WO2012041007A1 - 一种改善产物分布的催化转化方法 - Google Patents

一种改善产物分布的催化转化方法 Download PDF

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Publication number
WO2012041007A1
WO2012041007A1 PCT/CN2011/001613 CN2011001613W WO2012041007A1 WO 2012041007 A1 WO2012041007 A1 WO 2012041007A1 CN 2011001613 W CN2011001613 W CN 2011001613W WO 2012041007 A1 WO2012041007 A1 WO 2012041007A1
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Prior art keywords
catalyst
reaction
activity
oil
aging
Prior art date
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Ceased
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PCT/CN2011/001613
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English (en)
French (fr)
Inventor
许友好
崔守业
刘四威
姜楠
刘银亮
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
Original Assignee
Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Priority claimed from CN201010292906.7A external-priority patent/CN102417828B/zh
Priority claimed from CN201010292903.3A external-priority patent/CN102417827B/zh
Application filed by Sinopec Research Institute of Petroleum Processing , China Petroleum and Chemical Corp filed Critical Sinopec Research Institute of Petroleum Processing
Priority to JP2013529528A priority Critical patent/JP5947797B2/ja
Priority to RU2013119368/04A priority patent/RU2563637C2/ru
Priority to US13/825,975 priority patent/US9580664B2/en
Priority to KR1020137009245A priority patent/KR101672789B1/ko
Publication of WO2012041007A1 publication Critical patent/WO2012041007A1/zh
Anticipated expiration legal-status Critical
Ceased legal-status Critical Current

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Classifications

    • 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
    • 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
    • C10G11/187—Controlling or regulating
    • 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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions

Definitions

  • the present invention relates to a catalytic conversion process for improving product distribution, and more particularly to a catalytic conversion process for increasing the isobutylene content in liquefied gas and the olefin content in gasoline. Background technique
  • the conventional catalytic cracking process is mainly used for the production of gasoline, and the gasoline yield is as high as 50% by weight or more.
  • lead-free gasoline forced the catalytic cracking technology to develop high-octane gasoline.
  • the catalytic cracking process conditions and catalyst types have changed greatly.
  • the main purpose is to increase the reaction temperature, shorten the reaction time, increase the severity of the reaction, inhibit the hydrogen transfer reaction and the over-cracking reaction, and improve the contact efficiency of the oil and gas at the bottom of the riser.
  • the USY zeolite is combined with inertness. A catalyst of a matrix or an active matrix and a catalyst of a different type of zeolite composite.
  • the catalytic cracking technology has made the above progress, it satisfies the requirement of lead-free gasoline and improves the octane number of gasoline.
  • the current olefin content of the gasoline component is 35 - 65 wt%, which is far from the requirements of the new formula gasoline for olefin content.
  • the olefin content in the liquefied gas composition is higher, about 70% by weight, and the butene is several times that of isobutane, which is difficult to use as an alkylation raw material.
  • ZL99105904.2 discloses a catalytic conversion process for preparing isobutane and isoparaffin-rich gasoline by introducing the preheated feedstock oil into a reactor comprising two reaction zones in contact with a hot cracking catalyst.
  • the temperature of the first reaction zone is 530 ⁇ 620 ° C
  • the reaction time is 0.5 ⁇ 2.0 seconds
  • the temperature of the second reaction zone is 460 ⁇ 530 ° C
  • the reaction time is 2 ⁇ 30 seconds
  • the reaction product is separated, the catalyst is formed After being stripped into the regenerator, it is recycled and burned.
  • the liquefied gas obtained by the method provided by the invention has an isobutane content of 20 to 40% by weight, an isoparaffin content of the gasoline group composition of 30 to 45% by weight, and an olefin content of 30% by weight.
  • the research method has an octane number of 90 to 93 and a motor octane number of 80 to 84.
  • ZL99105905.0 discloses a catalytic conversion process for preparing propylene, isobutane and isoparaffin-rich gasoline by introducing the preheated feedstock oil into a reactor comprising two reaction zones, and thermally cracking Catalyst contact, the temperature of the first reaction zone is 550 ⁇ 650 °C, the reaction time is 0.5 ⁇ 2.5 seconds; the temperature of the second reaction zone is 480 ⁇ 550 °C, the reaction time is 2 ⁇ 30 seconds, the reaction product is separated, wait The biocatalyst is recycled after being stripped into the regenerator for charring.
  • the yield of liquefied gas obtained by the method provided by the invention can reach 25-40%, wherein the propylene content is about 30% by weight, the isobutane content is 20-40%, and the gasoline yield can reach 35-50. % by weight, the isoparaffin in the gasoline composition is 30 - 45 wt%.
  • ZL99105903.4 discloses a riser reactor for fluid catalytic conversion, which is a pre-lifting section which is coaxial with each other in the vertical direction from the bottom to the top, a first reaction zone, a second reaction zone having an enlarged diameter, The reduced diameter outlet zone has a horizontal tube at the end of the exit zone.
  • the reactor can control the process conditions of the first reaction zone and the second reaction zone differently, and can further crack the feedstock oil of different properties to obtain the desired product.
  • the present invention provides a catalytic conversion process for improving product distribution, wherein a high quality feedstock oil is reacted with a thermally regenerated catalyst having a lower activity (average activity) in a reactor to undergo a cracking reaction, a reaction product and a catalyst to be produced. Separation, the reaction product is sent to a separation system, and the spent catalyst is recycled after being stripped and regenerated.
  • the present invention provides a catalytic conversion process for improving product distribution
  • the Zhongyou shield feedstock oil reacts with the lower (reactive average) thermal regenerative catalyst in the lower part of the reactor, and the cracking reaction product and the carbon-containing catalyst ascend and undergo selective hydrogen transfer reaction and isomerization reaction.
  • the reaction product of the hydrogen transfer reaction and the isomerization reaction is separated from the catalyst to be produced, and the reaction product of the hydrogen transfer reaction and the isomerization reaction is sent to a separation system, and the catalyst to be produced is recycled after being stripped and regenerated.
  • the reactor used in the catalytic conversion process of the present invention refers to an industrial catalytic cracking unit, not a laboratory simulation unit.
  • the thermally regenerated catalyst having a lower activity (average activity) is added to or supplemented to an industrial catalytic converter for improving the distribution of products in industrial catalytic conversion processes, in particular for increasing the isobutene content and gasoline in the liquefied gas.
  • the reactor is selected from the group consisting of an equal diameter riser, a constant line riser, a variable diameter riser, a fluidized bed, or may be of equal diameter
  • a composite reactor consisting of a riser and a fluidized bed.
  • the variable diameter riser is a pre-lifting section that is coaxial with each other in the vertical direction from the bottom to the top, a first reaction zone, a second reaction zone having an enlarged diameter, and an outlet zone having a reduced diameter, in the exit zone.
  • a horizontal tube is connected to the end, wherein the ratio of the diameter of the second reaction zone to the diameter of the first reaction zone is 1.5 to 5.0:1.
  • the high quality feedstock oil is selected from one or more of atmospheric pressure overhead oil, gasoline, catalytic gasoline, diesel oil, samarium wax oil, and hydrogenated wax oil. .
  • the thermally regenerated catalyst has an activity (average activity) of from 35 to 55, preferably from 40 to 50.
  • the less active thermal regeneration catalyst has a relatively uniform activity profile.
  • the thermally regenerated catalyst having a relatively uniform activity profile means that the initial activity of the catalyst when added to the catalytic cracking unit does not exceed 80, preferably does not exceed 75, more preferably does not exceed 70, and the catalyst is self-balancing.
  • the time is from 0.1 hour to 50 hours, preferably from 0.2 to 30 hours, more preferably from 0.5 to 10 hours, and the equilibrium activity is from 35 to 60, preferably from 40 to 50.
  • the reaction conditions are: a reaction temperature of 450 to 620 V, preferably 500 to 600 ° C, a reaction time of 0.5 to 35.0 seconds, preferably 2.5 seconds to 15.0 seconds, a catalyst and a raw material.
  • the weight ratio of the oil is 3 ⁇ 15: 1, preferably 3 ⁇ 12: 1.
  • the cracking reaction conditions are: reaction temperature 490 ° C ⁇ 620 ° C, preferably 500 ° C ⁇ 600 ° C, reaction time 0.5 seconds ⁇ 2.0 seconds, preferably 0.8 seconds ⁇ 1.5 seconds, the weight ratio of catalyst to feedstock oil 3 ⁇ 15:1, preferably 3 - 12: 1.
  • the hydrogen transfer reaction and the isomerization reaction conditions are: a reaction temperature of 420 ° C to 550 ° C, preferably 460 ° C to 500, and a reaction time of 2 seconds to 30 seconds, preferably 3 seconds to 15 seconds.
  • the pressures of the cracking reaction, hydrogen transfer reaction and/or isomerization reaction in the first and second aspects are both 130 kPa to 450 kPa, and the weight ratio of water vapor to feedstock oil is 0.03 to 0.3:1.
  • the method provided by the present invention is embodied as follows:
  • preheated high-quality feedstock oil enters the reactor and the activity is 35 ⁇ 55, preferably 40 ⁇
  • Step (1) The pressure of the reaction is 130 kPa - 450 kPa, and the weight ratio of water vapor to feedstock oil (hereinafter referred to as water-oil ratio) is 0.03 to 0.3:1, preferably 0.05 to 0.3:1.
  • the method provided by the present invention is embodied as follows:
  • the preheated high-quality feedstock oil enters the reactor and is contacted with a thermally regenerated catalyst having an activity of 35 to 55, preferably 40 to 50, or a thermally regenerated catalyst having an activity of 35 to 55, preferably 40 to 50, and a relatively uniform activity distribution.
  • the cracking reaction occurs under the condition of 3 to 12:1;
  • the generated oil and gas and the used catalyst are ascending, at a reaction temperature of 420 ° C to 550 ° C, preferably 460 ° C to 500 ° C, and the reaction time is 2 seconds to 30 seconds, preferably 3 seconds to 15 seconds. Selective hydrogen transfer reaction and isomerization reaction occur;
  • the reaction product of the separation step (2) is obtained by obtaining a liquefied gas rich in isobutylene and a gasoline and other products having a moderate olefin content, and the catalyst to be produced is recycled by steaming into a regenerator for scorch regeneration.
  • Step (1) The cracking reaction, the pressure of the hydrogen transfer reaction and the isomerization reaction in the step (2) are both 130 kPa to 450 kPa, and the weight ratio of the water vapor to the feedstock oil (hereinafter referred to as the water-oil ratio) is 0.03 - 0.3 : 1 , preferably 0.05 ⁇ 0.3 : 1.
  • the process of the invention is particularly useful for increasing the isobutylene content of liquefied gases and the olefin content of gasoline.
  • the method provided by the present invention can be carried out in an equal diameter riser, a constant line riser or a fluidized bed reactor, wherein the equal diameter riser is the same as the conventional catalytic cracking reactor of the refinery, and the fluid in the line rate riser is equal
  • the line speeds are basically the same.
  • the equal-diameter riser and the equal-speed riser reactor are a pre-elevation section, a first reaction zone and a second reaction zone from bottom to top, and the fluidized bed reactor is a first reaction zone and a second reaction zone from bottom to top.
  • the ratio of the heights of the first reaction zone and the second reaction zone is 10 to 40:90 to 60.
  • one or more cold shock medium inlets are provided at the bottom of the second reaction zone, and/or a heat extractor is provided in the second reaction zone, The height of the heat extractor is 50% to 90% of the height of the second reaction zone.
  • the temperature and reaction time of each reaction zone were separately controlled.
  • the cold shock medium is a mixture of one or more selected from the group consisting of a cold shock agent, a cooled regenerated catalyst, and a cooled semi-regenerated catalyst.
  • the cold shock agent is a mixture of one or more selected from the group consisting of liquefied gas, crude gasoline, stabilized gasoline, diesel oil, heavy diesel oil or water;
  • the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be produced respectively After two stages of regeneration and a section of post-regeneration and cooling, the regenerated catalyst has a carbon content of 0.1% by weight or less, preferably 0.05% by weight or less, and a semi-regenerated catalyst having a carbon content of 0.1% by weight to 0.9% by weight, preferably having a carbon content of 0.15%.
  • the method provided by the present invention may also be carried out in a composite reactor consisting of an equal diameter riser and a fluidized bed, the lower equal diameter riser being the first reaction zone, the upper part
  • the fluidized bed is the second reaction zone, and the temperature and reaction time of each reaction zone are controlled separately.
  • One or more cold shock medium inlets are provided at the bottom of the fluidized bed, and/or a heat extractor is disposed in the second reaction zone, the height of the heat take-up being 50% to 90% of the height of the second reaction zone.
  • the temperature and reaction time of each reaction zone were separately controlled.
  • the cold shock medium is a mixture of one or more selected from the group consisting of a cold shock agent, a cooled regenerated catalyst, and a cooled semi-regenerated catalyst.
  • the cold shock agent is a mixture of one or more selected from the group consisting of liquefied gas, crude gasoline, stabilized gasoline, diesel oil, heavy diesel oil or water;
  • the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be produced respectively After two stages of regeneration and a section of post-regeneration and cooling, the regenerated catalyst has a carbon content of 0.1% by weight or less, preferably 0.05% by weight or less, and a semi-regenerated catalyst having a carbon content of 0.1% by weight to 0.9% by weight, preferably carbon content. It is 0. 15% by weight to 0.7% by weight.
  • the process provided by the present invention can also be carried out in a variable diameter riser reactor (see ZL99105903.4), the structural characteristics of which are shown in Figure 1: riser reaction In the vertical direction from bottom to top, the pre-lifting section a, the first reaction zone b, the second reaction zone c having an enlarged diameter, and the outlet zone d having a reduced diameter are connected in order from the bottom to the top, and a section is connected at the end of the exit zone.
  • Horizontal tube e The joint portion of the first and second reaction zones is in the shape of a truncated cone, and the apex angle cc of the isosceles trapezoid in the longitudinal section is 30.
  • the junction between the second reaction zone and the outlet zone is a truncated cone shape, and the base angle ⁇ of the isosceles trapezoid in the longitudinal section is 45° to 85°.
  • the sum of the heights of the pre-elevation section, the first reaction zone, the second reaction zone, and the outlet zone of the reactor is the total height of the reactor, generally from 10 m to 60 m.
  • the diameter of the pre-lift section is the same as that of a conventional equal-diameter riser reactor, typically 0.02 m to 5 m, which is 5% to 10% of the total height of the reactor.
  • the function of the pre-lift section is to move and accelerate the regenerated catalyst in the presence of a pre-lifting medium.
  • the pre-elevation medium used is the same as that used in conventional equal-diameter riser reactors, selected from water vapor or dry gas.
  • the structure of the first reaction zone is similar to that of a conventional equal-diameter riser reactor, and its diameter may be the same as that of the pre-lift section, or may be slightly larger than the pre-lift section, and the ratio of the diameter of the first reaction zone to the diameter of the pre-lift section is 1.0 - 2.0: 1 , its height accounts for 10% ⁇ 30% of the total height of the reactor.
  • the cracking reaction mainly occurs at a higher reaction temperature and ratio of solvent to oil, and a shorter residence time (generally 0.5 seconds to 2.5 seconds).
  • the second reaction zone is thicker than the first reaction zone, and the ratio of the diameter to the diameter of the first reaction zone is 1.5 to 5.0:1, and the height thereof is 30% to 60% of the total height of the reactor. Its role is to reduce the flow rate and reaction temperature of oil and gas and catalyst.
  • a method of lowering the reaction temperature in the region of 4 a cold shock medium may be injected from a joint portion of the region and the first reaction region, and/or a heat extractor may be disposed in the region to remove a portion of heat to lower the reaction temperature in the region. Thereby, the purpose of suppressing the secondary cracking reaction, increasing the isomerization reaction and the hydrogen transfer reaction is achieved.
  • the cold shock medium is a mixture of one or more selected from the group consisting of a cold shock agent, a cooled regenerated catalyst, and a cooled semi-regenerated catalyst.
  • the cold shock agent is a mixture of one or more selected from the group consisting of liquefied gas, crude gasoline, stabilized gasoline, diesel oil, heavy diesel oil or water;
  • the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be produced respectively After two stages of regeneration and a section of post-regeneration and cooling, the carbon content of the regenerated catalyst is 0.1% by weight or less, preferably 0.05% by weight or less, and the carbon content of the semi-regenerated catalyst is 0.1% by weight to 0.9% by weight, preferably the carbon content. 0.155% by weight ⁇ 0.7% by weight.
  • a heat extractor is provided, its height accounts for 50% to 90% of the height of the second reaction zone.
  • the residence time of the stream in the reaction zone can be long, from 1 second to 30 seconds.
  • the structure of the outlet zone is similar to the top outlet section of a conventional equal-diameter riser reactor.
  • the ratio of the diameter to the diameter of the first reaction zone is 0.8 to 1.5:1, and its height is 0 to 20% of the total height of the reactor.
  • the stream can be held in the zone for a certain period of time to inhibit the overcracking reaction and the thermal cracking reaction and increase the fluid flow rate.
  • One end of the horizontal pipe is connected to the outlet zone, and the other end is connected to the settler; when the height of the outlet zone is 0, that is, the riser reactor has no outlet zone, one end of the horizontal pipe is connected to the second reaction zone, and the other end is connected to the settler. .
  • the function of the horizontal pipe is to transport the product formed by the reaction and the catalyst to be produced to the separation system for gas-solid separation. The diameter is determined by those skilled in the art based on the specific circumstances.
  • the function of the pre-lifting section is to lift the regenerated catalyst into the first reaction zone in the presence of a pre-lifting medium.
  • the U.S. shield feedstock to which the process is applicable may be a petroleum fraction of a different boiling range.
  • the high-quality raw material oil is selected from one or more of an atmospheric pressure overhead oil, a gasoline oil, a catalytic gasoline, a diesel oil, a straight wax oil, and a hydrogenated wax oil.
  • the method can be applied to all catalysts of the same type, either an amorphous silica-alumina catalyst or a zeolite catalyst, and the active component of the zeolite catalyst is selected from the group consisting of Y-type zeolite, HY. a zeolite, an ultrastable Y zeolite, a ZSM-5 series zeolite or a mixture of one or more of a high silica zeolite, a ferrierite having a five-membered ring structure, which may contain rare earths and/or Phosphorus can also be free of rare earths and monuments.
  • different types of catalysts may also be employed in the process, and different types of catalysts may be catalysts having different particle sizes and/or catalysts having different apparent bulk densities.
  • the catalysts with different particle sizes and/or the active components on the catalyst with different apparent bulk densities are respectively selected from different types of zeolites.
  • the zeolite is selected from Y zeolite, HY zeolite, ultra stable Y zeolite, ZSM-5 series zeolite or has five One or more of a mixture of high silica zeolite and ferrierite having a ring structure, which may contain rare earth and/or phosphorus, or may contain no rare earth or phosphorus.
  • Catalysts of different particle sizes and/or catalysts of high and low apparent bulk density may enter different reaction zones, for example, a catalyst containing large particles of ultrastable Y-type zeolite enters the first reaction zone, increasing cracking reaction, containing rare earth Y-type The small particle catalyst of the zeolite enters the second reaction zone, increasing the hydrogen transfer reaction, and the catalysts of different particle sizes are stripped in the same stripper and regenerated in the same regenerator, and then separated into large The particulate and small particle catalyst, the small particle catalyst is cooled into the second reaction zone. Catalysts with different particle sizes are demarcated between 30 and 40 microns, and catalysts with different apparent bulk densities are demarcated between 0.6 and 0.7 g/cm 3 .
  • the less active catalyst useful in the process generally means a catalyst activity of from 35 to 55, preferably from 40 to 50.
  • a certain amount of a highly active catalyst e.g., fresh catalyst, or a catalyst having an activity greater than 60
  • the less active catalyst can be obtained in the reaction apparatus of the present invention by: reducing the catalyst replenishment rate of the apparatus (reducing the amount of replenishing catalyst); reducing the activity of the replenishing catalyst; or reducing the amount of the catalyst initially charged into the apparatus.
  • the less active catalyst may be treated by steam aging at a certain temperature (for example, 400-850 ° C) for a period of time (for example, 1 to 720 hours), or obtained by the following treatment methods 1, 2 or 3. .
  • the catalyst having a relatively uniform activity distribution as described in the present invention preferably means that the initial activity of the catalyst when added to the catalytic cracking unit is not more than 80, not more than 75, or not more than 70; the self-equilibration time of the catalyst is 0.1 hour - 50 hours, 0.2 ⁇ 30 hours, or 0.5 ⁇ 10 hours; balance activity is 35 - 60, or 40 ⁇ 50.
  • the catalyst having a relatively uniform activity distribution can be obtained by hydrothermal aging treatment. For example, it can be obtained by the following treatment methods 1, 2 and 3.
  • the catalyst activity (e.g., average activity, initial activity, equilibrium activity) is measured using prior art measurement methods.
  • the measurement methods in the prior art are: Enterprise Standard RIPP 92-90 Microreactor Activity Test Method for Catalytic Cracking "Petrochemical Analysis Method (RIPP Test Method)", Yang Cuiding et al., 1990, hereinafter referred to as RIPP 92-90.
  • the light oil micro-reverse device (refer to RIPP 92-90) is evaluated according to the following conditions: The catalyst is broken into particles with a diameter of 420 ⁇ 841 4, and the loading is 5 grams.
  • the reaction material is a distillation range of 235 ⁇ 337 °C. Straight-run light diesel oil, the reaction temperature is 460 ° C, the weight airspeed is 16 hours, the ratio of the agent to oil is 3.2.
  • the catalyst self-equilibration time refers to the catalyst at 800 ° C and 100% water vapor conditions. (Refer to RIPP 92-90) The time required for aging to reach equilibrium activity.
  • the thermally regenerated catalyst having a relatively uniform activity profile can be obtained by hydrothermal aging treatment.
  • hydrothermal aging treatment For example, it can be obtained by the following three methods:
  • the processing method 1 is embodied as follows:
  • the fresh catalyst is charged into the fluidized bed, preferably in the dense phase fluidized bed, water vapor is injected into the bottom of the fluidized bed, and the catalyst is fluidized by the action of water vapor, and the steam is aging the catalyst, and the aging temperature is 400°.
  • the apparent line speed of the fluidized bed is 0.1 m / sec - 0.6 m / sec, preferably 0.15 m /second - 0.5 m / sec, after aging for 1 hour - 720 hours, preferably 5 hours - 360 hours, the catalyst having a relatively uniform activity is obtained.
  • the catalyst having a relatively uniform activity is added to the regenerator of the industrial catalytic cracking unit as required in the industrial catalytic cracking unit to obtain a thermally regenerated catalyst having a relatively uniform activity distribution.
  • the catalyst is charged into a fluidized bed, preferably a dense phase fluidized bed, and a mixture of water vapor and other aging medium is injected at the bottom of the fluidized bed, and the catalyst is fluidized by a mixture of water vapor and other aging medium, and at the same time, water
  • the catalyst is aged with a mixture of steam and other aging medium at an aging temperature of from 400 ° C to 850 ° C, preferably from 500 ° C to 750 ° C, preferably from 600 ° C to 700 ° C, of the apparent line of the fluidized bed.
  • the speed is 0.1 m / sec - 0.6 m / sec, preferably 0.15 m / s - 0.5 m / sec
  • the weight ratio of water vapor to other aged shield is 0.20-0.9, preferably 0.40-0.60
  • aging 1 hour - 720 hours preferably 5 hours - 360 hours, to obtain a catalyst having a relatively uniform activity, and a catalyst having a relatively uniform activity according to the requirements of an industrial device.
  • the thermally regenerated catalyst having a relatively uniform activity distribution is obtained by being added to a regenerator of an industrial catalytic cracking unit.
  • the other aging medium includes air, dry gas, regenerated flue gas, air or dry gas burned gas or air and combustion oil burned gas, or other gases such as nitrogen.
  • the weight ratio of the water vapor to the aged medium is from 0.2 to 0.9, preferably from 0.40 to 0.60.
  • the fresh catalyst is conveyed to a fluidized bed, preferably a dense phase fluidized bed, while the hot regenerated catalyst of the regenerator is transferred to another fluidized bed for solid-solid heat exchange between the two fluidized beds.
  • a fluidized bed preferably a dense phase fluidized bed
  • the fresh catalyst is fluidized by a mixture of steam or water vapour and other aging medium, while steam or water vapour
  • the fresh catalyst is aged with a mixture of other aged media, and the aging temperature is 400 ° C - 850 ° C, preferably 500 ° C - 750 ° C, preferably 600 ° C - 70 (TC , apparent line of the fluidized bed
  • the speed is from 0.1 m/s to 0.6 m/s, preferably from 0.15 m/s to 0.5 m/s, aged from 1 hour to 720 hours, preferably from 5 hours to 360 hours, in a mixture of water vapor and other aging medium.
  • the weight ratio is more than 0-4, preferably 0.5-1.5, to obtain an aging catalyst with relatively uniform activity, and the aging catalyst is added to the industrial catalysis according to the requirements of the industrial catalytic cracking unit.
  • the steam after the cracking step enters the reaction system (as one of steam, anti-coke steam, atomized steam, and elevated steam, or one of several strippers, settlers, and raw material nozzles respectively entering the catalytic cracking unit)
  • the pre-elevation section) or the regeneration system, and the mixture of water vapor and other aging medium after the aging step enters the regeneration system, and the regenerated catalyst after the heat exchange is returned to the regenerator.
  • the other aging medium includes air, dry gas, regeneration Flue gas, air or dry gas burning gas or air and combustion oil burning gas, or other gases such as nitrogen.
  • the regenerative flue gas may come from the device or from other devices.
  • the reactor has the advantages of maintaining a higher reaction temperature and ratio of the agent to the oil at the bottom of the conventional riser reactor to increase the primary cracking reaction while suppressing the overcracking and thermal cracking reactions at the top. Further, the reaction time is prolonged at a lower reaction temperature in the upper portion of the reactor to increase the isomerization reaction and hydrogen transfer reaction of the olefin.
  • the content of isobutylene in the liquefied gas produced by the method of the present invention is increased by more than 30%.
  • the olefin content in the gasoline family composition can be increased to more than 30% by weight.
  • Figure 1 is a schematic view of a riser reactor, where a, b, c, d, and e represent a pre-elevation section, a first reaction zone, a second reaction zone, an outlet zone, and a horizontal pipe, respectively.
  • FIG. 2 is a schematic flow chart of a preferred embodiment of the second aspect of the present invention.
  • the numbers in the drawings are as follows:
  • 1, 3, 4, 6, 1 1, 13, 17, 18 are all pipelines; 2 is the pre-lift section of the riser; 5, 7 are the first reaction zone and the second reaction zone of the riser; The outlet area of the tube; 9 is the settler, 10 is the cyclone separator, 12 is the stripper, M is the inclined tube, 15 is the regenerator, and 16 is the regenerative inclined tube.
  • the invention has different embodiments, such as
  • the preheated feedstock oil is contacted with a less reactive hot regenerated catalyst or with a less reactive and thermally distributed catalyst with a relatively lower activity distribution.
  • the resulting oil and gas are used.
  • the catalyst is contacted with the regenerated catalyst injected into the cooling, followed by the isomerization reaction and the hydrogen transfer reaction, and the effluent enters the settler after the reaction; the reaction product is separated, and the catalyst to be produced is divided into two parts by steam stripping and regeneration, wherein One part enters the bottom of the reactor and the other part enters the lower middle of the reactor after cooling.
  • Embodiment 2 Embodiment 2:
  • the preheated feedstock oil is contacted with a less reactive hot regenerated catalyst or with a less reactive and thermally distributed catalyst with a relatively lower activity distribution.
  • the resulting oil and gas are used.
  • the catalyst is contacted with the cold-injecting agent and the cooled semi-regenerated catalyst, followed by the isomerization reaction and the hydrogen transfer reaction, and the effluent enters the settler after the reaction; the reaction product is separated, and the catalyst is stripped and then enters into two
  • the scintillator is scorched, and the semi-regenerated catalyst from the first stage regenerator is cooled to enter the lower middle of the reactor, and the regenerated catalyst from the second stage regenerator is directly returned to the bottom of the reactor without cooling.
  • the preheated conventional cracking feedstock enters from the lower portion of the riser to a less active heat regenerated catalyst or to a lower activity and a lower activity
  • the regenerated catalyst is contacted, and the oil generated after the reaction rises to the top of the riser, and continues to react with the catalyst after the temperature drop, and the effluent enters the settler after the reaction; the reaction product is separated, and the catalyst to be produced is divided into two after being stripped and regenerated. In part, one part enters the lower part of the riser and the other part enters the top of the riser after cooling.
  • Embodiment 4 is a diagrammatic representation of Embodiment 4:
  • This embodiment is a preferred embodiment of the invention.
  • the preheated conventional cracking feedstock enters from the lower portion of the first reaction zone of the reactor with a less active thermal regenerated catalyst or with a lower activity and a relatively uniform activity distribution.
  • a cracking reaction occurs, and the oil generated after the reaction rises to the lower portion of the second reaction zone of the reactor to contact the cooled catalyst for hydrogen transfer reaction and isomerization reaction, and the effluent enters the settler after the reaction;
  • the product, the catalyst to be produced is stripped, regenerated and then passed to the lower part of the second reaction zone.
  • the method provided by the present invention is not limited to this.
  • Fig. 2 is a flow chart of a catalytic conversion method for increasing the content of isobutylene and gasoline olefins in a liquefied gas by using a variable diameter riser reactor.
  • the shape and size of the equipment and piping are not limited by the drawings, but are determined according to specific conditions.
  • the pre-lifting steam enters from the riser pre-lift section 2 via line 1 and the lower activity heat a relatively uniform distribution of thermally regenerated catalyst via a regenerative inclined tube
  • the pre-lift section is lifted by pre-lift steam.
  • the preheated feedstock oil enters the preheating section of the riser through the pipeline 4 and the atomized steam from the pipeline 3, and is mixed with the hot catalyst to enter the first reaction zone 5, and the cracking reaction is carried out under certain conditions.
  • the reactant stream is mixed with a cold shock agent from line 6 and/or a cooled catalyst (not shown) into second reaction zone 7 for a second reaction, and the reacted stream enters outlet zone 8, which improves the flow.
  • the line speed causes the reactant stream to rapidly enter the settler 9 and the cyclone separator 10, and the reaction product is separated from the system via line 11.
  • the charcoal-containing catalyst enters the stripper 12, is stripped by the water vapor from the pipeline 13, and then enters the regenerator 15 by the inclined tube 14 to be produced.
  • the catalyst to be produced is scorched and regenerated in the air from the pipeline 17, and the smoke is regenerated.
  • the gas exits the regenerator via line 18, and the hot regenerated catalyst is returned to the bottom of the riser via the regeneration ramp 16 for recycling.
  • the invention is further illustrated by the following examples, which are not intended to limit the invention.
  • the properties of the feedstock oil and catalyst used in the examples and comparative examples are shown in Tables 1 and 2, respectively.
  • the catalysts in Table 2 are all produced by Qilu Catalyst Factory of China Petrochemical Corporation.
  • the ZCM-7 catalysts in Table 2 were aged at 800 ° C and 100% water vapor for 12 hours and 30 hours, respectively, to obtain two different activity levels of ZCM-7, ie, activities of 67 and 45; likewise, CGP in Table 2
  • the -1 catalyst was aged at 800 ° C, 100% water vapor for 12 hours and 30 hours, respectively, to obtain two different levels of activity of CGP-1, ie, activities of 62 and 50.
  • This example illustrates the use of the method provided by the present invention to increase the isobutylene content and the gasoline olefin content of the liquefied gas in a medium-sized variable-diameter riser reactor using different activity levels of the catalyst.
  • the total height of the pre-lift section, the first reaction zone, the second reaction zone, and the exit zone of the reactor is the total height of the pre-lift section, the first reaction zone, the second reaction zone, and the exit zone of the reactor.
  • the pre-lift section has a diameter of 0.025 m and a height of 1.5 m;
  • the first reaction zone has a diameter of 0.025 m and a height of 4 m;
  • the second reaction zone has a diameter of 0.1 m and a height of 6.5 m;
  • the diameter is 0.025 meters, and the height is 3 meters;
  • the longitudinal section of the first and second reaction zone joints has an apex angle of 45°; the second reaction zone and the exit zone have a longitudinal section of the isosceles trapezoid It is 60°.
  • the preheated feedstock B listed in Table 1 enters the reactor and is contacted with the hot catalyst ZCM-7 listed in Table 2 in the presence of steam.
  • the ZCM-7 catalyst activity is 45.
  • the reaction product is separated to obtain liquefied gas and gasoline and other products, and the catalyst to be produced is stripped into a regenerator, and the regenerated catalyst is recycled after being charred.
  • Example 2 The reactor type and operating conditions were exactly the same as in Example 1.
  • the feedstock oil used was also the feedstock B listed in Table 1, and the catalyst was also the catalyst ZCM-7 listed in Table 2, except that the ZCM-7 catalyst activity was 67 at this time. .
  • the operating conditions of the test, product distribution and properties of the gasoline are listed in Table 3.
  • the pre-lift section of the reactor, the first reaction zone, the second reaction zone, and the exit zone have a total height of 15 meters, the pre-lift section has a diameter of 0.025 meters, and the height is 1.5 meters; the first reaction zone has a diameter of 0.025.
  • the second reaction zone is 0.1 meters in diameter and its height is 6.5 meters; the diameter of the exit zone is 0.025 meters, and its height is 3 meters; the longitudinal section of the first and second reaction zone joints is isosceles The apex angle of the trapezoid is 45°; the longitudinal section of the joint portion of the second reaction zone and the exit zone has a base angle of 60°.
  • the preheated feedstocks listed in Table 1 entered the reactor and were contacted with the hot catalyst CGP-1 listed in Table 2 in the presence of steam.
  • the CGP-1 catalyst activity was 50 and the reaction product was isolated. Liquefied gas and gasoline and other products, the catalyst to be produced is stripped into the regenerator, and the regenerated catalyst is recycled after being charred.
  • the reactor type and operating conditions were identical to those in Example 2.
  • the feedstock oil used was also the feedstock oil listed in Table 1, and the catalyst was also the catalyst CGP-1 listed in Table 2, except that the CGP-1 catalyst activity was 62. .
  • the operating conditions of the test, product distribution and properties of the gasoline are listed in Table 4.
  • Table 4 the yield of isobutene increased from 3.0% by weight to 4.1% by weight with respect to the use of highly active CGP-1 (i.e., activity of 62), using low activity CGP-1 (i.e., activity of 50).
  • the gasoline olefin content increased from 18.2% to 27.9% by weight; in addition, the liquid yield still increased by 0.8%.
  • This example illustrates the use of the method of the present invention to increase the isobutene content and gasoline olefin content of a liquefied gas in a medium-sized variable-diameter riser reactor using different types of catalytic cracking feedstock oils.
  • the reactor, catalyst type, and catalyst activity used in this example were the same as those in Example 2 except that the feedstock oils were the feedstocks A and C listed in Table 1, respectively.
  • This example illustrates the use of the method provided by the present invention to increase the isobutylene content and the gasoline olefin content of the liquefied gas in a medium-sized variable-diameter riser reactor using different activity levels of the catalyst.
  • the pre-lift section, the first reaction zone, the second reaction zone, and the exit zone of the reactor have a total height of 15 meters, the pre-lift section has a diameter of 0.025 meters, and the height is 1.5 meters; the diameter of the first reaction zone is 0.025 meters, and the height thereof is 4 m; the second reaction zone has a diameter of 0.1 m and a height of 6.5 m; the outlet zone has a diameter of 0.025 m and a height of 3 m; the longitudinal section of the first and second reaction zone joints has an isosceles trapezoidal apex angle Is 45.
  • the longitudinal section of the junction between the second reaction zone and the outlet zone has a base angle of 60°.
  • the preheated feedstock B enters the reactor and is contacted with a hot catalyst ZCM-7 in the presence of water vapor.
  • the catalyst activity (average activity) is 45, and the reaction product is separated to obtain liquefied gas, gasoline and other products.
  • the catalyst to be produced is stripped into the regenerator, and the regenerated catalyst is recycled after being charred.
  • the ZCM-7 catalyst added to the apparatus is hydrothermally treated with fresh ZCM-7 (catalyst hydrothermal treatment method is treated by the catalyst treatment method 1 of the present invention: dense phase fluidized bed, aging temperature 650 ° C, apparent line of fluidized bed
  • the catalyst after the speed of 0.30 m / s, 100% water vapor, aging time 31 hours), the initial activity of 75, and then mixed with the equilibrium catalyst in the device, and then hydrothermally aged in the device, the added catalyst reaches the device
  • the self-equilibration time required for the catalyst balance activity at 45 (800 °C, 100% water) Vapour) is 30 hours.
  • the reactor type and operating conditions were identical to those in Example 5.
  • the feedstock oil used was also the feedstock B listed in Table 1, and the catalyst was also the catalyst ZCM-7 listed in Table 2, and the average catalyst activity was also 45.
  • the ZCM-7 catalyst added to the unit is a fresh ZCM-7 catalyst. It has an initial activity of 91 without hydrothermal treatment. It is mixed with the equilibrium catalyst in the unit and then hydrothermally aged in the unit until the catalyst in the unit. The equilibrium activity was 45.
  • the operating conditions of the test, product distribution and properties of the gasoline are listed in Table 6.
  • the dry gas yield decreased from 1.7 wt% to 1.5 wt%
  • the coke yield decreased from 3.2 wt% to 2.7 wt%
  • the liquid yield increased from 89.3 wt% to 89.8 wt%, an increase of 0.5 percentage points.
  • the isobutene yields of the two are substantially equivalent to the gasoline olefin content.
  • This example illustrates the use of the method provided by the present invention to increase the isobutylene content and the gasoline olefin content of the liquefied gas in a medium-sized variable-diameter riser reactor using different activity levels of the catalyst.
  • the pre-lift section, the first reaction zone, the second reaction zone and the exit zone of the reactor have a total height of 15 m, the pre-lift section has a diameter of 0.025 m and a height of 1.5 m; the first reaction zone has a diameter of 0.025 m and its height. 4 m; the second reaction zone has a diameter of 0.1 m and a height of 6.5 m; the outlet zone has a diameter of 0.025 m and a height of 3 m; the longitudinal section of the first and second reaction zone joints has an isosceles trapezoidal apex angle The angle of the isosceles trapezoid of the longitudinal section of the joint portion of the second reaction zone and the outlet zone is 60°.
  • the preheated feedstock oil B enters the reactor and is contacted with the hot catalyst CGP-1 in the presence of steam.
  • the average activity of the CGP-1 catalyst is 50, and the reaction product is separated to obtain liquefied gas, gasoline and other products.
  • the raw catalyst is stripped into the regenerator, and the regenerated catalyst is recycled after being charred.
  • the CGP-1 catalyst added to the apparatus is a hydrothermally treated catalyst of fresh CGP-1 (the catalyst hydrothermal treatment method is treated by the catalyst treatment method 1 of the present invention: dense phase fluidized bed, aging temperature 670 ° C, fluidized bed Apparent line speed 0.30m / s, 100% water vapor, aging time 28 hours), its initial activity is 72, and then mixed with the equilibrium catalyst in the device, and then hydrothermally aged in the device, the added catalyst reaches the device Self-equilibration time required for the catalyst to balance activity 50 (800 °C, 100% water vapor) It is 40 hours.
  • the catalyst hydrothermal treatment method is treated by the catalyst treatment method 1 of the present invention: dense phase fluidized bed, aging temperature 670 ° C, fluidized bed Apparent line speed 0.30m / s, 100% water vapor, aging time 28 hours
  • its initial activity is 72
  • the added catalyst reaches the device Self-equilibration time required for the catalyst to balance activity 50 (800 °C, 100% water vapor) It
  • the reactor type and operating conditions were identical to those in Example 6.
  • the feedstock oil used was also the feedstock B listed in Table 1, and the catalyst was also the catalyst listed in Table 2, CGP-1, and the average activity of the CGP-1 catalyst was also 50.
  • the CGP-1 catalyst added to the unit is a fresh CGP-1 catalyst. It has an initial activity of 95 without hydrothermal treatment. It is mixed with the equilibrium catalyst in the unit and then hydrothermally aged in the unit until the catalyst in the unit. The equilibrium activity is 50.
  • the operating conditions of the test, product distribution and properties of the gasoline are listed in Table 7.
  • the dry gas yield decreased from 2.0% to 1.9% by weight
  • the coke yield decreased from 3.0% to 2.55%
  • the liquid yield increased from 88.7% to 89.3. /. , an increase of 0.6 percentage points.
  • the isobutene yields of the two are substantially equivalent to the gasoline olefin content.
  • This example illustrates the use of the method provided by the present invention to improve product distribution using catalysts of different activity levels and medium conventional equal diameter riser reactors.
  • the preheated feedstock B listed in Table 1 enters the reactor and is contacted with the hot catalyst ZCM-7 listed in Table 2 in the presence of steam.
  • the ZCM-7 catalyst activity is 45, and the reaction product is isolated. Liquefied gas and gasoline and other products, the catalyst to be produced is stripped into the regenerator, and the regenerated catalyst is recycled after being charred.
  • Example 7 The reactor type and operating conditions were exactly the same as in Example 7.
  • the feedstock oil used was also the feedstock B listed in Table 1.
  • the catalyst was also the catalyst ZCM-7 listed in Table 2, except that the ZCM-7 catalyst activity was 67 at this time. .
  • the operating conditions of the test, the product distribution and the properties of the gasoline are listed in Table 8.
  • ZCM-7 catalyst activity 45 67 reaction temperature, °C
  • First reaction zone 550 550 Second reaction zone 500 500 Residence time, seconds 5.5 5.5 First reaction zone 2.0 2.0 Second reaction zone 3.5 3.5 Oil to oil ratio 5.0 5.0 Water to oil ratio 0.1 0. 1 Product distribution, weight %
  • First reaction zone 550 550
  • Second reaction zone 505 505
  • Agent oil ratio 6.0
  • Reaction time seconds
  • First reaction zone 1.3
  • Second reaction zone 4.7
  • Water to oil ratio 0.1

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Description

一种改善产物分布的催化转化方法 技术领域
本发明涉及改善产物分布的催化转化方法， 更具体地说， 是属于 提高液化气中异丁烯含量和汽油中烯烃含量的催化转化方法。 背景技术
催化裂化自 20世纪 40年代诞生以来， 一直是最主要的重质油轻 质化过程。 原因一是其原料来源广泛， 可以采用蜡油， 还可以采用常 压渣油、 减压渣油的脱沥青油或者部分掺入减压渣油； 二是其产品方 案灵活， 可以是燃料型， 也可以是燃料化工型， 如多产汽油、 多产柴 油、 多产丙烯等； 三是其产品性质可以通过催化剂配方的调整和工艺 参数的变化进行相应的调整， 如提高汽油辛烷值、 降低汽油烯烃含量 等。
常规的催化裂化工艺主要用于生产汽油， 汽油产率高达 50重％以 上。 20世纪 80年代初， 汽油无铅化迫使催化裂化技术向生产高辛烷值 汽油的方向发展， 为此， 催化裂化的工艺条件和催化剂类型发生了很 大变化。 在工艺方面， 主要是提高反应温度、 缩短反应时间、 提高反 应苛刻度、 抑制氢转移反应和过裂化反应和改善提升管底部油气和催 化剂的接触效率； 在催化剂方面， 开发了 USY型沸石结合惰性基质或 活性基质的催化剂以及不同类型的沸石复合的催化剂。
催化裂化技术虽已取得上述进展， 满足了汽油无铅化的要求， 提 高了汽油的辛烷值， 但无论是通过改变工艺条件， 还是使用新型的沸 石催化剂来提高汽油辛烷值， 都是以提高汽油组分中的烯烃含量来增 加汽油的辛烷值， 目前汽油组分中烯烃含量为 35 - 65重％, 这与新配 方汽油对烯烃含量的要求相差甚远。 液化气组成中烯烃含量更高， 大 约在 70重％左右， 其中丁烯是异丁烷的数倍， 难以作为烷基化原料。
ZL99105904.2 公开了一种制取异丁烷和富含异构烷烃汽油的催化 转化方法， 是将预热后的原料油进入一个包括两个反应区的反应器内， 与热的裂化催化剂接触， 第一反应区的温度为 530 ~ 620 °C、 反应时间 为 0.5 ~ 2.0秒； 第二反应区的温度为 460 ~ 530 °C、 反应时间为 2 ~ 30 秒， 分离反应产物， 待生催化剂经汽提进入再生器烧焦后循环使用。 采用该发明提供的方法制取的液化气中异丁烷含量为 20 - 40重％, 汽 油族组成中的异构烷烃含量为 30 ~ 45重％, 烯烃含量降低到 30重°/。以 下， 其研究法辛烷值为 90 ~ 93 , 马达法辛烷值为 80 ~ 84。
ZL99105905.0 公开了一种制取丙烯、 异丁烷和富含异构烷烃汽油 的催化转化方法， 是将预热后的原料油进入一个包括两个反应区的反 应器内， 与热的裂化催化剂接触， 笫一反应区的温度为 550 ~ 650 °C、 反应时间为 0.5 ~ 2.5秒； 第二反应区的温度为 480 ~ 550 °C、 反应时间 为 2 ~ 30 秒， 分离反应产物， 待生催化剂经汽提进入再生器烧焦后循 环使用。 采用本发明提供的方法制取的液化气产率可达 25 ~ 40重％ , 其中丙烯含量为 30重％左右， 异丁烷含量为 20 ~ 40重％, 汽油的产 率可达 35 ~ 50重％, 汽油组成中的异构烷烃为 30 - 45重％。
ZL99105903.4公开了一种用于流化催化转化的提升管反应器， 沿 垂直方向从下至上依次为互为同轴的预提升段、 第一反应区、 直径扩 大了的第二反应区、 直径缩小了的出口区， 在出口区末端有一水平管。 该反应器既可以控制第一反应区和第二反应区的工艺条件不同， 又可 以使不同性能的原料油进行分段裂化， 得到所需目的产品。
正是这些专利， 构成了多产异构烷烃的催化裂化工艺 （MIP )的基 础专利， 并得到广泛的应用， 目前已应用到近 50套催化裂化装置， 取 得巨大的经济效益和社会效益。 尽管现有技术可以得到富含异丁烷的 液化气和富含异构烷烃汽油， 但对处理优质的催化裂化原料油， 尤其 是加氢蜡油而言， 所得的汽油烯烃含量偏低， 液化气中的异丁烯含量 偏低， 产物分布不够优化， 石油资源未充分利用。 发明内容
本发明的目的是提供一种改善产物分布的催化转化方法， 特别是 提高液化气中的异丁烯含量， 同时提高汽油中烯烃含量， 降低干气和 焦炭的产率。
在第一方面， 本发明提供一种改善产物分布的催化转化方法， 其 中优质原料油与活性 （平均活性） 较低的热再生催化剂在反应器内接 触发生裂化反应， 将反应产物和待生催化剂分离， 该反应产物被送入 分离系统， 该待生催化剂经汽提、 再生后循环使用。
在第二方面， 本发明提供一种改善产物分布的催化转化方法， 其 中优盾原料油与活性 （平均活性） 较低的热再生催化剂在反应器的下 部接触发生裂化反应， 裂化反应产物和含炭的催化剂上行并且发生选 择性的氢转移反应和异构化反应， 将氢转移反应和异构化反应的反应 产物和待生催化剂分离， 氢转移反应和异构化反应的反应产物被送入 分离系统， 该待生催化剂经汽提、 再生后循环使用。
用于本发明的催化转化方法中的反应器是指工业催化裂化装置， 并非实验室的模拟装置。 换言之， 该活性 （平均活性） 较低的热再生 催化剂是加入到或补充到工业催化转化装置中， 用于改善工业催化转 化方法中的产物分布， 特别是用于提高液化气中异丁烯含量和汽油中 烯烃含量。
在第一方面和第二方面的一些实施方案中， 所述反应器选自等直 径提升管、 等线速提升管、 变径提升管、 流化床中的一种， 也可以是 由等直径提升管和流化床构成的复合反应器。 优选地， 所述变径提升 管沿垂直方向从下至上依次为互为同轴的预提升段、 第一反应区、 直 径扩大了的第二反应区、 直径缩小了的出口区， 在出口区末端连有一 段水平管， 其中第二反应区的直径与第一反应区的直径之比为 1.5 ~ 5.0: 1。
在第一方面和第二方面的一些实施方案中， 所述优质原料油选自 常压塔顶油、 汽油、 催化汽油、 柴油、 直镏蜡油、 加氢蜡油中的一种 或多种。
在第一方面和第二方面的一些实施方案中， 所述热再生催化剂活 性 (平均活性） 为 35 ~ 55 , 优选 40 ~ 50。
在第一方面和第二方面的一些实施方案中， 所述活性较低的热再 生催化剂具有相对均勾的活性分布。 在进一步的一些实施方案中， 所 述活性分布相对均匀的热再生催化剂是指加入到催化裂化装置内时催 化剂初始活性不超过 80 , 优选不超过 75 , 更优选不超过 70 , 该催化剂 的自平衡时间为 0.1小时 ~ 50小时， 优选 0.2 ~ 30小时， 更优选 0.5 ~ 10小时， 平衡活性为 35 ~ 60， 优选 40 ~ 50。
在第一方面的一些实施方案中， 所述反应条件为： 反应温度 450 - 620 V , 优选 500 °C ~ 600 °C , 反应时间 0.5秒 ~ 35.0秒， 优选 2.5 秒〜 15.0秒， 催化剂与原料油的重量比 3 ~ 15: 1， 优选 3 ~ 12: 1。
在第二方面的一些实施方案中， 所述裂化反应条件为： 反应温度 490 °C ~ 620 °C, 优选 500°C ~ 600°C , 反应时间 0.5秒 ~ 2.0秒， 优选 0.8 秒〜 1.5秒， 催化剂与原料油的重量比 3 ~ 15:1, 优选 3 - 12:1。
在第二方面的一些实施方案中， 所述氢转移反应和异构化反应条 件为：反应温度 420°C ~ 550°C,优选 460 °C ~ 500 ,反应时间为 2秒 ~ 30秒， 优选 3秒 ~ 15秒。
在第一和第二方面中的所述裂化反应、 氢转移反应和 /或异构化反 应的压力均为 130 kPa ~450kPa, 水蒸汽与原料油的重量比为 0.03 ~ 0.3:1。
在第一方面， 本发明提供的方法是这样具体实施的：
( 1 ) 、 预热的优质原料油进入反应器与活性为 35 ~ 55优选 40 ~
50的热再生催化剂或者活性为 35 ~ 55优选 40 ~ 50且活性分布相对均 匀的热再生催化剂接触，在反应温度 490 °C ~ 620°C优选 500 °C ~ 600 °C， 反应时间 0.5秒 - 35.0秒优选 2.5秒 ~ 15.0秒， 催化剂与原料油的重量 比 （以下简称剂油比） 3~ 15:1优选 3~ 12:1的条件下发生反应；
(2) 、 将生成的反应油气和待生催化剂分离；
(3 ) 、 分离反应油气得到富含异丁烯的液化气和烯烃含量适中的 汽油及其它反应产物， 待生催化剂经汽提进入再生器烧焦再生后循环 使用。
步骤 （ 1 ) 所述反应的压力为 130kPa - 450kPa, 水蒸汽与原料油 的重量比 （以下简称水油比） 为 0.03 ~0.3:1, 最好为 0.05 ~0.3:1。
在第二方面， 本发明提供的方法是这样具体实施的：
( 1 ) 、 预热的优质原料油进入反应器与活性为 35 ~ 55优选 40 ~ 50的热再生催化剂或者活性为 35 ~ 55优选 40 ~ 50且活性分布相对均 匀的热再生催化剂接触，在反应温度 490 °C - 620°C优选 500 °C ~ 600 °C , 反应时间 0.5秒 ~ 2.0秒优选 0.8秒 - 1.5秒， 催化剂与原料油的重量比 (以下简称剂油比） 3 ~ 15:1优选 3 ~ 12:1的条件下发生裂化反应；
(2)、 生成的油气和用过的催化剂上行， 在反应温度 420°C ~550 °C优选 460°C ~ 500 °C ,反应时间为 2秒 ~ 30秒优选 3秒 ~ 15秒的条件 下发生选择性的氢转移反应和异构化反应；
(3 ) 、 分离步骤 （2 ) 的反应产物得到富含异丁烯的液化气和烯 烃含量适中的汽油及其它产品， 待生催化剂经汽提进入再生器烧焦再 生后循环使用。 步骤 （ 1 ) 所述裂化反应、 步骤 （2 ) 所述氢转移反应和异构化反 应的压力均为 130 kPa ~ 450kPa, 水蒸汽与原料油的重量比 （以下简 称水油比） 为 0.03 - 0.3 : 1 , 最好为 0.05 ~ 0.3 : 1。
本发明的方法特别适用于提高液化气中异丁烯含量和汽油中烯烃 含量。 本发明提供的方法可以在等直径提升管、 等线速提升管或流化 床反应器中进行， 其中等直径提升管与炼厂常规的催化裂化反应器相 同， 等线速提升管中流体的线速基本相同。 等直径提升管、 等线速提 升管反应器从下至上依次为预提升段、 第一反应区、 第二反应区， 流 化床反应器从下至上依次为第一反应区、 第二反应区， 第一反应区、 第二反应区的高度之比为 10 ~ 40:90 ~ 60。 当使用等直径提升管、 等线 速提升管或流化床反应器时， 在笫二反应区底部设一个或多个冷激介 质入口， 和 /或在第二反应区内设置取热器， 取热器的高度占第二反应 区高度的 50% ~ 90%。分别控制每个反应区的温度和反应时间。冷激介 质是选自冷激剂、 冷却的再生催化剂和冷却的半再生催化剂中的一种 或一种以上的任意比例的混合物。 其中冷激剂是选自液化气、 粗汽油、 稳定汽油、 柴油、 重柴油或水中的一种或一种以上的任意比例的混合 物； 冷却的再生催化剂和冷却的半再生催化剂是待生催化剂分别经两 段再生和一段再生后冷却得到的， 再生催化剂碳含量为 0.1重％以下， 最好为 0.05重％以下， 半再生催化剂碳含量为 0.1 重％~ 0.9重％, 最 好碳含量为 0.15重％~ 0.7重％。
在第一和第二方面的一些方案中， 本发明提供的方法也可以在由 等直径提升管和流化床构成的复合反应器中进行， 下部的等直径提升 管为第一反应区， 上部的流化床为第二反应区， 分别控制每个反应区 的温度和反应时间。 在流化床的底部设一个或多个冷激介质入口， 和 / 或在第二反应区内设置取热器， 取热器的高度占第二反应区高度的 50% ~ 90%。 分别控制每个反应区的温度和反应时间。 冷激介质是选自 冷激剂、 冷却的再生催化剂和冷却的半再生催化剂中的一种或一种以 上的任意比例的混合物。 其中冷激剂是选自液化气、 粗汽油、 稳定汽 油、 柴油、 重柴油或水中的一种或一种以上的任意比例的混合物； 冷 却的再生催化剂和冷却的半再生催化剂是待生催化剂分别经两段再生 和一段再生后冷却得到的， 再生催化剂碳含量为 0.1重％以下， 最好为 0.05重％以下， 半再生催化剂碳含量为 0.1重％~ 0.9重％， 最好碳含量 为 0. 15重％~ 0.7重％。
在笫一和第二方面的一些方案中， 本发明提供的方法还可以在变 径提升管反应器 （参见 ZL99105903.4 ) 中进行， 该反应器的结构特征 如图 1 所示： 提升管反应器沿垂直方向从下至上依次为互为同轴的预 提升段 a、 第一反应区 b、 直径扩大了的第二反应区 c、 直径缩小了的 出口区 d, 在出口区末端连有一段水平管 e。 第一、 二反应区的结合部 位为圆台形， 其纵剖面等腰梯形的顶角 cc为 30。 ~ 80。 ； 第二反应区 与出口区的结合部位为圆台形， 其纵剖面等腰梯形的底角 β为 45° ~ 85° 。
该反应器的预提升段、 第一反应区、 第二反应区、 出口区的高度 之和为反应器的总高度， 一般为 10米〜 60米。
预提升段的直径与常规的等直径提升管反应器相同， 一般为 0.02 米〜 5米， 其高度占反应器总高度的 5% ~ 10%。 预提升段的作用是在 预提升介质的存在下使再生催化剂向上运动并加速， 所用的预提升介 质与常规的等直径提升管反应器所用的相同， 选自水蒸汽或干气。
第一反应区的结构类似于常规的等直径提升管反应器， 其直径可 与预提升段相同， 也可较预提升段稍大， 第一反应区的直径与预提升 段的直径之比为 1.0 - 2.0: 1 , 其高度占反应器总高度的 10% ~ 30%。 原 料油和催化剂在该区混合后， 在较高的反应温度和剂油比、 较短的停 留时间 （一般为 0.5秒~ 2.5秒） 下， 主要发生裂化反应。
第二反应区比第一反应区要粗， 其直径与第一反应区的直径之比 为 1.5 ~ 5.0: 1 , 其高度占反应器总高度的 30% ~ 60%。 其作用是降低油 气和催化剂的流速和反应温度。 降 4氏该区反应温度的方法， 可以从该 区与第一反应区的结合部位注入冷激介质， 和 /或通过在该区设置取热 器， 取走部分热量以降低该区反应温度， 从而达到抑制二次裂化反应、 增加异构化反应和氢转移反应的目的。 冷激介质是选自冷激剂、 冷却 的再生催化剂和冷却的半再生催化剂中的一种或一种以上的任意比例 的混合物。 其中冷激剂是选自液化气、 粗汽油、 稳定汽油、 柴油、 重 柴油或水中的一种或一种以上的任意比例的混合物； 冷却的再生催化 剂和冷却的半再生催化剂是待生催化剂分别经两段再生和一段再生后 冷却得到的，再生催化剂碳含量为 0. 1重％以下，最好为 0.05重％以下， 半再生催化剂碳含量为 0.1重％~ 0.9重％, 最好碳含量为 0.15重％~ 0.7重％。 若设置取热器， 则其高度占第二反应区高度的 50% ~ 90%。 物流在该反应区停留时间可以较长， 为 1秒〜 30秒。
出口区的结构类似于常规的等直径提升管反应器顶部出口部分， 其直径与第一反应区的直径之比为 0.8 ~ 1.5: 1 , 其高度占反应器总高度 的 0 ~ 20%。 物流可在该区停留一定时间， 以抑制过裂化反应和热裂化 反应， 提高流体流速。
水平管的一端与出口区相连， 另一端与沉降器相连； 当出口区的 高度为 0 即提升管反应器没有出口区时， 水平管的一端与第二反应区 相连， 另一端与沉降器相连。 水平管的作用是将反应生成的产物与待 生催化剂输送至分离系统进行气固分离。 其直径由本领域技术人员根 据具体情况确定。 预提升段的作用是在预提升介质的存在下， 将再生 后的催化剂进行提升， 进入第一反应区。
在第一和第二方面的一些方案中， 该方法适用的优盾原料油可以 是不同沸程的石油馏份。 具体地说， 优质原料油选自常压塔顶油、 汽 油、 催化汽油、 柴油、 直熘蜡油、 加氢蜡油中的一种或多种。
在第一和第二方面的一些方案中， 该方法可以适用所有同一类型 的催化剂， 既可以是无定型硅铝催化剂， 也可以是沸石催化剂， 沸石 催化剂的活性组分选自 Y型沸石、 HY型沸石、 超稳 Y型沸石、 ZSM-5 系列沸石或具有五元环结构的高硅沸石、 镁碱沸石中的一种或一种以 上的任意比例的混合物， 该沸石可以含稀土和 /或磷， 也可以不含稀土 和碑。
在第一和第二方面的一些方案中， 该方法中也可以适用不同类型 催化剂， 不同类型催化剂可以是颗粒大小不同的催化剂和 /或表观堆积 密度不同的催化剂。 颗粒大小不同的催化剂和 /或表观堆积密度不同的 催化剂上活性组分分别选用不同类型沸石， 沸石选自 Y型沸石、 HY型 沸石、超稳 Y型沸石、 ZSM-5系列沸石或具有五元环结构的高硅沸石、 镁碱沸石中的一种或一种以上的任意比例的混合物， 该沸石可以含稀 土和 /或磷， 也可以不含稀土和磷。 大小不同颗粒的催化剂和 /或高低表 观堆积密度的催化剂可以分别进入不同的反应区， 例如， 含有超稳 Y 型沸石的大颗粒的催化剂进入第一反应区，增加裂化反应，含有稀土 Y 型沸石的小颗粒的催化剂进入第二反应区， 增加氢转移反应， 颗粒大 小不同的催化剂在同一汽提器汽提和同一再生器再生， 然后分离出大 颗粒和小颗粒催化剂， 小颗粒催化剂经冷却进入第二反应区。 颗粒大 小不同的催化剂是以 30 ~ 40微米之间分界， 表观堆积密度不同的催化 剂是以 0.6 ~ 0.7g/cm3之间分界。
在第一和第二方面的一些方案中， 该方法适用的活性较低的催化 剂一般是指催化剂活性在 35 ~ 55 , 优选 40 ~ 50。 在以前的常规的工业 催化裂化操作中， 通常将一定量的高活性的催化剂 （例如新鲜催化剂， 或活性大于 60的催化剂）加入或补充到装置内。 例如， 可采用以下方 法在本发明反应装置内获得该活性较低的催化剂： 降低装置的催化剂 补充率 （减少补充催化剂的量） ； 降低补充催化剂的活性； 或降低初 始加入装置内的催化剂的量。 更具体而言， 所述活性较低的催化剂可 以通过一定温度（例如 400-850 °C ) 的水蒸汽老化处理一段时间 （例如 1 -720小时） , 或通过以下处理方法 1、 2或 3获得。
本发明中所述的活性分布相对均勾的催化剂优选是指加入到催化 裂化装置内时的催化剂初始活性不超过 80， 不超过 75， 或不超过 70; 该催化剂的自平衡时间为 0.1 小时- 50小时， 0.2 ~ 30小时， 或 0.5 ~ 10小时； 平衡活性为 35 - 60, 或 40 ~ 50。 所述活性分布相对均匀的催 化剂可经水热老化处理得到。 例如， 可经下述处理方法 1、 2和 3而得 到。
表述 "活性较低的催化剂" 或 "活性分布相对均勾的催化剂，， 中 的所述 "活性" 是指全部催化剂的平均微反活性， ， 并非单个催化剂 的活性。
所述的催化剂活性 （例如平均活性、 初始活性、 平衡活性） 都是 采用现有技术的测量方法测量。 现有技术中的测量方法是： 企业标准 RIPP 92-90 催化裂化的微反活性试验法 《石油化工分析方法 （RIPP 试验方法） 》 , 杨翠定等人， 1990， 下文简称为 RIPP 92-90。 所述催 化剂活性是由轻油微反活性 （MA ) 表示， 其计算公式为 MA = (产物 中低于 204 °C的汽油产量 +气体产量 +焦炭产量） /进料总量 * 100%=产物 中低于 204 °C的汽油产率 +气体产率 +焦炭产率。 轻油微反装置 （参照 RIPP 92-90 ) 的评价条件是： 将催化剂破碎成直径为 420 ~ 841 4啟米的 颗粒， 装量为 5克， 反应原料是馏程为 235 ~ 337°C的直馏轻柴油， 反 应温度为 460°C， 重量空速为 16小时 剂油比为 3.2。
所述的催化剂自平衡时间是指催化剂在 800°C和 100%水蒸气条件 (参照 RIPP 92-90) 下老化达到平衡活性所需的时间。
所述活性分布相对均勾的热再生催化剂可经水热老化处理得到。 例如可经下述 3种处理方法而得到：
催化剂处理方法 1:
( 1 )、 将新鲜催化剂装入流化床， 优选密相流化床， 与水蒸汽接 触， 在一定的水热环境下进行老化后得到活性相对均勾的催化剂；
(2) 、 将所述活性相对均勾的催化剂加入到工业催化裂化装置的 再生器内。
处理方法 1例如是这样具体实施的：
将新鲜催化剂装入流化床优选密相流化床内， 在流化床的底部注 入水蒸汽， 催化剂在水蒸汽的作用下实现流化， 同时水蒸汽对催化剂 进行老化， 老化温度为 400°C-850°C , 优选 500°C-750°C , 最好为 600°C-700°C, 流化床的表观线速为 0.1米 /秒 -0.6米 /秒， 最好为 0.15米 /秒 -0.5米 /秒， 老化 1小时 -720小时优选 5小时 -360小时后， 得到所述 的活性相对均匀的催化剂。 活性相对均匀的催化剂按工业催化裂化装 置内的要求， 加入到工业催化裂化装置的再生器内得到所述活性分布 相对均匀的热再生催化剂。
催化剂处理方法 2:
( 1 )、 将新鲜催化剂装入流化床优选密相流化床， 与水蒸汽与其 它老化介质的混合物接触， 在一定的水热环境下进行老化后得到活性 相对均勾的催化剂；
( 2 )、 将所述活性相对均勾的催化剂加入到工业催化裂化装置的 再生器内。
催化剂处理方法 2的技术方案例如是这样具体实施的：
将催化剂装入流化床优选密相流化床内， 在流化床的底部注入水 蒸汽与其它老化介质的混合物， 催化剂在水蒸汽与其它老化介质的混 合物作用下实现流化， 同时， 水蒸汽与其它老化介质的混合物对催化 剂进行老化， 老化温度为 400°C-850°C, 优选 500°C-750°C, 最好为 600°C-700°C, 流化床的表观线速为 0.1米 /秒 -0.6米 /秒， 最好为 0.15米 /秒 -0.5 米 /秒， 水蒸汽与其它老化介盾的重量比为 0.20-0.9, 最好为 0.40-0.60, 老化 1 小时 -720小时优选 5小时 -360小时后， 得到所述的 活性相对均匀的催化剂， 活性相对均匀的催化剂按工业装置的要求， 加入到工业催化裂化装置的再生器内得到所述活性分布相对均勾的热 再生催化剂。 所述其它老化介质包括空气、 干气、 再生烟气、 空气与 干气燃烧后的气体或空气与燃烧油燃烧后的气体、 或其它气体如氮气。 所述水蒸气与老化介质的重量比为 0.2-0.9 , 最好为 0.40-0.60。
催化剂处理方法 3:
( 1 )、 将新鲜催化剂输入到流化床优选密相流化床， 将再生器的 热再生催化剂输送到所述流化床， 在所述流化床内进行换热；
( 2 )、 换热后的新鲜催化剂与水蒸汽或水蒸汽与其它老化介质的 混合物接触， 在一定的水热环境下进行老化后得到活性相对均勾的催 化剂；
( 3 ) 、 将所述活性相对均勾的催化剂加入到工业催化裂化装置的 再生器内。
催化剂处理方法 3的技术方案例如是这样具体实施的：
将新鲜催化剂输送到流化床优选密相流化床内， 同时将再生器的 热再生催化剂输送到另一个流化床，在两个流化床之间进行固-固换热。 在含有新鲜催化剂的流化床的底部注入水蒸汽或水蒸汽与其它老化介 质的混合物， 新鲜催化剂在水蒸汽或水蒸汽与其它老化介质的混合物 作用下实现流化， 同时， 水蒸汽或水蒸汽与其它老化介质的混合物对 新鲜催化剂进行老化， 老化温度为 400 °C -850°C , 优选 500°C -750°C , 最好为 600°C -70(TC , 流化床的表观线速为 0.1米 /秒 -0.6米 /秒， 最好为 0. 15米 /秒 -0.5米 /秒， 老化 1小时 -720小时， 优选 5小时 -360小时， 在 水蒸汽与其它老化介质的混合物的情况下， 所述水蒸气与其它老化介 质的重量比为大于 0-4 , 最好为 0.5-1.5 , 得到活性相对均匀的老化催化 剂， 老化催化剂按工业催化裂化装置的要求， 加入到工业催化裂化装 步骤后的 蒸汽进入反应系统 （作为 提蒸汽、 防焦蒸汽、 雾化蒸汽、 提升蒸汽中的一种或几种分别进入催化裂化装置中的汽提器、 沉降器、 原料喷嘴、 预提升段） 或再生系统， 而老化步骤后的水蒸汽与其它老 化介质的混合物进入再生系统， 换热后的再生催化剂返回到该再生器 内。 所述其它老化介质包括空气、 干气、 再生烟气、 空气与干气燃烧 后的气体或空气与燃烧油燃烧后的气体、 或其它气体如氮气。 所述再 生烟气可以来自本装置， 也可以来自其它装置。 通过水热老化处理， 工业反应装置内的催化剂的活性和选择性分 布更加均匀， 催化剂的选择性得到明显改善， 从而干气产率和焦炭产 率明显降低。
本发明的优点在于：
1、 如果采用常规的等直径提升管或流化床反应器来实施本发明， 只需降低处理量， 延长反应时间就可以实施。
2、 如果采用变径提升管反应器， 该反应器的优点是既保留常规提 升管反应器底部较高的反应温度和剂油比来增加一次裂化反应， 同时 抑制顶部的过裂化和热裂化反应， 又在反应器中上部在较低的反应温 度下延长反应时间， 增加烯烃的异构化反应、 氢转移反应。
3、 用本发明提供的方法生产的液化气中异丁烯含量增加 30%以上 与常规方法相比， 汽油族组成中的烯烃含量可增加到 30重％以上。 附图说明
图 1为提升管反应器的示意图， 图中的 a、 b、 c、 d、 e分别代表预 提升段、 第一反应区、 第二反应区、 出口区、 水平管。
图 2 是本发明第二方面的最佳实施方式的流程示意图。 附图中各 编号说明如下：
1、 3、 4、 6、 1 1、 13、 17、 18 均代表管线； 2 为提升管的预提升 段； 5、 7 分别为提升管的第一反应区、 第二反应区； 8 为提升管的出 口区； 9为沉降器， 10为旋风分离器， 12为汽提器， M为待生斜管， 15为再生器， 16为再生斜管。 具体实施方式
本发明具有不同的实施方式， 例如
实施方式之一：
在常规等直径提升管反应器的底部， 预热的原料油与活性较低的 热再生催化剂或与活性较低且活性分布相对均勾的热再生催化剂接触 发生裂化反应， 生成的油气和用过的催化剂上行与注入冷却的再生催 化剂接触， 随之发生异构化反应和氢转移反应， 反应后流出物进入沉 降器； 分离反应产物， 待生催化剂经汽提、 再生后分为两部分， 其中 一部分进入该反应器底部， 另一部分经降温后进入该反应器中下部。 实施方式之二：
在常规等直径提升管反应器的底部， 预热的原料油与活性较低的 热再生催化剂或与活性较低且活性分布相对均勾的热再生催化剂接触 发生裂化反应， 生成的油气和用过的催化剂上行与注入冷激剂和冷却 的半再生催化剂接触， 随之发生异构化反应和氢转移反应， 反应后流 出物进入沉降器； 分离反应产物， 待生催化剂经汽提后， 进入两段再 生器中烧焦， 从第一段再生器中出来的半再生催化剂经降温后进入该 反应器中下部， 从第二段再生器中出来的再生催化剂不经降温直接返 回该反应器底部。
实施方式之三：
对于具有常规提升管 -流化床反应器的催化裂化装置， 预热后的常 规裂化原料从提升管的下部进入与活性较低的热再生催化剂或与活性 较低且活性分布相对均勾的热再生催化剂接触， 反应后生成的油气上 行至提升管的顶部， 与降温后的催化剂接触继续进行反应， 反应后流 出物进入沉降器； 分离反应产物， 待生催化剂经汽提、 再生后分为两 部分， 其中一部分进入提升管的下部， 另一部分经降温后进入提升管 的顶部。
实施方式之四：
该实施方式为本发明的最佳实施方式。
对于具有变径提升管反应器的催化裂化装置， 预热后的常规裂化 原料从反应器的第一反应区下部进入与活性较低的热再生催化剂或与 活性较低且活性分布相对均勾的热再生催化剂接触， 发生裂化反应， 反应后生成的油气上行至反应器的第二反应区下部与降温后的催化剂 接触进行氢转移反应和异构化反应， 反应后流出物进入沉降器； 分离 反应产物， 待生催化剂经汽提、 再生然后进入第二反应区下部。
本发明提供的方法并不局限于此。
下面结合附图进一步说明本发明所提供的方法， 但本发明并不因 此而受到任何限制。
图 2 是采用变径提升管反应器， 提高液化气中的异丁烯和汽油烯 烃含量的催化转化方法的流程， 设备和管线的形状、 尺寸不受附图的 限制， 而是根据具体情况确定。
预提升蒸汽经管线 1 从提升管预提升段 2进入， 活性较低的热再 分布相对均匀的热再生催化剂经再生斜管
16 进入提升管预提升段由预提升蒸汽进行提升。 预热后的原料油经管 线 4与来自管线 3 的雾化蒸汽按一定比例从提升管预提升段进入， 与 热催化剂混合后进入笫一反应区 5内， 在一定的条件下进行裂化反应。 反应物流与来自管线 6的冷激剂和 /或冷却的催化剂 （图中未标出） 混 合进入第二反应区 7， 进行二次反应， 反应后的物流进入出口区 8， 该 反应区提高物流的线速， 使反应物流快速进入沉降器 9 和旋风分离器 10， 反应产物经管线 1 1去分离系统。 反应后带炭的待生催化剂进入汽 提器 12 ,经来自管线 13的水蒸汽汽提后由待生斜管 14进入再生器 15 , 待生催化剂在来自管线 17的空气中烧焦再生， 烟气经管线 18 出再生 器， 热的再生催化剂经再生斜管 16返回提升管底部循环使用。 实施例
下面的实施例将对本发明予以进一步说明， 但并不因此而限制本 发明。 实施例、 对比例中所使用的原料油和催化剂的性质分别列于表 1 和表 2。表 2中的催化剂均由中国石油化工集团公司齐鲁催化剂厂生产。 表 2中的 ZCM-7催化剂经 800 °C , 100%水蒸汽分别老化 12小时和 30 小时， 得到两种不同活性水平的 ZCM-7 , 即活性为 67和 45 ; 同样， 表 2中的 CGP-1催化剂经 800 °C , 100%水蒸汽分别老化 12小时和 30 小时， 得到两种不同活性水平的 CGP- 1 , 即活性为 62和 50。
实施例 1
本实施例说明采用本发明提供的方法， 采用不同活性水平的催化 剂， 在中型变径提升管反应器上提高液化气中异丁烯含量和汽油烯烃 含量的情况。
反应器的预提升段、 第一反应区、 第二反应区、 出口区总高度为
15米， 预提升段直径为 0.025米， 其高度为 1.5米； 第一反应区直径为 0.025米， 其高度为 4米； 第二反应区直径为 0.1米， 其高度为 6.5米； 出口区的直径为 0.025米， 其高度为 3米； 第一、 二反应区结合部位的 纵剖面等腰梯形的顶角为 45° ； 第二反应区与出口区结合部位的纵剖 面等腰梯形的底角为 60° 。
预热的表 1 所列的原料油 B进入该反应器内， 在水蒸汽存在下， 与热的表 2所列的催化剂 ZCM-7接触反应， ZCM-7催化剂活性为 45 , 分离反应产物得到液化气和汽油及其它产品， 待生催化剂经汽提进入 再生器， 再生催化剂经烧焦后循环使用。
试猃的操作条件、 产品分布和汽油的性质列于表 3。
对比例 1
采用反应器类型和操作条件与实施例 1 完全相同， 所用的原料油 也是表 1 所列的原料油 B， 催化剂也是表 2所列的催化剂 ZCM-7 , 只 是此时 ZCM-7催化剂活性为 67。 试验的操作条件、 产品分布和汽油的 性质列于表 3。
从表 3可以看出， 相对于釆用高活性 ZCM-7 (即活性为 67 )， 采 用低活性 ZCM-7 (即活性为 45 ), 异丁烯产率由 1.4重％上升到 2.0重 %, 增加了 42.86%， 汽油烯烃含量由 16.3重％上升到 29.3重％; 此外， 液体收率仍增加 1.2个百分点。
实施例 2
本实施例说明采用本发明提供的方法， 釆用不同活性水平的催化 含量的情况。 a 、 、 、 反应器的预提升段、 第一反应区、 第二反应区、 出口区总高度为 15米， 预提升段直径为 0.025米， 其高度为 1.5米； 第一反应区直径为 0.025米， 其高度为 4米； 第二反应区直径为 0.1米， 其高度为 6.5米； 出口区的直径为 0.025米， 其高度为 3米； 第一、 二反应区结合部位的 纵剖面等腰梯形的顶角为 45° ； 第二反应区与出口区结合部位的纵剖 面等腰梯形的底角为 60° 。
预热的表 1 所列的原料油 Β进入该反应器内， 在水蒸汽存在下， 与热的表 2所列的催化剂 CGP- 1接触反应， CGP-1催化剂活性为 50 , 分离反应产物得到液化气和汽油及其它产品， 待生催化剂经汽提进入 再生器， 再生催化剂经烧焦后循环使用。
试验的操作条件、 产品分布和汽油的性质列于表 4。
^"比例 2
采用反应器类型和操作条件与实施例 2 完全相同， 所用的原料油 也是表 1 所列的原料油 Β， 催化剂也是表 2所列的催化剂 CGP- 1 , 只 是此时 CGP-1催化剂活性为 62。 试验的操作条件、 产品分布和汽油的 性质列于表 4。 从表 4可以看出， 相对于采用高活性 CGP-1 (即活性为 62 ), 采用 低活性 CGP-1 (即活性为 50 ), 异丁烯产率由 3.0重％上升到 4.1重％， 增加了 36.67% , 汽油烯烃含量由 18.2重％上升到 27.9重％; 此外， 液 体收率仍增加 0.8个百分点。
实施例 3和 4
本实施例说明采用本发明提供的方法， 使用不同类型的催化裂化 原料油， 在中型变径提升管反应器上提高液化气中异丁烯含量和汽油 烯烃含量的情况。
本实施例使用的反应器、 催化剂类型、 催化剂活性同实施例 2 , 只 是原料油分别为表 1所列出的原料油 A和 C。
操作条件、 产品分布和汽油性质列于表 5。 从表 5可以看出， 异丁 烯产率分别为 4.3重％和2.1重％，汽油烯烃含量分别为 30.2重％和 22.2 重0 /0。
实施例 5
本实施例说明采用本发明提供的方法， 采用不同活性水平的催化 剂， 在中型变径提升管反应器上提高液化气中异丁烯含量和汽油烯烃 含量的情况。
反应器的预提升段、 第一反应区、 第二反应区、 出口区总高度为 15米， 预提升段直径为 0.025米， 其高度为 1.5米； 笫一反应区直径为 0.025米， 其高度为 4米； 第二反应区直径为 0.1米， 其高度为 6.5米； 出口区的直径为 0.025米， 其高度为 3米； 第一、 二反应区结合部位的 纵剖面等腰梯形的顶角为 45。 ； 第二反应区与出口区结合部位的纵剖 面等腰梯形的底角为 60° 。
预热的原料油 B进入该反应器内， 在水蒸汽存在下， 与热的催化 剂 ZCM-7接触反应， 其催化剂活性（平均活性）为 45 , 分离反应产物 得到液化气和汽油及其它产品， 待生催化剂经汽提进入再生器， 再生 催化剂经烧焦后循环使用。 补充到装置内的 ZCM-7 催化剂是新鲜 ZCM-7 经水热处理 （催化剂水热处理方法采用本发明催化剂处理方法 1处理：密相流化床，老化温度 650°C ,流化床的表观线速 0.30m/s, 100% 水蒸气， 老化时间 31小时）后的催化剂， 其初始活性为 75 , 然后与装 置内的平衡催化剂混合， 再经装置内的水热老化， 该加入的催化剂达 到装置内的催化剂平衡活性 45时所需的自平衡时间 （ 800 °C , 100%水 蒸气） 为 30小时。
试验的操作条件、 产品分布和汽油的性质列于表 6。
实施例 5 A
采用反应器类型和操作条件与实施例 5 完全相同， 所用的原料油 也是表 1 所列的原料油 B , 催化剂也是表 2所列的催化剂 ZCM-7 , 其 催化剂平均活性也为 45。 只是补充到装置内的 ZCM-7 催化剂是新鲜 ZCM-7催化剂， 未经水热处理， 其初始活性为 91 , 与装置内的平衡催 化剂混合， 再经装置内的水热老化， 直到装置内的催化剂平衡活性为 45。 试验的操作条件、 产品分布和汽油的性质列于表 6。
从表 6可以看出， 相对于未处理 ZCM-7催化剂， 采用加入处理后
ZCM-7催化剂， 干气产率由 1.7重％下降到 1.5重％, 焦炭产率由 3.2 重％下降到 2.7重％， 液体收率由 89.3重％上升到 89.8重％, 增加 0.5 个百分点。 两者的异丁烯产率和汽油烯烃含量基本相当。
实施例 6
本实施例说明采用本发明提供的方法， 采用不同活性水平的催化 剂， 在中型变径提升管反应器上提高液化气中异丁烯含量和汽油烯烃 含量的情况。
反应器的预提升段、 笫一反应区、 第二反应区、 出口区总高度为 15米， 预提升段直径为 0.025米， 其高度为 1.5米； 第一反应区直径为 0.025米， 其高度为 4米； 第二反应区直径为 0.1米， 其高度为 6.5米； 出口区的直径为 0.025米， 其高度为 3米； 第一、 二反应区结合部位的 纵剖面等腰梯形的顶角为 45° ； 第二反应区与出口区结合部位的纵剖 面等腰梯形的底角为 60° 。
预热的原料油 B进入该反应器内， 在水蒸汽存在下， 与热的催化 剂 CGP-1接触反应， CGP-1催化剂平均活性为 50 , 分离反应产物得到 液化气和汽油及其它产品， 待生催化剂经汽提进入再生器， 再生催化 剂经烧焦后循环使用。 补充到装置内的 CGP-1催化剂是新鲜 CGP- 1经 水热处理后的催化剂 （催化剂水热处理方法采用本发明催化剂处理方 法 1处理： 密相流化床， 老化温度 670°C , 流化床的表观线速 0.30m/s , 100%水蒸气， 老化时间 28 小时） ， 其初始活性为 72 , 然后与装置内 的平衡催化剂混合， 再经装置内的水热老化， 该加入的催化剂达到装 置内的催化剂平衡活性 50时所需的自平衡时间（ 800 °C , 100%水蒸气） 为 40小时。
试验的操作条件、 产品分布和汽油的性质列于表 7。
实施例 6 A
采用反应器类型和操作条件与实施例 6 完全相同， 所用的原料油 也是表 1所列的原料油 B ,催化剂也是表 2所列的催化剂 CGP-1 , CGP-1 催化剂平均活性也为 50。 只是补充到装置内的 CGP-1 催化剂是新鲜 CGP-1催化剂， 未经水热处理， 其初始活性为 95， 与装置内的平衡催 化剂混合， 再经装置内的水热老化， 直到装置内的催化剂平衡活性为 50。 试验的操作条件、 产品分布和汽油的性质列于表 7。
从表 7可以看出， 相对于未处理 CGP-1催化剂， 采用加入处理后
CGP-1 催化剂， 干气产率由 2.0重％下降到 1.9重％, 焦炭产率由 3.0 重％下降到 2.5重％, 液体收率由 88.7重％上升到 89.3重。 /。， 增加 0.6 个百分点。 两者的异丁烯产率和汽油烯烃含量基本相当。
实施例 7
本实施例说明采用本发明提供的方法， 采用不同活性水平的催化 剂和中型常规等直径提升管反应器改善产物分布的情况。
预热的表 1 所列的原料油 B进入该反应器内， 在水蒸汽存在下， 与热的表 2所列的催化剂 ZCM-7接触反应， ZCM-7催化剂活性为 45， 分离反应产物得到液化气和汽油及其它产品， 待生催化剂经汽提进入 再生器， 再生催化剂经烧焦后循环使用。
试验的操作条件、 产品分布和汽油的性质列于表 8。
十比例 3
采用反应器类型和操作条件与实施例 7 完全相同， 所用的原料油 也是表 1 所列的原料油 B , 催化剂也是表 2所列的催化剂 ZCM-7, 只 是此时 ZCM-7催化剂活性为 67。 试验的操作条件、 产品分布和汽油的 性质列于表 8。
从表 8 可以看出， 相对于采用高活性 ZCM-7 (即活性为 67 ), 采 用低活性 ZCM-7 (即活性为 45 )， 异丁烯产率由 1.4重％上升到 1.9重 %, 增加了 0.5个百分点， 汽油烯烃含量由 25.6重％上升到 31.7重％; 此外， 液体收率仍增加 0.9个百分点。 表 1
Figure imgf000020_0001
表 2
Figure imgf000021_0001
实施例 1 对比例 1
ZCM-7催化剂活性 45 67 反应温度， °C
第一反应区 550 550 第二反应区 500 500 停留时间， 秒 5.5 5.5 第一反应区 2.0 2.0 第二反应区 3.5 3.5 剂油比 5.0 5.0 水油比 0.1 0. 1 产品分布， 重％
干气 1.4 1.8 液化气 17.3 17.5 其中异丁烯 2.0 1.4 汽油 55.0 56.0 柴油 17.8 15.4 重油 6.0 5.6 焦炭 2.5 3.7 液体收率， 重％ 90. 1 88.9 辛烷值
RON 91.0 90.6
MON 80.7 80.5 馏程， 。C
初留点 ~干点 38 - 200 37 ~ 200 汽油族组成， 重％
烷烃 40.5 50.6 环烷烃 7.3 8.2 烯烃 29.3 16.3 芳烃 22.9 24.9 表 4
Figure imgf000023_0001
实施例 3 实施例 4 原料油 A C 操作条件
反应温度， °c
第一反应区 550 550 第二反应区 505 505 剂油比 6.0 6.0 反应时间， 秒 6.0 6.0 第一反应区 1.3 1.3 第二反应区 4.7 4.7 水油比 0.1 0.1 产品分布， 重％
干气 1.7 2.0 液化气 28.0 24.0 其中丙烯 10.5 7.5 异丁烯 4.3 2.1 汽油 42.0 46.6 柴油 18.6 17.0 重油 6.9 7.5 焦炭 2.8 2.9 液体收率， 重％ 88.6 87.6 汽油辛燒值
RON 93.2 93.0
MON 81.2 81. 1 馏程， °C
初留点 ~干点 38 ~ 200 38 - 200 汽油族组成， 重％
烷烃 35.4 41.9 环烷烃 7.8 8.3 烯烃 30.2 22.2 芳烃 26.6 27.6 表 6
Figure imgf000025_0001
表 7
Figure imgf000026_0001
表 8
Figure imgf000027_0001

Claims

权 利 要 求
1. 一种改善产物分布的催化转化方法， 其特征在于优质原料油与 活性较低的热再生催化剂在反应器内接触发生裂化反应， 将反应产物 和待生催化剂分离， 该反应产物被送入分离系统， 该待生催化剂经汽 提、 再生后循环使用。
2. —种改善产物分布的催化转化方法， 其特征在于优质原料油与 活性较低的热再生催化剂在反应器的下部接触发生裂化反应， 裂化反 应产物和含炭的催化剂上行并且发生选择性的氢转移反应和异构化反 应， 将氢转移反应和异构化反应的反应产物和待生催化剂分离， 氢转 移反应和异构化反应的反应产物被送入分离系统， 该待生催化剂经汽 提、 再生后循环使用。
3. 按照权利要求 1或 2的方法， 其特征在于所述优质原料油选自 常压塔顶油、 汽油、 催化汽油、 柴油、 直馏蜡油、 加氢蜡油中的一种 或多种。
4. 按照权利要求 1或 2的方法， 其特征在于所述活性较低的热再 生催化剂活性为 35 ~ 55。
5. 按照权利要求 4的方法， 其特征在于所述活性较低的热再生催 化剂活性为 40 ~ 50。
6. 按照权利要求 1或 2的方法， 其特征在于所述活性较低的热再 生催化剂具有相对均勾的活性分布。
7. 按照权利要求 6的方法， 其特征在于所述活性分布相对均匀的 热再生催化剂在加入到催化裂化装置内时其初始活性不超过 80， 该催 化剂的自平衡时间为 0.1 小时 ~ 50小时， 平衡活性为 35 ~ 60。
8. 按照权利要求 7的方法， 其特征在于所述活性分布相对均匀的 热再生催化剂在加入到催化裂化装置内时其初始活性不超过 75 , 该催 化剂的自平衡时间为 0.2 ~ 30小时， 平衡活性为 40 ~ 50。
9. 按照权利要求 8的方法， 其特征在于所述活性分布相对均匀的 热再生催化剂在加入到催化裂化装置内时其初始活性不超过 70 , 该催 化剂的自平衡时间为 0.5 ~ 10小时。
10. 按照权利要求 1的方法， 其特征在于所述裂化反应条件为： 反 应温度 450 °C ~ 620 °C , 反应时间 0.5秒~ 35.0秒， 催化剂与原料油的 重量比 3 ~ 15:1。
11. 按照权利要求 2的方法， 其特征在于所述裂化反应条件为： 反 应温度 490°C ~62(TC, 反应时间 0.5秒~ 2.0秒， 催化剂与原料油的重 量比 3 ~ 15:1。
12. 按照权利要求 11 的方法， 其特征在于所述裂化反应条件为： 反应温度 500°C ~600°C, 反应时间 0.8秒~ 1.5秒， 催化剂与原料油的 重量比 3 ~ 12:1。
13. 按照权利要求 2的方法，其特征在于所述氢转移反应和异构化 反应条件为： 反应温度 420Ό ~ 550°C， 反应时间为 2秒~30秒。
14. 按照权利要求 13的方法， 其特征在于所述氢转移反应和异构 化反应条件为： 反应温度 460°C ~ 500°C， 反应时间为 3秒~ 15秒。
15. 按照权利要求 1或 2的方法， 其特征在于所述裂化反应、 氢转 移反应和 /或异构化反应的压力均为 130 kPa ~450kPa, 水蒸汽与原料 油的重量比为 0.03 ~ 0.3:1。
16. 按照权利要求 1或 2的方法，其特征在于所述反应器选自等直 径提升管、 等线速提升管、 流化床或变径提升管中之一， 或者是由等 直径提升管和流化床构成的复合反应器。
17. 按照权利要求 16的方法， 其特征在于所述变径提升管沿垂直 方向从下至上依次为互为同轴的预提升段、 第一反应区、 直径扩大了 的第二反应区、 直径缩小了的出口区， 在出口区末端连有一段水平管， 其中第二反应区的直径与第一反应区的直径之比为 1.5 ~ 5.0:1。
18. 按照权利要求 1或 2的方法， 用于提高工业 FCC工艺中的液 化气中异丁烯含量和汽油中烯烃含量。
19. 按照权利要求 1或 2的方法，其中所述活性较低的热再生催化 剂可通过以下方法获得：
1 ) 降低装置的催化剂补充率；
2) 降低补充催化剂的活性； 或
3) 降低初始加入到装置内的催化剂的活性。
20. 根据权利要求 6的方法，其特征在于所述具有相对均匀的活性 分布的热再生催化剂经下述处理方法而得到：
( 1 ) 、 将新鲜催化剂装入流化床， 与水蒸汽接触， 在一定的水热 环境下进行老化后得到活性相对均匀的催化剂； (2) 、 将所述活性相对均匀的催化剂加入到工业催化裂化装置的 再生器内以获得所述具有相对均勾的活性分布的热再生催化剂。
21. 按照权利要求 20的方法， 其特征在于水热环境包括老化温度 400°C-850°C, 流化床的表观线速 0.1米 /秒 -0.6米 /秒， 老化时间 1 小时 -720小时。
22. 根据权利要求 6的方法，其特征在于所述具有相对均勾的活性 分布的热再生催化剂经下述处理方法而得到：
( 1 ) 、 将新鲜催化剂装入流化床， 与水蒸汽与其它老化介质的混 合物接触， 在一定的水热环境下进行老化后得到活性相对均勾的催化 剂；
(2) 、 将所述活性相对均勾的催化剂加入到工业催化裂化装置的 再生器内以获得所述具有相对均勾的活性分布的热再生催化剂。
23. 按照权利要求 22的方法， 其特征在于水热环境包括老化温度 400oC-850°C, 流化床的表观线速 0.1米 /秒 -0.6米 /秒， 老化时间 1小时 -720 小时， 所述其它老化介质包括空气、 干气、 再生烟气、 空气与干 气燃烧后的气体、 空气与燃烧油燃烧后的气体、 或氮气。
24. 根据权利要求 6的方法，其特征在于所述具有相对均勾的活性 分布的热再生催化剂经下述处理方法而得到：
( 1 ) 、 将新鲜催化剂输入到流化床， 将再生器的热再生催化剂输 送到另一个流化床， 在两个流化床之间进行固-固换热；
(2) 、 换热后的新鲜催化剂与水蒸汽或水蒸汽与其它老化介质的 混合物接触， 在一定的水热环境下进行老化后得到活性相对均勾的催 化剂；
(3 ) 、 将所述活性相对均勾的催化剂加入到工业催化裂化装置的
25. 按照权利要求 24的方法， 其特征在于水热环境包括老化温度 400°C-850°C, 流化床的表观线速 0.1米 /秒 -0.6米 /秒， 老化时间 1小时 -720 小时， 所述其它老化介质包括空气、 干气、 再生烟气、 空气与干 气燃烧后的气体、 空气与燃烧油燃烧后的气体、 或氮气。
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"Laboratory Simulation of the Fcc Commercial Equilibrium Catalyst", PETROLEUM REFINERY ENGINEERING, vol. 33, no. 11, 2003, pages 15 - 17 *

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RU2563637C2 (ru) 2015-09-20
JP5947797B2 (ja) 2016-07-06
TW201224133A (en) 2012-06-16
KR101672789B1 (ko) 2016-11-04
KR20140000223A (ko) 2014-01-02
US9580664B2 (en) 2017-02-28
TWI525189B (zh) 2016-03-11
SA111320790B1 (ar) 2015-04-21
JP2013537926A (ja) 2013-10-07
RU2013119368A (ru) 2014-11-10

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