楚,不构成对本发明构思的任何具体实施方式的限制。
上述预烃池化反应或/和转化反应器的反应温度主要通过调节再生剂循环量(即调节上述冷再生剂输送管的控制元件如滑阀等)，或/和主要通过调节冷再生剂的温度进行控制；和/或调节所述的预烃池化反应器取热器的取热负荷，使其保持在最佳值，从而优化所述转化反应器的温度分布，更进一步地提高转化反应选择性，提高低碳烯烃等目的产品的产率。所述的冷再生剂温度通过调节进入再生剂冷却器的流化介质流量和/或返回再生器的冷催化剂的流量和/或其它参数进行控制。
为更好地控制所述的预烃池化反应器的反应温度，还可在预烃池化反应器内部或外部设置内取热器或/和外取热器以维持预烃池化反应系统的热量平衡。所述的预烃池化反应器的反应温度主要通过调节预烃池化再生剂的返回量，或/和取热介质的流量或/和流化介质的流量或/和其它参数进行控制。
所述的流化介质可以是空气、蒸汽或其他流体或者是它们的混合物，取热介质可以是水、蒸汽、空气或其他流体等或者是它们的混合物。
当然还可有许多其他控制设备和控制方法，不构成对本发明构思的任何具体实施方式的限制。
催化剂冷却器为成熟工业设备，本发明的方法及其装置可采用工业上(包括流化催化裂化装置、MTO装置等)使用的各种结构形式(如上流式、下流式等)，催化剂输送通道也可采用各种具体连接结构(如内循环管、Y型、U型外输送(循环)管等)，设置或不设脱气(平衡)管，本领域普通技术人员对其具体结构、连接型式、操作参数(表观线速等)的选取和控制非常清楚,不构成对本发明构思的任何具体实施方式的限制。
所述的转化反应器可以是工业上使用的各种反应器，包括流化床反应器、移动床、固定床反应器中的任意一种、两种或多种，或是它们的组合。所述的转化反应器可以是工业上(包括流化催化裂化装置、MTO装置等)使用的各种流化床(包括鼓泡床、湍流床、快速床等)反应器或提升管反应器等中的任意一种、两种或多种，或是它们的组合，所述的提升管反应器可以是工业上使用的各种等直径或变直径的提升管反应器。
流化床(包括提升管)反应再生装置为成熟工业过程，其使用的各种反应器、再生器、内或外催化剂冷却器(或称取热器，包括上流式、下流式、返混式外取热器等)、汽 (气)提器、催化剂分布器、汽(气)体分布器等，本发明均可使用，本领域普通技术人员对其具体结构、组合型式、操作和控制过程非常清楚，对其操作条件【如进料温度、反应温度、反应压力、接触时间、剂醇比(或剂油比，即催化剂与原料之比)、表观线速，等等】和催化剂的选用也非常清楚，均不构成对本发明构思的任何具体实施方那个式的限制。
采用本发明的方法及其装置，其转化反应条件、反应产物的分离及催化剂的再生均按常规方法进行，待生催化剂在再生器中于常规再生条件下进行烧焦再生，再生温度通常控制在550～800℃(优选600～730℃，更优选650～710℃)，转化反应温度通常为400～560℃(优选420～520℃，更优选450～500℃)。
本发明的方法及其装置可采用工业上(包括流化催化裂化装置、MTO装置等)使用的各种反应再生型式，本领域普通技术人员对其具体结构、组合型式、操作和控制过程非常清楚，不构成对本发明构思的任何具体实施方式的限制。
与现有技术相比有如下优点：
1、本发明的提高含氧化合物转化制低碳烯烃选择性的方法，通过设置预烃池化设施，对再生催化剂进行“预烃池化”处理，使再生催化剂在进入转化反应器前形成“烃池”活性物种和积炭，改善转化反应器内催化剂的“烃池”活性物种分布和积炭分布，从而缩短或消除转化反应的“诱导期”，提高再生剂含氧化合物制低碳烯烃反应的催化活性和选择性。
2、采用再生剂冷却技术，打破了反应再生系统的热量平衡，降低再生剂进入转化反应器的温度，消除再生剂温度过高而引起的转化反应器内的局部过热，使转化反应器的床层温度分布更加均匀，大大促进了含氧化合物转化制低碳烯烃等理想反应，抑制了低碳烯烃热聚等非理想反应，从而提高了反应选择性，进一步提高了低碳烯烃收率，降低了催化剂的结焦率(即再生剂与待生剂的碳差)。
3、采用低温再生剂循环，减轻了再生剂在输送过程中(到转化反应器前)的水热失活，提高再生剂活性，降低了催化剂消耗。
4、可设置催化剂混合缓冲空间，强化催化剂的混合，使再生剂温度达到均衡，均匀稳定，提高了预烃池化反应温度和下游转化反应温度控制的精准性和灵活性。
同时，还有效地提高了再生剂的密度和缓冲能力，提高再生剂循环系统的推动力，从而提高操作的可控性、可靠性、稳定性和灵活性，实现了转化反应温度的优化控制和反应深度的优化控制。
5、含氧化合物转化制烯烃的反应温度、催化剂循环量等操作条件的调节相对独立，更加灵活，可以根据市场情况灵活调整，以实现不同的产品分布。
6、低温再生剂可以用作冷激剂，直接进入快速分离设施包括入口或出口，实现转化反应的快速终止，从而抑制了低碳烯烃热聚等非理想反应，进一步提高了低碳烯烃收率，降低了催化剂的结焦率(即再生剂与待生剂的碳差)。同时，还可实现所述的反应气体在所述的再生剂上的预烃池化反应，形成烃池活性物种和积炭，缩短或消除反应“诱导期”，提高再生剂(制取低碳烯烃的反应)的活性和选择性，以进一步提高低碳烯烃收率。
附图说明
附图1～4：为本发明的提高含氧化合物转化制低碳烯烃选择性的方法及其装置的典型示意图。
下面结合附图详细说明本发明，附图是为了说明本发明而绘制的，不构成对本发明构思的任何具体实施方式的限制。
具体实施方式
附图1为本发明的提高含氧化合物转化制低碳烯烃选择性的方法及其装置的方块流程示意图(预烃池化设施单独设置在转化反应器外部)。
如附图1所示：本发明的提高含氧化合物转化制低碳烯烃选择性的方法及其装置包括(转化)反应器1、快速分离设施2、待生剂汽提器3、再生器4、再生剂汽提器6、预烃池化设施8。
含氧化合物原料11经预热升温后进入到(转化)反应器1，与催化剂接触，进行含氧化合物转化制烯烃反应。反应产物与夹带的部分失活催化剂上行至转化反应器1内的气固快速分离设施2将催化剂和反应产物快速分离，而后依次进入一、二级分离器(如旋风分离器等)进一步脱除催化剂，再经三级旋风分离器进一步除去微量催化剂细粉后，依次经过换热、冷却、水洗后的反应气体进入下游产品分离系统进行进一步分离，从而得到所需要的目的产品(如乙烯和丙烯等)，和副产品(乙烷、丙烷、混合碳四、和碳五以上馏份以及燃料气)。
转化反应器1主要操作条件如下：反应温度为400～560℃(优选420～520℃，更优选450～500℃)、反应压力为0.11～0.4MPa，(循环)再生剂与原料重量比(简称剂醇比)为0.1～0.5(优选0.12～0.3，更优选0.15～0.25)。
快速分离设施2和一、二级旋风分离器分离出的部分失活催化剂返回转化反应器1。来自转化反应器1反应区22的待生催化剂(简称待生剂)经待生剂汽提器3汽提后，进入再生器4进行烧焦再生，以便使转化反应器1内催化剂的活性满足反应要求。
再生温度通常控制在550～800℃(优选600～730℃，更优选650～710℃)。
来自再生器4的再生剂经再生剂汽提器6，后进入预烃池化设施8。在此与活化介质12接触、发生预烃池化等化学反应，形成“烃池”活性物种和一定的积炭，以缩短或消除含氧化合物转化制低碳烯烃反应的“诱导期”。离开预烃池化设施8后的“预烃池化”再生剂，进入转化反应器1循环使用。
所述的预烃池化设施(反应器)的主要操作条件如下：反应温度300～600℃(优选360～560℃，更优选400～530℃)、接触时间小于300分钟(优选0.001～200分钟，更优选10～150分钟)。
为更好地控制所述的预烃池化设施8的反应温度，还可在其内部和/或外部设置内取热器或/和外取热器(图中未画出)以维持预烃池化反应系统的热量平衡。所述的预烃池化设施8的反应温度主要通过调节预烃池化再生剂的返回量，或/和进入所述取热器的流化介质流量和/或取热介质流量和/或其它参数进行控制。
附图2为本发明的提高含氧化合物转化制低碳烯烃选择性的方法及其装置的一种工艺流程示意图。
如附图1所示：本发明的提高含氧化合物转化制低碳烯烃选择性的方法及其装置包括转化反应器1、快速分离设施2(包括旋流分配器24)、待生剂汽提器3、再生器4、再生剂冷却器5、再生剂汽提器6、预烃池化设施8。
设置隔板25将反应区22和沉降区9(包括稀相区9A、密相区9B)隔离分开。在气速较低的反应区22反应后的反应气体产物，与其夹带的催化剂一起进入缩径的输送管26上升，经旋流分配器24切线进入快速分离设施2，分离出绝大部分夹带的部分失活催化剂。快速分离设施2和一、二级旋风分离器20、21分离出的夹带催化剂进入沉降区9密相区9B。
含氧化合物原料11经预热升温后经分布器46进入到(转化)反应器1反应区22，与催化剂接触，进行含氧化合物转化制烯烃反应。反应产物与夹带的催化剂上行至转化反应器1内的气固快速分离设施2将催化剂和反应产物快速分离，而后进入一、二级旋风分离器20、21进一步脱除待生催化剂后，再经三级旋风分离器进一步除去微量催化剂细粉后，依次经过换热、冷却、水洗后的反应气体进入下游产品分离系统进行进一步分离，从而得到所需要的目的产品(如乙烯和丙烯等)，和副产品(乙烷、丙烷、混合碳四、和碳五以上馏份以及燃料气)。
转化反应器1主要操作条件如下：转化反应器1反应区22的表观线速小于1.2m/s(优选0.5～1.0m/s)，反应温度为400～560℃(优选420～520℃，更优选450～500℃)、 反应压力为0.11～0.4MPa，(循环)再生剂与原料重量比(简称剂醇比)为0.1～0.5(优选0.12～0.3，更优选0.15～0.25)。
为有效控制转化反应温度，设置内取热器或/和外取热器(图中未画出)移除转化反应过程的过剩热量。
来自转化反应器1反应区22的待生剂，经待生剂汽提器3汽提后，经过输送管37【包括控制阀38、催化剂分配器(图中未画出)】进入再生器4进行烧焦再生，以便使转化反应器1反应区22内催化剂的活性满足反应要求。
再生温度控制在600～730℃(优选650～710℃)。
来自再生器4的再生剂进入再生剂冷却器5冷却到200～630℃(优选300～550℃，更优选360～500℃，更优选420～480℃)经再生剂汽提器6(以便在进入预烃池化设施8前使再生剂温度达到均衡)后，经再生剂输送管35【包括控制阀36、催化剂分配器(图中未画出)】后进入预烃池化设施8。在此与活化介质12接触、发生预烃池化等化学反应，形成“烃池”活性物种和一定的积炭，以缩短或消除反应的“诱导期”。所述的活化介质12自下而上通过再生剂床层，预烃池化后的反应气体经管道10进入沉降区9上部稀相区9A。
所述的预烃池化设施(反应器)的主要操作条件如下：反应温度300～600℃(优选360～560℃，更优选400～530℃)、接触时间小于300分钟(优选0.001～200分钟，更优选10～150分钟)。
所述的预烃池化设施(反应器)优选低速密相流化床操作，其表观气速小于0.5m/s(优选0.0001～0.3m/s，更优选0.001～0.2m/s)。
离开预烃池化设施8后的“预烃池化”再生剂，经再生剂输送管33【包括控制阀34、催化剂分配器(图中未画出)】，进入转化反应器1沉降区9下部密相区9B(或过渡区28或输送管26)内与夹带催化剂混合后，再经催化剂循环管31(包括控制阀32)后进入转化反应器1反应区22，循环使用。
或者所述的“预烃池化”再生剂，经再生剂输送管33【包括控制阀34、催化剂分配器(图中未画出)】，直接进入转化反应器1反应区22，循环使用。
汽提介质13、14、16，流化介质15、19，提升介质17可以是蒸汽或其他流体(优选蒸汽)。提升介质18可以是空气或其他气体(优选空气)。
取热介质50可以是水、蒸汽或其他流体(优选水)。
为更好地控制所述的预烃池化设施8的反应温度，还可在其内部和/或外部设置内取热器或/和外取热器(图中未画出)以维持预烃池化反应系统的热量平衡。所述的预烃池化 设施8的反应温度主要通过调节预烃池化再生剂的返回量，或/和进入所述取热器的流化介质流量和/或取热介质流量和/或其它参数进行控制。
为实现“预烃池化”反应温度的优化控制和/或下游含氧化合物转化反应的反应温度优化控制，当再生剂汽提器6设置在再生剂冷却器5上游时，可在所述的再生剂冷却器5的下游设置催化剂混合缓冲空间，强化再生剂的混合，在进入预烃池化设施8前使再生剂温度达到均衡，以满足“预烃池化”反应和下游转化反应器1反应温度优化控制的要求。
再生剂汽提器6(或催化剂混合缓冲空间)可以通过管道相连接。
为实现节约空间和节省投资，再生剂汽提器6(或催化剂混合缓冲空间)也可以与再生剂冷却器5组成(等直径或变直径的)一体式结构(如图2所示)。再生剂汽提器6(或催化剂混合缓冲空间)采用低速密相流化床操作，其表观气速小于0.5m/s(优选0.0001～0.3m/s，更优选0.001～0.2m/s)。
附图3为本发明的一种催化剂预烃池化方法及其设备结构示意图(逆流接触)。
如附图3所示：本发明的催化剂预烃池化(或称预活化)设备包括再生剂进口101(包括催化剂分配器41)、再生剂出口102、活化介质进口103(包括分布器42)、活化介质出口104，或/和流化介质入口105(包括分布器43)等。
来自再生器的再生催化剂30经再生剂输送管33【包括控制阀、催化剂分配器，图中未画出】依次通过再生剂进口101、催化剂分配器41进入预烃池化反应器上部，向下流动与活化介质12逆流接触、发生预烃池化等化学反应，形成“烃池”活性物种和一定的积炭，以缩短或消除反应的“诱导期”。
所述的活化介质12依次通过活化介质进口103、分布器42进入预烃池化反应器底部，自下而上通过再生剂床层。预烃池化后的反应气体10经活化介质出口104，进入下游含氧混合物转化反应器(沉降区)。
可选择地，也可以在预烃池化反应器内设置一级或两级旋风分离器，除去预烃池化后的所述反应气体夹带的催化剂后，所述反应气体进入三级旋风分离器入口。
离开预烃池化反应器后的“预烃池化”再生剂40，经“预烃池化”再生剂出口102、经再生剂输送管35【包括控制阀、催化剂分配器，图中未画出】，进入转化反应器，循环使用。
为更好地控制所述的预烃池化反应器的反应温度，可在预烃池化反应器内部或/和外部设置内取热器或/和外取热器(图中未画出)以维持预烃池化反应系统的热量平衡。所述的预烃池化反应器的反应温度可以通过调节预烃池化再生剂的返回量，或/和进入所述取热 器的取热介质流量或/和流化介质流量或/和其它参数进行控制。
流化介质19可以是蒸汽或其他流体(优选蒸汽)。活化介质12优选反应气。取热介质可以是水、蒸汽或其他流体(优选水)。
所述催化剂分布器可以采用工业上使用的任何催化剂分配器，所述汽(气)体分布器可以采用工业上使用的任何气体分布器(包括分布板、分布管等)。
所述的预烃池化反应器的主要操作条件如下：反应温度300～600℃(优选360～560℃，更优选400～530℃)、接触时间小于300分钟(优选0.001～200分钟，更优选10～150分钟)。
所述的预烃池化反应器采用低速密相流化床操作，其表观气速小于0.5m/s(优选0.0001～0.3m/s，更优选0.001～0.2m/s)。
附图4为本发明的一种催化剂预烃池化(或称预活化)的方法及其设备的结构示意图(并流接触)。
如附图4所示：本发明的催化剂预烃池化(或称预活化)设备包括再生剂进口101(包括催化剂分配器41)、再生剂出口102、活化介质进口103(包括分布器42)、活化介质出口104，或/和流化介质入口105(包括分布器43)等。
来自再生器的再生催化剂30经再生剂输送管33【包括控制阀、催化剂分配器，图中未画出】依次通过再生剂进口101、催化剂分配器41进入预烃池化反应器底部；活化介质12依次通过活化介质进口103、分布器42进入预烃池化反应器底部，两者并流自下而上通过再生剂床层，发生预烃池化等化学反应，形成“烃池”活性物种和一定的积炭，以缩短或消除反应的“诱导期”。
预烃池化后的反应气体经活化介质出口104，进入下游含氧混合物转化反应器(沉降区)。
离开预烃池化反应器后的“预烃池化”再生剂40，经“预烃池化”再生剂出口102、经再生剂输送管35【包括控制阀、催化剂分配器，图中未画出】，进入转化反应器，循环使用。
为更好地控制所述的预烃池化反应器的反应温度，可在预烃池化反应器内部或/和外部设置内取热器或/和外取热器(图中未画出)以维持预烃池化反应系统的热量平衡。所述的预烃池化反应器的反应温度可以通过调节预烃池化再生剂的返回量，或/和进入所述取热器的取热介质流量或/和流化介质流量或/和其它参数进行控制。
流化介质19可以是蒸汽或其他流体(优选蒸汽)。活化介质12优选反应气。取热介 质可以是水、蒸汽或其他流体(优选水)。
所述催化剂分布器可以采用工业上使用的任何催化剂分配器，所述汽(气)体分布器可以采用工业上使用的任何气体分布器(包括分布板、分布管等)。
所述的预烃池化反应器的主要操作条件如下：反应温度300～600℃(优选360～560℃，更优选400～530℃)、接触时间小于300分钟(优选0.001～200分钟，更优选10～150分钟)。
所述的预烃池化反应器采用低速密相流化床操作，其表观气速小于0.5m/s(优选0.0001～0.3m/s，更优选0.001～0.2m/s)。
实例1
对于实例1，含氧化合物原料为甲醇，采用图2所示的工艺流程，催化剂为SAPO-34,甲醇转化反应器和再生器均采用常规操作条件，反应气为活化介质，预烃池化设施主要操作条件如下：表观线速0.1～0.2m/s，预烃池化反应温度为～460℃，接触时间60～80分钟。
在此，再生后的再生催化剂上的烃池活性物种基本恢复，并形成合理的积炭水平。模拟计算结果显示，增设本发明的预烃池化设施后，与现有MTO技术相比，本发明的低碳烯烃(C 2 ＝+C 3 ＝)选择性提高5.7个百分点。
主要参数与效果对比见表1。
表1
参数 现有MTO技术 本发明
转化反应温度℃ 480 480
再生温度℃ 680 680
剂/醇比重/重 0.23 0.23
再生剂定炭％ 2.3 2.3
再生剂冷却器 无 有
冷后再生剂温度   ℃   460
预烃池化设施 无 有
预烃池化反应温度 ℃   460
低碳烯烃(C 2 ＝+C 3 ＝)选择性％ 79.8 85.5
实例2
对于实例2，含氧化合物原料为甲醇，采用图2所示的工艺流程，催化剂为SAPO-34,甲醇转化反应器和再生器均采用常规操作条件，反应气为活化介质，预烃池化设施主要操作条件如下：表观线速0.1～0.2m/s，预烃池化反应温度为460℃，接触时间30～40分钟。
模拟计算结果显示，增设本发明的预烃池化设施后，与现有MTO技术相比，本发明的低碳烯烃(C 2 ＝+C 3 ＝)选择性提高3.1个百分点。
主要参数与效果对比见表2。
表2
参数 现有MTO技术 本发明
转化反应温度 ℃ 480 480
再生温度 ℃ 680 680
剂/醇比 重/重 0.23 0.23
再生剂定炭 ％ 2.3 2.3
再生剂冷却器 无 有
冷后再生剂温度 ℃   460
预烃池化设施 无 有
预烃池化反应温度 ℃   460
低碳烯烃(C 2 ＝+C 3 ＝)选择性％ 79.8 82.9

Claims (20)

  1. 一种提高含氧化合物转化制低碳烯烃选择性的方法，其特征在于，来自再生器的再生催化剂进入预烃池化设施，在此与活化介质接触、发生预烃池化反应，形成“烃池”活性物种，离开预烃池化设施的预烃池化再生催化剂，进入转化反应器循环使用。
  2. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施的主要操作条件如下：反应温度为300～600℃、接触时间小于300分钟。
  3. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施的主要操作条件如下：反应温度为360～560℃、接触时间0.001～200分钟。
  4. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施的主要操作条件如下：反应温度为400～530℃、接触时间10～150分钟。
  5. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施采用低速密相流化床反应器，其表观气速小于0.5m/s。
  6. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施采用低速密相流化床反应器，其表观气速为0.0001～0.3m/s。
  7. 按照权利要求1所述的方法，其特征在于，所述的预烃池化设施采用低速密相流化床反应器，其表观气速为0.001～0.2m/s。
  8. 按照权利要求1-7任一项所述的方法，其特征在于，所述的活化介质是含氧化合物原料、反应产物、各种烃类或其它含氧化合物中的任意一种、两种或多种，或者是它们的混合物；所述的含氧化合物原料是甲醇、乙醇、二甲醚、甲乙醚等的任意一种、两种或多种，或者是它们的混合物；所述的反应产物是未经分离或经一级或多级分离器除去部分或全部夹带催化剂的反应气体，或者是经过换热、急冷、水洗后的反应气，或者是来自下游烯烃产物分离装置反应气压缩机升压后的反应气，或者是来自污水汽提塔顶的汽提气，或者是它们中的任意一种、两种或多种，或者是它们的混合物；所述的各种烃类是来自下游烯烃产物分离装置包括乙烯、丙烯、乙烷、丙烷、混合碳四、C 5以上的馏分和燃料气在内的产品中的任意一种、两种或多种，或者是它们的混合物；或者是各种烯烃、芳烃或烷烃中的任意一种、两种或多种，或者是它们的混合物；所述的其它含氧化合物是任何有机含氧化合物中的任意一种、两种或多种，或者是它们的混合物。
  9. 按照权利要求1-8任一项所述的方法，其特征在于，所述预烃池化设施内部或/和外部设置有内取热器或/和外取热器；所述预烃池化设施的反应温度主要通过调节预烃池化再生剂的返回量，或/和进入所述取热器的取热介质流量或/和流化介质流量或/和其它参数进行控制。
  10. 按照权利要求8或9所述的方法，其特征在于，所述的再生剂进入所述的预烃池化设施前设置再生剂冷却器，将再生剂温度降低到200～630℃。
  11. 按照权利要求8或9所述的方法，其特征在于，所述的再生剂进入所述的预烃池化设施前设置再生剂冷却器，将再生剂温度降低到300～600℃。
  12. 按照权利要求8或9所述的方法，其特征在于，所述的再生剂进入所述的预烃池化设施前设置再生剂冷却器，将再生剂温度降低到360～560℃。
  13. 按照权利要求8或9所述的方法，其特征在于，所述的再生剂冷却器设置在再生剂汽提器前，再生剂冷却器与再生剂汽提器组成等直径或变直径的一体式结构。
  14. 按照权利要求10-13任一项所述的方法，其特征在于，所述的再生剂冷却器和/或再生剂汽提器均为低速密相流化床操作，其表观气速小于0.5m/s。
  15. 按照权利要求10-13任一项所述的方法，其特征在于，所述的再生剂冷却器和/或再生剂汽提器均为低速密相流化床操作，其表观气速为0.0001～0.3m/s。
  16. 按照权利要求10-13任一项所述的方法，其特征在于，所述的再生剂冷却器和/或再生剂汽提器均为低速密相流化床操作，其表观气速为0.001～0.2m/s。
  17. 按照权利要求10-13任一项所述的方法，其特征在于，所述预烃池化设施的反应温度或/和所述转化反应器的反应温度主要通过调节再生剂循环量，或/和主要通过调节冷再生剂的温度进行控制，所述冷再生剂温度通过调节进入再生剂冷却器的流化介质流量和/或返回再生器的冷催化剂的流量和/或取热介质流量进行控制；或者所述预烃池化设施的反应温度或/和所述转化反应器的反应温度主要通过调节预烃池化再生剂的返回量或/和进入所述预烃池化设施取热器的取热介质流量或/和流化介质流量或/和其它参数进行控制。
  18. 一种提高含氧化合物转化制低碳烯烃选择性的装置，其特征在于，包括预烃池化设施。
  19. 按照权利要求18所述的装置，其特征在于，还包括所述的再生剂冷却器和/或再生剂汽提器。
  20. 按照权利要求18或19所述的装置，其特征在于，所述预烃池化设施采用工业上使用的包括流化床反应器在内的各种反应器中的任意一种、两种或多种，或是它们的组合；所述的流化床反应器包括鼓泡床反应器、湍流床反应器、快速床反应器或提升管反应器等，所述的提升管反应器可以是工业上使用的各种等直径或变直径的提升管反应器。
PCT/CN2020/127548 2019-11-09 2020-11-09 一种提高含氧化合物转化制低碳烯烃选择性的方法及其装置 WO2021089049A1 (zh)

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