WO2021089048A1 - 一种催化剂预烃池化的方法及其设备 - Google Patents

一种催化剂预烃池化的方法及其设备 Download PDF

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WO2021089048A1
WO2021089048A1 PCT/CN2020/127547 CN2020127547W WO2021089048A1 WO 2021089048 A1 WO2021089048 A1 WO 2021089048A1 CN 2020127547 W CN2020127547 W CN 2020127547W WO 2021089048 A1 WO2021089048 A1 WO 2021089048A1
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hydrocarbon
pooling
reactor
reaction
catalyst
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PCT/CN2020/127547
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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 US17/775,516 priority Critical patent/US20220401943A1/en
Priority to EP20883716.1A priority patent/EP4056269A4/en
Publication of WO2021089048A1 publication Critical patent/WO2021089048A1/zh

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    • 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
    • B01J38/48—Liquid treating or treating in liquid phase, e.g. dissolved or suspended
    • B01J38/50—Liquid treating or treating in liquid phase, e.g. dissolved or suspended using organic liquids
    • B01J38/56—Hydrocarbons
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/82—Phosphates
    • B01J29/84—Aluminophosphates containing other elements, e.g. metals, boron
    • B01J29/85—Silicoaluminophosphates [SAPO compounds]
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11, as exemplified by patent documents US3702886, GB1334243 and US3709979, respectively
    • 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
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/90—Regeneration or reactivation
    • 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
    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • 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
    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • B01J37/08—Heat treatment
    • 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
    • B01J38/02—Heat treatment
    • 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
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • 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
    • 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
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00017—Controlling the temperature
    • 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
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00548—Flow
    • 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
    • B01J2208/00—Processes carried out in the presence of solid particles; Reactors therefor
    • B01J2208/00008—Controlling the process
    • B01J2208/00548—Flow
    • B01J2208/00557—Flow controlling the residence time inside the reactor vessel

Definitions

  • the invention relates to a method and equipment for improving the selectivity of oxygen-containing compound conversion to low-carbon olefins, and more specifically, to a method and equipment for catalyst pre-hydrocarbon pooling (or pre-activation).
  • Low-carbon olefins defined here as ethylene and propylene
  • Oxygenated compounds include methanol, ethanol, dimethyl ether, methyl ethyl ether and so on. Due to the wide range of sources of oxygenates, coupled with the economics of the process for preparing low-carbon olefins, the process of converting oxygenates to olefins (OTO), especially the process of converting methanol to olefins (MTO), has suffered More and more attention.
  • 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.
  • 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 deposited catalyst ie, spent catalyst, referred to as spent catalyst
  • the regenerated catalyst referred to as the regenerant
  • 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 four or more carbon atoms at high temperature for example: 530 ⁇ 600°C
  • MTO reaction conditions for example: 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
  • Selective "hydrocarbon pool” active species to improve the activity and selectivity of regenerating agent oxygenates to produce low-carbon olefins, improve the "hydrocarbon pool” active species distribution and coke deposit distribution of the catalyst in the conversion reactor, and use the regenerant at the same time
  • 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 the regeneration 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 reaction to produce low-carbon olefins.
  • 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 and equipment for pre-hydrocarbon pooling (or pre-activation) of a catalyst.
  • the regenerated catalyst (referred to as regenerator) from the regenerator enters the pre-hydrocarbon pooling facility, where it contacts with the activation medium to generate pre-hydrocarbon Chemical reactions such as pooling, forming a "hydrocarbon pool" of active species and a certain amount of coke.
  • the regenerated catalyst referred to as the pre-hydrocarbon pooling catalyst or the pre-hydrocarbon pooling regenerator
  • the temperature of the regenerated catalyst after cooling is 200-630°C (preferably 300-600°C, more preferably 360-560°C).
  • 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 adopts (equal diameter or variable diameter) low-speed dense phase fluidized bed operation, and its superficial gas velocity (the flow rate of the fluidizing medium and the cross-section of the empty tower of the equipment) The ratio) is less than 0.5 m/s (preferably 0.0001 to 0.3 m/s, more preferably 0.001 to 0.2 m/s).
  • the activation medium that enters the pre-hydrocarbon pooling facility can be any one, two or more of oxygen-containing compound raw materials, reaction products, various hydrocarbons or other oxygen-containing compounds, or they mixture.
  • 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 carbon four, fractions above C5, fuel gas, etc.) from downstream olefin product separation devices (units), Two or more, or a mixture of them; the various hydrocarbons can also be any one, two or more of various pure component 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 integrated with the conversion reactor or connected to it through a conveying pipe.
  • an internal heat extractor or/and an external heat extractor (not shown in the figure) can be installed inside or/and outside to maintain the pre-hydrocarbon pool The heat balance of the chemical reaction system.
  • the reaction temperature of the pre-hydrocarbon pooling reactor can be controlled by adjusting the amount of catalyst returned to the pre-hydrocarbon pooling reactor, or/and the flow rate of the heating medium or/and the flow rate of the fluidizing medium or/and other parameters.
  • 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 can be various reactors used in industry, including any one, two or more of fluidized bed reactor, moving bed and fixed bed reactor, or a combination thereof; preferably Any one or two of various fluidized bed (including bubbling bed, turbulent bed, fast bed, etc.) reactor or riser reactor used in industry (including fluidized catalytic cracking unit, MTO unit, etc.) Or multiple, 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 unit (including fluidized catalytic cracking unit, MTO unit, etc.) is a mature industrial process, which uses various reactors, internal or external catalyst coolers (or heat extractors, including Upflow type, downflow type, back-mixed external heat extractor, etc.), steam (gas) stripper, catalyst distributor, steam (gas) gas distributor, etc., all of which can be used in the present invention.
  • the structure, combination type, operation and control process are very clear, and its operating conditions (such as feed temperature, reaction temperature, reaction pressure, contact time, agent-to-alcohol ratio (or agent-to-oil ratio, that is, the ratio of catalyst to raw material), and appearance
  • agent-to-alcohol ratio or agent-to-oil ratio, that is, the ratio of catalyst to raw material
  • the line speed, etc.] and the selection of the catalyst are also very clear, and they do not constitute a restriction on any specific implementation of the concept of the present invention.
  • 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 burning in the regenerator under conventional regeneration conditions, 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 catalyst pre-hydrocarbon pooling method and equipment of the present invention are equipped with pre-hydrocarbon pooling facilities to perform "pre-hydrocarbon pooling" treatment on the regenerated catalyst, so that the regenerant forms "hydrocarbons" before entering the oxygenated compound conversion reactor. Pond" active species and coke deposits to improve the distribution of active species and coke deposits in the "hydrocarbon pool” of the catalyst in the conversion reactor, thereby shortening or eliminating the reaction "induction period", and improving the catalysis of the regenerant oxygenated compounds to light olefins Activity and selectivity.
  • the temperature of the regenerant after the pre-hydrocarbon pooling is reduced, which breaks the heat balance of the reaction regeneration system, reduces the temperature at which the regenerator enters the conversion reactor, and eliminates the local overheating in the conversion reactor caused by the overheating of the regenerator.
  • the bed temperature distribution of the conversion reactor 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 improving reaction selectivity and further improving low-carbon
  • the olefin yield reduces the coking rate of the catalyst (that is, the carbon difference between the regenerating agent and the spent agent).
  • the temperature of the regenerator after the pre-hydrocarbon pooling is reduced, which 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 regenerant temperature is lowered, and the adjustment of the reaction temperature of the oxygenated compound conversion reaction, the catalyst circulation volume and other operating conditions is relatively independent and more flexible. It can be flexibly adjusted according to the market situation to achieve different product distributions.
  • the low-temperature regenerant after pre-hydrocarbon pooling can be used as a cold shock agent and directly enters the rapid separation facility (including inlet or outlet) to achieve rapid termination of the reaction, thereby inhibiting non-ideal reactions such as thermal polymerization of low-carbon olefins, and further It improves the yield of low-carbon olefins and reduces the coking rate of the catalyst (that is, the carbon difference between the regenerant and the spent catalyst). At the same time, it can also realize the pre-hydrocarbon pooling reaction of the reaction gas on the regenerating agent, forming active species and coke deposits in the hydrocarbon pool, shortening or eliminating the "induction period" of the reaction, and increasing the regenerating agent (low production rate). Carbon olefin reaction) activity and selectivity to further improve the yield of low-carbon olefins.
  • Figures 1 to 2 are schematic diagrams of a method for pre-hydrocarbon pooling (or pre-activation) of a catalyst and its equipment according to the present invention.
  • Figure 1 is a schematic diagram of a method for pre-hydrocarbon pooling (or pre-activation) of a catalyst of the present invention and its equipment structure (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 11 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 to maintain the pre-hydrocarbon pool The heat balance of the chemical reaction system.
  • the reaction temperature of the pre-hydrocarbon pooling reactor can be controlled by adjusting the amount of catalyst returned to the pre-hydrocarbon pooling reactor, or/and the flow rate of the heating medium or/and the flow rate of the fluidizing medium or/and other parameters.
  • 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).
  • Figure 2 is a schematic diagram (cocurrent contact) of a catalyst pre-hydrocarbon pooling (or pre-activation) method and equipment of the present invention.
  • 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 will 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 to maintain the pre-hydrocarbon pool The heat balance of the chemical reaction system.
  • the reaction temperature of the pre-hydrocarbon pooling reactor can be controlled by adjusting the amount of catalyst returned to the pre-hydrocarbon pooling reactor, or/and the flow rate of the heating medium or/and the flow rate of the fluidizing medium or/and other parameters.
  • 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 ⁇ 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 structure shown in Figure 1 is adopted, the methanol conversion reactor and the regenerator are both under normal operating conditions, the catalyst is SAPO-34, the reaction gas is the activation medium, and the pre-hydrocarbon pooling reactor
  • the main operating conditions are as follows: the apparent linear velocity is 0.1-0.2m/s, the reaction temperature is -460°C, and the contact time is 60-80 minutes.
  • the active species in the hydrocarbon pool on the regenerated regenerant are basically restored, and a reasonable level of coke deposition is formed.
  • the oxygenate raw material is methanol
  • the structure shown in Figure 1 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, and the pre-hydrocarbon pooling reactor
  • 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

公开了一种催化剂预烃池化的方法和预烃池化设备，属于低碳烯烃制备技术领域。再生催化剂进入预烃池化反应器，与活化介质发生预烃池化反应，形成"烃池"活性物种。通过设置预烃池化设施，对再生催化剂进行"预烃池化"处理，使再生剂在进入含氧化合物转化反应器前形成"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，从而缩短或消除反应"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

一种催化剂预烃池化的方法及其设备 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设备，更具体地来说，涉及一种催化剂预烃池化(或称预活化)的方法及其设备。
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。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。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)，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。
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)混合后返回到转化反应区继续反应。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。
本发明提供一种催化剂预烃池化(或称预活化)的方法及其设备，来自再生器的再生催化剂(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。离开预烃池化设施后的再生催化剂(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。
所述的冷却后再生催化剂温度200～630℃(优选300～600℃，更优选360～560℃)。
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(等直径或变直径的)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(反应器)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、C5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纯组分烯烃、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(反应器)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。
为更好地控制所述的预烃池化反应器的反应温度，可在其内部或/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返回预烃池化反应器的催化剂量，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。
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；优选工业上(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(包括流化催化裂化装置、MTO装置等)为成熟工业过程，其使用的各种反应器、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设备，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℃)。
本发明的方法及其设备可应用于工业上(包括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、本发明的催化剂预烃池化方法及其设备，通过设置预烃池化设施，对再生催化剂进行“预烃池化”处理，使再生剂在进入含氧化合物转化反应器前形成“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，从而缩短或消除反应“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、预烃池化后的再生剂温度降低，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、预烃池化后的再生剂温度降低，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、预烃池化后的再生剂温度降低，含氧化合物转化反应的反应温度、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。
5、预烃池化后的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，缩短或消除反应的“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1～2：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。
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为本发明的一种催化剂预烃池化(或称预活化)的方法及其设备结构示意图(逆流接触)。
如附图1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11、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，可在其内部或/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返回预烃池化反应器的催化剂量，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。
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2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2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，可在其内部或/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返回预烃池化反应器的催化剂量，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。
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)。
实施例
实例1
对于实例1，含氧化合物原料为甲醇，采用图1所示的结构，甲醇转化反应器和再生器均采用常规操作条件，催化剂为SAPO-34,反应气为活化介质，预烃池化反应器主要操作条件如下：表观线速0.1～0.2m/s，反应温度为～460℃，接触时间60～80分钟。在此，再生后的再生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，采用本发明的方法及其设备与现有MTO技术相比，本发明的低碳烯烃(C 2 ＝+C 3 ＝)的选择性提高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  ℃   460～500
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1所示的结构，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，预烃池化反应器主要操作条件如下：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 (12)

  1. 一种催化剂预烃池化的方法，其特征在于，再生催化剂进入预烃池化反应器，与活化介质发生预烃池化反应，形成“烃池”活性物种。
  2. 按照权利要求1所述的方法，其特征在于，所述的预烃池化反应器的主要操作条件为反应温度为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，所述的预烃池化反应器的主要操作条件为反应温度为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，所述的预烃池化反应器的主要操作条件为反应温度为400～530℃、接触时间10～150分钟。
  5. 按照权利要求1-4任一项所述的方法，其特征在于，所述的预烃池化反应器为低速密相流化床操作，其表观气速小于0.5m/s。
  6. 按照权利要求1-4任一项所述的方法，其特征在于，所述的预烃池化反应器为低速密相流化床操作，其表观气速为0.0001～0.3m/s。
  7. 按照权利要求1-4任一项所述的方法，其特征在于，所述的预烃池化反应器为低速密相流化床操作，其表观气速为0.001～0.2m/s。
  8. 按照权利要求1-7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的产品，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，所述的预烃池化反应器设置内取热器或/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. 一种催化剂预烃池化设备，其特征在于，所述预烃池化反应器包括再生剂进口、再生剂出口、活化介质进口和活化介质出口。
  11. 按照权利要求10的设备，其特征在于，所述预烃池化反应器还包括内取热器或/和外取热器。
  12. 按照权利要求10或11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。
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