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WO2022048440A1 - 一种石油烃催化转化制丙烯和乙烯的方法及其装置 - Google Patents

一种石油烃催化转化制丙烯和乙烯的方法及其装置 Download PDF

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WO2022048440A1
WO2022048440A1 PCT/CN2021/113118 CN2021113118W WO2022048440A1 WO 2022048440 A1 WO2022048440 A1 WO 2022048440A1 CN 2021113118 W CN2021113118 W CN 2021113118W WO 2022048440 A1 WO2022048440 A1 WO 2022048440A1
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catalyst
reaction
zone
regeneration
reaction zone
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PCT/CN2021/113118
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English (en)
French (fr)
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石宝珍
李荻
郭江伟
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青岛京润石化设计研究院有限公司
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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/08—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by splitting-off an aliphatic or cycloaliphatic part from the molecule
    • 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
    • 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
    • B01J8/0015—Feeding of the particles in the reactor; Evacuation of the particles out of the reactor
    • B01J8/0035—Periodical feeding or evacuation
    • 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
    • B01J8/005—Separating solid material from the gas/liquid stream
    • B01J8/007—Separating solid material from the gas/liquid stream by sedimentation
    • 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
    • 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
    • B01J8/1809—Controlling processes
    • 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
    • B01J8/1845—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with particles moving upwards while fluidised
    • 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
    • B01J8/1872—Details of the fluidised bed 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
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/04—Ethylene
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
    • C07C11/02—Alkenes
    • C07C11/06—Propene
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/02—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by cracking a single hydrocarbon or a mixture of individually defined hydrocarbons or a normally gaseous hydrocarbon fraction
    • C07C4/06—Catalytic processes
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/16—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "moving bed" method
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts

Definitions

  • the invention belongs to the technical field of catalytic conversion of petroleum hydrocarbons, and in particular relates to a method for catalytic conversion of petroleum hydrocarbons to propylene and ethylene.
  • the invention also provides a device for realizing the above method.
  • Propylene is one of the most important petrochemical raw materials. 70% of the propylene is produced from petroleum hydrocarbons by the tubular furnace cracking method, and the other 30% is provided by the catalytic cracking process. With reference to the operation and design experience of conventional heavy oil catalytic cracking reaction-regeneration system, domestic and foreign researchers have developed a series of process technologies for heavy oil catalytic cracking to produce propylene.
  • TMP technology based on two-stage riser catalytic cracking technology.
  • the technology uses heavy oil as raw material, utilizes the process characteristics of two-stage riser catalytic cracking process, staged reaction, catalyst relay and large agent-oil ratio, optimizes the combination of feeding methods for different properties of reaction materials, and controls the appropriate materials for different materials.
  • the reaction conditions to achieve the purpose of improving the propylene yield;
  • the form of the layer reactor is different in that the DCC-PLUS process is additionally equipped with a light gasoline and C4 return refining riser, and the stream after the light gasoline and C4 reaction is introduced into the fluidized bed reactor.
  • Both DCC and DCC-PLUS divide the feedstock reaction into riser and fluidized bed reaction; but both DCC and DCC-PLUS control the reaction in the fluidized bed zone by the amount of regenerated catalyst entering the feedstock riser reaction zone
  • the temperature that is, the whole reaction process, is controlled according to the conditions of the fluidized bed catalytic cracking reaction zone, which inevitably causes the catalytic cracking conditions of the riser reaction zone, that is, the heavy oil reaction zone, to deviate from the ideal catalytic cracking reaction conditions of feedstock oil, especially the increase in thermal reaction;
  • the space velocity of the fluidized bed reaction zone can only be controlled by the change of the catalyst material level in the fluidized bed; due to the requirements of catalyst entrainment and gas-solid separation, the fluidized
  • the prior art focuses on the production of propylene, which is divided into two categories.
  • the first category is a riser plus fluidized bed series reaction
  • the second category is a double riser parallel reaction.
  • the researchers believe that the propylene in the heavy oil catalytic cracking reaction is indirectly generated by the secondary cracking of the gasoline fraction generated by the primary cracking of heavy hydrocarbons, and the C5-C8 olefins in the gasoline fraction are the main precursors of propylene.
  • the prior art has many common characteristics, and in operation, it adopts higher reaction temperature, agent-oil ratio and steam injection amount than conventional FCC process, so as to improve the depth of cracking reaction and the selectivity of propylene.
  • the object of the present invention is to provide a method for the catalytic conversion of petroleum hydrocarbons to produce propylene and ethylene, which can significantly reduce the yields of by-products such as dry gas and coke while increasing the yields of propylene and ethylene.
  • Another object of the present invention is to provide a device for realizing the above-mentioned method for catalytic conversion of petroleum hydrocarbons to propylene and ethylene.
  • the catalytic conversion of petroleum hydrocarbons to propylene and ethylene is also a strong endothermic and coke-producing reaction. No matter whether it is a catalyst, a reactant, or a product, the reaction process changes significantly and is complex; control the actual catalyst conditions in the reaction process in the reactor, Temperature conditions and space velocity conditions are extremely important to control the chemical reaction process. Different conditions in the reactor actually form different chemical reactions and obtain different products, especially when propylene is used as the product. Targeted technology is required.
  • the goal of catalytic cracking of petroleum hydrocarbon reaction feedstocks to produce propylene and ethylene is to maximize propylene and ethylene.
  • a method for preparing propylene and ethylene by catalytic conversion of petroleum hydrocarbons wherein the reaction process of the method for preparing propylene and ethylene through catalytic conversion of petroleum hydrocarbons is as follows: the heavy component macromolecules are firstly transported in a gas-solid fluidized reaction condition, that is, in a riser reactor, the reaction process is relatively high.
  • the catalytic cracking conversion is carried out under the condition of active regeneration catalyst environment and lower temperature to form C4-C12-based intermediate molecules, and then in the form of gas-solid fast fluidized bed that increases catalyst density and reduces catalyst space velocity, carbon-containing catalyst environment and Catalytic cracking conversion is carried out under high severity conditions to produce propylene and ethylene;
  • the petroleum hydrocarbon catalytic conversion method is carried out in a two-stage reactor connected in series, and the lower part is the fluidized form of gas-solid transport bed, that is, the fluidized form of riser
  • the upper part is in the form of gas-solid rapid fluidization.
  • the upper and lower parts of the reactor are independently controlled two-stage heating and two-way catalyst circulation from the regenerator.
  • the reactor is divided into two upper and lower reaction zones, namely the riser reaction zone and the fast fluidized bed reaction zone.
  • the lower part of the reactor is the riser reaction zone, or the low temperature reaction zone.
  • the heavy components of macromolecules The catalytic cracking reaction to convert C5-C12 intermediate components is carried out under the environment of regenerated catalyst and lower reaction temperature to provide intermediate raw materials for the production of propylene and ethylene; above the reactor is a fast fluidized bed reaction zone, or high temperature reaction zone , the catalyst entering from the upper entry point of the reactor, that is, the upper catalyst provides heat to further increase the reaction temperature of the fast fluidized bed reaction zone, forming a high temperature reaction zone, and the intermediate components mainly C5 ⁇ C12 are at higher temperatures and larger reagents.
  • the catalytic cracking reaction is carried out under the harsher conditions of oil ratio and lower weight hourly space velocity, so that the petroleum hydrocarbon reaction raw material is converted into propylene; the high boiling point petroleum hydrocarbon reaction raw material is realized in two-stage heating, two-stage catalyst supply, and two-stage control.
  • Selective reaction is carried out in the two-zone reactor, and the reaction method of gradually increasing the reaction temperature adapts to the change of the molecular weight of the reactant and the change of the requirements of the reaction conditions, so as to improve the efficiency of producing propylene and ethylene and the selectivity of the target product;
  • the reaction raw materials are sequentially reacted in the low-temperature riser reaction zone and the high-temperature fast fluidized-bed reaction zone to realize the low-temperature catalytic cracking reaction and the high-temperature propylene and ethylene catalytic cracking reaction.
  • the catalyst of the regenerator enters the riser reaction zone and the fast-flow reaction zone respectively.
  • the gradual heating of the reaction process is realized by means of graded heat supply; the petroleum hydrocarbon reaction raw materials are firstly subjected to the catalytic cracking reaction of high-boiling heavy components in the low-temperature reaction zone, and the catalytic cracking conversion and decarbonization of heavy components and macromolecules are initially completed.
  • reaction process includes the following steps:
  • the reaction raw material after the reaction raw material is atomized by steam, it first enters the riser reaction zone at the lower part of the reactor, and carries out catalytic cracking reaction under the lower catalyst environment introduced from the regenerator through the lower regeneration standpipe; the riser reaction zone That is, the low-temperature reaction zone is carried out under the conditions favorable for the catalytic conversion of high-boiling heavy components, the catalyst is a regenerated catalyst, the reaction temperature is 515°C to 620°C, and the reaction time is 0.5 to 1.5s. ⁇ 550°C, the reaction time is 1.0 ⁇ 2.5s; the actual reaction temperature and agent-oil ratio in the riser reaction zone are independently controlled by the amount of lower catalyst entering the low-temperature reaction zone;
  • the reaction temperature of the fast fluidized bed reaction zone is 530 °C ⁇ 720 °C
  • the reaction time is 0.5 ⁇ 4.0s
  • the reaction temperature is 530 °C ⁇ 580 °C when the goal of producing propylene is 530 °C ⁇ 580 °C
  • the reaction time is 1.0 ⁇ 5.0s
  • the absolute pressure of the reaction pressure is 0.20MPa ⁇ 0.40MPa
  • the actual reaction temperature and agent-oil ratio in the fast fluidized bed reaction zone are controlled by the amount of catalyst entering the fast fluidized bed reaction zone;
  • the upper catalyst enters the reactor from the upper catalyst inlet above the reaction temperature control point of the riser reaction zone, or directly enters the fast fluidized bed reaction zone, or the introduced upper catalyst first enters the reactor at the outlet of the riser reaction zone and then enters the reactor. It is transported to the fast fluidized bed reaction zone, and further provides heat to the fast fluidized bed reaction zone, increases the temperature of the stream and the ratio of agent to oil, and further conducts the catalytic cracking reaction;
  • the reaction raw materials are preferably vacuum wax oil, atmospheric residual oil, coking wax oil, deasphalted oil, hydrogenated wax oil (after the hydrotreatment).
  • mixed C4 also known as C4
  • C4 first enters the riser reaction zone under the reaction raw material inlet to react, and then together with the reaction raw materials in the riser reaction zone and rapid fluidization.
  • bed reaction zone reaction; or C4 enters the reactor above the upper catalyst inlet or directly enters the fast fluidized bed reaction zone for reaction.
  • the method for the above-mentioned catalytic conversion of petroleum hydrocarbons to produce propylene and ethylene further, from the stripping section to lead a part of the catalyst to be generated to return to the reactor, or to extract the partially reacted catalyst from the reactor above the fast fluidized bed reaction zone from the return catalyst.
  • the inlet is returned to the reactor; the catalyst to be produced or the reacted catalyst is returned to the reactor from the catalyst return conveying pipe and the return catalyst inlet, and the return catalyst inlet is set at the outlet of the riser reaction zone or at the bottom of the fast fluidized bed reaction zone; the to-be-generated catalyst Or the amount of the catalyst returned after the reaction controls the catalyst storage and the weight hourly space velocity of the reaction catalyst in the fast fluidized bed reaction zone, and the catalyst return slide valve controls the amount of the catalyst to be produced or the catalyst returned after the reaction.
  • the riser reaction zone adopts the fluidized form of gas-solid pneumatic conveying, and the average gas flow rate is 5.0m/s ⁇ 20m/s; the diameter of the fast fluidized bed reaction zone is larger than that of the lifter
  • the tube reaction zone adopts fast gas-solid fluidized bed conditions, the average gas velocity is 1.8m/s ⁇ 5.0m/s, and the weight hourly space velocity of the catalyst is 10(1/H) ⁇ 35(1/H).
  • two upper and lower regeneration zones are arranged below the dilute phase zone of the regenerator of the regenerator, that is, the lower regeneration zone and the upper regeneration zone.
  • the upper regeneration zone is the dense-phase fluidized bed regeneration zone;
  • the lower regeneration zone is the first regeneration zone,
  • the upper regeneration zone is the second regeneration zone, and the catalyst to be produced from the stripping section enters the lower regeneration zone, and the upper regeneration zone
  • the regeneration zone is provided with the lower catalyst outlet and the upper catalyst outlet at the same time; the lower catalyst and the upper catalyst are supplied to the reactor from the upper regeneration zone.
  • the lower catalyst temperature is 660-740°C, and the carbon content of the catalyst is lower than 0.40%; the upper catalyst temperature is 680-730°C, and the carbon content of the catalyst is lower than 0.1%. Both the lower catalyst and the upper catalyst are regenerated catalysts.
  • two upper and lower regeneration zones are set under the dilute phase of the regenerator of the regenerator, that is, the lower regeneration zone and the upper regeneration zone, when implementing, the upper regeneration zone and the lower regeneration zone.
  • Both are dense-phase fluidized bed regeneration zones; the upper regeneration zone is the first regeneration zone, the lower regeneration zone is the second regeneration zone, the catalyst to be produced from the stripping section enters the upper regeneration zone, and the lower catalyst outlet is set in the lower regeneration zone.
  • the upper regeneration zone is provided with an upper catalyst outlet, and the lower catalyst provided by the lower regeneration standpipe (drawn from the lower catalyst outlet) to the riser reaction zone through the lower regeneration zone is a regeneration catalyst, and the upper regeneration riser (drawn from the upper catalyst outlet) is a regeneration catalyst.
  • the upper catalyst provided to the reactor is a moderately carbon-containing semi-regenerated catalyst.
  • the lower catalyst temperature is 660-740°C, and the carbon content of the catalyst is lower than 0.10%; the upper catalyst temperature is 680-730°C, and the carbon content of the catalyst is lower than 0.5%.
  • the regenerator adopts In the conventional regeneration method of coke drum, the lower regeneration zone is a fast fluidized bed, and the upper regeneration zone is a dense-phase fluidized bed, and the catalyst is supplied to the reactor from the dense-phase fluidized bed;
  • the regenerator adopts the upper and lower double density In the form of phase fluidized bed regeneration, the amount of catalyst or the amount of charred oxygen or air entering the two fluidized bed regeneration zones is adjusted and controlled, so that the catalyst temperature and carbon content in the two regeneration zones are different, and different dense phase
  • the fluidized bed provides catalysts with different carbon contents and temperatures to the reactor;
  • the catalyst in the upper regeneration zone is a “semi-regenerated catalyst” with moderate carbon content
  • the catalyst in the lower regeneration zone is a regenerated catalyst.
  • the catalyst is supplied from the lower regeneration zone to the riser reaction zone.
  • the riser reaction zone of the method of the present invention is mainly used for the catalytic conversion of macromolecular heavy components; the fast fluidized bed reaction zone is mainly used for the reaction of cracking to produce propylene, ethylene and aromatic hydrocarbons.
  • the present invention also provides a device for catalytic conversion of petroleum hydrocarbons to produce propylene and ethylene, specifically, the device is a device for realizing the method for catalytic conversion of petroleum hydrocarbons to produce propylene and ethylene described in the present invention.
  • the device includes: a reactor, a regenerator, a settler and a stripping section, a lower regeneration standpipe and a lower regeneration slide valve, an upper regeneration standpipe and an upper catalyst slide valve, and a regeneration standpipe and the spool valve of the catalyst to be produced; the regenerator and the settler are arranged in parallel;
  • the reactor is arranged in the form of an upper and lower zoned reactor with an upper and lower two-way catalyst circulation and two heat supply, and is divided into two sections or two zones connected in series, including a lower riser reaction zone and an upper fast fluidized bed reaction zone,
  • the riser reaction zone is in the form of gas-solid pneumatic conveying fluidization with high gas flow rate
  • the fast fluidized bed reaction zone is in the form of fast fluidized bed with reduced gas flow rate
  • the diameter of the fast fluidized bed reaction zone is larger than that of the riser reaction zone;
  • the reactor is provided with two paths of catalyst circulation from the regenerator, and supplies heat to the reactor twice, up and down, the riser reaction zone is used for the low-temperature catalytic cracking reaction of raw oil macromolecules, and the fast fluidized bed reaction zone is used for high-temperature production of propylene and propylene.
  • Ethylene catalytic cracking reaction a reaction raw material inlet is set at the lower part of the riser reaction zone;
  • the lower catalyst inlet at the lower part of the reaction zone of the riser is arranged below the inlet of the reaction raw material, and the lower catalyst inlet is communicated with the lower catalyst outlet of the regenerator through the lower regeneration standpipe; at the bottom of the fast fluidized bed reaction zone or the riser
  • the outlet of the reaction zone is provided with an upper catalyst inlet, and the upper catalyst inlet is communicated with the upper catalyst outlet of the regenerator through the upper regeneration standpipe;
  • the lower regeneration standpipe and the upper regeneration standpipe are respectively provided with a lower regeneration slide valve and an upper catalyst slide valve;
  • the stripping section is communicated with the regenerator through the outlet of the catalyst to be grown, the standpipe to be grown and the inlet of the catalyst to be grown, and a slide valve of the catalyst to be grown is arranged on the standpipe to be grown.
  • the device for the above-mentioned catalytic conversion of petroleum hydrocarbons to produce propylene and ethylene preferably, below the reaction raw material inlet, or above the upper catalyst inlet or at the bottom of the fast fluidized bed reaction zone, set the C4 inlet, or/and above the upper catalyst inlet or fast
  • the liquid light hydrocarbon inlet is arranged at the bottom of the fluidized bed reaction zone or the outlet of the riser reaction zone.
  • a return catalyst inlet is set at the outlet of the riser reaction zone at the bottom of the fast fluidized bed reaction zone or above the temperature detection control point of the riser reaction zone.
  • the stripping section is provided with a return catalyst outlet, and the return catalyst inlet is communicated with the return catalyst outlet through the catalyst return conveying pipe, so as to provide the catalyst to be produced to the fast fluidized bed reaction zone through the stripping section; in the catalyst return conveying pipe Set the catalyst return spool valve on.
  • a return catalyst inlet is set at the outlet of the riser reaction zone at the bottom of the fast fluidized bed reaction zone or above the temperature detection control point of the riser reaction zone.
  • the return catalyst outlet is arranged above the reaction zone of the riser pipe of the reactor, and the catalyst return conveying pipe is communicated with the return catalyst inlet, so that the reacted catalyst can be refluxed in the reactor, and the reacted catalyst is provided to the fast fluidized bed reaction zone, so as to realize the reaction of the catalyst.
  • Control of catalyst space velocity in fast fluidized bed reaction zone; catalyst return slide valve is set on catalyst return conveying pipe.
  • two upper and lower regeneration zones namely a lower regeneration zone and an upper regeneration zone, are arranged below the dilute phase zone of the regenerator of the regenerator;
  • the lower catalyst outlet and the upper catalyst outlet are both arranged in the upper regeneration zone, so as to provide the regenerated catalyst to the reactor through the upper regeneration zone, and the inlet of the waiting agent is arranged in the lower regeneration zone;
  • the upper regeneration zone is a dense-phase fluidized bed regeneration zone
  • the lower catalyst outlet is set in the lower regeneration zone
  • the regeneration catalyst is provided to the riser reaction zone through the lower regeneration zone, that is, the lower re-agent
  • the upper catalyst outlet is It is arranged in the upper regeneration zone to realize the provision of catalyst, namely the upper catalyst, to the fast fluidized bed reaction zone through the fluidized bed regeneration zone of the upper regeneration zone.
  • the steam in the low temperature reaction zone accounts for 5-50% of the mass ratio of the reaction raw materials
  • the steam in the high-temperature reaction zone accounts for 15-50% of the mass ratio of the reaction raw materials
  • a steam generator can be installed after the product outlet of the reaction settler, and the heat of the high-temperature product stream can be used to generate steam to realize the cooling or rapid cooling of the product stream, which is controlled by the engineering design unit of the steam generator.
  • the invention proposes a method for producing propylene and ethylene by catalytic conversion with gradual temperature increase, two-stage temperature gradient and independent control of weight hourly space velocity.
  • the catalytic cracking process of heavy oil can be regarded as a parallel sequential reaction.
  • the heavy component macromolecules are first cracked to generate medium molecular (C5-C12) products, and a lower cracking temperature can highlight the catalytic cracking reaction.
  • the suitable temperature for the catalytic cracking reaction is It is 490°C ⁇ 530°C; some gasoline and diesel oil continue to crack into C3 ⁇ C4 at 540°C ⁇ 580°C.
  • the present invention follows this reaction law, and sets up two-stage temperature gradients in series with gradually increasing temperature: a low temperature zone and a high temperature zone; the present invention realizes the adjustment and control of the catalyst-oil ratio and the catalyst space velocity in the reaction process by returning the catalyst to the reactor. Does not affect thermal balance. Under the premise of lower energy consumption, the invention reduces the yield of low-value target products, such as coke and dry gas, and improves the yield of high-value target products, propylene and ethylene.
  • low-value target products such as coke and dry gas
  • the method of the invention controls the ratio of agent to oil, space velocity and temperature by grading in the reaction process, especially realizes that as the reaction proceeds, the ratio of agent to oil and temperature are gradually increased, the weight hourly space velocity is decreased, and the reaction severity is gradually increased, so that the reaction conditions and the reaction conditions are gradually increased.
  • the petroleum hydrocarbon molecules gradually become smaller in the cracking process of the petroleum hydrocarbon reaction raw material and are compatible with the reaction chemical conditions that the required reaction severity gradually increases; the invention also well optimizes the joint conversion effect of the heavy components and the light hydrocarbon raw materials with different properties, avoiding small Excessive cracking of molecular light hydrocarbons not only ensures the conditions for recombination and cracking, but also ensures the conditions for cracking light hydrocarbons; the method improves the efficiency and increases the selectivity of the target product.
  • Fig. 1 is a process schematic diagram of Embodiment 1 of the present invention.
  • FIG. 2 is a schematic diagram of the second embodiment of the present invention.
  • FIG. 3 is a schematic diagram of the third embodiment of the present invention.
  • FIG. 4 is a schematic diagram of a fourth process according to Embodiment 4 of the present invention.
  • FIG. 5 is a schematic diagram of a fifth embodiment of the present invention.
  • R10 reactor R11 catalyst lift gas; R11A catalyst lift gas inlet, R12 reaction raw material, R12A reaction raw material inlet, R13 raw material atomization steam, R14A lower catalyst inlet (the catalyst inlet of the riser reaction zone); R15 reaction supplementary steam, R15A first Two-reaction supplementary steam, R17 riser reaction zone (or low temperature reaction zone), R18 fast fluidized bed reaction zone (or high temperature reaction zone); R24A upper catalyst inlet (fast fluidized bed reaction zone catalyst inlet); R34A returns to the catalyst inlet ; R32 liquid light hydrocarbon; R32A liquid light hydrocarbon inlet; R22A C4 inlet; C4 carbon four-component petroleum hydrocarbon;
  • S10 stripping section S11 stripping component; S12 standpipe; S12A ready catalyst outlet (unborn agent outlet); VS12 unborn catalyst spool valve (unborn agent spool valve); S13 stripping steam; S14 catalyst return Delivery pipe; S14A return catalyst outlet; VS14 catalyst return spool valve;
  • TI temperature detection signal TC temperature control signal
  • DPC differential pressure control signal TI temperature detection signal
  • Embodiment 1 is a diagrammatic representation of Embodiment 1:
  • a reactor R10 As shown in Figure 1, a reactor R10, a regenerator G10, a settler D10 and a stripping section S10 are provided, and petroleum hydrocarbons are used as raw materials.
  • the regenerator G10 and the settler D10 are arranged side by side, the outlet of the reactor R10 is connected to the settling cyclone D11 in the settler D10, the stripping section S10 is set below the settler D10, and the stripping section S10 is provided with a stripping member S11, which is directed to the settling cyclone D11.
  • the stripping section S10 introduces stripping steam S13 to realize catalyst stripping;
  • the reactor R10 is set as an upper and lower zoned reactor with upper and lower two-way catalyst circulation and two heat supply, including the lower riser reaction zone R17 and the upper fast fluidized bed reaction zone R18.
  • the diameter of the fast fluidized bed reaction zone R18 is larger than The riser reaction zone R17, the riser reaction zone R17 is used for low-temperature catalytic cracking reaction, and the fast fluidized bed reaction zone R18 is used for high-temperature propylene production and ethylene cracking reaction; the lower catalyst inlet R14A at the lower part of the riser reaction zone R17 passes through the lower regeneration stand.
  • the pipe G14 is communicated with the lower catalyst outlet G14A of the regenerator G10, and the upper catalyst inlet R24A at the lower part of the fast fluidized bed reaction zone R18 is communicated with the upper catalyst outlet G24A of the regenerator G10 through the upper regeneration standpipe G24; at the lower part of the reactor R10
  • the reaction raw material inlet R12A is set to introduce the reaction raw material R12 and the raw material atomization steam R13, the catalyst lifting gas inlet R11A is set at the bottom of the reactor R10 to introduce the catalyst lifting gas R11; the top of the settler D10 is provided with a settler reaction product outlet D12A, to The reaction product D12 is drawn;
  • the regenerator G10 adopts two-zone regeneration arranged up and down in series.
  • Two upper and lower regeneration zones namely the lower regeneration zone G13 and the upper regeneration zone G12, are set below the dilute phase zone G15 of the regenerator.
  • the lower regeneration zone G13 and the upper regeneration zone G12 All adopt the form of dense phase fluidized bed, the lower regeneration zone G13 is the second regeneration zone, and the upper regeneration zone G12 is the first regeneration zone.
  • the lower catalyst outlet G14A is arranged in the lower regeneration zone G13, and the upper catalyst outlet G24A is arranged in the upper regeneration zone G12, so that the upper catalyst provided by the upper regeneration zone G12 to the fast fluidized bed reaction zone R18 is a semi-regenerated catalyst, and the upper catalyst through the lower regeneration zone
  • the lower catalyst provided by G13 to the riser reaction zone R17 is a regenerated catalyst;
  • the lower part of the stripping section S10 is connected to the upper regeneration zone G12 of the regenerator G10 through the to-be-grown riser inlet G12A through the to-be-grown standpipe S12.
  • the upper regeneration zone G12 of the regenerator G10 is connected to the lower catalyst inlet R14A at the lower part of the riser reaction zone R17 from the lower catalyst outlet G14A through the lower regeneration riser G14, and the lower regeneration riser G14 is provided with a lower regeneration slide valve VG14; the upper regeneration zone G12
  • the upper catalyst outlet G24A is connected to the upper catalyst inlet R24A at the lower part of the fast fluidized bed reaction zone R18 or the outlet of the riser reaction zone R17 through the upper regeneration standpipe G24, and the upper regeneration standpipe G24 is provided with an upper catalyst slide valve VG24;
  • a regeneration cyclone G16 is installed in the regenerator dilute phase zone G15 of the regenerator G10. After the regenerator is scorched, the flue gas G17 is discharged from the flue gas outlet G17A at the top of the regenerator G10, and the catalyst regeneration gas G11 is regenerated by the bottom of the regenerator G10.
  • the gas inlet G11A is introduced; the air G18 is introduced into the regenerator G10 from the lower part of the upper regeneration zone G12;
  • Reaction make-up steam R15 is introduced into the riser reaction zone R17.
  • the catalyst of the regenerator G10 enters the riser reaction zone R17 and the fast fluidized bed reaction zone R18 respectively, and realizes the gradual heating of the reaction process by means of graded heat supply;
  • the specific implementation process includes the following steps:
  • the preheated reaction raw material R12 is atomized by the raw material atomizing steam R13 and then first enters the riser reaction zone R17 at the lower part of the reactor R10, and the lower catalyst from the lower regeneration riser G14 from the upper regeneration zone G12 of the regenerator from the bottom
  • the catalyst inlet R14A enters the riser reaction zone R17, and is transported upward under the action of the catalyst lifting gas R11 to contact with the raw material.
  • the reaction raw material R12 undergoes catalytic cracking conversion under mild conditions under the catalyst environment to form intermediate products mainly composed of C5-C12.
  • the lower catalyst temperature is preferably 660 °C ⁇ 740 °C, and the carbon content of the catalyst is less than 0.10%;
  • the reaction temperature of the riser reaction zone R17 is 515 °C ⁇ 620 °C, and the reaction time is 0.5 ⁇ 1.5s; when propylene is produced, the riser reacts
  • the reaction temperature of zone R17 is 515°C ⁇ 550°C, and the reaction time is 1.0s ⁇ 2.5s;
  • the riser reaction zone R17 adopts the fluidized form of gas-solid pneumatic conveying, the average gas velocity is 5.0m/s ⁇ 20m/s, the catalyst weight hourly space velocity 10(1/H) ⁇ 35(1/H);
  • the product generated in the riser reaction zone R17 and the catalyst flow upward together into the fast fluidized bed reaction zone R18, and the new catalyst is introduced from the regenerator G10 through the upper regeneration riser G24.
  • the upper catalyst enters the rapid fluidized bed reaction zone R18, and is transported to the rapid fluidized bed reaction zone R18 by the stream from the riser reaction zone R17, and the new catalyst, namely the upper catalyst, further provides heat to the rapid fluidized bed reaction zone R18, improving the flow rate.
  • the temperature and the ratio of catalyst to oil form more severe high-temperature cracking reaction conditions.
  • the products from the riser reaction zone continue to undergo a combined reaction of catalytic cracking and thermal cracking, and low-carbon small molecule products such as propylene and ethylene;
  • the preferred temperature of the upper catalyst is 680 ⁇ 730°C, the carbon content of the catalyst is less than 0.5%
  • the reaction temperature of the rapid fluidized bed reaction zone R18 is 530°C ⁇ 720°C, and the reaction time is 0.5 ⁇ 4.0s, when propylene is produced more, the rapid fluidized bed reaction zone R18
  • the reaction temperature is 530°C ⁇ 580°C
  • the reaction time is 1.0 ⁇ 5.0s
  • the absolute pressure of the reaction pressure is 0.20MPa ⁇ 0.40Mpa
  • the actual reaction temperature is controlled by the amount of catalyst entering the rapid fluidized bed reaction zone R18;
  • the rapid fluidized bed reaction zone The diameter of R18 is larger than that of R17 in the riser reaction zone, and the conditions of fast gas-solid fluidized bed are adopted, and the average gas velocity is 1.8m/s ⁇ 5.0m/s
  • the reacted stream enters the settler D10 for gas-solid separation to obtain the reaction product D12, which is sent out from the settler reaction product outlet D12A and enters the follow-up treatment part; after the reacted catalyst is separated by the sedimentation cyclone D11, the After the section S10 is stripped, it enters the upper regeneration zone G12 of the regenerator G10 through the standpipe S12 and the inlet G12A for the growth agent, and is regenerated and recycled.
  • reaction product D12 leaves the catalytic conversion device shown in FIG. 1, product fractionation is performed; fractionation and the like are well known to engineers and technicians.
  • the raw material for the reaction is vacuum wax oil, the density is 0.89, the hydrogen content is 13.2% by weight, the residual carbon is 4.0%, and the saturated hydrocarbon is 60%; the preheating temperature of the raw material oil is 220°C;
  • the reaction device is arranged in parallel with the settler and the regenerator, and the regenerator adopts the serial regeneration form of the fast fluidized bed and the dense phase fluidized bed of the coke tank;
  • reaction temperature is controlled by TIC at 530°C, the average gas flow rate is 6.5m/s, and the reaction time is 1.7s (seconds);
  • catalyst transport gas is steam, and the quantity is 3% of the reaction raw materials, and the raw material atomization steam is the raw material
  • the supplementary steam is 20% of the raw material, the catalyst entering from the lower catalyst inlet is the regenerated catalyst, the carbon content is 0.02%, and the lower catalyst temperature is 680°C;
  • the reaction conditions of the fast fluidized bed reaction zone the catalyst entering from the upper catalyst inlet is a semi-regenerated catalyst, the carbon content is 0.1%, the temperature is 705°C, the reaction temperature is controlled at 560°C, the average gas flow rate is 3.5m/s, and the reaction time 1.5 seconds;
  • the reaction process is:
  • the raw material After the raw material is atomized with steam, it enters the riser reaction zone, and the catalytic cracking conversion of heavy oil is carried out under the heat provided by the lower catalyst and the catalyst environment, so as to realize the cracking and conversion of heavy oil macromolecules to intermediate molecules, and obtain intermediate component raw materials with a molecular weight of 100-200 as much as possible.
  • the gas stream and catalyst generated in the riser reaction zone continue to flow upward into the high-temperature reaction zone; the high-temperature catalyst from the regenerator, the upper catalyst, enters the fast fluidized bed reaction zone.
  • the gas in the low temperature reaction zone is transported upwards into the high temperature reaction zone, further provides heat to the high temperature reaction zone, increases the reaction temperature in the high temperature reaction zone, and realizes the conversion reaction to propylene combining the catalytic reaction of the intermediate components and the thermal reaction; the reaction in the high temperature reaction zone
  • the material flows through the gas-solid separator in the settler for gas-solid separation, and the gas from which the catalyst is separated flows out of the settler and enters the subsequent treatment system;
  • the catalyst to be produced is separated in the settler and stripped in the stripping section, and then enters the regenerator for catalyst regeneration. It first enters the upper regeneration zone to be regenerated as a semi-regenerated catalyst, and then the semi-regenerated catalyst enters the lower regeneration zone to continue regeneration. Into the reactor, recycle;
  • Embodiment 2 is a diagrammatic representation of Embodiment 1:
  • a reactor R10, a regenerator G10, a settler D10 and a stripping section S10 are set, and petroleum hydrocarbons are used as raw materials, and C4 is simultaneously introduced into the riser reaction zone R17 to participate in the catalytic reaction; the mixed C4 reacts at the inlet of the reaction raw materials
  • the bottom of the raw material inlet R12A first enters the riser reaction zone R17 for reaction, and then reacts with the reaction raw material R12 in the riser reaction zone R17 and the fast fluidized bed reaction zone R18;
  • the reaction supplementary steam R15 is introduced into the riser reaction zone R17, and the second reaction supplementary steam R15A is introduced above the upper catalyst inlet R24A at the lower part of the fast fluidized bed reaction zone R18;
  • the riser reaction zone R17 that is, the low-temperature reaction zone steam accounts for the proportion of the reaction raw material R12.
  • the mass ratio is 5-50%, and the mass ratio of the steam in the rapid fluidized bed reaction zone R18, that is, the high-temperature reaction zone, accounts for 15-50% of the reaction raw material R12;
  • the regenerator G10 adopts the two-zone regeneration arranged up and down in series, the lower regeneration zone G13 adopts the form of a fast fluidized bed in a coke tank and is the first regeneration zone, and the upper regeneration zone G12 adopts the form of a dense phase turbulent fluidized bed and is the second regeneration zone.
  • the regenerator G10 is in the form of rapid fluidized bed and dense-phase fluidized bed series regeneration of the coke tank, and the lower part of the stripping section S10 is connected to the lower regeneration area G13 of the regenerator G10 through the standpipe S12 to be generated from the inlet G12A of the regenerator G10,
  • the to-be-grown catalyst first enters the lower regeneration zone G13 for regeneration, and then enters the upper dense-phase fluidized bed regeneration zone, that is, the upper regeneration zone G12 for regeneration;
  • the lower catalyst outlet G14A and the upper catalyst outlet G24A are both located in the upper regeneration zone G12, and the upper regeneration zone G12 of the regenerator G10 is connected to the lower catalyst inlet R14A at the lower part of the riser reaction zone R17 from the lower catalyst outlet G14A through the lower regeneration standpipe G14,
  • the upper regeneration zone G12 provides the lower catalyst to the riser reaction zone R17;
  • the upper regeneration zone G12 is connected to the lower part of the fast fluidized bed reaction zone R18 or the upper catalyst inlet at the outlet of the riser reaction zone R17 through the upper regeneration riser G24 through the upper catalyst outlet G24A R24A, supply the upper catalyst from the upper regeneration zone G12 to the fast fluidized bed reaction zone R18;
  • the lower catalyst temperature is 660°C ⁇ 740°C, the carbon content of the catalyst is less than 0.40%;
  • the upper catalyst temperature is 680 ⁇ 730°C, the carbon content of the catalyst is low at 0.1%.
  • Both the lower catalyst and the upper catalyst are regenerated
  • the reactor R10 is set as the upper and lower zone reactors with upper and lower three-way catalyst circulation and heating twice;
  • the return catalyst inlet R34A is set at the lower part of the fast fluidized bed reaction zone 18, and the return catalyst outlet S14A is set at the stripping section S10, which returns the catalyst
  • the inlet R34A is communicated with the return catalyst outlet S14A through the catalyst return delivery pipe S14, and the catalyst return spool valve VS14 is set on the catalyst return delivery pipe S14;
  • the regenerated catalyst introduced through the lower catalyst inlet R14A namely the lower catalyst
  • the regenerated catalyst introduced through the upper catalyst inlet R24A, the upper catalyst is the second-path catalyst
  • the regenerated catalyst introduced through the return catalyst inlet R34A is the catalyst of the second path.
  • the raw agent is the third-path catalyst;
  • the regenerated catalyst of the regenerator G10 enters the riser reaction zone R17 and the fast fluidized bed reaction zone R18 respectively, and realizes the gradual heating of the reaction process by means of graded heat supply; part of the catalyst to be grown in the stripping section S10 returns to the fast fluidized bed reaction zone R18, the catalyst storage and reaction weight hourly space velocity in the rapid fluidized bed reaction zone R18 are controlled by the return amount of the unborn agent; part of the unborn catalyst is drawn from the stripping section S10 and returned to the rapid fluidized bed reaction zone R18, the unborn agent Return to the fast fluidized bed reaction zone R18 from the catalyst delivery pipe S14 and the return catalyst inlet R34A, and the return amount of the unborn agent controls the catalyst storage and reaction weight hourly space velocity in the fast fluidized bed reaction zone R18, which is controlled by the unborn catalyst slide valve VS14
  • the return amount of the unborn catalyst; the weight hourly space velocity of the catalyst in the fast fluidized bed reaction zone R18 is 10(1/H)-35(1/H).
  • the return catalyst outlet can also be arranged above the riser reaction zone of the reactor, and the catalyst return conveying pipe is communicated with the return catalyst inlet, so that the reacted catalyst can be refluxed in the reactor, and the reaction zone can be supplied to the fast fluidized bed.
  • the catalyst after the reaction realizes the control of the space velocity of the catalyst in the fast fluidized bed reaction zone.
  • the reactor R10, the regenerator G10, the settler D10 and the stripping section S10 are set, and petroleum hydrocarbons are used as raw materials;
  • the reactor R10 is set as an upper and lower zoned reactor with upper and lower three-way catalyst circulation and heating twice; a return catalyst inlet R34A is set in the lower part of the fast fluidized bed reaction zone R18, and a return catalyst outlet S14A is set in the stripping section S10, which returns to the catalyst
  • the inlet R34A is communicated with the return catalyst outlet S14A through the catalyst return delivery pipe S14, and the catalyst return spool valve VS14 is set on the catalyst return delivery pipe S14;
  • the catalyst of the regenerator G10 enters the riser reaction zone R17 and the fast fluidized bed reaction zone R18 respectively, and realizes the gradual temperature rise of the reaction process by means of graded heat supply; part of the catalyst to be grown in the stripping section S10 returns to the fast fluidized bed reaction zone R18 , the catalyst storage and reaction weight hourly space velocity in the fast fluidized bed reaction zone R18 are controlled by the return amount of the unborn agent;
  • the lower catalyst is supplied from the lower regeneration zone G13 to the riser reaction zone R17, the catalyst temperature is 680°C-700°C, and the carbon content is 0.02%;
  • the upper regeneration zone G12 is connected to the riser reaction zone R17 outlet from the upper catalyst outlet G24A through the upper regeneration standpipe G24 At the upper catalyst inlet R24A, the upper catalyst is provided from the upper regeneration zone G12 to the fast fluidized bed reaction zone R18, the upper catalyst temperature is 740 ° C, and the carbon content is 0.15%;
  • the reactor R10, the regenerator G10, the settler D10 and the stripping section S10 are set, and petroleum hydrocarbons are used as raw materials;
  • Reactor R10 is set up as an upper and lower zoned reactor with upper and lower three-way catalyst circulation and heating twice; a liquid light hydrocarbon inlet R32A is set at the outlet of the riser reaction zone R17, and a C4 inlet R22A is set at the bottom of the fast fluidized bed reaction zone R18, The liquid light hydrocarbon R32 enters the reactor R10 at the outlet of the riser reaction zone R17, and the C4 directly enters the fast fluidized bed reaction zone R18 for reaction;

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Abstract

本发明涉及一种石油烃催化转化制丙烯和乙烯的方法，属于石油烃类催化转化技术领域。本发明提出了逐渐升温、两级温度梯度、重时空速独立控制的催化转化制丙烯和乙烯的方法，将反应器分成上下串联的两段或两区，包括下部的提升管反应区和上部的快速流化床反应区，反应原料按分子结构依次进行大分子催化裂化反应和小分子催化裂解反应，实现反应过程从大分子到丙烯和乙烯的分级转化。本发明在能耗更低的前提下，降低了低价值目的产品，如焦炭和干气的收率；提高了高价值目的产品丙烯和乙烯的收率。

Description

一种石油烃催化转化制丙烯和乙烯的方法及其装置 技术领域
本发明属于石油烃类催化转化技术领域，特别涉及一种石油烃催化转化制丙烯和乙烯的方法，本发明同时提供了实现上述方法的装置。
背景技术
丙烯是最重要的石油化工原料之一。70％的丙烯是石油烃经管式炉裂解法生产的，另外30％的丙烯由催化裂化工艺过程提供。借鉴常规重油催化裂化反应-再生系统的操作、设计经验，国内外研究人员开发了一系列重油催化裂化生产丙烯的工艺技术。
中国石油大学(华东)在两段提升管催化裂化技术基础上开发了TMP技术。该技术以重质油为原料，利用两段提升管催化裂化工艺分段反应、催化剂接力和大剂油比的工艺特点，针对不同性质的反应物料进行进料方式的优化组合，控制不同物料适宜的反应条件，以达到提高丙烯产率的目的；
上世纪90年代中国石化石油化工科学研究院开发了以重油为原料、丙烯为目的产品的DCC技术。该技术采用提升管加湍流流化床床层串联式反应器，在重时空速4(1/H)-6(1/H)气固流态化条件下进行重油制催化制丙烯和乙烯。石科院在DCC工艺基础上又开发了采用新型组合式反应器体系的增强型催化裂解技术(DCC-PLUS)，该技术与DCC工艺相同之处为都采用提升管反应器加流化床床层反应器的形式，不同之处在于DCC-PLUS工艺另外设置轻汽油和C4回炼提升管，轻汽油和C4反应后的物流引入流化床床层反应器。无论DCC还是DCC-PLUS都将原料油反应分为提升管和流化床床层反应；但DCC和DCC-PLUS都是通过进入原料油提升管反应区的再生催化剂量控制流化床区的反应温度，即全部反应过程都是按流化床催化裂解反应区的条件控制，必然使提升管反应区即重油反应区的催化裂化条件偏离理想的原料油催化裂化反应条件，尤其是热反应增加；另外对固定的原料量流化床床层反应区的空速只能靠流化床床层内的催化剂料位变化控制；由于催化剂携带和气固分离的要求，流化床床层反应区与气固分离器之间全部是稀相空间，并且油气离开流化床床层时仍然携带大量催化剂，而油气离开流化床床层到气固分离器的停留时间20秒以上，采用流化床床层反应时催化剂料位上方催化剂的携带以及油气停留时间都必然造成进一步的副反应止，丙烯进一步热裂解，影响产品分布及丙烯选择性，反应难以及时终止，必然导致催化裂化反应受到限制， 热反应增加，大幅度降低丙烯的选择性，导致干气和焦炭产率偏高。
已有技术的侧重点均为生产丙烯，分为两类，第一类为提升管加流化床串联式反应，第二类为双提升管并联式反应。研究者认为，重油催化裂化反应过程中的丙烯是由重质烃类一次裂化生成的汽油馏分经二次裂解间接生成的，汽油馏分中的C5-C8烯烃是丙烯主要的前身物。现有技术具有许多共同的特点，在操作上都采用比常规FCC工艺更高的反应温度、剂油比和蒸汽注入量，以提高裂化反应深度和丙烯的选择性。
发明内容
本发明的目的在于提供一种石油烃催化转化制丙烯和乙烯的方法，增加丙烯和乙烯产率的同时，显著降低干气和焦炭等副产品的收率。
本发明的另一目的在于提供一种实现上述石油烃催化转化制丙烯和乙烯方法的装置。
石油烃尤其是重油催化裂化反应似乎是熟知的，表面看似乎“差不多”，不过都是再生催化剂和原料油经过管式反应器反应。由于石油烃反应原料化学组成复杂，反应过程复杂，反应过程反应物化学成分的变化复杂，产品要求不同，导致实际化学反应各有不同。通过改进反应器内催化剂活性、空速等条件，以及改进热量分配、反应过程的温度分布、时间分布等工艺条件改进反应结果，正是石油烃催化裂化技术的进步途径。石油烃催化转化制丙烯和乙烯又是强烈的吸热、生焦大的反应，无论是催化剂还是反应物、产品在反应过程都变化显著而且复杂；控制反应器内反应过程中实际的催化剂条件、温度条件、空速条件对控制化学反应过程极其重要，反应器内不同的条件实际形成不同的化学反应，得到不同的产品，尤其是对以丙烯为产品时需要针对性的技术。石油烃反应原料催化裂解制丙烯和乙烯的目标是最大可能增加丙烯和乙烯。
本发明技术方案如下：
一种石油烃催化转化制丙烯和乙烯的方法，所述石油烃催化制丙烯和乙烯方法的反应过程为：重组分大分子先在气固输送流态化反应条件即提升管反应器内较高活性再生催化剂环境和较低的温度条件下进行催化裂化转化，形成C4-C12为主的中间分子，然后在提高催化剂密度、降低催化剂空速的气固快速流化床形态、含碳催化剂环境及高苛刻度条件下进行催化裂解转化，生产丙烯和乙烯；该石油烃催化转化方法在一个上下串联的两级反应器内进行，下方为气固输送床流态化形式即提升管流态化形式，上方为气固快速流态化形式，该反应器上下分独立控制的两级供热和两路来自再生器的催化剂循环，由上催化剂即上再生催化剂热量提供位置(即下文的上催化剂入口)将该反应器分 成上下两个反应区即提升管反应区和快速流化床反应区，反应器下方为提升管反应区，或称低温反应区，在提升管反应区内，重组分大分子在再生催化剂环境和较低的反应温度下进行向C5～C12中间组分转化的催化裂化反应，为制丙烯和乙烯提供中间原料；反应器上方为快速流化床反应区，或称高温反应区，由该反应器上方进入点进入的催化剂即上催化剂提供热量进一步提高快速流化床反应区的反应温度，形成高温反应区，C5～C12为主的中间组分在更高温度、更大剂油比、更低重时空速的更苛刻条件下进行催化裂解反应，使石油烃反应原料转化为丙烯；实现高沸点石油烃反应原料在两级供热、两级供催化剂、两级控制的上下两区反应器内进行选择性反应，用反应温度的逐渐升高的反应方式适应反应物分子量逐渐变小分子结构和对反应条件要求的变化，提高制丙烯和乙烯的效率和目的产品选择性；
反应原料依次在低温的提升管反应区和高温的快速流化床反应区反应，实现低温催化裂化反应和高温制丙烯和乙烯催化裂解反应，再生器的催化剂分别进入该提升管反应区和快速流化床反应区，通过分级供热方式实现反应过程的逐渐升温；石油烃反应原料先在低温反应区进行高沸点重组分的催化裂化反应，初步完成重组分、大分子的催化裂化转化和脱碳，生成以高烯烃汽油和柴油组分为主的中间组分，这些中间产物和催化剂一起向上进入反应器上方的高温反应区，同时通过另一路来自再生器的催化剂即上催化剂继续向该区提供热量和催化剂，提高反应物温度和剂油比，实现重组分大分子催化裂化反应和中间组分、小分子裂解转化；反应过程包含以下步骤：
(1)所述反应原料被蒸汽雾化后先进入反应器下部的提升管反应区，在来自再生器经下再生立管引入的下催化剂环境下，进行催化裂化反应；所述提升管反应区即低温反应区按有利于高沸点重组分催化转化条件进行，催化剂为再生后的催化剂，反应温度515℃～620℃，反应时间0.5～1.5s，以多产丙烯为目标时，反应温度515℃～550℃，反应时间1.0～2.5s；提升管反应区内的实际反应温度和剂油比由进入低温反应区的下催化剂量独立控制；
(2)所述反应原料在提升管反应区完成低温催化裂化反应后，生成的产物和催化剂向上流动进入快速流化床反应区，在快速流化床反应区内降低气体和催化剂流速，增加催化剂密度，降低催化剂重时空速；来自再生器经上再生立管引入的上催化剂进入反应器的快速流化床反应区，向快速流化床反应区提供热量，提高温度和剂油比，继续进行催化裂解反应，生成丙烯和乙烯产品；所述快速流化床反应区的反应温度为530℃～720℃，反应时间0.5～4.0s，以多产丙烯为目标时反应温度530℃～580℃，反应时间1.0～ 5.0s，反应压力绝压0.20MPa～0.40MPa，快速流化床反应区内的实际反应温度和剂油比由进入快速流化床反应区的催化剂量控制；
所述上催化剂，在提升管反应区反应温度控制点上方从上催化剂入口进入反应器，或直接进入快速流化床反应区，或者引入的上催化剂先在提升管反应区出口进入反应器后再输送到快速流化床反应区，进一步向快速流化床反应区提供热量，提高物流温度和剂油比，进一步进行催化裂解反应；
(3)反应后的物流进入沉降器进行气固分离，得到反应产物，分离出的催化剂在汽提段汽提后进入再生器再生，循环使用。
上述的石油烃原料油催化转化制丙烯和乙烯的方法，具体实施时，反应原料优选减压蜡油、常压渣油、焦化蜡油、脱沥青油、加氢蜡油(加氢处理后的蜡油)、加氢渣油(加氢处理后的渣油)、(直馏)柴油中的一种或混合物，终馏点高于320℃。
上述的石油烃催化转化制丙烯和乙烯的方法，进一步地，混合C4(又称C4)在反应原料入口下方先进入提升管反应区反应，然后和反应原料一起在提升管反应区和快速流化床反应区反应；或者C4在上催化剂入口上方进入反应器或直接进入快速流化床反应区反应。
上述的石油烃催化转化制丙烯和乙烯的方法，进一步地，从汽提段引部分待生催化剂返回反应器，或者在快速流化床反应区上方从反应器引出部分反应后的催化剂从返回催化剂入口返回反应器；待生催化剂或反应后的催化剂从催化剂返回输送管和返回催化剂入口返回反应器，返回催化剂入口设在提升管反应区出口或设在快速流化床反应区底部；待生催化剂或反应后的催化剂返回量控制快速流化床反应区内催化剂藏量和反应催化剂重时空速，由催化剂返回滑阀控制待生催化剂或反应后的催化剂返回量。
上述的石油烃催化转化制丙烯和乙烯的方法，进一步地，质量比90％以上沸点或终馏点低于360℃的液体轻烃在提升管反应区出口进入反应器或直接进入快速流化床反应区。
上述石油烃催化转化制丙烯和乙烯的方法，进一步地，提升管反应区采用气固气力输送流态化形式，气体平均流速5.0m/s～20m/s；快速流化床反应区直径大于提升管反应区，采用快速气固流化床条件，气体平均流速1.8m/s～5.0m/s，催化剂重时空速10(1/H)～35(1/H)。
上述的石油烃催化转化制丙烯和乙烯的方法，进一步地，所述再生器的再生器稀相区下方设置上下两个再生区即下再生区和上再生区，具体实施时，下再生区为烧焦罐再 生区，上再生区为密相流化床再生区；下再生区为第一再生区，上再生区为第二再生区，来自汽提段的待生催化剂进入下再生区，上再生区同时设置下催化剂出口和上催化剂出口；从上再生区向反应器提供下催化剂和上催化剂。优选地，下催化剂温度660℃～740℃，催化剂含碳量低于0.40％；上催化剂温度680～730℃，催化剂含碳量低于0.1％。下催化剂和上催化剂均为再生催化剂。
上述的石油烃催化转化制丙烯和乙烯的方法，进一步地，再生器的再生器稀相下方设置上下两个再生区即下再生区和上再生区，具体实施时，上再生区和下再生区均为密相流化床再生区；上再生区为第一再生区，下再生区为第二再生区，来自汽提段的待生催化剂进入上再生区，在下再生区设置下催化剂出口，在上再生区设置上催化剂出口，经下再生区由下再生立管(由下催化剂出口引出的)向提升管反应区提供的下催化剂为再生催化剂，经上再生立管(由上催化剂出口引出的)向反应器提供的上催化剂为适度含碳的半再生催化剂。具体实施时，通过调整再生器两个再生区催化剂量和再生氧量或空气量的分配，调节两个再生区催化剂的含碳量和温度，实现向高温反应区供热和向两个反应区提供不同含碳量的催化剂；当向反应器提供不同含碳量的催化剂时，向上方高温裂解反应区提供含碳量高的催化剂。优选地，下催化剂温度660℃～740℃，催化剂含碳量低于0.10％；上催化剂温度680～730℃，催化剂含碳量低于0.5％。
本发明中，在具体实施时，根据反应器内反应物的不同分子结构对催化剂温度和含碳量要求的不同，需用不同的再生方式。当进入反应器不同反应区的催化剂不需要分别控制含碳量或不需要不同反应区使用不同含碳量的催化剂反应时，或不需要进入不同反应区的催化剂温度有明显区别时，再生器采用常规烧焦罐再生方式，下再生区是快速流化床，上再生区是密相流化床，从密相流化床向反应器提供催化剂；
当进入反应器不同反应区的催化剂需要分别控制含碳量或需要不同反应区使用不同含碳量的催化剂反应时，或需要进入不同反应区的催化剂温度有明显区别时，再生器采用上下双密相流化床再生形式，调节控制两个流化床再生区的催化剂量或进入的烧焦氧气或空气量，使两个再生区的催化剂温度和含碳量不同，根据需要从不同的密相流化床向反应器提供不同含碳量和温度的催化剂；
汽提后的待生催化剂进入上再生区时，上再生区内的催化剂为适度含碳的“半再生催化剂”，下再生区催化剂为再生催化剂，从上再生区向反应器快速流化床反应区提供催化剂，从下再生区向提升管反应区提供催化剂。
本发明方法提升管反应区主要用于大分子重组分催化转化；快速流化床反应区主要 进行裂解制丙烯、乙烯和芳烃的反应。
本发明还提供了一种石油烃催化转化制丙烯和乙烯的装置，具体而言，该装置为用于实现本发明所述的石油烃催化转化制丙烯和乙烯方法的装置。
根据本发明的具体实施方案，所述装置包括：反应器、再生器、沉降器和汽提段，下再生立管和下再生滑阀，上再生立管和上催化剂滑阀，待生立管和待生催化剂滑阀；再生器和沉降器并列设置；
所述反应器设置成上下两路催化剂循环、两次供热的上下分区反应器形式，分成上下串联的两段或两区，包括下部的提升管反应区和上部的快速流化床反应区，提升管反应区为气体流速高的气固气力输送流态化形式，快速流化床反应区为气体流速降低的快速流化床形式，快速流化床反应区直径大于提升管反应区；
所述反应器设置两路来自再生器的催化剂循环、上下两次向反应器供热，提升管反应区用于原料油大分子低温催化裂化反应，快速流化床反应区用于高温制丙烯和乙烯催化裂解反应；在所述提升管反应区下部设置反应原料入口；
所述提升管反应区下部的下催化剂入口设于反应原料入口下方，所述下催化剂入口通过下再生立管与所述再生器的下催化剂出口连通；在快速流化床反应区底部或者提升管反应区出口设置上催化剂入口，所述上催化剂入口通过上再生立管与所述再生器的上催化剂出口连通；在下再生立管、上再生立管分别设置下再生滑阀和上催化剂滑阀；
所述汽提段通过待生催化剂出口、待生立管和待生剂入口与再生器连通，在待生立管上设置待生催化剂滑阀。
上述的石油烃催化转化制丙烯和乙烯的装置，优选地，在反应原料入口下方，或在上催化剂入口上方或快速流化床反应区底部设置C4入口，或/和在上催化剂入口上方或快速流化床反应区底部或提升管反应区出口设置液体轻烃入口。
上述的石油烃催化转化制丙烯和乙烯的装置，优选地，在所述快速流化床反应区下部或所述提升管反应区温度检测控制点上方的提升管反应区出口设置返回催化剂入口，在所述汽提段设置返回催化剂出口，所述返回催化剂入口通过催化剂返回输送管与所述返回催化剂出口连通，实现经汽提段向快速流化床反应区提供待生催化剂；在催化剂返回输送管上设置催化剂返回滑阀。
上述的石油烃催化转化制丙烯和乙烯的装置，优选地，在所述快速流化床反应区下部或所述提升管反应区温度检测控制点上方的提升管反应区出口设置返回催化剂入口，在所述反应器提升管反应区上方设置返回催化剂出口，通过催化剂返回输送管与返回催 化剂入口连通，实现反应后的催化剂在反应器回流，向快速流化床反应区提供反应后的催化剂，实现对快速流化床反应区催化剂空速的控制；在催化剂返回输送管上设置催化剂返回滑阀。
上述的石油烃催化转化制丙烯和乙烯的装置，优选地，所述再生器的再生器稀相区下方设置上下两个再生区即下再生区和上再生区；
所述下催化剂出口和上催化剂出口均设于上再生区，实现经上再生区向反应器提供再生催化剂，所述待生剂入口设于下再生区；
或者，所述上再生区为密相流化床再生区，所述下催化剂出口设于下再生区，实现经下再生区向提升管反应区提供再生催化剂即下再剂，所述上催化剂出口设于上再生区，实现经上再生区的流化床再生区向快速流化床反应区提供催化剂即上催化剂，所述待生剂入口设于上再生区。
本发明中，低温反应区蒸汽占反应原料的质量比5～50％，高温反应区蒸汽占反应原料的质量比为15～50％。
具体实施时，可以在反应沉降器产物出口后设置蒸汽发生器，利用高温产品物流的热量产生蒸汽，实现产品物流的降温或急冷，该蒸汽发生器工程设计单位掌握。
本发明的有益效果：
本发明从催化裂化原理出发，提出了逐渐升温、两级温度梯度、重时空速独立控制的催化转化制丙烯和乙烯的方法。本领域人员周知，重油催化裂化过程可视为平行顺序反应，重组分大分子首先裂化生成中等分子(C5～C12)产物，较低的裂解温度能突出催化裂化反应，通常催化裂化反应适宜的温度是490℃～530℃；部分汽、柴油在540℃～580℃时继续裂解为C3～C4。本发明遵循该反应规律，设置了逐渐升温的两级温度梯度串联：低温区、高温区；本发明通过待生催化剂返回反应器，实现对反应过程剂油比和催化剂空速的调节和控制而不影响热平衡。本发明在能耗更低的前提下，降低了低价值目的产品，如焦炭和干气的收率；提高了高价值目的产品丙烯和乙烯的收率。
本发明方法通过对反应过程剂油比、空速、和温度的分级控制，尤其是实现随反应进行，剂油比和温度逐渐提高、重时空速降低、反应苛刻度逐渐提高，使反应条件与石油烃反应原料裂化过程石油烃分子逐渐变小和需要的反应苛刻度逐渐提高的反应化学条件相适应；本发明也很好的优化了重组分和轻烃不同性质原料的共同转化效果，避免小分子轻烃过度裂化，即保证了重组分裂化条件，又保证了轻烃裂化条件；本方法提高了效率,增加了目的产品选择性。
附图说明
图1为本发明实施方式一工艺示意图；
图2为本发明实施方式二工艺示意图；
图3为本发明实施方式三工艺示意图；
图4为本发明实施方式四工艺示意图；
图5为本发明实施方式五工艺示意图。
图中编号标记内容如下：
R10反应器；R11催化剂提升气体；R11A催化剂提升气体入口，R12反应原料，R12A反应原料入口，R13原料雾化蒸汽，R14A下催化剂入口(提升管反应区催化剂入口)；R15反应补充蒸汽，R15A第二反应补充蒸汽，R17提升管反应区(或低温反应区)，R18快速流化床反应区(或高温反应区)；R24A上催化剂入口(快速流化床反应区催化剂入口)；R34A返回催化剂入口；R32液体轻烃；R32A液体轻烃入口；R22A C4入口；C4碳四组分石油烃；
S10汽提段，S11汽提构件；S12待生立管；S12A待生催化剂出口(待生剂出口)；VS12待生催化剂滑阀(待生剂滑阀)；S13汽提蒸汽；S14催化剂返回输送管；S14A返回催化剂出口；VS14催化剂返回滑阀；
D10沉降器，D11沉降旋风分离器；D12反应产物，D12A沉降器反应产品出口；
G10再生器，G11催化剂再生气体，G11A再生气体入口，G12上再生区(上催化剂再生区)，G12A待生剂入口，G13下再生区(下催化剂再生区)；G14下再生立管(下催化剂输送管)，G14A下催化剂出口，G15再生器稀相区，G16再生旋风分离器，G17烧焦后烟气，G17A烟气出口，G18空气，G24上再生立管(上催化剂输送管)，G24A上催化剂出口(上催化剂出口)；VG14下再生滑阀(下再生催化剂滑阀)，VG24上催化剂滑阀；
TI温度检测信号，TC温度控制信号，DPC压差控制信号。
具体实施方式
以下结合附图及实施例详细说明本发明的技术方案，但本发明的保护范围并不限于此。
实施方式一：
如图1所示，设置反应器R10、再生器G10、沉降器D10和汽提段S10，采用石油 烃为原料，石油烃反应原料可以为减压蜡油、常压渣油、加氢处理后的蜡油、加氢处理后的渣油、原油中的一种或混合物；
再生器G10和沉降器D10并列设置，反应器R10出口连通沉降器D10内的沉降旋风分离器D11，汽提段S10设于沉降器D10下方，汽提段S10内设置有汽提构件S11，向汽提段S10引入汽提蒸汽S13实现催化剂汽提；
反应器R10设置成上下两路催化剂循环、两次供热的上下分区反应器形式，包括下部的提升管反应区R17和上部的快速流化床反应区R18，快速流化床反应区R18直径大于提升管反应区R17，提升管反应区R17用于低温催化裂化反应，快速流化床反应区R18用于高温制丙烯和乙烯裂解反应；提升管反应区R17下部的下催化剂入口R14A通过下再生立管G14与再生器G10的下催化剂出口G14A连通，快速流化床反应区R18下部的上催化剂入口R24A通过上再生立管G24与所述再生器G10的上催化剂出口G24A连通；在反应器R10下部设置反应原料入口R12A，以引入反应原料R12及原料雾化蒸汽R13，反应器R10底部设置催化剂提升气体入口R11A，以引入催化剂提升气体R11；沉降器D10顶部设置有沉降器反应产品出口D12A，以引出反应产物D12；
再生器G10采用上下串联布置的两区再生，在再生器稀相区G15下方设置上下两个再生区即下再生区G13和上再生区G12，具体实施时，下再生区G13和上再生区G12均采用密相流化床形式，下再生区G13为第二再生区，上再生区G12为第一再生区，待生催化剂在上再生区G12再生后，回流进入下再生区G13继续再生；所述下催化剂出口G14A设于下再生区G13，上催化剂出口G24A设于上再生区G12，实现经上再生区G12向快速流化床反应区R18提供的上催化剂为半再生催化剂，经下再生区G13向提升管反应区R17提供的下催化剂为再生催化剂；汽提段S10下部通过待生立管S12由待生剂入口G12A连通再生器G10的上再生区G12，待生立管S12上设有待生催化剂滑阀VS12；
再生器G10的上再生区G12由下催化剂出口G14A通过下再生立管G14连通提升管反应区R17下部的下催化剂入口R14A，下再生立管G14上设有下再生滑阀VG14；上再生区G12由上催化剂出口G24A通过上再生立管G24连通快速流化床反应区R18下部或提升管反应区R17出口的上催化剂入口R24A，上再生立管G24上设有上催化剂滑阀VG24；
再生器G10的再生器稀相区G15内设置有再生旋风分离器G16，再生器烧焦后烟气G17由再生器G10顶部的烟气出口G17A排出，催化剂再生气体G11由再生器G10 底部的再生气体入口G11A引入；空气G18由上再生区G12下部引入再生器G10；
在提升管反应区R17引入反应补充蒸汽R15。
再生器G10的催化剂分别进入提升管反应区R17和快速流化床反应区R18，通过分级供热方式实现反应过程的逐渐升温；具体实施过程包含以下步骤：
(1)预热后的反应原料R12被原料雾化蒸汽R13雾化后先进入反应器R10下部的提升管反应区R17，从下再生立管G14来自再生器上再生区G12的下催化剂从下催化剂入口R14A进入提升管反应区R17，在催化剂提升气体R11作用下向上输送与原料接触，反应原料R12在催化剂环境下以较缓和的条件进行催化裂化转化，形成以C5-C12为主的中间产物；下催化剂温度优选660℃～740℃，催化剂含碳量低于0.10％；提升管反应区R17的反应温度515℃～620℃，反应时间0.5～1.5s；当多产丙烯时，提升管反应区R17的反应温度515℃～550℃，反应时间1.0s～2.5s；提升管反应区R17采用气固气力输送流态化形式，气体平均流速5.0m/s～20m/s，催化剂重时空速10(1/H)～35(1/H)；
(2)反应原料R12完成低温催化裂化反应后，然后提升管反应区R17生成的产物和催化剂一起向上流动进入快速流化床反应区R18，新催化剂即来自再生器G10经上再生立管G24引入的上催化剂进入快速流化床反应区R18，被来自提升管反应区R17的物流输送到快速流化床反应区R18，新催化剂即上催化剂进一步向快速流化床反应区R18提供热量，提高物流温度和剂油比，形成苛刻度更高的高温裂解反应条件，来自提升管反应区的产物继续进行催化裂解和热裂解组合反应，丙烯、乙烯等低碳数小分子产品；上催化剂优选温度680～730℃，催化剂含碳量低于0.5％，快速流化床反应区R18的反应温度为530℃～720℃，反应时间0.5～4.0s，当多产丙烯时，快速流化床反应区R18的反应温度为530℃～580℃，反应时间1.0～5.0s，反应压力绝压0.20MPa～0.40Mpa，实际反应温度由进入快速流化床反应区R18的催化剂量控制；快速流化床反应区R18直径大于提升管反应区R17，采用快速气固流化床条件，气体平均流速1.8m/s～5.0m/s；
(3)反应后的物流进入沉降器D10进行气固分离，得到反应产物D12，从沉降器反应产品出口D12A送出进入后续处理部分；反应后的催化剂经沉降旋风分离器D11分离后，在汽提段S10汽提后经待生立管S12和待生剂入口G12A进入再生器G10的上再生区G12，再生并循环利用。
反应产物D12离开图1所示的催化转化装置后，进行产品分馏；分馏等都是工程技术人员熟知的。
实施例1
本实施例使用的装置见图1，实施参数如下：
反应原料为减压蜡油，密度0.89，氢含量13.2％(重量)，残碳4.0％，饱和烃60％；原料油预热温度220℃；
反应装置为沉降器和再生器并列布置，再生器采用烧焦罐快速流化床和密相流化床串联再生形式；
提升管反应区反应条件：反应温度TIC控制530℃，气体平均流速6.5m/s，反应时间1.7s(秒)；催化剂输送气体为蒸汽，数量为反应原料的3％，原料雾化蒸汽为原料的7％；补充蒸汽为原料的20％，从下催化剂入口进入的催化剂即下催化剂为再生催化剂，含碳量0.02％，下催化剂温度680℃；
快速流化床反应区反应条件：从上催化剂入口进入的催化剂即上催化剂为半再生催化剂，含碳量0.1％，温度705℃，反应温度控制560℃，气体平均流速3.5m/s，反应时间1.5秒；
反应过程为：
原料用蒸汽雾化后进入提升管反应区，在下催化剂提供的热量和催化剂环境下进行重油催化裂化转化，实现重油大分子向中间分子的裂化转化，尽可能得到分子量100～200的中间组分原料，为进一步转化成丙烯和乙烯提供中间原料；在提升管反应区生成的气体物流和催化剂继续向上流动进入高温反应区；来自再生器的高温催化剂即上催化剂进入快速流化床反应区，被来自低温反应区的气体向上输送进入高温反应区，进一步向高温反应区提供热量、提高高温反应区内的反应温度，实现中间组分催化反应和热反应结合的向丙烯转化反应；高温反应区的反应物流在沉降器内经过气固分离器进行气固分离，分离出催化剂的气体流出沉降器，进入后续处理系统；
沉降器内分离出得到待生催化剂在汽提段内汽提后进入再生器的进行催化剂再生，先进入上再生区再生为半再生催化剂，半再生催化剂再进入下再生区继续再生，催化剂再生后进入反应器，循环使用；
催化剂的再生、气固分离、以及后续的油气处理是常用技术，技术人员熟知，不再赘述。
实施方式二：
如图2所示，设置反应器R10、再生器G10、沉降器D10和汽提段S10，采用石油烃为原料，在提升管反应区R17同时引入C4参与催化反应；混合C4在反应原料入口反应原料入口R12A下方先进入提升管反应区R17反应，然后和反应原料R12一起在提 升管反应区R17和快速流化床反应区R18反应；
在提升管反应区R17引入反应补充蒸汽R15，在快速流化床反应区R18下部的上催化剂入口R24A上方引入第二反应补充蒸汽R15A；提升管反应区R17即低温反应区蒸汽占反应原料R12的质量比5～50％，快速流化床反应区R18即高温反应区蒸汽占反应原料R12的质量比为15～50％；
再生器G10采用上下串联布置的两区再生，下再生区G13采用烧焦罐快速流化床形式，为第一再生区，上再生区G12采用密相湍流流化床形式，为第二再生区，即再生器G10为烧焦罐快速流化床和密相流化床串联再生形式，汽提段S10下部通过待生立管S12由待生剂入口G12A连通再生器G10的下再生区G13，待生催化剂先进入下再生区G13再生，再进入上方的密相流化床再生区即上再生区G12再生；
下催化剂出口G14A和上催化剂出口G24A均设于上再生区G12，再生器G10的上再生区G12由下催化剂出口G14A通过下再生立管G14连通提升管反应区R17下部的下催化剂入口R14A，从上再生区G12向提升管反应区R17提供下催化剂；上再生区G12由上催化剂出口G24A通过上再生立管G24连通快速流化床反应区R18下部或提升管反应区R17出口处的上催化剂入口R24A，从上再生区G12向快速流化床反应区R18提供上催化剂；下催化剂温度660℃～740℃，催化剂含碳量低于0.40％；上催化剂温度680～730℃，催化剂含碳量低于0.1％。下催化剂和上催化剂均为再生催化剂；
其他部分装置结构与实施方式一相同。
实施方式三：
如图3所示，设置反应器R10、再生器G10、沉降器D10和汽提段S10，采用石油烃为原料；
反应器R10设置成上下三路催化剂循环、两次供热的上下分区反应器形式；在快速流化床反应区18下部设置返回催化剂入口R34A，在汽提段S10设置返回催化剂出口S14A，返回催化剂入口R34A通过催化剂返回输送管S14与返回催化剂出口S14A连通，在催化剂返回输送管S14上设置催化剂返回滑阀VS14；
在反应器R10内，经下催化剂入口R14A引入的再生催化剂即下催化剂为第一路催化剂，经上催化剂入口R24A引入的再生催化剂即上催化剂为第二路催化剂，经返回催化剂入口R34A引入的待生剂为第三路催化剂；
再生器G10的再生催化剂分别进入提升管反应区R17和快速流化床反应区R18，通过分级供热方式实现反应过程的逐渐升温；汽提段S10的待生催化剂部分返回快速流 化床反应区R18，通过待生剂返回量来控制快速流化床反应区R18内催化剂藏量和反应重时空速；从汽提段S10引部分待生催化剂返回快速流化床反应区R18内，待生剂从催化剂输送管S14和返回催化剂入口R34A返回快速流化床反应区R18，待生剂返回量控制快速流化床反应区R18内催化剂藏量和反应重时空速，由待生催化剂滑阀VS14控制待生催化剂返回量；快速流化床反应区R18内的催化剂重时空速为10(1/H)-35(1/H)。
其他部分装置结构同实施方式二。
在具体实施时，返回催化剂出口也可以设置在反应器的提升管反应区上方，通过催化剂返回输送管与返回催化剂入口连通，实现反应后的催化剂在反应器回流，向快速流化床反应区提供反应后的催化剂，实现对快速流化床反应区催化剂空速的控制。
实施方式四：
如图4所示，设置反应器R10、再生器G10、沉降器D10和汽提段S10，采用石油烃为原料；
反应器R10设置成上下三路催化剂循环、两次供热的上下分区反应器形式；在快速流化床反应区R18下部设置返回催化剂入口R34A，在汽提段S10设置返回催化剂出口S14A，返回催化剂入口R34A通过催化剂返回输送管S14与返回催化剂出口S14A连通，在催化剂返回输送管S14上设置催化剂返回滑阀VS14；
再生器G10的催化剂分别进入提升管反应区R17和快速流化床反应区R18，通过分级供热方式实现反应过程的逐渐升温；汽提段S10的待生催化剂部分返回快速流化床反应区R18，通过待生剂返回量来控制快速流化床反应区R18内催化剂藏量和反应重时空速；
从下再生区G13向提升管反应区R17提供下催化剂，催化剂温度680℃-700℃，含碳0.02％；上再生区G12由上催化剂出口G24A通过上再生立管G24连通提升管反应区R17出口处的上催化剂入口R24A，从上再生区G12向快速流化床反应区R18提供上催化剂，上催化剂温度740℃，含碳0.15％；
其他部分装置结构与实施方式一相同。
实施方式五：
如图5所示，设置反应器R10、再生器G10、沉降器D10和汽提段S10，采用石油烃为原料；
反应器R10设置成上下三路催化剂循环、两次供热的上下分区反应器形式；在提升管反应区R17出口设置液体轻烃入口R32A，在快速流化床反应区R18底部设置C4入 口R22A，液体轻烃R32在提升管反应区R17出口进入反应器R10，C4直接进入快速流化床反应区R18反应；
其他部分装置结构与实施方式一相同。

Claims (15)

  1. 一种石油烃催化转化制丙烯和乙烯的方法，该方法包括：反应原料(R12)按分子结构依次进行大分子催化裂化反应和小分子催化裂解反应；反应方法包含以下步骤：
    (1)所述反应原料(R12)被蒸汽雾化后从反应原料入口(R12A)进入反应器(R10)下部的提升管反应区(R17)，在再生后的催化剂环境下进行大分子低温催化裂化反应，下再生立管(G14)输送来自再生器(G10)的再生催化剂也称下催化剂从下催化剂入口(R14A)在反应原料(R12)下方进入反应器(R10)，催化剂提升气体(R11)从催化剂提升气体入口(R11A)进入反应器(R10)，将催化剂提升到提升管反应区(R17)，实现反应原料(R12)中大分子组分的催化裂化反应；提升管反应区(R17)内的反应温度由来自下再生立管(G14)的再生催化剂流量控制，来自下再生立管(G14)的再生催化剂流量由下再生滑阀(VG14)控制；
    (2)所述反应原料(R12)完成大分子催化裂化反应后，提升管反应区(R17)生成的产物和催化剂向上流动进入快速流化床反应区(R18)，在快速流化床反应区(R18)内降低气体和催化剂流速，增加催化剂密度，降低催化剂重时空速；来自再生器(G10)、经上再生立管(G24)引入的催化剂也称上催化剂，从上催化剂入口(R24A)进入反应器(R10)，或直接进入快速流化床反应区(R18)，或者引入的催化剂先在提升管反应区(R17)出口进入反应器(R10)后再输送到快速流化床反应区(R18)，进一步向快速流化床反应区(R18)提供热量，提高物流温度和剂油比，进一步进行催化裂解反应；快速流化床反应区(R18)的反应温度由来自上再生立管(G24)的催化剂流量控制，来自上再生立管(G24)的催化剂剂流量由上催化剂滑阀(VG24)控制；
    (3)反应后的物流进入沉降器(D10)进行气固分离，得到反应产物(D12)，分离出的催化剂在汽提段(S10)汽提后进入再生器(G10)再生，循环使用。
  2. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，混合C4在反应原料入口(R12A)下方先进入提升管反应区(R17)反应，然后和反应原料(R12)一起在提升管反应区(R17)和快速流化床反应区(R18)反应；或者，C4在上催化剂入口(R24A)上方进入反应器(R10)或C4直接进入快速流化床反应区(R18)反应。
  3. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，从汽提段(S10)引部分待生催化剂返回反应器(R10)，或者在快速流化床反应区(R18)上方从反应器(R10)引出部分反应后的催化剂返回反应器(R10)；待生催化剂或反应后的催化剂从催化剂返回输送管(S14)和返回催化剂入口(R34A)返回反应器(R10)，返回催化剂入口(R34A)设在提升管反应区(R17)出口或设在快速流化床反应区(R18)底部；待生催化剂或反应 后的催化剂返回量控制快速流化床反应区(R18)内催化剂藏量和反应催化剂重时空速，由催化剂返回滑阀(VS14)控制待生催化剂或反应后的催化剂返回量。
  4. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，质量比90％以上沸点或终馏点低于360℃的液体轻烃(R32)在提升管反应区(R17)出口进入反应器(R10)或直接进入快速流化床反应区(R18)。
  5. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，所述再生器(G10)的再生器稀相区(G15)下方设置上下两个再生区即下再生区(G13)和上再生区(G12)，下再生区(G13)为第一再生区，上再生区(G12)为第二再生区，来自汽提段(S10)的待生催化剂进入下再生区(G13)，从上再生区(G12)向反应器(R10)提供再生催化剂即下催化剂和上催化剂。
  6. 如权利要求5所述的石油烃催化转化制丙烯和乙烯的方法，其中，下催化剂温度优选660℃～740℃，催化剂含碳量低于0.40％；上催化剂优选温度680～730℃，催化剂含碳量低于0.1％。
  7. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，所述再生器(G10)的再生器稀相区(G15)下方设置上下两个再生区即下再生区(G13)和上再生区(G12)，上再生区(G12)为第一再生区，下再生区(G13)为第二再生区，来自汽提段(S10)的待生催化剂进入上再生区(G12)，经下再生立管(G14)向提升管反应区(R17)提供的下催化剂为再生催化剂，经上再生立管(G24)向反应器(R10)提供的上催化剂为半再生催化剂；通过调整下再生区(G13)和上再生区(G12)的催化剂量和再生氧量的分配，来调节两个再生区催化剂的含碳量和温度。
  8. 如权利要求7所述的石油烃催化转化制丙烯和乙烯的方法，其中，下催化剂温度优选660℃～740℃，催化剂含碳量低于0.10％；上催化剂优选温度680～730℃，催化剂含碳量低于0.5％。
  9. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，所述提升管反应区(R17)的反应温度515℃～620℃，反应时间0.5～1.5s；所述快速流化床反应区(R18)的反应温度为530℃～720℃，反应时间0.5～4.0s，反应压力绝压0.20MPa～0.26Mpa。
  10. 如权利要求9所述的石油烃催化转化制丙烯和乙烯的方法，其中，所述提升管反应区(R17)的反应温度515℃～550℃，反应时间1.0～2.5s；所述快速流化床反应区(R18)的反应温度为530℃～580℃，反应时间1.2～5.0s，反应压力绝压0.20MPa～0.40Mpa。
  11. 如权利要求1所述的石油烃催化转化制丙烯和乙烯的方法，其中，提升管反应区(R17)采用气固气力输送流态化形式，气体平均流速5.0m/s～20m/s；快速流化床反应区(R18)直径大于提升管反应区(R17)，采用快速气固流化床条件，气体平均流速1.8m/s～ 5.0m/s。
  12. 一种石油烃催化转化制丙烯和乙烯的装置，该装置包括：反应器(R10)、再生器(G10)、沉降器(D10)和汽提段(S10)，下再生立管(G14)和下再生滑阀(VG14)，上再生立管(G24)和上催化剂滑阀(VG24)，待生立管(S12)和待生催化剂滑阀(VS12)；再生器(G10)和沉降器(D10)并列设置；
    所述反应器(R10)分成上下串联的两段或两区，包括下部的提升管反应区(R17)和上部的快速流化床反应区(R18)，提升管反应区(R17)为气体流速高的气固气力输送流态化形式，快速流化床反应区(R18)为气体流速降低的快速流化床形式，快速流化床反应区(R18)直径大于提升管反应区(R17)；在所述提升管反应区(R17)下部设置反应原料入口(R12A)；
    所述提升管反应区(R17)下部的下催化剂入口(R14A)设于反应原料入口(R12A)下方，所述下催化剂入口(R14A)通过下再生立管(G14)与所述再生器(G10)的下催化剂出口(G14A)连通；在快速流化床反应区(R18)底部或者提升管反应区(R17)出口设置上催化剂入口(R24A)，所述上催化剂入口(R24A)通过上再生立管(G24)与所述再生器(G10)的上催化剂出口(G24A)连通；在下再生立管(G14)上设置下再生滑阀(VG14)，在上再生立管(G24)上设置上催化剂滑阀(VG24)；
    所述汽提段(S10)通过待生催化剂出口(S12A)、待生立管(S12)和待生剂入口(G12A)与再生器(G10)连通，在待生立管(S12)上设置待生催化剂滑阀(VS12)。
  13. 如权利要求12所述的石油烃催化转化制丙烯和乙烯的装置，其中，在所述快速流化床反应区(R18)下部或所述提升管反应区(R17)出口设置返回催化剂入口(R34A)，在所述汽提段(S10)或快速流化床反应区(R18)上方设置返回催化剂出口(S14A)，所述返回催化剂入口(R34A)通过催化剂返回输送管(S14)与所述返回催化剂出口(S14A)连通；在催化剂返回输送管(S14)上设置催化剂返回滑阀(VS14)。
  14. 如权利要求12所述的石油烃催化转化制丙烯和乙烯的装置，其中，所述再生器(G10)的再生器稀相区(G15)下方设置上下两个再生区即下再生区(G13)和上再生区(G12)；
    所述下催化剂出口(G14A)和上催化剂出口(G24A)均设于上再生区(G12)，所述待生剂入口(G12A)设于下再生区(G13)；
    或者，所述下催化剂出口(G14A)设于下再生区(G13)，所述上催化剂出口(G24A)设于上再生区(G12)，所述待生剂入口(G12A)设于上再生区(G12)。
  15. 如权利要求12所述的石油烃催化转化制丙烯和乙烯的装置，其中，在反应原料油入口(R12A)下方或在快速流化床反应区(R18)底部设置C4入口(R22A)，或/和在快速流化床反应区(R18)底部或提升管反应区(R17)出口设置液体轻烃入口(R32A)。
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CN115155466B (zh) * 2022-08-05 2023-10-13 中国石油大学(北京) 一种甲烷氧化偶联制乙烯的耦合反应系统和方法
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Citations (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1986505A (zh) * 2005-12-23 2007-06-27 中国石油化工股份有限公司 一种增产低碳烯烃的催化转化方法
CN101293806A (zh) * 2007-04-28 2008-10-29 中国石油化工股份有限公司 一种提高低碳烯烃产率的催化转化方法
CN102295510A (zh) * 2010-06-24 2011-12-28 中国石油化工股份有限公司 石脑油催化转化为低碳烯烃的方法
CN103131464A (zh) * 2011-11-23 2013-06-05 中国石油化工股份有限公司 一种生产低碳烯烃和轻芳烃的烃类催化转化方法
CN110240925A (zh) * 2018-05-29 2019-09-17 青岛京润石化设计研究院有限公司 一种流化催化裂化反应再生方法

Family Cites Families (12)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN101440014B (zh) * 2007-11-22 2012-05-30 中国石油化工股份有限公司 一种生产低碳烯烃的方法
CN101323798B (zh) * 2008-07-31 2012-10-31 石宝珍 一种催化转化方法及其专用装置
CN101760228B (zh) * 2008-12-25 2013-06-26 中国石油化工股份有限公司 制取丙烯和高辛烷值汽油的催化转化方法
CN102453500A (zh) * 2010-10-21 2012-05-16 中国石油化工股份有限公司 一种烃油催化裂化方法和烃油催化裂化设备
CN102875285B (zh) * 2011-07-12 2015-12-09 中国石油化工股份有限公司 甲醇和轻石脑油催化转化为低碳烯烃的方法
CN104478648A (zh) * 2014-12-30 2015-04-01 江苏健神生物农化有限公司 一种以石脑油为原料催化裂解制备低碳烯烃的方法
CN107286972B (zh) * 2017-07-21 2019-11-22 石宝珍 一种多产丙烯的催化转化方法
CN108794292B (zh) * 2017-07-21 2020-03-17 青岛京润石化设计研究院有限公司 一种多产丙烯的催化转化方法
CN109420468B (zh) * 2017-08-31 2023-10-31 中国石油化工股份有限公司 反应设备及其用途
CN109833834B (zh) * 2018-05-29 2024-05-31 青岛京润石化设计研究院有限公司 一种石油烃催化裂解反应方法和反应器
CN110240932B (zh) * 2018-05-29 2020-07-21 石宝珍 一种石油烃多级流化催化反应方法及反应器
CN110194967B (zh) * 2018-05-29 2021-07-23 青岛京润石化设计研究院有限公司 一种多产丙烯的催化反应再生方法

Patent Citations (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN1986505A (zh) * 2005-12-23 2007-06-27 中国石油化工股份有限公司 一种增产低碳烯烃的催化转化方法
CN101293806A (zh) * 2007-04-28 2008-10-29 中国石油化工股份有限公司 一种提高低碳烯烃产率的催化转化方法
CN102295510A (zh) * 2010-06-24 2011-12-28 中国石油化工股份有限公司 石脑油催化转化为低碳烯烃的方法
CN103131464A (zh) * 2011-11-23 2013-06-05 中国石油化工股份有限公司 一种生产低碳烯烃和轻芳烃的烃类催化转化方法
CN110240925A (zh) * 2018-05-29 2019-09-17 青岛京润石化设计研究院有限公司 一种流化催化裂化反应再生方法

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