WO2021023172A1 - 一种从催化柴油生产轻质芳烃的全转化方法和装置 - Google Patents

一种从催化柴油生产轻质芳烃的全转化方法和装置 Download PDF

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WO2021023172A1
WO2021023172A1 PCT/CN2020/106710 CN2020106710W WO2021023172A1 WO 2021023172 A1 WO2021023172 A1 WO 2021023172A1 CN 2020106710 W CN2020106710 W CN 2020106710W WO 2021023172 A1 WO2021023172 A1 WO 2021023172A1
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stream
aromatics
selective
reaction zone
saturation
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PCT/CN2020/106710
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English (en)
French (fr)
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郑均林
姜向东
宋奇
孔德金
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中国石油化工股份有限公司
中国石油化工股份有限公司上海石油化工研究院
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Priority to BR112022002033A priority Critical patent/BR112022002033A2/pt
Priority to EP20850749.1A priority patent/EP4012006A4/en
Priority to CA3149654A priority patent/CA3149654A1/en
Priority to MX2022001612A priority patent/MX2022001612A/es
Priority to JP2022507440A priority patent/JP2022543288A/ja
Priority to US17/632,347 priority patent/US20220275294A1/en
Priority to KR1020227007224A priority patent/KR20220044541A/ko
Publication of WO2021023172A1 publication Critical patent/WO2021023172A1/zh

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    • 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
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • C10G67/14—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only including at least two different refining steps 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/02—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing
    • 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • C10G45/46—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used
    • C10G45/52—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing platinum group metals or compounds thereof
    • C—CHEMISTRY; METALLURGY
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    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • C10G45/46—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used
    • C10G45/54—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing crystalline alumino-silicates, e.g. molecular sieves
    • 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
    • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
    • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
    • C10G47/10—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used with catalysts deposited on a carrier
    • C10G47/12—Inorganic carriers
    • C10G47/16—Crystalline alumino-silicate carriers
    • C10G47/20—Crystalline alumino-silicate carriers the catalyst containing other metals or compounds thereof
    • 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
    • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
    • C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
    • C10G65/12—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including cracking steps and other hydrotreatment steps
    • 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
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • C10G67/12—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only including oxidation as the refining step 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
    • C10G2300/1048—Middle distillates
    • 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
    • C10G2300/1048—Middle distillates
    • C10G2300/1055—Diesel having a boiling range of about 230 - 330 °C
    • 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
    • C10G2300/1048—Middle distillates
    • C10G2300/1059—Gasoil having a boiling range of about 330 - 427 °C
    • 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
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • 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
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/207—Acid gases, e.g. H2S, COS, SO2, HCN
    • 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
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4081—Recycling aspects
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/30—Aromatics
    • 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 relates to a technology for preparing light aromatic hydrocarbons in the field of petroleum catalytic cracking, in particular to a method and device for producing light aromatic hydrocarbons from catalytic diesel.
  • Light aromatic hydrocarbons such as benzene, toluene and xylene are important basic organic chemical raw materials, which are widely used in synthetic materials and other fields, and are closely related to the development of the national economy and people's food, clothing, housing and transportation.
  • aromatic raw materials there are mainly two process routes for the sources of aromatics: one is to obtain aromatic raw materials by catalytic reforming of naphtha and aromatics extraction; the second is to obtain aromatic raw materials from pyrolysis gasoline, a byproduct of the ethylene plant, through hydrogenation and aromatics extraction.
  • LCO catalyzed diesel
  • Catalytic diesel hydrofining is to carry out olefin hydrogenation saturation, desulfurization, denitrification and partial saturation of aromatics under medium and low pressure conditions, which can improve its color and stability.
  • the hydrorefining process is far from meeting the product's cetane number requirements.
  • Hydro-upgrading processes such as UOP's Unicracking process (US5026472), whose target product is high cetane number diesel.
  • the process has good aromatics hydrogenation saturation performance and ring-opening selectivity, high aromatics conversion depth, large cetane number increase range and high diesel yield.
  • Light oil-type hydrocracking is the process of refining light diesel components and then vigorously saturating and hydrogenating them to obtain reformate of naphtha fraction or gasoline fraction. This process also has the problem of low yield of raw materials converted into aromatics. If the naphtha fraction is used for the reforming of aromatics raw materials, the cycloalkanes and paraffins generated after being over-saturated have to be converted into aromatics in the reformer, which is not an economical route.
  • the light oil type hydrocracking method described in the CN101684415 patent does not directly produce aromatics, and the aromatic potential of heavy naphtha is only 57%.
  • Document CN1955262A describes a two-stage hydrocracking method, the hydrocracking catalyst contains Pt and /Pd precious metals and non-precious metals, as well as Y zeolite and alumina, and the raw material is catalytic diesel.
  • the highest aromatic potential value of its naphtha product is only 76.8%, and the purity of aromatics is not high, which cannot meet the requirements of the aromatics combined device.
  • the document CN103897731A describes a method for producing light aromatics by mixing catalytically cracked diesel oil and C 10 + distillate oil. The product is cut through hydrorefining and hydrocracking. The fraction greater than 195°C is used as a clean diesel blending component, less than 195 The °C fraction enters the aromatics plant to produce light aromatics and clean gasoline blending components, and the yield of aromatics products is relatively low.
  • the heavy tail oil is discharged as a diesel component or partially recycled back to the hydrorefining reactor, and cannot be fully utilized to increase the production of light aromatics.
  • the hydrorefining reaction on the metal sulfide type hydrorefining catalyst needs to be carried out under severe operating conditions of high temperature and high pressure.
  • the reaction is limited by the thermodynamic equilibrium, and the selectivity to the partial saturation reaction of fused ring aromatics is not good.
  • LCO undergoes hydrogenation.
  • the aromatics retention rate after refining is less than 90%.
  • the heavy tail oil produced by catalytic diesel to produce light aromatics contains more than 90% fused-ring aromatics and low sulfur and nitrogen content. Recycling to the hydrorefining reactor will cause oversaturation and aromatics loss.
  • the inventors conducted a series of studies and found that the catalytic diesel stream is subjected to a selective conversion reaction including hydrocracking after hydrorefining and separation of impurities, and the mixed aromatics produced are separated into benzene in turn.
  • the heavy tail oil at the bottom of the tower enters the post-saturation selective reactor under low temperature and low pressure conditions After highly selective hydrogenation and saturation to obtain a product with a benzene ring, it is then sent to the selective conversion reaction, so as to realize the conversion of the whole fraction of light aromatics from catalytic diesel oil, and has a good light aromatics yield.
  • the light aromatic hydrocarbons mentioned in the present invention refer to aromatic hydrocarbons with carbon number less than or equal to 10, including C6 aromatic hydrocarbons, such as benzene; C7 aromatic hydrocarbons, such as toluene; C8 aromatic hydrocarbons, such as ethylbenzene and xylene; C9 aromatic hydrocarbons, such as methyl ethyl benzene and propylene Benzene, trimethylbenzene; C10 aromatic hydrocarbons, such as tetramethylbenzene, dimethylethylbenzene, diethylbenzene, etc.
  • the heavy aromatic hydrocarbons above C 10 in the present invention refer to aromatic hydrocarbons with a carbon number greater than 10.
  • One of the objectives of the present invention is to provide a full conversion method for producing light aromatics from catalytic diesel.
  • the full conversion method for producing light aromatics from catalytic diesel according to the present invention includes the following steps:
  • the catalytic diesel enters the first reaction zone for hydrorefining to obtain the first stream;
  • the first stream enters the second reaction zone for selective conversion to obtain the second stream, wherein optionally, the impurity separation is performed in the second separation zone before the first stream enters the second reaction zone;
  • the full conversion method for producing light aromatics from catalytic diesel fuel includes the following steps:
  • Catalyzed diesel fuel enters the first reaction zone and contacts with the hydrorefining catalyst under hydrogen conditions to obtain the first stream; the first reaction zone performs the hydrorefining reaction;
  • the selective conversion includes a hydrocracking reaction
  • the third stream enters the post-saturation selective reaction zone and contacts with the post-saturation selective catalyst under hydrogen conditions to obtain the fourth stream; the post-saturation selective reaction is carried out by hydrogenation saturation reaction;
  • step 1) of the method of the present invention the catalytic diesel as the feedstock is subjected to hydrorefining in the first reaction zone under hydrogen conditions, wherein the catalytic diesel stream and hydrogen are combined with
  • the hydrorefining catalyst is contacted to perform desulfurization and denitrification, and a selective saturation reaction of condensed aromatic hydrocarbons with one aromatic ring occurs.
  • the hydrorefining can be carried out in any manner and any method conventionally known in the art, as long as the catalytic diesel is desulfurized and denitrogenated, and the fused-ring aromatic hydrocarbons are hydrogenated and saturated to retain one aromatic ring. Special restrictions.
  • the first stream obtained by hydrorefining the catalytic diesel mainly contains refined catalytic diesel from which most of the sulfur and nitrogen impurities have been removed, and a gas phase containing hydrogen sulfide and ammonia.
  • step 1) of the method of the present invention the catalytic diesel as the feedstock and hydrogen are contacted with the hydrorefining catalyst in the first reaction zone to perform the hydrorefining reaction.
  • the hydrorefining reaction is a well-known catalytic diesel hydrorefining technology in the art.
  • the hydrorefining reaction conditions can be the reaction conditions known in the art for catalytic diesel hydrorefining; the hydrorefining catalyst can be any type of hydrorefining catalyst available in the art, as long as the steps can be achieved 1) The purpose of catalytic diesel hydrofining is sufficient.
  • step 1) of the method of the present invention the hydrorefining reaction conditions in the first reaction zone preferably include:
  • volume ratio of hydrogen to oil 500 ⁇ 3000Nm 3 / m 3, preferably 800 ⁇ 2000Nm 3 / m 3, more preferably 1000 ⁇ 1500Nm 3 / m 3;
  • the inlet temperature of the reactor is 280-420°C, preferably 300-410°C, more preferably 310-390°C;
  • the hydrogen partial pressure is 5-10 MPa, preferably 5-8 MPa, more preferably 6-7 MPa; and/or
  • the space velocity is 0.5-2.0 hr -1 , preferably 0.6-1.5 hr -1 , more preferably 0.8-1.2 hr -1 .
  • the hydrorefining catalyst in step 1) may preferably be as follows:
  • parts by weight it includes: a1) 60 to 99.9 parts, preferably 65 to 99.9 parts, preferably 70 to 99.9 parts, more preferably 75 to 99.9 parts of support; and b1) hydrogenated metal oxide, wherein the hydrogenated metal oxide
  • the parts by weight are 0.1-40 parts, preferably 0.1-35 parts, preferably 0.1-30 parts, more preferably 0.1-25 parts; based on the total weight parts of the carrier and the hydrogenated metal oxide.
  • the carrier in terms of parts by weight, includes: 60-100 parts of alumina; 0-40 parts of silica; based on the total parts by weight of the alumina and the silica.
  • the hydrogenation metal is at least one selected from the group consisting of nickel, cobalt, molybdenum, tungsten, and iron.
  • the hydrogenated metal is vulcanized after being loaded.
  • the hydrorefining catalyst of the present invention can be prepared by any method known in the art, for example, the carrier can be prepared by extrusion, rolling, or oil column forming methods in the art. In one embodiment, the catalyst can be prepared by forming the support and then impregnating the metal.
  • the first stream obtained by the hydrorefining in step 1) is subjected to impurity separation, and after the impurities such as hydrogen sulfide and ammonia contained therein are separated, the first stream from which the impurities have been separated enters the second reaction zone.
  • the impurity separation preferably includes gas-liquid separation and hydrogen sulfide stripping, so as to obtain the first stream of separated impurities in a liquid phase separated from impurities such as hydrogen sulfide and ammonia. More specifically, conventional separation techniques in the art can be used, such as gas-liquid separation containing gas-phase water injection to wash ammonia, liquid-phase stripping and hydrogen sulfide removal.
  • step 2) of the method of the present invention the first stream after separation of impurities is selectively converted in the second reaction zone under hydrogen conditions through reactions including hydrocracking .
  • the selective conversion includes a hydrocracking reaction, which selectively converts a first stream obtained after hydrofining into a second stream.
  • the second stream obtained in step 2) mainly contains dry gas (including methane and ethane), C 3 -C 5 light hydrocarbons, benzene-toluene fraction, xylene fraction, C 9 -C 10 fraction and heavy tail oil.
  • One of the purposes of selective conversion in step 2) is to carry out hydrocracking under the premise of retaining one aromatic ring of polycyclic aromatic hydrocarbons in the heavy aromatics in the first stream, effectively controlling the saturation depth and ring opening position, and at the same time making the first
  • the macromolecular non-aromatic hydrocarbons in the logistics are isomerized and cracked; the light aromatic hydrocarbons are maximized under economic hydrogen consumption.
  • the selective conversion reaction in this step can be carried out according to any known method of conventional hydrogenation reactions in the art, as long as the first stream can be selectively converted into the second stream.
  • the reaction conditions of the second reaction zone in step 2) can adopt the reaction conditions of conventional hydrocracking reactions in the art.
  • reaction conditions of the second reaction zone preferably include:
  • the hydrogen oil volume ratio is 800-5000 Nm 3 /m 3 , preferably 1000-4000 Nm 3 /m 3 , more preferably 1500-3000 Nm 3 /m 3 ;
  • the reactor inlet temperature is 280-450°C, preferably 300-430°C, more preferably 310-400°C
  • the partial pressure of hydrogen is 5-10 MPa, preferably 5-9 MPa, more preferably 6-8 MPa; and/or
  • the space velocity is 0.5-2.0 hr -1 , preferably 0.6-1.5 hr -1 , more preferably 0.8-1.2 hr -1 .
  • the selective conversion catalyst in step 2) can be any type of hydrocracking catalyst available in the art, as long as it can achieve the purpose of step 2).
  • the selective conversion catalyst described in the present invention is preferably the catalyst provided in Chinese patent application ZL201810153543.5.
  • the content of Chinese patent application ZL201810153543.5 is incorporated herein by reference in its entirety.
  • the preferred selective conversion catalyst is as follows:
  • the selective conversion catalyst includes: a 2 ) 5 to 80 parts of solid acid zeolite; b 2 ) 0.05 to 8 parts of group VIII metal; c 2 ) 3 to 25 parts of group VIB metal oxide; d 2 ) 0.1 to 2 parts of VIB group metal sulfide; e 2 ) 20 to 95 parts of the first binder; the weight parts of the above components are based on the total weight parts of the catalyst.
  • the selective conversion catalyst of the present invention may also include other auxiliary agents commonly used in catalysts in the field, such as diatomaceous earth, activated clay and the like.
  • the dosage can be a conventional dosage.
  • the solid acid zeolite is at least one of mordenite, beta zeolite, ZSM zeolite, EU-1 zeolite, SAPO zeolite and Y zeolite.
  • the crystal grain diameter of the solid acid zeolite is less than 500 nanometers, preferably less than 400 nanometers, more preferably less than 300 nanometers, and more preferably less than 200 nanometers.
  • the silicon-to-aluminum molecular ratio of the solid acid zeolite is 10 to 500, preferably 10 to 200, more preferably 11 to 80, and more preferably 20 to 60.
  • the group VIII metal is at least one of platinum, palladium, cobalt, nickel and iridium.
  • the VIB group metal oxide is at least one of molybdenum oxide and tungsten oxide.
  • the VIB group metal sulfide is at least one of molybdenum sulfide and tungsten sulfide.
  • the first binder is at least one of alumina, silica-alumina composite, titania-alumina composite, and magnesia-alumina composite.
  • the selective conversion catalyst of the present invention can be prepared by any known method in the art.
  • the carrier can be prepared by extrusion, rolling, or oil column forming methods in the art.
  • the catalyst can be prepared by forming the support and then impregnating the metal.
  • the selective conversion catalyst may be prepared by a method including the following steps:
  • the solid acid zeolite is mixed with the first binder, then kneaded, extruded, dried at 60-150°C, and calcined in an air atmosphere at 500-600°C for 3-6 hours to obtain the desired catalyst carrier.
  • a composite metal aqueous solution is prepared with a group VIII metal compound and a group VIB metal compound, the catalyst support is impregnated by an equal volume impregnation method, dried at 60-150°C and calcined in an air atmosphere at 450-520°C for 1 to 4 hours to obtain a catalyst precursor.
  • the catalyst precursor is reduced to 400 ⁇ 500°C under hydrogen condition and kept for 2 ⁇ 24 hours (pre-reduction), and then cooled to 300 ⁇ 380°C, after injecting sulfiding agent to vulcanize for 4 ⁇ 24 hours, you can get the required hydrogenation Cracking catalyst.
  • step 3) of the method of the present invention the second stream is subjected to first separation in the first separation zone, and the obtained C 6 -C 8 aromatic hydrocarbon stream at least includes Benzene, toluene, xylene and other fractions.
  • the first separation of the second stream preferably includes gas-liquid separation and rectification of the second stream; more preferably the benzene-toluene obtained after rectification
  • the fractions are extracted and separated.
  • the second stream undergoes gas-liquid separation to separate dry gas and liquid phase, wherein the dry gas is discharged outside, and the liquid phase is sent to the depentane tower for depentane; depentane separates the discharged C3-C5
  • the light hydrocarbon fraction and the bottom stream of the depentanizer tower are sent to the deheptane tower; the deheptane tower separates the stream rich in the benzene-toluene fraction and the bottom stream of the deheptane tower ,
  • the bottom stream of the deheptane tower is sent to the xylene tower; the top of the xylene tower separates the mixed xylene product and the bottom stream of the dexylene tower, and the bottom stream of the dexylene tower performs heavy aromatics removal; heavy aromatics removal Separate the C9-C10 sent out and the third stream separated from the bottom of the tower.
  • the third stream is heavy tail oil containing heavy aromatics above C10.
  • the heavy tail oil is sent to the post-saturation selective reactor.
  • the above deheptane tower separates a stream rich in benzene-toluene fractions, this stream is preferably extracted to separate pure benzene-toluene mixed aromatics, and the extracted non-aromatics are sent out.
  • the above-mentioned gas-liquid separation and rectification can be carried out by the extraction and rectification methods commonly used in this field.
  • the content of aromatics in the third stream obtained after separation of the second stream obtained by the selective conversion of the present invention is preferably higher than the content of non-aromatics; the third stream of the present invention is more preferably that the content of aromatics can reach more than 80% by weight, and most preferably reach 90% by weight. %the above.
  • step 4) of the method of the present invention the third stream containing heavy aromatic hydrocarbons above C 10 obtained in step 3) is subjected to high pressure in the post-saturation selective reaction zone under hydrogen, low temperature and low pressure conditions.
  • the selective hydrogenation saturation reaction results in a product with a benzene ring, forming a fourth stream containing the product, that is, a fraction with a boiling point greater than 210°C.
  • the hydrogenation saturation can be carried out according to any known method conventionally in the art, as long as the effect of the post-saturation selective reaction described above can be achieved.
  • the hydrogenation saturation of the post-saturation selective reaction zone in step 4) of the method of the present invention is preferably a liquid hydrogenation reaction in order to simplify the process, reduce equipment, and save energy consumption.
  • the reaction conditions can be the reaction conditions of conventional hydrogenation saturation reactions in the art, and preferably include:
  • the reactor inlet temperature is 100-300°C, preferably 120-280°C, more preferably 150-250°C;
  • the hydrogen partial pressure is 1.0-4.0 MPa, preferably 1.2-3.0 MPa; and/or
  • the space velocity is 0.1 to 5.0 h -1 , preferably 0.5 to 4.0 h -1 , more preferably 0.6 to 2.0 h -1 .
  • the third stream contacts the post-saturation selective catalyst in the post-saturation selective reaction zone to carry out the hydrogenation saturation reaction.
  • the post-saturation selective catalyst can be an existing hydrogenation saturation catalyst in the art, as long as it can be realized.
  • the above step 4) is sufficient for the purpose of hydrogenation saturation, such as the aromatic hydrocarbon hydrogenation saturation catalyst described in Chinese Patent CN103041832A.
  • the post-saturation selective catalyst in step 4) of the present invention may preferably be:
  • the post-saturation selective catalyst includes: a 3 ) 10 to 90 parts of amorphous silicon aluminum, wherein the content of silicon oxide is between 3 to 20 wt%; b 3 ) 0.1 to 5.0 parts of group VIII metal; c 3 ) 5 to 80 parts of the second binder; based on the total weight parts of the amorphous silicon-aluminum, the group VIII metal and the second binder.
  • the Group VIII metal is at least one selected from the group consisting of platinum, palladium, cobalt, nickel, and iridium.
  • the second binder is selected from alumina.
  • the post-saturation selective catalyst of the present invention can be prepared by any method known in the art.
  • the carrier can be prepared by extrusion, rolling, or oil column forming methods in the art.
  • the catalyst can be prepared by forming the support and then impregnating the metal.
  • the catalytic diesel used as feedstock can be from a catalytic cracking unit in this field, and its initial boiling point under normal pressure is between 160-210°C.
  • the composition of the catalyzed diesel is not particularly limited, and the composition of catalyzed diesel derived from crude oil from different places is not the same.
  • the catalytic diesel mainly contains components such as alkanes, cycloalkanes, alkenes, sulfur-containing hydrocarbons, nitrogen-containing hydrocarbons, C 11 + alkyl benzene, and fused ring aromatic hydrocarbons.
  • the content of C 11 + alkylbenzene ranges from 10 to 40 wt%
  • the content of fused-ring aromatic hydrocarbons ranges from 15 to 50 wt%
  • the content of sulfur ranges from 200 to 15000 wt ppm
  • the content of nitrogen ranges from 100 to 1500 wt ppm.
  • high boiling point alkanes cycloalkanes and alkenes.
  • Another object of the present invention is to provide a device for the full conversion method for producing light aromatic hydrocarbons from catalyzed diesel oil.
  • the device for producing light aromatics from catalytic diesel according to the present invention includes:
  • the first reaction zone for hydrorefining it is configured to receive the catalyzed diesel and emit the first stream;
  • a second reaction zone for selective conversion (including hydrocracking); it is configured to receive the first stream and discharge the second stream;
  • First separation zone configured to receive the second stream; discharge the third stream at the bottom;
  • a post-saturation selective reaction zone for hydrogenation saturation it is configured to receive the third stream and discharge the fourth stream;
  • the first pipe it is configured to circulate the fourth stream to the second reaction zone.
  • the device for producing light aromatics from catalytic diesel includes:
  • a first reaction zone it is configured to receive the catalyzed diesel and emit a first stream;
  • a second reaction zone it is configured to receive the first stream and discharge the second stream;
  • the first separation zone it is configured to receive the second stream; the discharge includes the C 6 ⁇ C 8 aromatics stream, the stream containing C 9 aromatics and C 10 aromatics, and the third stream containing heavy aromatics above C 10 Fraction within
  • Post-saturation selective reaction zone which is configured to receive the third stream and discharge the fourth stream
  • the first pipe it is configured to circulate the fourth stream to the second reaction zone.
  • the first reaction zone is equipped with a hydrorefining device, and the hydrorefining reactor used is a fixed bed reaction system.
  • the hydrorefining reactor used is a fixed bed reaction system.
  • an existing fixed bed reaction system in the field can be used, and a fixed bed reaction system equipped with a circulating hydrogen system is more preferred.
  • the inlet temperature of the hydrofining reactor may be 250-450°C.
  • the second reaction zone is configured with a hydrocracking reaction device for selective conversion
  • the hydrocracking reactor used is a fixed bed reaction system.
  • a fixed bed reaction system existing in the art can be used, and a fixed bed reaction system equipped with a circulating hydrogen system is more preferred.
  • the inlet temperature of the selective conversion (hydrocracking) reactor may be 280-450°C.
  • the post-saturation selective reaction zone is equipped with a hydrogenation saturation device, wherein the post-saturation selective reactor used is a fixed bed reaction system; more preferably, it is a liquid hydrogenation fixed without a circulating hydrogen system. Bed reaction system. Specifically, a fixed bed reaction system existing in the field can be used.
  • the inlet temperature of the post-saturation (hydrogenation saturation) reactor is between 100-300°C, and the reaction hydrogen partial pressure is between 1.0-4.0 MPa.
  • the first separation zone includes a gas-liquid separator and a rectification tower optionally connected in sequence
  • the rectification tower preferably includes a depentanizer optionally connected in sequence (the first Rectification tower), deheptane tower (second rectification tower), dexylene tower (third rectification tower), and de-heavy aromatics tower (fourth rectification tower) for sequential separation to obtain benzene-rich - stream of toluene fraction, xylene stream, the stream containing C 9 aromatics and C 10 aromatics, and a third stream containing the C 10 or more heavy aromatics, distillate.
  • the second stream passes through a gas-liquid separator to separate dry gas and liquid phase streams
  • the liquid phase stream passes through a depentane tower to separate the C3-C5 light hydrocarbon stream at the top of the tower and the depentane tower Bottom stream, which is sent to the deheptane tower.
  • the top of the deheptane tower separates a stream rich in benzene-toluene fraction and a stream at the bottom of the deheptane tower.
  • the stream rich in benzene-toluene fraction is preferably passed through an extraction device to separate pure benzene-toluene mixed aromatics, The non-aromatics separated by extraction are sent out.
  • the deheptane tower bottom stream enters the xylene tower to directly separate the mixed xylene product and the dexylene tower bottom stream.
  • the bottom stream of the xylene tower is sent to the heavy aromatics tower, and the C9-C10 sent from the top of the tower is separated from the third stream separated from the bottom of the tower.
  • the third stream is sent to the post-saturation selective reactor.
  • the extraction and rectification can use the extraction and rectification methods commonly used in this field.
  • the gas-liquid separator, rectification tower and extraction device can also use conventional equipment in the field.
  • a second separation zone is arranged between the first reaction zone and the second reaction zone to separate sulfides and/or nitrides in the first stream. ⁇ impurities.
  • the second separation zone is configured to receive the first stream, and discharge gas phase, hydrogen sulfide and ammonia streams, and a first stream from which impurities have been separated.
  • the separation device of the second separation zone can use conventional separation devices in the field, such as a gas-liquid separator (with gas-phase water injection to wash ammonia), a stripping device (such as a liquid-phase stripping and hydrogen sulfide removal device, etc.) Wait.
  • the method of the present invention removes the sulfur and nitrogen impurities in the catalytic diesel stream by passing the catalytic diesel stream of the catalytic cracking unit through the first reaction zone for hydrorefining, and makes the polycyclic aromatic hydrocarbons and polycyclic aromatic hydrocarbons selective. Hydrogenation saturation reaction, hydrogenation to products with only one aromatic ring, such as tetrahydronaphthalene, indene and polyalkylbenzene, etc., and then the stream is optionally separated from impurities and sent to the second reaction zone for selective conversion.
  • Hydrocracking reaction produces a stream rich in light aromatics such as benzene, toluene, xylene, C 9 aromatics, C 10 aromatics, and then the product stream passes through the rectification tower to separate benzene after the light components before the removal of benzene -Toluene, xylene, C 9 aromatics, C 10 aromatics and heavy tail oil at the bottom of the tower (mainly containing heavy aromatics); the heavy tail oil at the bottom of the tower enters the saturated selective reactor, and high selection occurs under low temperature and low pressure conditions Hydrogenation is saturated to obtain a product that retains an aromatic ring, which is sent to the second reaction zone for selective conversion of the hydrocracking reaction to realize the full conversion process of producing light aromatics from catalytic diesel, which improves the yield of light aromatics and reduces Reduce the loss of aromatics and reduce hydrogen consumption.
  • the existing technical problems are better solved, and good technical effects have been achieved for increasing the production of aromatic products.
  • the saturation rate of the fused ring aromatic hydrocarbons in the catalytic diesel stream is greater than 50% after the hydrorefining in the first reaction zone, the sulfur content is reduced to 100 ppm, the nitrogen content is reduced to 15 ppm, and the final boiling point Reduce by more than 10°C; after the catalytic diesel stream passes through the hydrotreating unit in the first reaction zone and the selective conversion device in the second reaction zone in turn, it is converted into monocyclic aromatic hydrocarbons with carbon ten and below, and the conversion rate is greater than 50%.
  • the hydrogenation saturation selectivity of the stream after passing through the post-saturation selective reactor is high, and the aromatics retention rate is greater than 98%.
  • the technical scheme of the present invention adopts a two-stage hydrorefining-selective conversion process, and a two-stage dual catalyst (hydrorefining catalyst and selective conversion catalyst) series scheme, including hydrorefining, selective conversion and heavy After the quality tail oil is saturated. It mainly solves the technical problems in the prior art that the full-cut catalytic diesel cannot be completely converted and the yield of light aromatics in the conversion process is not high.
  • the heavy tail oil of light aromatics made from catalyzed diesel is sent to the post-treatment reactor, where a selective saturation reaction occurs under mild pressure and temperature conditions.
  • the selectivity of hydrogenation saturation is greatly increased to more than 98% or even higher, which solves the problem of excessive hydrogenation saturation; it also helps to reduce the cracking hydrogen consumption reaction that occurs when the non-aromatic generated by excessive hydrogenation enters the selective conversion reactor.
  • This method improves the technical and economic indicators of the overall process for preparing light aromatics from catalytic diesel, and realizes the conversion of full fractions of catalytic diesel. Comparing the hydrorefining-selective conversion two-stage catalytic diesel to light aromatics technology process, the yield of benzene, toluene, xylene, and C 9 and C 10 monocyclic light aromatics of the present invention can be increased by at least 2%. It is preferably increased by 5% or more.
  • Figure 1 is a schematic process flow diagram of the full conversion method for producing light aromatics from catalytic diesel according to the present invention.
  • 5 is a gaseous stream containing hydrogen sulfide and ammonia
  • the first separation zone 12 is the first separation zone, for example including gas-liquid separator, depentanizer, deheptane tower, xylene tower, heavy aromatics tower and other distillation towers and benzene-toluene fraction extraction device
  • 17 is the heavy tail oil stream separated in the first separation zone-the third stream
  • the pressure mentioned in this manual is gauge pressure.
  • Figure 1 is a schematic process flow diagram of an exemplary embodiment of the method for producing light aromatics from catalyzed diesel according to the present invention.
  • Many conventional equipment such as pumps, compressors, heat exchangers, extraction devices, hydrogen pipelines, etc. are omitted in the figure. However, these devices are well known to those of ordinary skill in the art.
  • the flow of an exemplary implementation of the method of the present invention is described in detail as follows:
  • the catalytic diesel 1 as the feedstock enters the hydrorefining unit of the first reaction zone 2 to obtain hydrorefined catalytic diesel containing hydrogen sulfide and ammonia, that is, the first reaction zone exit stream 3 (first stream);
  • the stream passes through the gas-liquid separator 4 and the hydrogen sulfide stripper 7 in the second separation zone 20 to separate hydrogen sulfide and ammonia obtained from the denitrification and desulfurization in the hydrorefining process (through the gaseous stream 5 containing hydrogen sulfide and ammonia and the sulfur-containing After the hydrogen stripped stream 8), a first stream 9 from which impurities have been separated is obtained.
  • This stream enters the selective conversion device of the second reaction zone 10.
  • the exit stream 11 (second stream) of the second reaction zone rich in light aromatics such as benzene, toluene, xylene, C9A and C10A fractions, and heavy tail oil enters the first separation zone 12 to obtain dry gas and C3- C5 light hydrocarbon stream 13 benzene - 14 toluene stream, xylene stream 15 containing C 9 aromatics and C 10 aromatics stream 16, and C 10 containing the heavy aromatics or more heavy tail third stream 17.
  • the third stream 17 enters the post-saturation selective reactor 18 of the post-saturation selective reaction zone, and the post-saturation selective reactor outlet stream 19 (the fourth stream) is recycled to the selective conversion device of the second reaction zone 10 without being separated.
  • the first separation zone 12 includes a rectification tower such as a gas-liquid separator, a depentanizer, a deheptane tower, a xylene tower, and a heavy aromatics tower, which are connected in sequence, and a benzene-toluene fraction extraction device (in the drawings) Not shown).
  • a rectification tower such as a gas-liquid separator, a depentanizer, a deheptane tower, a xylene tower, and a heavy aromatics tower, which are connected in sequence, and a benzene-toluene fraction extraction device (in the drawings) Not shown).
  • composition analysis of the catalysts involved in the present invention all adopt analysis methods known in the art.
  • the composition of the catalyst can be analyzed by ICP (inductively coupled plasma) and XRF (X-ray fluorescence) methods.
  • the XPS (X-ray Photoelectron Spectroscopy) method was used to determine the composition ratio of VIB group metal oxides and metal sulfides.
  • the ICP test is performed with a Varian 700-ES series XPS instrument.
  • the XRF test is carried out using Rigaku ZSX 100e XRF instrument.
  • XPS test conditions include: Perkin Elmer PHI 5000C ESCA X-ray photoelectron spectrometer, using Mg K excitation light source, operating voltage 10kV, current 40mA, vacuum degree 4.0 ⁇ 10-8Pa.
  • the comprehensive two-dimensional gas chromatography/high-throughput time-of-flight mass spectrometer (GC ⁇ GC-TOFMS) of American LECO Company is used to analyze (multi-dimensional chromatographic analysis) the family composition of catalytic diesel and hydrorefined catalytic diesel. Analyze the group composition of heavy tail oil and select saturated heavy tail oil.
  • the composition of the reactant stream (such as selective conversion products, etc.) is determined by gas chromatography.
  • the chromatographic model is Agilent 7890A, equipped with FID detector, FFAP capillary chromatographic column for separation, the chromatographic column adopts temperature program, the initial temperature is 90°C, keep for 15 minutes, and then increase to 220°C at a rate of 15°C/min for 45 minutes.
  • the catalyst raw materials of the examples and comparative examples of the present invention are all commercially available.
  • a two-stage hydrorefining-selective conversion method is used to process catalytic diesel, which means that the catalytic diesel used as feedstock is hydrorefined and separated from impurities, and then hydrocracked, and then the hydrocracked product is passed through a gas-liquid separation and rectification system , And separated products such as benzene-toluene, xylene, C 9 A aromatics, C 10 A aromatics and heavy tail oil.
  • the process flow of Comparative Example 1 does not include sending heavy tail oil >210°C into the post-saturation selective reaction zone for selective hydrogenation saturation treatment.
  • the analysis data of the catalytic diesel raw material and the hydrofining product are shown in Table 1.
  • the aromatic content of the catalytic diesel is 87.15wt%.
  • Table 2 lists the hydrorefining catalysts, selective conversion (hydrocracking) catalysts and their reaction conditions used.
  • the preparation of the used hydrorefining catalyst A1 Add 2g of Sesbania powder, 9ml of nitric acid and 60ml of water to 100g of pseudo-boehmite, knead it into a dough, extrude it, cure for 24h at room temperature, and dry it at 100°C for 12h. It is calcined at 550°C for 3 hours to obtain a hydrorefining catalyst carrier. 7.90g nickel nitrate hexahydrate, 8.71g ammonium molybdate, 9.18g ammonium metatungstate and 10ml ammonia are dissolved in water to obtain 50ml clear solution.
  • the composition of the catalyst A1 is 3.0 wt% NiO-10.5 wt% MoO 3 -12.7 wt% WO 3 /73.8 wt% Al 2 O 3 , that is, it contains nickel, molybdenum, and tungsten.
  • the catalytic diesel After the catalytic diesel is mixed with hydrogen, it enters the hydrorefining reactor to remove most of the sulfur and nitrogen impurities, and the fused-ring aromatic hydrocarbons are saturated to be converted into hydrocarbons containing only one aromatic ring.
  • Table 1 also lists the sulfur and nitrogen content, density, aromatic hydrocarbon content, and fraction distribution of the hydrorefined products.
  • the first stream after the hydrorefining of catalyzed diesel is subjected to impurity separation, including the first stream is subjected to gas-liquid separation, and stripped with nitrogen under normal pressure for 3 hours to fully remove hydrogen sulfide dissolved in the first stream.
  • the sulfur content and nitrogen content of the hydrofining product were 87 ppm and 8.6 ppm, respectively.
  • the retention rate of the fused-ring aromatics in the hydrorefining process is 89.04wt%.
  • Table 2 also lists the composition of the selective conversion catalyst B1 used in hydrocracking and the reaction conditions used.
  • USY zeolite and alumina are kneaded and extruded to obtain selective conversion catalyst carrier.
  • an appropriate amount of chloroplatinic acid is prepared into a clear solution, which is immersed in an equal volume and then dried and calcined in air at 500°C for 2 hours to obtain a selective conversion catalyst precursor.
  • the selective conversion catalyst precursor is reduced to 450° C. under hydrogen conditions to obtain the required selective conversion catalyst B1, the composition of which is: 0.1 parts Pt-60 parts USY zeolite-39.9 parts Al 2 O 3 .
  • the catalyst bed is cooled to 340°C, and the stripped hydrofining product (the first stream from which the impurities have been separated) is mixed with hydrogen and enters the selective conversion reactor, and the reaction product is sent to the gas-liquid separation and rectification system.
  • the yield of the obtained heavy tail oil at >210°C is 38.27wt%, the specific gravity is 0.935, and its sulfur and nitrogen content are 19.5ppm and 1.5ppm respectively;
  • the group composition of the third stream can be obtained by multi-dimensional chromatographic analysis: non-aromatics account for 41.98wt% , Monocyclic aromatic hydrocarbons accounted for 26.38wt%, fused ring aromatic hydrocarbons accounted for 31.64wt%.
  • the process of catalyzing the full conversion of diesel to produce light aromatics is shown in FIG. Including the hydrorefining of catalytic diesel, separation of impurities, selective conversion (hydrocracking), and then sending the heavy tail oil >210°C obtained after selective conversion to the post-saturation selective reaction zone for selective hydrosaturation, specifically as follows: :
  • the raw materials, hydrorefining catalyst, and hydrorefining reaction conditions are the same as those in Comparative Example 1, and the selected conversion catalyst B2 (hydrocracking catalyst) and reaction conditions are shown in Table 3.
  • Table 3 lists the composition of the selective conversion catalyst B2 and the reaction conditions used.
  • the first stream after the hydrorefining of catalyzed diesel is subjected to impurity separation, including the first stream is subjected to gas-liquid separation, and stripped with nitrogen under normal pressure for 3 hours to fully remove hydrogen sulfide dissolved in the first stream.
  • the stripped hydrofining product (the first stream from which impurities have been separated) is mixed with hydrogen and enters the selective conversion reactor, and the reaction product is sent to the gas-liquid separation and rectification system.
  • benzene-toluene, xylene, C 9 A aromatics and C 10 A aromatics are separated and obtained by calculation, benzene-toluene, xylene, C 9 A aromatics and C 10 A aromatics, etc.
  • the yield of monocyclic light aromatics is 32.27wt%.
  • the obtained heavy tail oil (third stream) at >210°C has a yield of 24.75wt%, a specific gravity of 0.957, and a sulfur and nitrogen content of 25.4ppm and 1.6ppm, respectively;
  • the group composition of the third stream can be obtained by multi-dimensional chromatographic analysis as shown in the table 4 shows: non-aromatic hydrocarbons accounted for 8.54 wt%, monocyclic aromatic hydrocarbons accounted for 37.56 wt%, and fused ring aromatic hydrocarbons accounted for 53.90 wt%.
  • the post-saturation selective catalyst C2 is prepared as follows: a commercial amorphous silica-alumina material with a SiO 2 content of 20 wt% is mixed with pseudo-boehmite, nitric acid peptizer, sesbene powder extrusion aid and appropriate amount of water are added, and then extruded after kneading Shaped, dried in air at 100°C for 24 hours, and calcined in air at 550°C for 4 hours to obtain a catalyst carrier.
  • the process of catalyzing the full conversion of diesel to produce light aromatics is shown in FIG. Including the hydrorefining of catalytic diesel, separation of impurities, selective conversion (hydrocracking), and then sending the heavy tail oil >210°C obtained after selective conversion to the post-saturation selective reaction zone for selective hydrosaturation, specifically as follows: :
  • the raw materials, hydrorefining catalyst, and hydrorefining reaction conditions are the same as those in Comparative Example 1, and the selected conversion catalyst B3 (hydrocracking catalyst) and reaction conditions are shown in Table 6.
  • Table 6 lists the composition of the selective conversion catalyst B3 and the reaction conditions used.
  • a trimetal solution was prepared with palladium chloride, nickel nitrate and ammonium molybdate, and the selective conversion catalyst carrier was impregnated by an equal volume impregnation method, dried at 120°C and calcined in an air atmosphere at 500°C for 2 hours to obtain a selective conversion catalyst precursor.
  • the selective conversion catalyst precursor is reduced to 450° C.
  • the composition of catalyst B3 is: 0.2 parts Pd-6.5 parts Ni-4.2 parts MoO 2 -7.9 parts MoO 3 -1.1 parts MoS 2 -35 parts mordenite-10 parts ⁇ -zeolite-11 Parts of ZSM-5-24.1 parts of Al 2 O 3 .
  • the first stream after the hydrorefining of catalytic diesel oil is subjected to impurity separation the first stream is subjected to gas-liquid separation, and stripped with nitrogen at normal pressure for 3 hours to fully remove the hydrogen sulfide dissolved in the first stream.
  • the stripped hydrofining product (the first stream from which impurities have been separated) is mixed with hydrogen and enters the selective conversion reactor, and the reaction product is sent to the gas-liquid separation and rectification system.
  • the obtained heavy tail oil (third stream) at >210°C has a yield of 33.15 wt%, a specific gravity of 0.961, and its sulfur and nitrogen content of 16.4 ppm and 0.8 ppm, respectively;
  • the group composition of the third stream can be obtained by multi-dimensional chromatographic analysis as shown in the table As shown in 7: non-aromatic hydrocarbons accounted for 7.58wt%, monocyclic aromatic hydrocarbons accounted for 38.12wt%, and fused ring aromatic hydrocarbons accounted for 54.30wt%.
  • C and saturated catalyst selected heavy tail composition C3 of 0.10wt% Pt-0.30% Pd- 4.0wt% Ni-6.0wt% SiO 2 -89.6wt% Al 2 O 3.
  • the post-saturation selective catalyst C3 is prepared as follows: a commercial amorphous silicon-alumina material with a SiO 2 content of 9% is mixed with pseudo-boehmite, nitric acid peptizer, sesbene powder extrusion aid and appropriate amount of water are added, and then extruded after kneading Shaped, dried in air at 100°C for 24 hours, and calcined in air at 550°C for 4 hours to obtain a catalyst carrier.

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Abstract

本发明涉及一种从催化柴油生产轻质芳烃的全转化方法和装置。本发明的技术方案包括将催化柴油物流经过加氢精制、分离杂质后进行选择转化反应，产生的混合芳烃经过分离依次分离出苯-甲苯、二甲苯等轻质芳烃、C 9A芳烃、C 10A芳烃和塔底重质尾油；塔底重质尾油进入后饱和选择反应器，在低温低压条件下高选择性加氢饱和得到一个苯环的产物后，送回选择转化反应器。实现从催化柴油生产轻质芳烃的全馏分转化，具有苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等单环芳烃收率较高的技术效果。

Description

一种从催化柴油生产轻质芳烃的全转化方法和装置 技术领域
本发明涉及石油催化裂化领域中制备轻质芳烃的技术，具体涉及一种从催化柴油生产轻质芳烃的方法和装置。
背景技术
苯、甲苯和二甲苯等轻质芳烃是重要的基本有机化工原料，广泛应用于合成材料等领域，与国民经济发展及人们的衣食住行密切相关。目前芳烃的来源主要有两条工艺路线：一是石脑油经过催化重整、芳烃抽提得到芳烃原料；二是将乙烯装置的副产品-裂解汽油经过加氢、芳烃抽提得到芳烃原料。
催化柴油(LCO)的主要成分是C 11 +烷基苯和稠环芳烃，由于含有大量稠环芳烃等原因，将其加工成柴油的经济性不佳，部分企业只能将其作为燃料油使用。随着柴油需求增长停滞，亟需开发高效转化技术，将催化柴油通过加氢裂化反应转化为轻质芳烃，通过炼化一体实现芳烃产业的降本增效。
当前普遍采用的催化柴油改质手段是加氢精制、加氢改质和轻油型加氢裂化。催化柴油加氢精制是在中、低压的条件下，进行烯烃加氢饱和、脱硫、脱氮及芳烃部分饱和反应，可改善其颜色和安定性。但是对于加工劣质原料的催化装置得到的催化柴油，通过加氢精制还远不能满足产品对十六烷值的要求。加氢改质工艺，如UOP公司的Unicracking工艺(US5026472)，其目标产物是高十六烷值柴油。该工艺具有良好芳烃加氢饱和性能和开环选择性，芳烃转化深度很高，保证较大的十六烷值提高幅度和较高的柴油收率。轻油型加氢裂化则是将轻柴油组分经过精制后，剧烈饱和加氢，得到石脑油馏分的重整料或汽油馏分，此过程也存在原料转化为芳烃收率较低的问题。若石脑油馏分用于重整制芳烃原料，过度饱和后生成的环烷烃和链烷烃还要在重整装置中转化为芳烃，不是一条经济的路线。如CN101684415专利描述的轻油型加氢裂化方法，不直接产芳烃，重石脑油的芳潜最高只有57％。
文献CN1955262A描述了一种两段加氢裂化方法，其加氢裂化催化剂含Pt和/Pd贵金属及非贵金属，以及Y沸石和氧化铝，原料是催化柴油。但是，其石脑油产品的芳潜值最高只有76.8％，芳烃的纯度不高， 达不到芳烃联合装置的要求。文献CN103897731A描述了一种催化裂化柴油和C 10 +馏分油混合生产轻质芳烃的方法，通过加氢精制和加氢裂化，产品进行切割，大于195℃馏分用作清洁柴油调和组分，小于195℃馏分进芳烃装置生产轻质芳烃和清洁汽油调和组分，芳烃产品的收率相对不高。
已有的催化柴油转化技术中，重质尾油作为柴油组分外排或者部分循环回加氢精制反应器，未能有效地全部利用于增产轻质芳烃。
而且，金属硫化物型加氢精制催化剂上的加氢精制反应需要在高温高压的苛刻操作条件下进行，反应受热力学平衡限制，对稠环芳烃部分饱和反应的选择性不佳，LCO经过加氢精制后芳烃保留率低于90％。催化柴油生产轻质芳烃的重质尾油中稠环芳烃含量大于90％，硫氮含量都较低，循环回加氢精制反应器处理会带来过度饱和及芳烃损失问题。
发明内容
针对现有技术的问题，本发明人进行一系列研究发现，将催化柴油物流经过加氢精制、分离杂质后进行包括加氢裂化在内的选择转化反应，产生的混合芳烃经过分离依次分离出苯-甲苯、二甲苯、含C 9芳烃、C 10芳烃物流和富含C 10以上重芳烃的塔底重质尾油；其中塔底重质尾油进入后饱和选择反应器，在低温低压条件下高选择性加氢饱和得到具有一个苯环的产物后，再送往选择转化反应，从而实现从催化柴油生产轻质芳烃的全馏分转化，具有较好的轻芳烃收率。
本发明所述的轻质芳烃是指碳数小于或等于10的芳烃，包括C6芳烃，例如苯；C7芳烃，例如甲苯；C8芳烃，例如乙苯、二甲苯；C9芳烃，例如甲乙苯、丙苯、三甲苯；C10芳烃，例如四甲苯、二甲基乙基苯、二乙苯等。相应地，本发明所述C 10以上重芳烃是指碳数大于10的芳烃。
本发明的目的之一是提供一种从催化柴油生产轻质芳烃的全转化方法。
本发明所述的从催化柴油生产轻质芳烃的全转化方法，包括以下步骤：
1)使催化柴油进入第一反应区进行加氢精制，得到第一物流；
2)使所述第一物流进入第二反应区进行选择转化，得到第二物流，其中任选地，所述第一物流进入第二反应区前在第二分离区中进行杂质 分离；
3)使所述第二物流在第一分离区中进行第一分离，在该第一分离区底部得到含C10以上重芳烃的第三物流；
4)使所述第三物流进入后饱和选择反应区进行加氢饱和，得到第四物流；
5)使所述第四物流循环至所述第二反应区。
在一个例示实施方案中，本发明提供的从催化柴油生产轻质芳烃的全转化方法包括以下步骤：
1)催化柴油进入第一反应区，在临氢条件下与加氢精制催化剂接触，得到第一物流；所述第一反应区进行加氢精制反应；
2)所述第一物流脱除杂质后进入第二反应区，在临氢条件下与选择转化催化剂接触，得到第二物流；所述选择转化包括加氢裂化反应；
3)所述第二物流经分离后，得到包括C 6～C 8芳烃物流、含C 9芳烃和C 10芳烃物流、以及含C 10以上重芳烃的第三物流在内的馏分；
4)所述第三物流进入后饱和选择反应区，在临氢条件下与后饱和选择催化剂接触，得到第四物流；所述后饱和选择利用加氢饱和反应进行；
5)所述第四物流循环至所述第二反应区。
根据本发明的一个方面：在本发明所述方法的步骤1)中，作为原料油的所述催化柴油在临氢条件下在第一反应区进行加氢精制，其中，催化柴油物流和氢气与加氢精制催化剂接触，进行脱硫脱氮，并发生保留一个芳环的稠环芳烃的选择饱和反应。所述加氢精制可以按照本领域常规已知的任何方式和任何方法进行，只要将所述催化柴油实现脱硫脱氮，并将其中的稠环芳烃加氢饱和保留一个芳环即可，并没有特别的限定。催化柴油经过加氢精制后得到的第一物流，主要包含脱除了绝大部分硫氮杂质的精制催化柴油、含硫化氢和氨的气相。
在本发明所述方法的步骤1)中，作为原料油的催化柴油与氢气在第一反应区与加氢精制催化剂接触进行加氢精制反应。
所述的加氢精制反应为本领域公知的催化柴油加氢精制技术。其加氢精制反应条件可采用本领域中已知的催化柴油加氢精制的反应条件；其所述加氢精制催化剂可以采用本领域中已有的任何类型的加氢精制催化剂，只要能实现步骤1)的催化柴油加氢精制目的即可。
在本发明所述方法的步骤1)中，第一反应区加氢精制反应条件优选包括：
氢油体积比500～3000Nm 3/m 3，优选800～2000Nm 3/m 3，更优选1000～1500Nm 3/m 3；
反应器入口温度280-420℃，优选300～410℃，更优选310～390℃；
氢气分压力为5～10MPa，优选5～8MPa，更优选6～7MPa；和/或
空速0.5～2.0小时 -1，优选0.6～1.5小时 -1，更优选0.8～1.2小时 -1。
在本发明的方法中，步骤1)的加氢精制催化剂可优选为以下：
以重量份数计包括：a1)60～99.9份，优选65～99.9份，优选70～99.9份，更优选75～99.9份载体；和b1)加氢金属氧化物，其中加氢金属氧化物的重量份数为0.1～40份，优选0.1～35份，优选0.1～30份，更优选0.1～25份；基于所述载体和所述加氢金属氧化物的总重量份数。
在一个例示实施方案中，以重量份数计，所述载体包括：60～100份的氧化铝；0～40份的氧化硅；基于所述氧化铝、所述氧化硅的总重量份数。
在一个例示实施方案中，所述加氢金属选自由镍、钴、钼、钨和铁组成的组中的至少一种。所述加氢金属在负载后进行硫化。
本发明所述加氢精制催化剂可采用本领域已知的任何方法进行制备，例如载体可通过本领域的挤条、滚球或油柱成型等方法制备。在一个实施方案中，催化剂可通过载体成型然后浸渍金属的方法制备。
优选地，将经过步骤1)所述加氢精制得到的第一物流进行杂质分离，分离其中包含的硫化氢和氨等杂质后，使已分离出杂质的第一物流再进入第二反应区。所述的杂质分离优选包括气液分离和硫化氢汽提，以得到分离了硫化氢和氨等杂质的液相的所述已分离出杂质的第一物流。更具体地可采用本领域常规的分离技术，比如含气相注水洗氨的气液分离、液相汽提脱硫化氢等。
根据本发明的一个方面：在本发明所述方法的步骤2)中，对分离杂质后的第一物流在临氢条件下于第二反应区中通过包括加氢裂化在内的反应进行选择转化。例如，所述选择转化包括加氢裂化反应，其将加氢精制后得到的第一物流选择转化为第二物流。步骤2)得到的第二物流主要包含干气(包括甲烷和乙烷)、C 3-C 5轻烃、苯-甲苯馏分、二甲苯馏分、C 9-C 10馏分和重质尾油。步骤2)中选择转化的目的之一是 在保留第一物流中重芳烃中多环芳烃的一个芳环的前提下发生加氢裂化，有效控制饱和深度和开环位置，同时还能使得第一物流中的大分子非芳烃发生异构化及裂解；在经济氢耗下最大化生产轻质芳烃。本步骤的选择转化反应可以按照本领域常规的加氢反应的任何已知方法进行，只要能将所述第一物流选择转化成所述的第二物流即可。
在本发明的方法中，所述步骤2)第二反应区的反应条件可采用本领域中常规加氢裂化反应的反应条件。
在本发明中，所述第二反应区的反应条件优选包括：
氢油体积比800～5000Nm 3/m 3，优选1000～4000Nm 3/m 3，更优选1500～3000Nm 3/m 3；
反应器入口温度280-450℃，优选300～430℃，更优选310～400℃
氢气分压力5～10MPa，优选5～9MPa，更优选6～8MPa；和/或
空速0.5～2.0小时 -1，优选0.6～1.5小时 -1，更优选0.8～1.2小时 -1。
所述步骤2)中的选择转化催化剂可以采用本领域中已有的任何类型的加氢裂化催化剂，只要能实现上述步骤2)目的即可。
为了更好的实现本发明所述的第一物流向第二物流的转化，本发明所述的选择转化催化剂优选为中国专利申请ZL201810153543.5中提供的催化剂。中国专利申请ZL201810153543.5的内容在此通过引用的方式全文并入本文中。
优选的选择转化催化剂具体如下：
以重量份数计，所述选择转化催化剂包括：a 2)5～80份固体酸沸石；b 2)0.05～8份VIII族金属；c 2)3～25份VIB族金属氧化物；d 2)0.1～2份VIB族金属硫化物；e 2)20～95份第一粘结剂；以上各组分的重量份数基于催化剂总重量份数。本发明所述选择转化催化剂除以上主要组分外还可以包括本领域催化剂常用的其他助剂，例如硅藻土、活性粘土等。用量可为常规用量。
优选所述固体酸沸石为丝光沸石、β沸石、ZSM沸石、EU-1沸石、SAPO沸石和Y沸石中的至少一种。
优选所述固体酸沸石的晶粒直径小于500纳米，优选小于400纳米，更优选小于300纳米，更优选小于200纳米。
优选所述固体酸沸石的硅铝分子比为10～500，优选10～200，更优选11～80，更优选20～60。
优选所述VIII族金属为铂、钯、钴、镍和铱中的至少一种。
优选所述VIB族金属氧化物为氧化钼和氧化钨中的至少一种。
优选所述VIB族金属硫化物为硫化钼和硫化钨中的至少一种。
优选所述第一粘结剂为氧化铝、氧化硅-氧化铝复合物、氧化钛-氧化铝复合物和氧化镁-氧化铝复合物中的至少一种。
本发明所述选择转化催化剂可采用本领域的任何已知方法进行制备，例如载体可通过本领域的挤条、滚球或油柱成型等方法制备。在一个实施方案中，催化剂可通过载体成型然后浸渍金属的方法制备。在一个实施方案中，所述选择转化催化剂可以用包括以下步骤的方法制备：
将固体酸沸石与第一粘合剂混合，然后混捏、挤条、60～150℃烘干，在500～600℃空气气氛中焙烧3～6小时，即得所需的催化剂载体。以VIII族金属化合物和VIB族金属化合物配制复合金属水溶液，通过等体积浸渍方法浸渍催化剂载体，60～150℃烘干后在450～520℃空气气氛中焙烧1～4小时，得到催化剂前体。催化剂前体在氢气条件下还原到400～500℃并保持2～24小时(预还原)，之后降温到300～380℃后，注入硫化剂硫化4～24小时，即可得到所需的加氢裂化催化剂。
根据本发明的一个方面，在本发明所述的方法的步骤3)中，使所述第二物流在第一分离区中进行第一分离，得到的所述C 6～C 8芳烃物流至少包括苯、甲苯、二甲苯等馏分。
对于本发明所述的方法，在所述步骤3)中，第二物流的所述第一分离优选包括对第二物流进行气液分离、精馏；更优选对精馏后得到的苯-甲苯馏分进行抽提分离。
具体地，所述第二物流经过气液分离，分离出干气和液相，其中干气外放，液相送入脱戊烷塔进行脱戊烷；脱戊烷分离出外放的C3-C5的轻烃馏分和脱戊烷塔塔底物流，该脱戊烷塔塔底物流送入脱庚烷塔；脱庚烷塔分离出富含苯-甲苯馏分的物流和脱庚烷塔塔底物流，该脱庚烷塔塔底物流送入二甲苯塔；二甲苯塔顶分离出混合二甲苯产品和脱二甲苯塔塔底物流，该脱二甲苯塔塔底物流进行脱重芳烃；脱重芳烃分离出外送的C9-C10和塔底分离出的第三物流。所述的第三物流即含C10以上重芳烃的重质尾油。该重质尾油送入后饱和选择反应器。以上脱庚烷塔分离出富含苯-甲苯馏分的物流，这股物流优选经过抽提后分离出纯的苯-甲苯混合芳烃，抽提分离出的非芳烃外送。以上所述的气液分离 和精馏都可采用本领域中常用的抽提和精馏方法进行。本发明经所述选择转化得到的第二物流经分离后得到的第三物流中芳烃含量优选高于非芳烃含量；本发明第三物流更优选其中芳烃含量可达80wt％以上，最优选达到90wt％以上。
根据本发明的一个方面，在本发明方法所述步骤4)中将步骤3)得到的含C 10以上重芳烃的第三物流在后饱和选择反应区中于临氢、低温低压条件下进行高选择性的加氢饱和反应，从而得到具有一个苯环的产物，形成包含该产物的第四物流，即馏点大于210℃的馏分。所述加氢饱和可以按照本领域常规的任何已知方法进行，只要能实现以上所述后饱和选择反应的效果即可。
本发明方法所述步骤4)的后饱和选择反应区的加氢饱和优选为液态加氢反应，以便简化流程，减少设备，节省能耗。反应条件可采用本领域中常规加氢饱和反应的反应条件，优选包括：
氢油体积比200～3000Nm 3/m 3，优选300～1500Nm 3/m 3，更优选300～1000Nm 3/m 3；
反应器入口温度100～300℃，优选120～280℃，更优选150～250℃；
氢气分压力1.0～4.0MPa，优选1.2～3.0MPa；和/或
空速0.1～5.0小时 -1，优选0.5～4.0小时 -1，更优选0.6～2.0小时 -1。
所述步骤4)中第三物流在后饱和选择反应区与后饱和选择催化剂接触进行加氢饱和反应，其所述后饱和选择催化剂可以采用本领域中已有的加氢饱和催化剂，只要能实现上述步骤4)加氢饱和的目的即可，如中国专利CN103041832A所记载的芳烃加氢饱和催化剂。
本发明所述步骤4)的后饱和选择催化剂可优选为：
以重量份数计，所述后饱和选择催化剂包括：a 3)10～90份无定型硅铝，其中氧化硅含量介于3-20wt％；b 3)0.1～5.0份VIII族金属；c 3)5～80份第二粘结剂；基于无定型硅铝、VIII族金属和第二粘结剂总重量份数。
在一个实施方案中，所述VIII族金属选自由铂、钯、钴、镍和铱组成的组中的至少一种。
在一个实施方案中，所述第二粘结剂选自氧化铝。
本发明所述后饱和选择催化剂可采用本领域的任何已知方法进行制备，例如载体可通过本领域的挤条、滚球或油柱成型等方法制备。在一个实施方案中，催化剂可通过载体成型然后浸渍金属的方法制备。
对于本发明所述的催化柴油生产轻质芳烃的全转化方法，其中作为原料油的催化柴油可以来自本领域的催化裂化装置，其常压下初馏点介于160～210℃。对所述催化柴油的组成没有特别的限定，可源自不同产地原油的催化柴油，组成不尽相同。但作为举例，所述催化柴油主要含有烷烃、环烷烃、烯烃、含硫烃类、含氮烃类、C 11 +烷基苯和稠环芳烃等组分。其中，C 11 +烷基苯的含量范围为10～40wt％，稠环芳烃的含量范围为15～50wt％，硫的含量范围为200～15000wt ppm，氮的含量范围为100～1500wt ppm，其它为高沸点烷烃、环烷烃和烯烃。
本发明的另一目的是提供所述催化柴油生产轻质芳烃的全转化方法的装置。
本发明所述的一种从催化柴油生产轻质芳烃的装置包括：
进行加氢精制的第一反应区；其配置成接收所述催化柴油、以及排放第一物流；
进行选择转化(包括加氢裂化)的第二反应区；其配置成接收所述第一物流、以及排放第二物流；
第一分离区；其配置成接收所述第二物流；在底部排放所述第三物流；
进行加氢饱和的后饱和选择反应区；其配置成接收所述第三物流、以及排放第四物流；
第一管道；其配置成将所述第四物流循环至所述第二反应区。
在一个实施方案中，所述从催化柴油生产轻质芳烃的装置包括：
第一反应区；其配置成接收所述催化柴油、以及排放第一物流；
第二反应区；其配置成接收所述第一物流、以及排放第二物流；
第一分离区；其配置成接收所述第二物流；排放包括所述C 6～C 8芳烃物流、含C 9芳烃和C 10芳烃的物流、以及含C 10以上重芳烃的第三物流在内的馏分；
后饱和选择反应区；其配置成接收所述第三物流、以及排放第四物流；
第一管道；其配置成将所述第四物流循环至所述第二反应区。
具体地，在一个实施方案中：
对于本发明所述的装置的一个实施方案，第一反应区配置加氢精制装置，其中使用的加氢精制反应器为固定床反应***。具体可采用本领 域中已有的固定床反应***，更优选配置有循环氢***的固定床反应***。所述加氢精制反应器入口温度可在250-450℃。
对于本发明所述的装置的一个实施方案，第二反应区配置进行选择转化的加氢裂化反应装置，其中使用的加氢裂化反应器为固定床反应***。具体可采用本领域中已有的固定床反应***，更优选配置有循环氢***的固定床反应***。所述选择转化(加氢裂化)反应器入口温度可在280-450℃。
对于本发明所述的装置的一个实施方案，后饱和选择反应区配置加氢饱和装置，其中使用的后饱和选择反应器为固定床反应***；更优选为未配置循环氢***的液态加氢固定床反应***。具体可采用本领域中已有的固定床反应***。后饱和(加氢饱和)反应器的入口温度在100-300℃之间，反应氢分压在1.0-4.0MPa之间。
对于本发明所述的装置的一个实施方案，第一分离区包括任选顺序联接的气液分离器、精馏塔，所述精馏塔优选包括任选顺序联接的脱戊烷塔(第一精馏塔)、脱庚烷塔(第二精馏塔)、脱二甲苯塔(第三精馏塔)和脱重芳烃塔(第四精馏塔)，用以顺序分离得到包括富含苯-甲苯馏分的物流、二甲苯物流、所述含C 9芳烃和C 10芳烃的物流、以及所述含C 10以上重芳烃的第三物流在内的馏分。
进一步优选，所述第二物流经过气液分离器，分离出干气和液相物流，所述液相物流经脱戊烷塔分离出塔顶的C3-C5轻烃物流和脱戊烷塔塔底物流，该塔底物流送入脱庚烷塔。脱庚烷塔顶分离出富含苯-甲苯馏分的物流和脱庚烷塔塔底物流，所述富含苯-甲苯馏分的物流优选经过抽提装置后分离出纯的苯-甲苯混合芳烃，抽提分离出的非芳烃外送。脱庚烷塔底物流进入二甲苯塔直接分离出混合二甲苯产品和脱二甲苯塔塔底物流。该二甲苯塔塔底物流送入重芳烃塔，塔顶分离出外送的碳九-碳十和塔底分离出的第三物流。所述第三物流送入后饱和选择反应器。所述的抽提和精馏都可采用本领域中常用的抽提和精馏方法。所述的气液分离器、精馏塔及抽提装置也都可采用本领域中常规的设备。
在本发明所述的装置的一个实施方案中，第一反应区和第二反应区之间配置有第二分离区，用以分离所述第一物流中包括硫化物和/或氮化物在内的杂质。所述第二分离区配置成接收所述第一物流、以及排放气相、硫化氢和氨物流以及已分离出杂质的第一物流。所述第二分离区的 分离装置可采用本领域中常规的分离装置，例如气液分离器(带气相注水洗氨)、汽提装置(如液相汽提脱硫化氢装置等汽提塔)等。
本发明的方法通过将催化裂化装置的催化柴油物流经过第一反应区加氢精制，将催化柴油物流中的杂质硫和氮进行脱除，而且使其中的稠环芳烃和多环芳烃发生选择性加氢饱和反应，加氢至只保留一个芳环的产物，如四氢萘、茚和多烷基苯等，然后该物流任选分离了杂质后送往第二反应区，发生选择转化的加氢裂化反应，产生富含苯、甲苯、二甲苯、C 9芳烃、C 10芳烃等轻质芳烃的物流，然后产物物流经过脱除苯之前的轻组分后，依次经过精馏塔分离出苯-甲苯、二甲苯、C 9芳烃、C 10芳烃和塔底重质尾油(主要含重芳烃)；塔底重质尾油物料进入后饱和选择反应器，在低温低压条件下，发生高选择性加氢饱和，得到保留一个芳环的产物，送往第二反应区进行选择转化的加氢裂化反应，实现从催化柴油生产轻质芳烃的全转化工艺，提高了轻质芳烃收率，减小了芳烃损失，降低氢耗。较好地解决了现有技术问题，用于增产芳烃产品取得了较好的技术效果。
本发明所述的技术方案中，经过第一反应区的加氢精制，催化柴油物流中稠环芳烃的饱和率大于50％，硫含量降低到100ppm以下，氮含量降低到15ppm以下，终馏点降低10℃以上；催化柴油物流依次经过第一反应区的加氢精制装置和第二反应区选择转化装置后，转化为碳十及以下单环芳烃，其转化率大于50％。经过后饱和选择反应器后物流的加氢饱和选择性高，芳烃保留率大于98％。
与现有技术相比，本发明的技术方案采用加氢精制-选择转化两段法工艺，双段双催化剂(加氢精制催化剂及选择转化催化剂)串联方案，包括加氢精制、选择转化及重质尾油后饱和过程。主要解决现有技术中存在的全馏分催化柴油不能完全转化，转化过程轻质芳烃收率不高的技术问题。催化柴油制轻质芳烃的重质尾油送入后处理反应器中，在缓和压力、温度条件下发生选择饱和反应。加氢饱和的选择性大幅提高到98％以上甚至更高，解决过度加氢饱和的问题；也有助于降低过度加氢生成的非芳进入选择转化反应器中发生的裂解耗氢反应。此方法提升了催化柴油制轻质芳烃整体工艺的技术经济指标，实现了催化柴油的全馏分转化。对比加氢精制-选择转化两段法催化柴油制轻质芳烃技术工艺，本发明的苯、甲苯、二甲苯及C 9和C 10等单环轻质芳烃的收率可至少提 高2％以上，优选提高5％以上。
附图说明
图1为本发明从催化柴油生产轻质芳烃的全转化方法的示意性的工艺流程图。
附图标记说明：
1为原料油-催化柴油
2为第一反应区
3为第一反应区出口物流-第一物流
4为气液分离器
5为含硫化氢和氨的气相物流
6为气液分离后的液相物流
7为硫化氢汽提塔
8为含硫化氢的汽提物流
9为脱除硫化氢后的加氢精制催化柴油-已分离出杂质的第一物流
10为选择转化反应器
11为选择转化反应产物-第二物流
12为第一分离区，例如包括气液分离器、脱戊烷塔、脱庚烷塔、二甲苯塔和重芳烃塔等精馏塔和苯-甲苯馏分抽提装置
13为第一分离区分离出的干气和C3-C5轻烃物流
14为第一分离区分离出的苯-甲苯物流
15为第一分离区分离出二甲苯物流
16为第一分离区分离出的含C 9芳烃和C 10芳烃的物流
17为第一分离区分离出的重质尾油物流-第三物流
18为后饱和选择反应器
19为后饱和选择反应器出口物流-第四物流
20为第二分离区(虚线框内)
具体实施方式
下面结合具体附图及实施例对本发明进行具体的描述，有必要在此指出的是以下实施例只用于对本发明的进一步说明，不能理解为对本发明保护范围的限制，本领域技术人员根据本发明内容对本发明做出的一些非本质的改进和调整仍属本发明的保护范围。
本说明书提到的所有出版物、专利申请、专利和其它参考文献全都 通过引用的方式并入本文。除非另有定义，本说明书所用的所有技术和科学术语都具有本领域技术人员常规理解的含义。在不一致或者有冲突的情况下，以本说明书的定义为准。
当本说明书以词头“本领域技术人员公知”、“现有技术”或其类似用语来导出材料、物质、方法、步骤、装置或部件等时，该词头导出的对象涵盖本申请提出时本领域常规使用的那些，但也包括目前还不常用，却将变成本领域公认为适用于类似目的的那些。
在本申请文件中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。在下文中，各个技术方案之间原则上可以相互组合而得到新的技术方案，这也应被视为在本文中具体公开。
以下详细描述本发明的具体实施方式，但是，本发明并不限于所述实施方式中的具体细节，在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，这些简单变型均属于本发明的保护范围。
另外需要说明的是，在以下具体实施方式中所描述的各个具体技术特征，在不矛盾的情况下，可以通过任何合适的方式进行组合。为了避免不必要的重复，本发明对各种可能的组合方式不再另行说明。
此外，本发明的各种不同的实施方式之间也可以进行任意组合，只要其不违背本发明的思想，由此而形成的技术方案属于本说明书原始公开内容的一部分，同时也落入本发明的保护范围。
在没有明确指明的情况下，本说明书内所提到的压力为表压。
在没有明确指明的情况下，本说明书内所提到的空速为液时空速LHSV。
在没有明确指明的情况下，本说明书内所提到的所有百分数、份数、比率等都是以重量为基准的，除非以重量为基准时不符合本领域技术人员的常规认识。
图1是本发明催化柴油生产轻质芳烃的方法的一个例示性实施方式的工艺流程示意图，图中省略了许多常规设备，如泵、压缩机、换热器、抽提装置、氢气管线等，但这些设备对本领域普通技术人员是公知的。 如图1所示，本发明所述的方法的一个例示性实施方式的流程详细描述如下：
作为原料油的催化柴油1进入第一反应区2的加氢精制装置，得到含硫化氢和氨的加氢精制后的催化柴油，即第一反应区出口物流3(第一物流)；第一物流经过第二分离区20的气液分离器4和硫化氢汽提塔7分离出加氢精制过程中脱氮脱硫得到的硫化氢和氨(通过含硫化氢和氨的气相物流5和含硫化氢的汽提物流8)后，得到已分离出杂质的第一物流9。该物流进入第二反应区10的选择转化装置。富含苯、甲苯、二甲苯等轻质芳烃、C9A和C10A馏分、重质尾油的第二反应区出口物流11(第二物流)进入第一分离区12，经分离得到干气和C3-C5轻烃物流13、苯-甲苯物流14、二甲苯物流15、含C 9芳烃和C 10芳烃的物流16、以及含C 10以上重芳烃的重质尾油第三物流17。第三物流17进入后饱和选择反应区的后饱和选择反应器18，后饱和选择反应器出口物流19(第四物流)不经分离，循环至第二反应区10的选择转化装置。
具体地，第一分离区12包括顺序联接的气液分离器、脱戊烷塔、脱庚烷塔、二甲苯塔和重芳烃塔等精馏塔和苯-甲苯馏分抽提装置(附图中未示)。
本发明中涉及的催化剂的组成分析均采用本领域已知的分析方法。例如，对于所述选择转化催化剂可通过ICP(电感耦合等离子体)和XRF(X射线荧光)方法分析催化剂的组成。用XPS(X射线光电子能谱)方法确定VIB族金属氧化物和金属硫化物的组成比例。ICP测试使用Varian 700-ES系列XPS仪进行。XRF测试使用Rigaku ZSX 100e型XRF仪进行。XPS测试条件包括：Perkin Elmer PHI 5000C ESCA型X射线光电子能谱仪，使用Mg K激发光源，操作电压l0kV，电流40mA，真空度4.0×10-8Pa。
本发明中，采用美国LECO公司的全二维气相色谱/高通量飞行时间质谱仪(GC×GC-TOFMS)分析(多维色谱分析)催化柴油和加氢精制催化柴油的族组成，也用来分析重质尾油和选择饱和重质尾油的族组成。
本发明中，通过气相色谱法测定反应物流(如选择转化产物等)组成。色谱型号为Agilent 7890A，配备FID检测器，FFAP毛细管色谱柱进行分离，色谱柱采用程序升温，初始温度90℃，保持15分钟，然后以15℃/分钟速率升温至220℃，维持45分钟。
本发明所述的方法中，加氢精制和选择饱和(后饱和)过程中的芳 烃保留率计算公式如下：
Figure PCTCN2020106710-appb-000001
苯-甲苯、二甲苯、C9A芳烃和C10A芳烃等单环轻芳烃收率计算公式为：
Figure PCTCN2020106710-appb-000002
本发明实施例和对比例的催化剂原料均可市售而得。
对比例1
采用加氢精制-选择转化两段法方案加工催化柴油，即将作为原料油的催化柴油经过加氢精制、分离杂质后进行加氢裂化，之后将加氢裂化产物经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃及重质尾油等产物。对比例1的工艺流程不包括将＞210℃重质尾油送入后饱和选择反应区进行选择加氢饱和处理。
催化柴油原料和加氢精制产物分析数据见表1，催化柴油的芳烃含量为87.15wt％。表2列出了所用的加氢精制催化剂、选择转化(加氢裂化)催化剂及其反应条件。
表1
项目 催化柴油原料 加氢精制产物
密度(4℃) 0.953 0.932
硫(wtppm) 1070 87
氮(wtppm) 632 8.6
非芳香烃(wt) 10.85 20.62
单环芳烃(wt％) 37.40 53.71
稠环芳烃(wt％) 51.75 25.67
蒸馏试验(D-86) ℃ ℃
初馏点 193 188
5％ 212 210
10％ 235 232
30％ 246 237
50％ 288 275
70％ 315 313
90％ 345 337
终馏点 372 363
表2
Figure PCTCN2020106710-appb-000003
所用的加氢精制催化剂A1的制备：在100g拟薄水铝石中加入2g 田菁粉、9ml硝酸和60ml水，混捏成团、挤条，室温养生24h，100℃条件下12h烘干，空气气氛中550℃焙烧3h得到加氢精制催化剂载体。将7.90g六水合硝酸镍、8.71g钼酸铵、9.18g偏钨酸铵和10ml氨水溶于水中得到50ml澄清溶液。取所述加氢精制催化剂载体50g，以等体积浸渍方式加入50ml溶液浸泡3小时，110℃条件下12h烘干，空气气氛中500℃焙烧4h得到加氢精制催化剂A1。该催化剂A1的组成为3.0wt％NiO-10.5wt％MoO 3-12.7wt％WO 3/73.8wt％Al 2O 3，即含有镍、钼、钨三种金属。
将含0.5％二硫化碳的环己烷溶液注入装有加氢精制催化剂A1的固定床反应器，从室温按10℃/h程序升温到硫化终点温度360℃，并在此温度保持12h，完成加氢精制催化剂的预硫化，得到硫化的加氢精制催化剂A1’，其组成为：3.1wt％NiS-10.2wt％MoS 2-13.2wt％WS 2/73.5wt％Al 2O 3，其中VIB族和VIII族金属以硫化态形式存在。
将催化柴油与氢气混合后，进入加氢精制反应器脱除其中大部分的硫、氮杂质，其中的稠环芳烃发生饱和以转化为只含一个芳环的烃类。表1还列出了加氢精制产品的硫氮含量、密度、芳香族烃类含量、以及馏分分布。
催化柴油经加氢精制之后的第一物流进行杂质分离，包括将第一物流经过气液分离，在常压下用氮气汽提3h，充分去除溶解在第一物流中的硫化氢。加氢精制产物(液相的已分离出杂质的第一物流)的硫含量和氮含量分别为87ppm和8.6ppm。根据芳烃族组成数据计算可得，加氢精制过程的稠环芳烃保留率为89.04wt％。
表2也列出了加氢裂化所用的选择转化催化剂B1组成和所采用的反应条件。USY沸石与氧化铝经混捏、挤条成型后得选择转化催化剂载体。再将适量氯铂酸配制成澄清溶液，通过等体积浸渍后再烘干、500℃空气中焙烧2h，得到选择转化催化剂前体。该选择转化催化剂前体在氢气条件下还原到450℃，得到所需的选择转化催化剂B1，其组成为：0.1份Pt-60份USY沸石-39.9份Al 2O 3。催化剂床层降温到340℃，将汽提后的加氢精制产物(已分离出杂质的第一物流)与氢气混合，进入选择转化反应器，反应生成物送往气液分离和精馏***。
经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃，计算可得，苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等 单环轻芳烃收率21.48wt％。所得＞210℃重质尾油收率为38.27wt％，比重为0.935，其硫氮含量分别为19.5ppm和1.5ppm；多维色谱分析可得第三物流的族组成为：非芳烃占41.98wt％，单环芳烃占26.38wt％，稠环芳烃占31.64wt％。
实施例1
本实施例进行催化柴油全转化制备轻芳烃的流程如图1所示。包括将催化柴油经过加氢精制、分离杂质、选择转化(加氢裂化)之后，将选择转化之后得到的＞210℃重质尾油送入后饱和选择反应区进行选择加氢饱和处理，具体为：
其中原料、加氢精制催化剂、及加氢精制反应条件同对比例1，选择转化催化剂B2(加氢裂化催化剂)及反应条件见表3。
表3
Figure PCTCN2020106710-appb-000004
表3列出了选择转化催化剂B2的组成和所采用的反应条件。
所述选择转化催化剂B2如下制备：70wt％的β沸石(硅铝分子比SAR＝25)与30wt％的氧化铝经混捏、挤条成型后得选择转化催化剂载体。再将适量硝酸镍和钨酸铵配制成澄清溶液，通过等体积浸渍后再100℃烘干、500℃空气中焙烧2小时，得到选择转化催化剂前体。该选择转化催化剂前体在氢气条件下还原到450℃还原时间4小时，降温到330℃后再注入二甲基二硫硫化4小时，可得到所需的选择转化催化剂B2。以催化剂总重为100重量份数计，催化剂B2的组成为：3.5份Ni-5.0份WO 3-0.27份WS 2-50份β沸石-41.23份Al 2O 3。
催化柴油经加氢精制之后的第一物流进行杂质分离，包括将第一物流经过气液分离，在常压下用氮气汽提3h，充分去除溶解在第一物流中的硫化氢。将汽提后的加氢精制产物(已分离出杂质的第一物流)与氢气混合，进入选择转化反应器，反应生成物送往气液分离和精馏***。
经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃 和C 10A芳烃，计算可得，苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等单环轻芳烃收率32.27wt％。所得＞210℃重质尾油(第三物流)收率为24.75wt％，比重为0.957，其硫氮含量分别为25.4ppm和1.6ppm；多维色谱分析可得第三物流的族组成为如表4所示：非芳烃占8.54wt％，单环芳烃占37.56wt％，稠环芳烃占53.90wt％。
表4
  ＞210℃重馏分
密度(4℃) 0.957
硫(wtppm) 25.4
氮(wtppm) 1.6
非芳香烃(wt％) 8.54
单环芳烃(wt％) 37.56
稠环芳烃(wt％) 53.90
处理＞210℃重质尾油的后饱和选择催化剂C2组成为：0.05wt％Pt-0.15wt％Pd-4.5wt％SiO 2-95.3wt％Al 2O 3。该后饱和选择催化剂C2如下制备：SiO 2含量为20wt％的商品无定型硅铝材料与拟薄水铝石混合，加入硝酸胶溶剂、田菁粉助挤剂和适量的水，捏合后挤条成型，在100℃空气中烘干24小时，再在550℃空气条件下焙烧4小时，得到催化剂载体。将适量氯铂酸和氯化钯溶于水中，得到金属浸渍液，通过等体积法浸渍催化剂载体，在80℃空气中烘干48小时，再在480℃空气条件下焙烧2小时，得到后饱和选择催化剂C2。将后饱和选择催化剂C2在氢气条件下还原，还原终点温度为450℃并保持两小时。
通过氢气混合器将过饱和量的氢气溶解在＞210℃重质尾油中，送入选择饱和反应器，反应条件为：氢油体积比450Nm 3/m 3，反应器入口温度180℃，氢气分压力1.5MPa，进料体积空速1.0小时 -1。待整个反应***建立平衡后，选择饱和产物的分析结果见表5，硫氮含量分别为16.8ppm和1.2ppm。根据芳烃族组成数据计算可得，选择饱和过程的稠环芳烃保留率为99.43wt％。
表5
  选择饱和＞210℃重馏分
密度(4℃) 0.939
硫(wtppm) 16.8
氮(wtppm) 1.2
非芳香烃(wt％) 8.96
单环芳烃(wt％) 59.82
稠环芳烃(wt％) 31.22
经过选择饱和的＞210℃重质尾油返回选择转化反应器，建立稳定的物流平衡后，经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃，计算可得，苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等单环轻芳烃收率46.35wt％。
实施例2
本实施例进行催化柴油全转化制备轻芳烃的流程如图1所示。包括将催化柴油经过加氢精制、分离杂质、选择转化(加氢裂化)之后，将选择转化之后得到的＞210℃重质尾油送入后饱和选择反应区进行选择加氢饱和处理，具体为：
其中原料、加氢精制催化剂、及加氢精制反应条件同对比例1，选择转化催化剂B3(加氢裂化催化剂)及反应条件见表6。
表6
Figure PCTCN2020106710-appb-000005
表6列出了选择转化催化剂B3的组成和所采用的反应条件。
所述选择转化催化剂B3如下制备：氢型丝光沸石(SAR＝45)、氢型β沸石(SAR＝25)、氢型ZSM-5(SAR＝27)与拟薄水铝石充分混合后，混捏、挤条、120℃烘干后在550℃空气气氛中焙烧4小时，即得所需的选择转化催化剂载体。以氯化钯、硝酸镍和钼酸铵配制三金属溶液， 通过等体积浸渍方法浸渍选择转化催化剂载体，120℃烘干后在500℃空气气氛中焙烧2小时，得到选择转化催化剂前体。该选择转化催化剂前体在氢气条件下还原到450℃并保持8小时，降温到330℃后注入二甲基二硫硫化4小时，可得到所需的选择转化催化剂B3。以催化剂总重量为100重量份计，催化剂B3的组成为：0.2份Pd-6.5份Ni-4.2份MoO 2-7.9份MoO 3-1.1份MoS 2-35份丝光沸石-10份β沸石-11份ZSM-5-24.1份Al 2O 3。
催化柴油经加氢精制之后的第一物流进行杂质分离：包括将第一物流经过气液分离，在常压下用氮气汽提3h，充分去除溶解在第一物流中的硫化氢。将汽提后的加氢精制产物(已分离出杂质的第一物流)与氢气混合，进入选择转化反应器，反应生成物送往气液分离和精馏***。
经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃，计算可得，苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等单环轻芳烃收率30.08wt％。所得＞210℃重质尾油(第三物流)收率为33.15wt％，比重为0.961，其硫氮含量分别为16.4ppm和0.8ppm；多维色谱分析可得第三物流的族组成为如表7所示：非芳烃占7.58wt％，单环芳烃占38.12wt％，稠环芳烃占54.30wt％。
表7
  ＞210℃重馏分
密度(4℃) 0.951
硫(wtppm) 16.4
氮(wtppm) 0.8
非芳香烃(wt％) 7.58
单环芳烃(wt％) 38.12
稠环芳烃(wt％) 54.30
处理＞210℃重质尾油的后饱和选择催化剂C3组成为0.10wt％Pt-0.30％Pd-4.0wt％Ni-6.0wt％SiO 2-89.6wt％Al 2O 3。该后饱和选择催化剂C3如下制备：SiO 2含量为9％的商品无定型硅铝材料与拟薄水铝石混合，加入硝酸胶溶剂、田菁粉助挤剂和适量的水，捏合后挤条成型，在100℃空气中烘干24小时，再在550℃空气条件下焙烧4小时，得到催化剂载体。将适量氯铂酸、氯化钯和醋酸镍溶于水中，得到金属浸渍液，通过等体积法浸渍催化剂载体，在100℃空气中烘干18小时， 再在500℃空气条件下焙烧2小时，得到后饱和选择催化剂C3。将后饱和选择催化剂C3在氢气条件下还原，还原终点温度为450℃并保持两小时。
通过氢气混合器将过饱和量的氢气溶解在＞210℃重质尾油中，送入选择饱和反应器，反应条件为：氢油体积比600Nm 3/m 3，反应器入口温度150℃，氢气分压力2.0MPa，进料体积空速1.5小时 -1。待整个反应***建立平衡后，选择饱和产物的分析结果见表8，硫氮含量分别为11.3ppm和0.6ppm。根据芳烃族组成数据计算可得，选择饱和过程的稠环芳烃保留率为99.63wt％。
表8
Figure PCTCN2020106710-appb-000006
经过选择饱和的＞210℃重质尾油返回选择转化反应器，建立稳定的物流平衡后，经过气液分离和精馏***后，分离得到苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃，计算可得，苯-甲苯、二甲苯、C 9A芳烃和C 10A芳烃等单环轻芳烃收率47.98wt％。

Claims (17)

  1. 一种从催化柴油生产轻质芳烃的方法，包括以下步骤：
    1)使催化柴油进入第一反应区进行加氢精制，得到第一物流；
    2)使所述第一物流进入第二反应区进行选择转化，得到第二物流，其中任选地，所述第一物流进入第二反应区前在第二分离区中进行杂质分离；
    3)使所述第二物流在第一分离区中进行第一分离，在该第一分离区底部得到含C10以上重芳烃的第三物流；
    4)使所述第三物流进入后饱和选择反应区进行加氢饱和，得到第四物流；
    5)使所述第四物流循环至所述第二反应区。
  2. 根据权利要求1所述的方法，其特征在于：
    除所述第三物流外，所述步骤3)还得到包括C6～C8芳烃物流、以及含C9芳烃和C10芳烃物流在内的馏分，其中所述C 6～C 8芳烃物流至少包括苯、甲苯、二甲苯中的一种。
  3. 根据权利要求1或2所述的方法，其特征在于：
    在所述步骤2)中，进行所述杂质分离，包括使所述第一物流经历气液分离和硫化氢汽提。
  4. 根据前述权利要求中任一项所述的方法，其特征在于：
    在所述步骤3)中，所述第二物流的第一分离包括气液分离、精馏；
    所述精馏优选包括脱戊烷、脱庚烷、脱二甲苯、脱重芳烃；其中优选对脱庚烷得到富含苯-甲苯馏分的物流进行抽提分离。
  5. 根据前述权利要求中任一项所述的方法，其特征在于：
    所述第一反应区的反应条件包括：
    氢油体积比500～3000Nm 3/m 3，优选800～2000Nm 3/m 3，更优选1000～1500Nm 3/m 3；和/或：
    反应器入口温度280-420℃，优选300～410℃，更优选310～390℃；和/或：
    氢气分压力为5～10MPa，优选5～8MPa，更优选6～7MPa；和/或：
    空速0.5～2.0小时 -1，优选0.6～1.5小时 -1，更优选0.8～1.2小时 -1。
  6. 根据前述权利要求中任一项所述的方法，其特征在于，
    在所述步骤2)中，所述选择转化在选择转化催化剂的存在下进行，所述选择转化催化剂以重量份数计包括：a 2)5～80份固体酸沸石；b 2)0.05～8份VIII族金属；c 2)3～25份VIB族金属氧化物；d 2)0.1～2份VIB族金属硫化物；e 2)20～95份第一粘结剂。
  7. 根据权利要求6所述的方法，其特征在于，
    所述固体酸沸石为丝光沸石、β沸石、ZSM沸石、EU-1沸石、SAPO沸石和Y沸石中的至少一种；
    所述VIII族金属为铂、钯、钴、镍和铱中的至少一种；
    所述VIB族金属氧化物为氧化钼和氧化钨中的至少一种；
    所述VIB族金属硫化物为硫化钼和硫化钨中的至少一种；和
    所述第一粘结剂为氧化铝、氧化硅-氧化铝复合物、氧化钛-氧化铝复合物和氧化镁-氧化铝复合物中的至少一种。
  8. 根据前述权利要求中任一项所述的方法，其特征在于：
    所述第二反应区的反应条件包括：
    氢油体积比800～5000Nm 3/m 3，优选1000～4000Nm 3/m 3，更优选1500～3000Nm 3/m 3；和/或：
    反应器入口温度280-450℃，优选300～430℃，更优选310～400℃；和/或：
    氢气分压力5～10MPa，优选5～9MPa，更优选6～8MPa；和/或：
    空速0.5～2.0小时 -1，优选0.6～1.5小时 -1，更优选0.8～1.2小时 -1。
  9. 根据前述权利要求中任一项所述的方法，其特征在于：
    在所述步骤4)中，所述加氢饱和在后饱和选择催化剂的存在下进行，所述后饱和选择催化剂包括以重量份数计的：a 3)10～90份无定型硅铝，所述无定型硅铝的氧化硅含量为3-20wt％；b 3)0.1～5.0份VIII族金属；c 3)5～80份第二粘结剂；
    所述VIII族金属优选选自由铂、钯、钴、镍和铱中的至少一种；和
    所述第二粘结剂选自氧化铝。
  10. 根据前述权利要求中任一项所述的方法，其特征在于：
    所述后饱和选择反应区的反应条件包括：
    氢油体积比200～3000Nm 3/m 3，优选300～1500Nm 3/m 3，更优选300～1000Nm 3/m 3；和/或：
    反应器入口温度100～300℃，优选120～280℃，更优选150～250℃； 和/或：
    氢气分压力1.0～4.0MPa，优选1.2～3.0MPa；和/或：
    空速0.1～5.0小时 -1，优选0.5～4.0小时 -1，更优选0.6～2.0小时 -1。
  11. 用于实施根据权利要求1～10之任一项所述的方法，以从催化柴油生产轻质芳烃的装置，包括：
    进行加氢精制的第一反应区；其配置成接收所述催化柴油、以及排放第一物流；
    进行选择转化的第二反应区；其配置成接收所述第一物流、以及排放第二物流；
    第一分离区；其配置成接收所述第二物流；在底部排放所述第三物流；
    进行加氢饱和的后饱和选择反应区；其配置成接收所述第三物流、以及排放第四物流；
    第一管道；其配置成将所述第四物流循环至所述第二反应区。
  12. 根据权利要求11所述的装置，其特征在于：
    所述第一反应区的反应器为固定床反应***；和/或
    所述第二反应区的反应器为固定床反应***；和/或
    所述后饱和选择反应区的反应器为固定床反应***。
  13. 根据权利要求12所述的装置，其特征在于：
    所述第一反应区的固定床反应***配置有循环氢***；和/或
    所述第二反应区的固定床反应***配置有循环氢***；和/或
    所述后饱和选择反应区的固定床反应***为不配置循环氢***的液相加氢反应***。
  14. 根据权利要求11-13中任一项所述的装置，其特征在于：
    所述第一分离区包括，任选顺序联接的，气液分离器、精馏塔，用以顺序分离得到包括苯-甲苯物流、二甲苯物流、含C 9芳烃和C 10芳烃的物流、以及所述含C 10以上重芳烃的第三物流在内的馏分；
    所述精馏塔优选包括，任选顺序联接的，脱戊烷塔、脱庚烷塔、二甲苯塔和重芳烃塔。
  15. 根据权利要求14所述的装置，其特征在于：
    所述第一分离区包括脱庚烷塔，并在其下游包括苯-甲苯馏分抽提装置，将脱庚烷塔分离出的富含苯-甲苯馏分的物流进行分离。
  16. 根据权利要求11-15中任一项所述的装置，其特征在于：
    所述第一反应区和所述第二反应区之间配置有第二分离区，用以分离所述第一物流中包括的硫化氢和氨在内的杂质；
    所述第二分离区配置成接收所述第一物流、以及排放气相、硫化氢和氨物流以及已分离出杂质的第一物流。
  17. 根据权利要求16所述的装置，其特征在于：
    所述第二分离区包括气液分离器和汽提装置。
PCT/CN2020/106710 2019-08-05 2020-08-04 一种从催化柴油生产轻质芳烃的全转化方法和装置 WO2021023172A1 (zh)

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