WO2022007736A1 - 一种馏分油超/亚临界流体强化加氢方法 - Google Patents

一种馏分油超/亚临界流体强化加氢方法 Download PDF

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WO2022007736A1
WO2022007736A1 PCT/CN2021/104430 CN2021104430W WO2022007736A1 WO 2022007736 A1 WO2022007736 A1 WO 2022007736A1 CN 2021104430 W CN2021104430 W CN 2021104430W WO 2022007736 A1 WO2022007736 A1 WO 2022007736A1
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oil
hydrogenation
distillate
hydrogen
stage
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PCT/CN2021/104430
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English (en)
French (fr)
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陈振涛
赵锁奇
徐春明
许志明
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中国石油大学(北京)
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Publication of WO2022007736A1 publication Critical patent/WO2022007736A1/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
    • 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
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/002—Apparatus for fixed bed hydrotreatment processes
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G3/00—Production of liquid hydrocarbon mixtures from oxygen-containing organic materials, e.g. fatty oils, fatty acids
    • C10G3/42—Catalytic treatment
    • C10G3/44—Catalytic treatment characterised by the catalyst used
    • C10G3/45—Catalytic treatment characterised by the catalyst used containing iron group metals or compounds thereof
    • C10G3/46—Catalytic treatment characterised by the catalyst used containing iron group metals or compounds thereof in combination with chromium, molybdenum, tungsten 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
    • C10G3/00—Production of liquid hydrocarbon mixtures from oxygen-containing organic materials, e.g. fatty oils, fatty acids
    • C10G3/50—Production of liquid hydrocarbon mixtures from oxygen-containing organic materials, e.g. fatty oils, fatty acids in the presence of hydrogen, hydrogen donors or hydrogen generating compounds
    • 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
    • C10G45/04—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 characterised by the catalyst used
    • C10G45/06—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 characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof
    • C10G45/08—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 characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof in combination with chromium, molybdenum, or tungsten 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
    • 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
    • C10G45/48—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof
    • C10G45/50—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof in combination with chromium, molybdenum or tungsten metal, 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
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/02—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00 characterised by the catalyst used
    • C10G49/04—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00 characterised by the catalyst used containing nickel, cobalt, chromium, molybdenum, or tungsten 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
    • 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
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/20—Technologies relating to oil refining and petrochemical industry using bio-feedstock

Definitions

  • the invention belongs to the petrochemical field, relates to a distillate oil hydrogenation technology, and in particular relates to a distillate oil super/subcritical fluid enhanced hydrogenation method.
  • the catalytic hydrogenation of distillate raw materials generally adopts a fixed-bed hydrogenation process.
  • hydrogen and distillate raw materials are mixed and heated to a suitable temperature, then passed into a fixed-bed reactor, contacted with a solid catalyst, and a catalytic hydrogenation reaction occurs. , and then the hydrogenation reaction effluent is separated to obtain hydrogen-containing gas and hydrogenated oil.
  • the hydrogen since the hydrogen is in the gas phase and the distillate raw material is mainly in the liquid phase, the hydrogen needs to dissolve into the liquid phase first in the reaction system, and then reach the active site of the solid catalyst through diffusion and mass transfer to carry out the hydrogenation reaction .
  • the mass transfer rate at which hydrogen dissolves in the oil phase and reaches the active center of the catalyst has a significant effect on the distillate hydrogenation process.
  • the traditional trickle-bed process usually uses excess high-pressure hydrogen circulation to supply hydrogen for the reaction, resulting in high investment and operating costs of the device.
  • Feedstock oil with high hydrogen consumption Chinese patent application CN201310523200.0 discloses a continuous liquid-phase wax oil hydrotreating method.
  • the wax oil raw material and the reaction effluent are mixed with hydrogen in a mixer to dissolve the hydrogen in the hydrocarbon oil,
  • the volume ratio of hydrogen to liquid is only 20-30v/v, and hydrogen needs to be supplemented in multiple catalyst beds between the reactors
  • Chinese patent application CN201510036537.8 discloses a continuous liquid-phase diesel hydrogenation treatment method. There are multiple beds in the device, and mixers are arranged between the beds.
  • the core of the above-mentioned liquid-phase hydrogenation technology of distillate oil such as diesel oil and wax oil is to supply hydrogen to the reaction system by circulating the hydrogenated diesel oil, wax oil and other distillate oil and strengthening the mixing effect of hydrogen and distillate oil to achieve the elimination of hydrogenation.
  • the purpose of hydrogen circulation system is to reduce equipment investment costs and operating costs.
  • the present invention provides a super/subcritical fluid enhanced hydrogenation method for distillate oil, which can improve the solubility of hydrogen in the reaction system by introducing super/subcritical fluid to achieve efficient hydroconversion of distillate oil raw materials.
  • the distillate super/subcritical fluid enhanced hydrogenation method provided by the invention at least includes:
  • the distillate raw material, solvent and hydrogen are preheated and passed into the first-stage fixed-bed reactor, and the first-stage catalytic hydrogenation is carried out in a supercritical or subcritical fluid state to obtain a catalytic hydrogenation reaction effluent;
  • the above-mentioned catalytic hydrogenation reaction effluent is separated, the hydrogen-containing gas therein is separated, and the hydrogenated oil product and the solvent-containing light fraction are obtained; the solvent-containing light fraction is returned to the first-stage fixed-bed reactor to replace the solvent for recycling;
  • the aforementioned solvent is selected from at least one of the following materials: paraffins with 4 to 10 carbon atoms; naphthenic hydrocarbons with 5 to 10 carbon atoms; aromatic hydrocarbons with 6 to 8 carbon atoms; gasoline, gasoline fractions , naphtha, naphtha fractions.
  • the super/subcritical fluid enhanced hydrogenation method provided by the present invention is not particularly limited for the distillate oil raw material, and the applicable distillate oil raw material can be selected from at least one of the following materials: kerosene, diesel oil, wax oil, lubricating oil, catalytic Cracking back to refinery, catalytically cracked slurry and bio-oil.
  • the hydrogenation method can also be applied to the catalytic hydrogenation of biomass oil and has a good catalytic hydrogenation effect.
  • the present invention does not specifically limit the source of the above-mentioned distillate oil raw materials, for example, kerosene, diesel oil, and wax oil from petroleum, shale oil, biomass oil, etc. can be used as the distillate oil raw material.
  • the present invention also does not specifically limit the specific acquisition process of kerosene, diesel oil, wax oil, and lubricating oil as raw materials of distillate oil.
  • the diesel fuel may be straight-run diesel fuel, catalytically cracked diesel fuel, coking diesel fuel, or the like.
  • the wax oil may specifically be straight-run wax oil, coking wax oil, wax oil mixed with an appropriate amount of deasphalted oil, and the like.
  • Biomass oil can specifically be animal or vegetable oil obtained by chemical, pyrolysis and biological methods.
  • distillate oil can also be understood as an oil product that is not too heavy compared to the recognized heavy oil in the petrochemical industry, that is, the distillation range of the raw material oil is concentrated in It belongs to the above-mentioned range of oil products; or although the raw oil contains or mixes "heavy" fractions, it is usually not suitable for heavy oil processing equipment; or, these raw materials are cut into fractions to separate the heavy fractions.
  • the pressure wax oil fraction is the lighter fraction obtained by vacuum distillation and cutting of atmospheric residual oil to separate the heavier fraction.
  • the super/subcritical fluid enhanced hydrogenation method of distillate oil is very suitable for distillate oil such as catalytic cracked diesel, coker diesel, straight-run diesel, straight-run wax oil, coker wax oil, catalytic cracking oil and biomass oil. and its mixtures for hydroconversion, and finally obtain good-quality distillate hydrogenation products or provide high-quality raw materials for subsequent processing.
  • the solvent used in the present invention can be, in particular, linear alkanes having 5 to 8 carbon atoms, branched alkanes having 5 to 8 carbon atoms, cycloalkanes having 5 to 8 carbon atoms, and 6 carbon atoms.
  • the solvent used is at least one of n-butane, n-pentane, n-hexane, n-heptane, n-octane and cyclohexane, or light naphtha.
  • hydrogen can also be appropriately supplemented according to the properties of the distillate feedstock and the process conditions of the catalytic hydrogenation.
  • the amount of hydrogen supplementation in the first-stage catalytic hydrogenation process can be specifically determined by hydrogen consumption, phase equilibrium results, and reaction conditions.
  • the conditions of the first-stage catalytic hydrogenation reaction carried out in a supercritical or subcritical state may specifically be: the average reaction temperature is 240-460° C., the reaction pressure is 3-30 MPa; The ratio is 0.1-6:1; the volume ratio of hydrogen to oil is 50-1200:1, and the liquid hourly volume space velocity is 0.2-12.0h -1 .
  • the recycled solvent-containing light fractions can be
  • the catalytic hydrogenation reaction of the first stage is basically stable, so a small amount of fresh solvent can be added according to the actual situation or even no need to add fresh solvent.
  • the solvent-containing light ends passed into the first-stage fixed-bed reactor are understood to be equivalent to the fresh solvent.
  • the mass ratio of agent oil refers to the mass ratio between the fresh solvent introduced into the first stage fixed bed reactor and the raw material of distillate oil, while in continuous production, the mass ratio of agent oil is Refers to the mass ratio of solvent-containing light ends and possibly supplemental fresh solvent to distillate feedstock.
  • the liquid-phase materials introduced into the first-stage fixed-bed reactor refer to the solvent and distillate raw materials; when the production is stable, the liquid-phase materials refer to the solvent-containing light fractions, which may be supplemented.
  • Fresh solvent and distillate feedstocks The liquid hourly volume space velocity refers to the volume space velocity of the liquid material (unit is hour -1 ).
  • the volume ratio of hydrogen to oil refers to the volume ratio between hydrogen and distillate raw materials in a standard state.
  • the inventor's research has found that appropriately increasing the reaction temperature, reaction pressure and reducing the liquid hourly space velocity, and selecting appropriate agent-oil ratio and hydrogen-oil volume ratio are beneficial to improve the catalytic hydrogenation effect of distillate raw materials, especially can improve the desulfurization rate, desulfurization rate, desulfurization rate Nitrogen rate and deoxidation rate.
  • the optimal process conditions can also be determined according to factors such as the specific composition of the distillate feedstock, energy consumption, production cost, production efficiency, and production goals.
  • the process conditions of the first-stage catalytic hydrogenation reaction are usually controlled at: average reaction temperature of 260-430° C., reaction pressure of 4-25 MPa; agent-oil mass ratio of 0.1-4.0: 1; hydrogen oil
  • the volume ratio is 80 ⁇ 1000:1; the liquid hourly volume space velocity is 0.3 ⁇ 6.0h -1 .
  • the process conditions of the first-stage catalytic hydrogenation reaction can be reasonably adjusted according to the specific composition of the distillate feedstock.
  • the raw material of distillate oil is kerosene, diesel oil, light distillate of bio-oil, or a mixture of any of them
  • the average temperature of the first-stage catalytic hydrogenation reaction carried out in a supercritical or subcritical state can be specifically 240 ⁇ 440°C, preferably 260-420°C, more preferably 280-420°C
  • the reaction pressure can be specifically 3-25MPa, preferably 4-20MPa
  • the agent-oil mass ratio can be specifically 0.1-5.0:1, preferably 0.1- 3.0:1
  • the volume ratio of hydrogen to oil can be 50-1000:1 , preferably 80-800:1, further 80-600:1
  • the liquid hourly volume space velocity can be specifically 0.5-12.0h -1 , preferably 0.5 -6.0h -1 , further 0.5 - 5.0h -1 .
  • the first stage is carried out in a supercritical or subcritical state.
  • the average temperature of the catalytic hydrogenation reaction can be specifically 300-460 °C, preferably 320-430 °C, further 340-420 °C; the reaction pressure can specifically be 4-30 MPa, preferably 5-25 MPa, further 6-20 MPa;
  • the mass ratio of the agent to oil can be specifically 0.1-6.0:1, preferably 0.1-5.0:1, and further 0.2-4.0:1; the hydrogen-oil volume ratio can specifically be 50-1200:1, preferably 80-1000:1, Further, it is 100-800:1; the liquid hourly volume space velocity can be specifically 0.2-8.0 h -1 , preferably 0.3-5.0 h -1 , and further 0.5-4.0 h -1 .
  • the distillate raw material to be treated is subjected to a super/subcritical fluid enhanced hydrogenation process, also known as the first-stage catalytic hydrogenation, which can effectively improve the effect of catalytic hydrogenation.
  • the present invention does not specifically limit the specific reaction type of the first-stage catalytic hydrogenation, such as hydrodesulfurization, hydrodenitrogenation, hydrodeoxygenation, olefin hydrosaturation, aromatics hydrosaturation, hydrodemetallization, Hydrocracking and other catalytic hydrogenation reactions.
  • the catalytic hydrogenation process provided by the present invention can be applied to hydrorefining for obtaining finished products, hydroprocessing for providing high-quality raw materials for subsequent processing, and for obtaining light oils and chemicals. Processes such as hydroupgrading and hydrocracking.
  • the corresponding hydrogenation catalyst should be selected, such as a hydrofinishing catalyst, a hydroprocessing catalyst, a hydroupgrading catalyst, and a hydrocracking catalyst. one or more.
  • the above-mentioned hydrogenation catalysts are commercially available, and can also be prepared according to conventional methods in the art.
  • the catalyst needs to be sulfided before use to convert oxides of active metal components into sulfides with higher activity, and pre-sulfide treatment is a conventional method in the field.
  • the selected hydrogenation catalyst is an industrial-grade hydrogenation catalyst, which may specifically be an industrial-grade hydrofinishing catalyst, an industrial-grade hydrotreating catalyst, an One or more of the hydrocracking catalysts.
  • the hydrogenation catalyst is obtained by supporting the active metal component on the carrier, wherein the active metal can be selected from at least one of the metals of Group VIB and Group VIII; the carrier can be determined according to the processing requirements and the characteristics of the corresponding hydrogenation reaction.
  • the carriers of the hydrogen refining catalyst and the hydrotreating catalyst are selected from one or more of alumina, alumina-silica, alumina-titanium oxide, etc., and the carriers of the hydro-upgrading catalyst and the hydrocracking catalyst are those containing molecular sieves.
  • alumina alumina-silica, alumina-titanium oxide, etc.
  • the carriers of the hydro-upgrading catalyst and the hydrocracking catalyst are those containing molecular sieves.
  • fluorine, phosphorus, titanium, zirconium and boron can be optionally added to the above-mentioned carrier as auxiliary components.
  • auxiliary components Those skilled in the art can select appropriate active metal, carrier and auxiliary components and their contents as required.
  • the above-mentioned industrial hydrogenation catalysts for petroleum-based oil products can be selected for the catalytic hydrogenation of biological oils and fats, and other hydrogenation catalysts can also be selected.
  • the super/subcritical fluid enhanced hydrogenation process of distillate oil needs to set up a fixed bed reactor and at least one catalyst bed.
  • the mixture of hydrogen, solvent, and fresh distillate raw materials can be heated to a suitable temperature and then enter the first-stage fixed-bed reactor.
  • the first-stage catalytic hydrogenation reaction occurs in the fluid state.
  • the materials entering the first-stage fixed-bed reactor that is, hydrogen, solvent, and fresh distillate raw materials, can be simultaneously fed into the feed pipeline to enter the fixed-bed reactor together, or they can be fully mixed in the mixer before entering the reactor.
  • fixed-bed reactor or fully mixed in a mixer arranged between the catalyst beds in the fixed-bed reactor, which is not particularly limited in this embodiment.
  • the present invention does not specifically limit how to separate the effluent of the catalytic hydrogenation reaction, and conventional separation means in the art can be used.
  • the effluent of the catalytic hydrogenation reaction is subjected to water injection and heat exchange, and the three-phase separation of oil, gas and water is carried out, and the gas phase and oil phase obtained by separation are further separated to obtain hydrogen-containing gas and solvent-containing light fractions.
  • hydrogenated oil products such as gasoline fraction, kerosene fraction, diesel fraction and tail oil after hydrogenation.
  • the solvent-containing light fractions are returned to the first-stage fixed-bed reactor mixture material to replace the solvent for recycling; the hydrogen-containing gas can be directly discharged or discharged after harmless treatment, or the hydrogen-containing gas can be purified and used continuously.
  • one-stage fixed-bed catalytic hydrogenation can be carried out on the distillate feedstock in the state of supercritical or subcritical fluid, and the second-stage catalytic hydrogenation can also be carried out on this basis, that is, two-stage fixed-bed hydrogenation can be carried out.
  • Hydrogen process Specifically, a one-stage fixed bed hydrogenation process or a two-stage fixed bed hydrogenation process can be selected and implemented according to the composition of the distillate feedstock and the production target.
  • the first-stage catalytic hydrogenation reaction effluent can be directly passed into the second-stage fixed-bed reactor for the second-stage catalytic hydrogenation to obtain The second stage catalytic hydrogenation reaction effluent (ie, the catalytic hydrogenation reaction effluent).
  • the first-stage catalytic hydrogenation reaction effluent is not separated, but is directly introduced into the second-stage fixed-bed reactor to carry out the second-stage catalytic hydrogenation to obtain the second-stage catalytic hydrogenation reaction effluent (that is, the catalytic hydrogenation reaction effluent). hydrogen reaction effluent).
  • the effluent of the catalytic hydrogenation reaction is separated to obtain hydrogenated oil, hydrogen-containing gas and solvent-containing light fraction; the solvent-containing light fraction is returned to the first-stage fixed-bed reactor to replace the solvent for recycling.
  • the catalytic hydrogenation reaction effluent obtained by the first-stage catalytic hydrogenation can also be separated to obtain hydrogen-containing gas, solvent-containing light fractions and the first-stage hydrogenated oil product, and then the first-stage hydrogenated oil product is passed through. into the second-stage fixed bed reactor to carry out the second-stage catalytic hydrogenation to obtain the second-stage catalytic hydrogenation reaction effluent.
  • the catalytic hydrogenation reaction effluent obtained by the first-stage catalytic hydrogenation reaction is first subjected to separation treatment, the hydrogen-containing gas therein is separated, and the first-stage hydrogenated oil products and solvent-containing light ends are collected, and then the hydrogen-containing gas is separated.
  • the solvent-containing light fraction is returned to the first-stage fixed-bed reactor to replace the solvent for recycling, and the first-stage hydrogenated oil product is mixed with hydrogen and then passed into the second-stage fixed-bed reactor for the second-stage catalytic hydrogenation to obtain The second-stage catalytic hydrogenation reaction effluent; finally, the second-stage catalytic hydrogenation reaction effluent is separated, the hydrogen-containing gas therein is separated, and the second-stage hydrogenated oil product therein is collected.
  • the above-mentioned two-stage fixed-bed hydrogenation process can use two-stage fixed-bed reactors connected in series to carry out the super/subcritical fluid enhanced hydrogenation reaction, or it can be based on the super/subcritical fluid intensified hydrogenation process.
  • Distillate Hydrogenation Process It is not difficult to understand that by adopting the two-stage fixed bed hydrogenation process, the composition of the finally obtained second-stage hydrogenated oil is different from that of the first-stage hydrogenated oil. Therefore, the second-stage hydrogenated oil can be adjusted according to the actual situation. Implement a suitable separation process.
  • the present invention does not specifically limit the specific type of the fixed bed reactor, it can be a descending bed (that is, the material enters from the top of the fixed bed reactor and flows out from the bottom), or it can be an ascending bed (that is, the material enters from the bottom of the fixed bed reactor, top out).
  • the distillate oil super/subcritical fluid enhanced hydrogenation method provided by the present invention, by controlling the distillate oil raw material, solvent-containing light ends (or solvent) and hydrogen ternary system in the ultra-supercritical state.
  • Catalytic hydrogenation in a critical or subcritical fluid state has the following advantages:
  • the present invention can greatly improve the solubility and diffusion mass transfer performance of hydrogen in the distillate oil hydrogenation reaction system, thereby greatly improving the supply of effective hydrogen in the catalytic hydrogenation reaction process, and then significantly improving the distillate oil hydrogenation reaction. Efficiency and utilization of hydrogen and catalyst for efficient hydroconversion of distillate feedstocks.
  • the super/subcritical fluid can take away the heat released by the catalytic hydrogenation reaction, avoid catalyst deactivation caused by local overheating and avoid the occurrence of side reactions, thereby extending the life of the catalyst and the entire distillate super/subcritical fluid. Strengthen the operation cycle of the hydrogenation system.
  • the method can greatly improve the supply of effective hydrogen, so compared with the conventional distillate oil hydrogenation technology, the distillate oil super/subcritical fluid enhanced hydrogenation method provided by the present invention can also greatly reduce the hydrogen circulation amount and Equipment investment and production operating costs such as hydrogen compressors.
  • the present invention can be applied to hydrorefining, hydrotreating, hydro-upgrading and hydrocracking of distillate oils such as kerosene, diesel oil and wax oil, and can also be used as a front-end process for distillate hydrotreating. Process adjustments can be made flexibly based on distillate feedstock properties and production goals.
  • distillate super/subcritical fluid enhanced hydrogenation process of the present invention can be obtained by simple transformation on the basis of the conventional distillate oil hydrogenation process, so it has the advantages of low transformation and investment cost, and is convenient for promotion and application in actual production. application.
  • FIG. 1 is a schematic flow sheet of a distillate super/subcritical fluid enhanced hydrogenation embodiment provided by the invention.
  • FIG. 2 is a schematic flow diagram of another distillate super/subcritical fluid enhanced hydrogenation embodiment provided by the present invention.
  • FIG. 3 is a schematic flow chart of yet another embodiment of distillate super/subcritical fluid enhanced hydrogenation provided by the present invention.
  • the invention introduces a solvent into the catalytic hydrogenation reaction of the distillate raw material and constructs a super/subcritical fluid enhanced hydrogenation reaction system, aiming to promote the catalytic hydrogenation reaction of the distillate raw material and realize the efficient hydrogenation conversion of the distillate raw material.
  • a visual high temperature and high pressure phase balance instrument can be used as a critical point measuring device, and the phase behavior of the distillate raw material-hydrogen-solvent ternary system is investigated through the change of the phase interface and the critical opalescence phenomenon in the phase balance experiment.
  • the critical point parameters for reaching the supercritical fluid state are thus determined. According to the experimental results of the above critical point, the reaction condition range of the super/subcritical fluid enhanced hydrogenation of distillate oil is set.
  • Fig. 1 is the distillate super/subcritical fluid enhanced hydrogenation method provided by the present invention, wherein a schematic flow diagram of an embodiment, for the sake of clarity, many general equipment or devices are omitted in the figure, such as hydrogen compressor, heat exchange However, the functions, settings and specific selection of these general equipment or devices are well known to those of ordinary skill in the art.
  • distillate oil raw material 1, solvent 2 or solvent-containing light ends 2 and hydrogen 3 are mixed, and the obtained mixed material 4 is heated in the heat exchanger after heat exchange, and then enters the heating furnace and continues to be warmed up to the required process temperature, Then entering the fixed bed reactor 5, the mixed material 4 is contacted with the pre-sulfided catalyst, and the catalytic hydrogenation reaction takes place under the condition of supercritical or subcritical fluid.
  • the hydrogenation reaction effluent 6 enters the high-pressure separator 7 after water injection and heat exchange to carry out the three-phase separation of oil, gas and water, and the separated high fraction gas 8 can carry out subsequent recovery, and the separated water-phase product (not shown in the figure) ) out of the device, the separated high-separation oil 9 enters the low-pressure separator 10 for gas-liquid separation, the obtained low-separation gas 11 is discharged from the top outlet of the low-pressure separator 10, and the low-separation oil 12 is discharged from the bottom outlet of the low-pressure separator 10.
  • the fractionation system 13 separates the solvent-containing light fraction 2 and obtains a hydrogenated oil product (not shown) after hydrogenation.
  • the hydrogenated oil product may be one or more of gasoline fraction, kerosene fraction, diesel fraction and tail oil.
  • the solvent-containing light fraction 2 is returned and merged into the mixed material 4 to replace the solvent for recycling.
  • the low-separation oil 12 discharged from the bottom of the low-pressure separator 10 is passed into the stripper 14, and the flash oil 15 discharged from the bottom of the stripper 14 enters the fractionation system 13 for separation to obtain hydrogenation Oil product (not shown), the hydrogenated oil product can be specifically hydrogenated gasoline, hydrogenated kerosene, hydrogenated diesel oil and hydrogenated tail oil, etc.
  • the solvent-containing light fraction 2 discharged from the top of the stripper 14 is returned and incorporated into the mixture
  • the substitute solvent in feed 4 is recycled.
  • Fig. 3 The difference between Fig. 3 and the embodiment of Fig. 1 is: the high-divided gas 8 separated from the top of the high-pressure separator 7 is further processed, and the high-divided gas 8 is passed into the hydrogen sulfide removal tower 16 after heat exchange, and the lower part of the hydrogen sulfide removal tower 16 is processed.
  • the solvent-containing light fraction 2 from which gas such as hydrogen sulfide is removed is discharged, and the solvent-containing light fraction 2 is returned and incorporated into the mixed material 4 to replace the solvent for recycling.
  • No. 1 diesel oil is catalytically cracked diesel oil
  • No. 2 diesel oil is a mixed diesel oil of straight-run diesel oil and catalytically cracked diesel oil
  • No. 1 wax oil is a mixed wax oil of straight-run wax oil and coking wax oil
  • No. 1 bio-oil is palm oil. .
  • the properties of the hydrogenation catalysts used in the following examples and comparative examples are shown in Table 2.
  • the hydrofinishing catalysts and hydrocracking catalysts in Table 2 are both technical grades.
  • reaction temperature is the average temperature of the catalyst bed
  • liquid hourly space velocity is all based on the distillate raw material.
  • Embodiments 1-5 adopt the technological process shown in Figure 1 to carry out the super/subcritical fluid enhanced hydrogenation reaction of No. 1 diesel oil (see Table 1 for properties), wherein the solvent used is a mixture of n-pentane and cyclohexane of equal quality , the used hydrogenation catalyst is an industrial hydrotreating catalyst (see Table 2 for properties), and the industrial hydrotreating catalyst is packed in a fixed bed reactor to form a catalyst bed.
  • the hydrogen, the solvent and the fresh No. 1 diesel oil are mixed in proportion and heated to a suitable temperature, and then enter the fixed-bed reactor.
  • the mixed material is contacted with an industrial hydrofinishing catalyst and undergoes a catalytic hydrogenation reaction under supercritical or subcritical fluid conditions to obtain Catalytic hydrogenation reaction effluent; the catalytic hydrogenation reaction effluent is separated to obtain hydrogen-containing gas, solvent-containing light fractions and hydrogenated oil products.
  • the super/subcritical fluid enhanced hydrogenation process conditions adopted in Examples 1-5 are shown in Table 3, and the diesel hydrogenation evaluation results are shown in Table 5.
  • Comparative examples 1-4 adopt conventional fixed-bed hydrogenation reaction process: the fresh No. 1 diesel oil (see Table 1 for properties) is mixed with hydrogen and heated to a certain temperature, and then fed into a fixed-bed reactor (same as Examples 1-5) Catalytic hydrogenation was carried out (the used hydrogenation catalysts were the same as those in Examples 1 to 5, and the specific properties were shown in Table 2), and the effluent of the catalytic hydrogenation reaction was separated, and the hydrogen-containing gas therein was separated to obtain a hydrogenated oil product.
  • the hydrogenation reaction conditions in Comparative Examples 1 to 4 are shown in Table 4, and the evaluation results of diesel hydrogenation are shown in Table 5.
  • Example 1 Example 2 Example 3 Example 4 Example 5 Reaction temperature, °C 370 300 370 370 370 Reaction pressure, MPa 7.5 7.5 7.5 7.5 Hydrogen oil volume ratio 100 100 400 300 100 Liquid hourly space velocity, h -1 1.0 1.0 1.0 0.5 1.0 agent oil ratio 0.2 1.5 1.5 1.5 2.0
  • Comparative Example 1 Comparative Example 2 Comparative Example 3 Comparative Example 4 Reaction temperature, °C 370 300 370 370 Reaction pressure, MPa 7.5 7.5 7.5 Hydrogen oil volume ratio 100 100 400 300 Liquid hourly space velocity, h -1 1.0 1.0 1.0 0.5
  • Example 1 Example 2 Example 3 Example 4 Example 5 Sulfur content, ⁇ g/g 678.4 1120.6 24.6 16.2 461.3 Nitrogen content, ⁇ g/g 110.5 189.2 27.2 19.3 105.1 Comparative Example 1 Comparative Example 2 Comparative Example 3 Comparative Example 4 Sulfur content, ⁇ g/g 732.6 1339.4 120.5 111.3 Nitrogen content, ⁇ g/g 140.4 198.5 53.9 49.3
  • Embodiment 6 and embodiment 7 adopt the technological process shown in Figure 1 to carry out No. 2 diesel oil (see Table 1 for properties) super/subcritical fluid enhanced hydrogenation reaction, wherein the solvent used is the normal hexane and cyclohexane mixture of equal quality , the used hydrogenation catalyst is an industrial hydrofinishing catalyst (see Table 2 for properties), and this hydrogenation catalyst is packed in a fixed bed reactor to form a catalyst bed.
  • Comparative examples 6-7 adopt conventional fixed-bed hydrogenation reaction process: No. 2 fresh mixed diesel oil (see Table 1 for properties) is mixed with hydrogen and heated to a certain temperature, and then enters the fixed-bed reactor (identical to Example 6 and Example 7). ) to carry out catalytic hydrogenation (the hydrogenation catalysts used are the same as those in Examples 6 to 7, and the specific properties are shown in Table 2), and the catalytic hydrogenation reaction effluent is separated to obtain hydrogen-containing gas and hydrogenated oil products.
  • the hydrogenation reaction conditions in Comparative Example 6 and Comparative Example 7 are shown in Table 6, and the evaluation results of diesel hydrogenation are shown in Table 7.
  • Example 6 Comparative Example 6
  • Example 7 Comparative Example 7 Reaction temperature, °C 330 370 340 340 Reaction pressure, MPa 7.5 7.5 4.5 7.5 Hydrogen oil volume ratio 200 200 200 200 Liquid hourly space velocity, h -1 1.0 1.0 1.0 1.0 agent oil ratio 1.5 / 1.5 /
  • Example 6 The sulfur content and nitrogen content in the hydrogenated diesel products obtained in Example 6 (average reaction temperature 330° C.) and Comparative Example 6 (average reaction temperature 370° C.) were relatively close, indicating that compared with conventional catalytic hydrogenation, the mixture of No.
  • the average reaction temperature can be reduced by 40 °C, so the reaction energy consumption caused by material heating can be significantly reduced.
  • the reduction in reaction temperature also reduces the rate of catalyst deactivation, thereby extending the operating cycle of the plant.
  • Example 6 Comparative Example 6
  • Example 7 Comparative Example 7 Desulfurization rate, % 95.2 94.9 94.7 94.5 Nitrogen removal rate, % 93.3 93.7 93.1 93.5
  • Example 7 The sulfur content and nitrogen content in the hydrogenated diesel products obtained in Example 7 (reaction pressure 4.5 MPa) and Comparative Example 7 (reaction pressure 7.5 MPa) were relatively close, indicating that compared with conventional catalytic hydrogenation, the No. Supercritical fluid enhanced hydrogenation, the reaction pressure can be reduced by 3.0MPa, so the reaction energy consumption caused by hydrogen compression can be significantly reduced.
  • Embodiment 8 and embodiment 9 adopt the technological process shown in Figure 1 to carry out the super/subcritical fluid enhanced hydrogenation reaction of No. 1 wax oil (see Table 1 for properties), wherein the solvent used is n-heptane: cyclohexane: The toluene mass ratio is a mixture of 2:2:1, the hydrogenation catalyst used is an industrial hydrofinishing catalyst (see Table 2 for properties), and the hydrogenation catalyst is packed in a fixed bed reactor to form a catalyst bed.
  • a certain proportion of hydrogen and solvent are mixed with fresh No. 1 wax oil and heated to a suitable temperature before entering a fixed-bed reactor.
  • the mixed material is contacted with an industrial hydrofinishing catalyst and catalyzed under supercritical or subcritical fluid conditions.
  • the hydrogenation reaction is performed to obtain a catalytic hydrogenation reaction effluent; the catalytic hydrogenation reaction effluent is separated to obtain a hydrogen-containing gas, a solvent-containing light fraction and a hydrogenated oil product.
  • the super/subcritical fluid enhanced hydrogenation process conditions adopted in Example 8 and Example 9 are shown in Table 8, and the evaluation results of wax oil hydrogenation are shown in Table 9.
  • Comparative example 8 and comparative example 9 adopt conventional fixed-bed hydrogenation reaction flow process: No. 1 wax oil (see Table 1 for properties) is mixed with hydrogen and heated to a certain temperature, and sent to a fixed-bed reactor (identical to Examples 8-9). ) to carry out catalytic hydrogenation (the hydrogenation catalysts used are the same as those in Examples 8-9, and the specific properties are shown in Table 2). The catalytic hydrogenation reaction effluent is separated to obtain hydrogen-containing gas and hydrogenated oil.
  • the hydrogenation reaction conditions used in Comparative Examples 8 to 9 are shown in Table 8, and the hydrogenation evaluation results of the mixed wax oil are shown in Table 9.
  • Example 8 Comparative Example 8 Example 9 Comparative Example 9 Reaction temperature, °C 370 370 360 360 Reaction pressure, MPa 10.0 10.0 7.0 7.0 Hydrogen oil volume ratio 200 200 200 400 Liquid hourly space velocity, h -1 1.5 1.5 1.5 1.5 1.5 agent oil ratio 2 / 1 /
  • Examples 10 to 12 carried out the super/subcritical fluid-enhanced hydrogenation reaction of No. 1 wax oil (see Table 1 for properties) with reference to the process flow shown in FIG. 1 .
  • the solvent used is a mixture of n-heptane: cyclohexane: toluene mass ratio of 2:2:1
  • the hydrogenation catalyst used is an industrial hydrofinishing catalyst and an industrial hydrocracking catalyst (see Table 2 for properties).
  • the two hydrogenation catalysts were packed in a fixed-bed reactor according to a volume ratio of 1:1 to form a catalyst bed; the catalytic hydrogenation reaction effluent was separated to obtain hydrogen-containing gas, solvent-containing light fractions and hydrogenated oils.
  • the adopted super/subcritical fluid enhanced hydrogenation process conditions are shown in Table 10, and the evaluation results of wax oil hydrogenation are shown in Table 12.
  • Example 10 Example 11 Example 12 Refining reaction section temperature, °C 390 390 375 Cracking reaction section temperature, °C 395 397 385 Reaction pressure, MPa 15.0 15.0 15.0 Hydrogen oil volume ratio 600 450 600 Liquid hourly space velocity, h -1 1.0 1.0 1.0 agent oil ratio 2 2 2
  • Comparative Example 10 Comparative Example 11 Comparative Example 12 Refining reaction section temperature, °C 390 390 375 Cracking reaction section temperature, °C 395 397 385 Reaction pressure, MPa 15.0 15.0 15.0 Hydrogen oil volume ratio 600 450 600 Liquid hourly space velocity, h -1 1.0 1.0 1.0
  • a conventional fixed-bed hydrogenation reaction process is adopted: No. 1 wax oil (see Table 1 for properties) is mixed with hydrogen and heated to a certain temperature, and then enters a fixed-bed reactor (same as in Examples 10-12) for catalytic hydrogenation (used in The hydrogenation catalyst is the same as in Examples 10-12, and the specific properties are shown in Table 2) to obtain the effluent of the catalytic hydrogenation reaction; the effluent of the catalytic hydrogenation reaction is separated to obtain the hydrogen-containing gas and the hydrogenated oil product.
  • the hydrogenation reaction conditions in Comparative Examples 10 to 12 are shown in Table 11, and the evaluation results of wax oil hydrogenation are shown in Table 12.
  • Example 13 The super/subcritical fluid enhanced hydrogenation reaction of No. 1 bio-oil (ie, palm oil, see Table 1 for properties) was carried out with reference to the process flow shown in FIG. 1 .
  • the solvent used is light naphtha with a boiling point lower than 150°C
  • the hydrogenation catalysts used are industrial hydrofining catalysts and industrial hydrocracking catalysts (both properties are shown in Table 2), and these two hydrogenation catalysts are used.
  • the catalysts are packed in a fixed bed reactor according to the volume ratio of 1:1 to form a catalyst bed (respectively corresponding to the "refining reaction section” and the “cracking reaction section”); the catalytic hydrogenation reaction effluent is separated to obtain a catalyst containing Hydrogen gas, solvent-containing light ends and hydrogenated oils.
  • the adopted super/subcritical fluid enhanced hydrogenation process conditions are shown in Table 13, and the evaluation results of palm oil hydrogenation are shown in Table 14.
  • a conventional fixed-bed hydrogenation reaction process is adopted: No. 1 bio-oil (i.e. palm oil, see Table 1 for properties) is mixed with hydrogen and heated to a certain temperature, and then enters a fixed-bed reactor (same as Example 13) for catalytic hydrogenation (The hydrogenation catalyst used is the same as in Example 13, and the specific properties are shown in Table 2) to obtain a catalytic hydrogenation reaction effluent; the catalytic hydrogenation reaction effluent is separated to obtain a hydrogen-containing gas and a hydrogenated oil product.
  • the hydrogenation reaction conditions used in Comparative Example 13 are shown in Table 13, and the evaluation results of palm oil hydrogenation are shown in Table 14.
  • Example 13 Comparative Example 13 Reaction temperature, °C 360 360 Reaction pressure, MPa 10.0 10.0 Hydrogen oil volume ratio 800 800 Liquid hourly space velocity, h -1 0.6 0.6 agent oil ratio 2 /
  • Example 13 Comparative Example 13 Nitrogen removal rate, % 97.9 93.5 Deoxidation rate, % 98.5 92.6
  • the introduction of super/subcritical fluid greatly promotes the solubility of hydrogen in the distillate hydrogenation reaction system, greatly improves the supply of effective hydrogen, and thus promotes the improvement of the distillate oil. Efficient hydroconversion of sulfides, nitrides and oxides.
  • the introduction of super/subcritical fluid also reduces the energy consumption of the reaction and prolongs the operation period of the device.

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Abstract

一种馏分油超/亚临界流体强化加氢方法，包括：将馏分油原料、溶剂和氢气的混合物料预热后通入第一段固定床反应器中，在超临界或亚临界流体状态下进行第一段催化加氢，对反应流出物进行分离，收集其中的加氢油品并获得含氢气体以及含溶剂轻馏分；将含溶剂轻馏分返回第一段固定床反应器中循环使用。该方法改善了氢气在馏分油加氢反应体系中的溶解性能和扩散传质性能，提高馏分油的加氢反应效率。

Description

一种馏分油超/亚临界流体强化加氢方法 技术领域
本发明属于石油化工领域，涉及馏分油加氢技术，尤其涉及一种馏分油超/亚临界流体强化加氢方法。
背景技术
当前世界范围内原油重质化和劣质化日趋加剧，而全球环保意识日益增强，柴油、蜡油和脱沥青油等馏分油的高效清洁转化已成为炼油企业面临的重要课题。上述各种馏分油中通常含有大量的硫、氮、氧等非烃组分和芳烃化合物，加氢技术是将其转化为清洁燃料和化工原料的主要手段。加氢技术的主要优点是收率高、加氢后油品质量好。但随着原油劣质化趋势的不断加剧，二次加工馏分油如催化裂化柴油、焦化柴油和蜡油的比例不断增加，硫、氮、氧和金属等杂原子的脱除压力日益增大。清洁燃料和化工原料的生产面临着巨大挑战。
目前，馏分油原料的催化加氢普遍采用固定床加氢工艺，一般是将氢气和馏分油原料混合并加热到适宜温度后通入固定床反应器，与固体催化剂进行接触并发生催化加氢反应，然后对加氢反应流出物进行分离，得到含氢气体和加氢油品。在固定床加氢工艺中，由于氢气呈气相，馏分油原料主要呈液相，因此氢气在反应体系中需要首先溶解进入液相，然后通过扩散传质到达固体催化剂的活性位从而进行加氢反应。氢气溶解于油相并到达催化剂活性中心的传质速率对馏分油加氢过程存在显著影响。传统的滴流床工艺通常采用过量的高压氢气循环为反应供氢，造成装置的投资费用和操作成本居高不下。
为此，国内外的企业以及科研机构陆续开发出了液相加氢技术。比如美国专利申请US2013068657A1中公开的柴油液相加氢技术中，是依靠新鲜进料和循环液相产品携带饱和溶解氢供给加氢反应所需；中国专利申请CN200910165119.3和CN201010001103.1分别公开了一种液相循环加氢处理 方法，通过液相产物循环以增加溶解氢；中国专利申请CN201410506339.9中公开的液相加氢方法，是通过液相加氢产物循环和两段反应器相结合实现超低硫柴油的生产；美国专利US6123825公开的加氢处理方法，将部分加氢后馏分油作为稀释剂与反应原料和氢气混合进入反应器，但由于混合物料溶氢能力有限，难以处理化学氢耗大的原料油；中国专利申请CN201310523200.0中公开了连续液相蜡油加氢处理方法，蜡油原料和反应流出物在混合器中与氢气混合，以将氢气溶解于烃油中，但氢气与液体体积比仅为20～30v/v，需要在反应器间的多个催化剂床层补充氢气；中国专利申请CN201510036537.8中公开了一种连续液相柴油加氢的处理方法，反应器内有多个床层，床层间设有混合器，加氢反应中根据原料性质和产品要求从不同混合器补氢；中国专利申请CN201320384424.3提出一种液相循环加氢反应系统，将新鲜柴油引入循环油泵借以降低其操作温度；中国专利申请CN201410506173.0提出一种上行床液相加氢方法，通过引入高压汽提分离器和循环氢脱硫罐，降低循环油中的硫化氢，实现国Ⅴ车用柴油的生产。
上述柴油、蜡油等馏分油液相加氢技术的核心是，通过对加氢后的柴油、蜡油等馏分油进行循环以及强化氢气和馏分油混合效果从而给反应体系供氢，以达到取消氢气循环系统，降低设备投资费用和运行成本的目的。但是这些改进会带来新的问题：反应体系中大量反应产物循环造成装置处理量降低，并且对于硫、氮等杂质和多环芳烃含量高的劣质馏分油，馏分油和循环油中的溶解氢很难供给加氢反应所需的足够氢气，造成催化剂严重失活和产品质量明显下降，导致加氢装置运转周期缩短，加氢效率难以达到预期。因此，如何提供一种馏分油加氢工艺，通过提高氢气的溶解性能从而促进馏分油的高效加氢转化，仍旧是目前有待解决的难题。
发明内容
为解决上述问题，本发明提供一种馏分油超/亚临界流体强化加氢方法，能够通过引入超/亚临界流体提高反应体系中氢气的溶解性能而实现馏分油原料的高效加氢转化。
为实现上述目的，本发明提供的馏分油超/亚临界流体强化加氢方法，至少包括：
将馏分油原料、溶剂和氢气预热后通入第一段固定床反应器中，在超临界或亚临界流体状态下进行第一段催化加氢，获得催化加氢反应流出物；
对上述催化加氢反应流出物进行分离，分出其中的含氢气体并获得加氢油品以及含溶剂轻馏分；将含溶剂轻馏分返回第一段固定床反应器中替代溶剂循环使用；
其中，前述溶剂选自如下材料中的至少一种：碳原子数为4～10的链烷烃；碳原子数为5～10的环烷烃；碳原子数为6～8的芳烃；汽油、汽油馏分、石脑油、石脑油馏分。
本发明提供的超/亚临界流体强化加氢方法对于馏分油原料不做特别限定，适用的馏分油原料具体可以选自如下材料中的至少一种：煤油、柴油、蜡油、润滑油、催化裂化回炼油、催化裂化油浆和生物油。除此之外，该加氢方法同样能够适用于生物质油的催化加氢并具有很好的催化加氢效果。
本发明对于上述馏分油原料的来源不做特别限定，比如可以采用来自于石油、页岩油、生物质油等的煤油、柴油、蜡油作为馏分油原料。本发明对于作为馏分油原料的煤油、柴油、蜡油、润滑油的具体获取工艺也不做特别限定。比如柴油具体可以是直馏柴油、催化裂化柴油、焦化柴油等。蜡油具体可以是直馏蜡油、焦化蜡油、被掺入了适量脱沥青油的蜡油等。生物质油具体可以是化学、热解和生物等方法得到的动植物油脂。
基于以上说明，本发明方法所适用的原料之所以称为“馏分油”，也可理解为相对于石化工业领域公认的重油而言不太重的油品，即，原料油的馏程集中于属于上述油品范围；或者原料油中虽然含有或混有“偏重”的馏分，通常不适于重油处理装置；又或者，将这些原料油进行馏分切割，分出偏重馏分后的处理产物，比如常压蜡油馏分就是对常压渣油进行减压蒸馏切割、分出偏重馏分后所获得的偏轻馏分。
尤其是，该馏分油超/亚临界流体强化加氢方法，非常适合对催化裂化柴油、焦化柴油、直馏柴油、直馏蜡油、焦化蜡油、催化裂化回炼油和生物质油等馏分油及其混合物进行加氢转化，并最终获得品质良好的馏分油加氢产品或为后续加工提供优质原料。
本发明中所用的溶剂，尤其可以是碳原子数为5～8的直链烷烃、碳原子数为5～8的支链烷烃、碳原子数为5～8的环烷烃以及碳原子数为6～8的芳 烃中的一种或多种，或者是主要成分为上述化合物的汽油、汽油馏分、石脑油、石脑油馏分。在本发明具体实施过程中，所用的溶剂是正丁烷、正戊烷、正己烷、正庚烷、正辛烷和环己烷中的至少一种，或者也可以是轻石脑油。
当然，在具体实施过程中，还可以根据馏分油原料的性质和催化加氢的工艺条件适当补氢。第一段催化加氢过程中的补氢量具体可由氢耗、相平衡结果以及反应条件等决定。
在本发明的实施方案中，在超临界或亚临界状态下进行的第一段催化加氢反应的条件具体可以为：平均反应温度为240～460℃，反应压力为3～30MPa；剂油质量比为0.1～6：1；氢油体积比为50～1200：1，液时体积空速为0.2～12.0h -1。
实际上，在第一段催化加氢反应的开始阶段，需要向第一段固定床反应器中通入溶剂，而当第一段催化加氢反应达到稳定后，循环利用的含溶剂轻馏分能够维持第一段催化加氢反应基本稳定，因此可以根据实际情况补加少量新鲜溶剂甚至无需补加新鲜溶剂。本发明根据行业的惯常操作和认知，将通入第一段固定床反应器的含溶剂轻馏分与新鲜溶剂等同理解。
不难理解，在生产开始时，剂油质量比指的是向第一段固定床反应器中通入的新鲜溶剂与馏分油原料之间的质量比，而在连续生产中，剂油质量比指的是含溶剂轻馏分以及可能补充的新鲜溶剂与馏分油原料之间的质量比。
同样，在生产开始阶段，通入第一段固定床反应器中的液相物料，指的是溶剂和馏分油原料；在生产稳定时，液相物料指的是含溶剂轻馏分、可能补充的新鲜溶剂以及馏分油原料。液时体积空速，指的是液相物料的体积空速(单位为小时 -1)。氢油体积比，指的是氢气与馏分油原料之间的标准状态下体积比。
发明人研究发现，适当提高反应温度、反应压力以及降低液时空速，并选择合适的剂油比和氢油体积比，有利于提高馏分油原料的催化加氢效果，尤其可以提高脱硫率、脱氮率和脱氧率。当然在实际生产中，还可根据馏分油原料的具体组成、能耗、生产成本、生产效率以及生产目标等因素确定最优的工艺条件。
在本发明优选的实施方案中，通常将第一段催化加氢反应的工艺条件控制在：平均反应温度260～430℃，反应压力4～25MPa；剂油质量比0.1～4.0： 1；氢油体积比80～1000：1；液时体积空速0.3～6.0h -1。
当然，在实际生产中，可以根据馏分油原料的具体组成合理调整第一段催化加氢反应的工艺条件。比如馏分油原料为煤油、柴油、生物油脂轻馏分、或者它们之中任意几种的混合物，则在超临界或亚临界状态下进行的第一段催化加氢反应的平均温度具体可以为240～440℃，优选为260～420℃，进一步优选为280～420℃；反应压力具体可以为3～25MPa，优选为4～20MPa；剂油质量比具体可以为0.1～5.0：1，优选为0.1～3.0：1；氢油体积比可以为50～1000：1，优选为80～800：1，进一步为80～600：1；液时体积空速具体可以为0.5～12.0h -1，优选为0.5～6.0h -1，进一步为0.5～5.0h -1。
再比如馏分油原料为蜡油、润滑油、催化裂化回炼油、催化裂化油浆、生物油脂重馏分或者它们之中任意几种的混合物，则在超临界或亚临界状态下进行的第一段催化加氢反应的平均温度具体可以为300～460℃，优选为320～430℃，进一步为340～420℃；反应压力具体可以为4～30MPa，优选为5～25MPa，进一步为6～20MPa；剂油质量比具体可以为0.1～6.0：1，优选为0.1～5.0：1，进一步为0.2～4.0：1；氢油体积比具体可以为50～1200：1，优选为80～1000：1，进一步为100～800：1；液时体积空速具体可以为0.2～8.0h -1，优选为0.3～5.0h -1，进一步为0.5～4.0h -1。
根据本发明提出的方案，待处理的馏分油原料经过超/亚临界流体强化加氢过程，也称第一段催化加氢，可有效提升催化加氢的效果。本发明对于第一段催化加氢的具体反应类型不做特别限定，比如可以是加氢脱硫、加氢脱氮、加氢脱氧、烯烃加氢饱和、芳烃加氢饱和、加氢脱金属、加氢裂化等催化加氢反应。或者说，本发明提供的催化加氢工艺，可以适用于为获得成品所进行的加氢精制、为后续加工提供优质原料所进行的加氢处理、以及为获得轻质油品和化学品所进行的加氢改质和加氢裂化等工艺过程。
可以理解，根据馏分油原料的组成以及催化加氢反应类型的不同，应选择相应的加氢催化剂，比如加氢精制催化剂、加氢处理催化剂、加氢改质催化剂、加氢裂化催化剂等的一种或多种。上述加氢催化剂可以商购，也可以根据本领域的常规方法制备。催化剂在使用前需要进行硫化处理，将活性金属组分的氧化物转化为活性更高的硫化物，预硫化处理为本领域的常规方法。
在本发明具体实施过程中，所选用的加氢催化剂为工业级的加氢催化剂， 具体可以是工业级加氢精制催化剂、工业级加氢处理催化剂、工业级加氢改质催化剂和工业级加氢裂化催化剂中的一种或多种。加氢催化剂由载体负载活性金属成分得到，其中活性金属可以选自ⅥB族金属和Ⅷ族金属中的至少一种；载体可以根据处理要求和相应加氢反应的特点而确定，具体方案中，加氢精制催化剂和加氢处理催化剂的载体选自氧化铝、氧化铝—氧化硅、氧化铝—氧化钛等中的一种或几种，加氢改质催化剂和加氢裂化催化剂的载体为含有分子筛的氧化铝、氧化铝—氧化硅、氧化铝—氧化钛等中的一种或几种，其中，前述氧化铝具体可以为γ-Al 2O 3、δ-Al 2O 3、θ-Al 2O 3和η-Al 2O 3中的至少一种。上述载体中还可选择性地添加氟、磷、钛、锆和硼中的一种或多种作为助剂组分。本领域技术人员可以根据需要选择适宜的活性金属、载体和助剂组分及其含量。
在本发明具体实施过程中，生物油脂的催化加氢可以选用上述针对石油基油品的工业加氢催化剂，也可以选用其它加氢催化剂。
在本发明某些具体实施方式中，馏分油超/亚临界流体强化加氢工艺需要设置一段固定床反应器和至少一级催化剂床层。具体而言，可将氢气、溶剂、新鲜馏分油原料的混合物料先加热到适宜温度后进入第一段固定床反应器，混合物料与加氢催化剂在催化剂床层接触并在超临界或亚临界流体状态下发生第一段催化加氢反应。
具体的，进入第一段固定床反应器的物料，即氢气、溶剂、新鲜馏分油原料，可同时汇入进料管路共同进入固定床反应器，也可以先在混合器中充分混合后进入固定床反应器；或者还可以在固定床反应器中的催化剂床层之间设置的混合器内充分混合，本实施例在此不做特别限定。
本发明对于如何对催化加氢反应流出物进行分离不做特别限定，可采用本领域常规分离手段。在本发明具体实施过程中，是将催化加氢反应流出物经过注水和换热后进行油、气、水三相分离，分离得到的气相和油相进一步分离得到含氢气体、含溶剂轻馏分，以及加氢后的汽油馏分、煤油馏分、柴油馏分和尾油等加氢油品的一种或多种。其中，含溶剂轻馏分返回第一段固定床反应器混合物料中替代溶剂循环使用；含氢气体可以直接排放或经无害化处理后排放，或者也可以对含氢气体提纯后继续使用。
本发明中，可以在超临界或亚临界流体状态下对馏分油原料实施一段式 固定床催化加氢，也可以在此基础上继续实施第二段催化加氢，即采取两段式固定床加氢工艺。具体可以根据馏分油原料的组成以及生产目标选择实施一段式固定床加氢工艺或两段式固定床加氢工艺。
具体的，在对第一段固定床加氢反应流出物进行分离之前，可将第一段催化加氢反应流出物直接通入第二段固定床反应器中进行第二段催化加氢，得到第二段催化加氢反应流出物(即催化加氢反应流出物)。或者说，对第一段催化加氢反应流出物不进行分离，而是直接引入第二段固定床反应器进行第二段催化加氢，得到第二段催化加氢反应流出物(即催化加氢反应流出物)。然后对催化加氢反应流出物进行分离，获得加氢油品、含氢气体以及含溶剂轻馏分；其中含溶剂轻馏分返回到第一段固定床反应器中替代溶剂循环利用。
或者，也可以先将第一段催化加氢得到的催化加氢反应流出物经分离得到含氢气体、含溶剂轻馏分和第一段加氢油品，然后将第一段加氢油品通入第二段固定床反应器中进行第二段催化加氢，获得第二段催化加氢反应流出物。或者说，对第一段催化加氢反应得到的催化加氢反应流出物先实施分离处理，分出其中的含氢气体并收集其中的第一段加氢油品和含溶剂轻馏分，然后将含溶剂轻馏分返回到第一段固定床反应器中替代溶剂循环利用，将第一段加氢油品与氢气混合后通入第二段固定床反应器中进行第二段催化加氢，获得第二段催化加氢反应流出物；最后对第二段催化加氢反应流出物进行分离，分出其中的含氢气体并收集其中的第二段加氢油品。
上述两段式固定床加氢工艺可以应用串联的两段固定床反应器进行超/亚临界流体强化加氢反应，也可以是在超/亚临界流体强化加氢工艺基础上引入现有常规的馏分油加氢工艺。不难理解，采用两段式固定床加氢工艺，所最终得到的第二段加氢油品的成分区别于第一段加氢油品，因此可根据实际情况对第二段加氢油品实施适宜的分离工艺。
本发明对于固定床反应器的具体类型不做特别限定，可以是下行床(即物料从固定床反应器顶部进入，底部流出)，也可以是上行床(即物料从固定床反应器底部进入，顶部流出)。
与现有常规馏分油加氢技术相比，本发明提供的馏分油超/亚临界流体强化加氢方法，通过将馏分油原料、含溶剂轻馏分(或溶剂)和氢气三元体系 控制在超临界或亚临界流体状态下进行催化加氢反应，具有如下优点：
(1)本发明可以极大地改善氢气在馏分油加氢反应体系中的溶解性能和扩散传质性能，从而大幅度提高催化加氢反应过程中有效氢的供给，进而明显提高馏分油加氢反应效率以及氢气和催化剂的利用效率，实现馏分油原料的高效加氢转化。
(2)超/亚临界流体能够带走催化加氢反应所释放的热量，避免因局部过热造成的催化剂失活以及避免副反应的发生，从而延长催化剂的寿命和整个馏分油超/亚临界流体强化加氢系统的运转周期。
(3)本方法可以大幅度提高有效氢的供给，因此相较于常规馏分油加氢技术，本发明提供的馏分油超/亚临界流体强化加氢方法，还可大幅度降低氢气循环量以及氢气压缩机等设备投资和生产运行成本。
(4)本发明可以适用于煤油、柴油和蜡油等馏分油的加氢精制、加氢处理、加氢改质和加氢裂化，也可作为馏分油加氢处理的前端工艺，具体可以根据馏分油原料性质和生产目标灵活地进行工艺调整。
(5)本发明的馏分油超/亚临界流体强化加氢工艺，可在常规馏分油加氢工艺基础上进行简单改造得到，因此具有改造和投资成本低的优势，便于在实际生产中推广和应用。
附图说明
图1为本发明提供的一种馏分油超/亚临界流体强化加氢实施方式的流程示意图；
图2为本发明提供的另一种馏分油超/亚临界流体强化加氢实施方式的流程示意图；
图3为本发明提供的再一种馏分油超/亚临界流体强化加氢实施方式的流程示意图。
附图标记说明：
1-馏分油原料；                 2-溶剂或含溶剂轻馏分；
3-氢气；                       4-混合物料；
5-固定床反应器；               6-加氢反应流出物；
7-高压分离器；                 8-含氢气体；
9-高分油；                     10-低压分离器；
11-低分气；                    12-低分油；
13-分馏系统；                  14-汽提塔；
15-闪蒸油；                    16-脱硫化氢塔。
具体实施方式
本发明在馏分油原料的催化加氢反应中引入溶剂并构建超/亚临界流体强化加氢反应体系，旨在促进馏分油原料的催化加氢反应，实现馏分油原料的高效加氢转化。
为使本发明的目的、技术方案和优点更加清楚，下面将结合本发明实施例中的附图和实施例，对本发明实施例中的技术方案进行描述。显然，所描述的实施例是本发明一部分实施例，而不是全部的实施例。本发明保护范围不限于实施例，基于本发明中的实施例，本领域普通技术人员在没有做出创造性劳动前提下所获得的所有其他实施例，都属于本发明保护的范围。
本发明具体实施过程中，可利用可视高温高压相平衡仪作为临界点测定装置，通过相平衡实验中相界面的变化和临界乳光现象考察馏分油原料-氢气-溶剂三元体系相行为，从而确定达到超临界流体状态的临界点参数。根据上述临界点的实验结果设定馏分油超/亚临界流体强化加氢的反应条件范围。
图1是本发明提供的馏分油超/亚临界流体强化加氢方法中，其中一种实施方式的流程示意图，为清晰起见，图中省略了许多通用设备或装置，如氢气压缩机、换热器、加热炉、泵、分离器、蒸馏塔和储罐等，但这些通用设备或装置的作用、设置以及具体选择，对本领域的普通技术人员是公知的。
参考图1，将馏分油原料1、溶剂2或含溶剂轻馏分2以及氢气3混合，得到的混合物料4在换热器中换热升温后，再进入加热炉中继续升温至工艺要求温度，然后进入固定床反应器5中，混合物料4与经预硫化处理的催化剂接触，并在超临界或亚临界流体条件下发生催化加氢反应。加氢反应流出物6经注水和换热后进入高压分离器7进行油、气、水三相分离，分离出的高分气8可进行后续的回收，分离出的水相产物(未图示)出装置，分离出的高分油9进入低压分离器10进行气液分离，得到的低分气11自低压分离器10顶部出口排出，低分油12自低压分离器10底部出口排出后进入分馏系 统13，分离出含溶剂轻馏分2，并获得加氢后的加氢油品(未图示)。该加氢油品具体可以是汽油馏分、煤油馏分、柴油馏分和尾油等中的一种或多种。含溶剂轻馏分2返回并入混合物料4中替代溶剂循环使用。
图2与图1实施方式的区别是：将从低压分离器10底部排出的低分油12通入汽提塔14，汽提塔14底部排出的闪蒸油15进入分馏系统13分离得到加氢油品(未图示)，该加氢油品具体可以是加氢汽油、加氢煤油、加氢柴油以及加氢尾油等，汽提塔14顶部排出的含溶剂轻馏分2返回并入混合物料4中替代溶剂循环使用。
图3与图1实施方式的区别是：对高压分离器7顶部分离出的高分气8进一步实施处理，高分气8经换热后通入脱硫化氢塔16，脱硫化氢塔16下部排出脱除硫化氢等气体的含溶剂轻馏分2，含溶剂轻馏分2返回并入混合物料4中替代溶剂循环使用。
下面通过实施例和比较例将对本发明实施例提供的方法和效果予以进一步的说明，但并不因此而限制本发明。
以下实施例和比较例中所用馏分油原料的性质见表1。其中，1号柴油为催化裂化柴油，2号柴油为直馏柴油与催化裂化柴油的混合柴油，1号蜡油为直馏蜡油与焦化蜡油的混合蜡油，1号生物油为棕榈油。
表1馏分油原料的性质
项目 1号柴油 2号柴油 1号蜡油 1号生物油
密度(20℃)，g/cm 3 0.9535 0.8665 0.9126 0.8865
硫含量，μg/g 4432.2 6700.0 6567.0 506.0
氮含量，μg/g 819.4 756.1 2156.0 1200.0
氧含量，m％ / / / 12.0
芳烃含量，m％ 82.3 45.2 48.8 /
IBP 200.1 180.1 291.0 197
10％ 235.8 195.8 378.0 332
50％ 277.7 258.2 443.0 415
90％ 346.8 328.0 499.0 /
95％ 364.8 351.3 523.0 /
FBP 369.9 360.4 541.0 /
金属，μg/g / / 10.1 /
注：“/”代表未检测。
以下实施例和比较例中所用加氢催化剂的性质见表2。表2中的加氢精制催化剂和加氢裂化催化剂均为工业级。
表2加氢催化剂的性质
Figure PCTCN2021104430-appb-000001
注 #1：Y型分子筛性质：比表面积712m 2/g；孔容0.51cm 3/g；平均孔径8.1nm
以下实施例和比较例的工艺条件中，“反应温度”均为催化剂床层的平均温度，“液时空速”均以馏分油原料计。
实施例1～5
实施例1～5采用图1所示的工艺流程进行1号柴油(性质见表1)的超/亚临界流体强化加氢反应，其中所用的溶剂为等质量的正戊烷与环己烷混合物，所用的加氢催化剂为工业加氢精制催化剂(性质见表2)，并将此工业加氢精制催化剂装填于一个固定床反应器中形成一个催化剂床层。
将氢气、溶剂与新鲜1号柴油按比例混合并加热到适宜温度后进入固定床反应器，混合物料与工业加氢精制催化剂接触并在超临界或亚临界流体条件下发生催化加氢反应，获得催化加氢反应流出物；对催化加氢反应流出物进行分离，获得含氢气体、含溶剂轻馏分和加氢油品。实施例1～5所采用的超/亚临界流体强化加氢工艺条件详见表3，柴油加氢评价结果见表5。
比较例1～4
比较例1～4采用常规固定床加氢反应流程：将新鲜1号柴油(性质见表 1)与氢气混合后加热到一定温度，然后通入固定床反应器(与实施例1～5相同)进行催化加氢(所用加氢催化剂与实施例1～5相同，具体性质见表2)，对催化加氢反应流出物进行分离，分出其中的含氢气体并获得加氢油品。比较例1～4中的加氢反应条件见表4，柴油加氢评价结果见表5。
表3实施例1～5中超/亚临界流体强化加氢工艺条件
  实施例1 实施例2 实施例3 实施例4 实施例5
反应温度，℃ 370 300 370 370 370
反应压力，MPa 7.5 7.5 7.5 7.5 7.5
氢油体积比 100 100 400 300 100
液时空速，h -1 1.0 1.0 1.0 0.5 1.0
剂油比 0.2 1.5 1.5 1.5 2.0
表4比较例1～4的加氢反应操作条件
  比较例1 比较例2 比较例3 比较例4
反应温度，℃ 370 300 370 370
反应压力，MPa 7.5 7.5 7.5 7.5
氢油体积比 100 100 400 300
液时空速，h -1 1.0 1.0 1.0 0.5
表5催化裂化柴油加氢评价结果
  实施例1 实施例2 实施例3 实施例4 实施例5
硫含量，μg/g 678.4 1120.6 24.6 16.2 461.3
氮含量，μg/g 110.5 189.2 27.2 19.3 105.1
  比较例1 比较例2 比较例3 比较例4  
硫含量，μg/g 732.6 1339.4 120.5 111.3  
氮含量，μg/g 140.4 198.5 53.9 49.3  
由表5中实施例1～4和比较例1～4的加氢反应结果对比可以看出，在相同的柴油原料进料条件和相同的反应条件下，对1号柴油进行加氢精制，实施例1～4所获得的加氢柴油产物中硫含量和氮含量均明显低于对应的比较例；并且经催化加氢后，实施例1～4中硫含量的降低效果尤为明显。这表 明，采用超/亚临界流体强化加氢工艺，能够实现催化裂化柴油中硫化物和氮化物的有效脱除，脱硫率和脱氮率均明显优于常规固定床加氢工艺的结果，从而获得品质更高的加氢精制柴油。
进一步对比实施例1～5的结果可知，适当提高反应温度、氢油体积比，以及适当降低液时空速，加氢精制柴油的硫含量明显降低，同时也得到较为优异的脱氮效果。
实施例6～7
实施例6和实施例7采用图1所示的工艺流程进行2号柴油(性质见表1)超/亚临界流体强化加氢反应，其中所用的溶剂为等质量的正己烷与环己烷混合物，所用的加氢催化剂为工业加氢精制催化剂(性质见表2)，并将此加氢催化剂装填于固定床反应器中形成一个催化剂床层。
具体的，将氢气、溶剂与新鲜2号柴油按比例混合并加热到适宜温度后进入固定床反应器，混合物料与工业加氢精制催化剂接触并在超临界或亚临界流体条件下发生催化加氢反应，获得催化加氢反应流出物；对催化加氢反应流出物进行分离，获得含氢气体、含溶剂轻馏分和加氢油品。实施例6～7所采用的超/亚临界流体强化加氢工艺条件见表6，柴油加氢评价结果见表7。
比较例6～7
比较例6～7采用常规固定床加氢反应流程：2号新鲜混合柴油(性质见表1)与氢气混合后加热到一定温度，然后进入固定床反应器(与实施例6和实施例7相同)进行催化加氢(所用加氢催化剂与实施例6～7相同，具体性质见表2)，对催化加氢反应流出物进行分离，获得含氢气体和加氢油品。比较例6和比较例7中的加氢反应条件见表6，柴油加氢评价结果见表7。
表6实施例6～7和比较例6～7中加氢反应操作条件
  实施例6 比较例6 实施例7 比较例7
反应温度，℃ 330 370 340 340
反应压力，MPa 7.5 7.5 4.5 7.5
氢油体积比 200 200 200 200
液时空速，h -1 1.0 1.0 1.0 1.0
剂油比 1.5 / 1.5 /
由表7中实施例6和比较例6的加氢反应结果对比可以看出，在相同的氢气和柴油原料进料条件以及相同的反应压力下，对混合柴油进行加氢精制，
实施例6(平均反应温度330℃)和比较例6(平均反应温度370℃)所获得的加氢柴油产物中硫含量和氮含量较为接近，表明与常规催化加氢相比，对2号混合柴油进行超临界流体强化加氢，平均反应温度可以降低40℃，因此可以明显降低因物料加热所带来的反应能耗。此外，反应温度的降低也会降低催化剂的失活速率，从而延长装置的运转周期。
表7混合柴油加氢评价结果
  实施例6 比较例6 实施例7 比较例7
脱硫率，％ 95.2 94.9 94.7 94.5
脱氮率，％ 93.3 93.7 93.1 93.5
由表7中实施例7和比较例7的加氢反应结果对比可以看出，在相同的氢气和柴油原料进料条件以及相同的反应温度下，对混合柴油进行加氢精制，
实施例7(反应压力4.5MPa)和比较例7(反应压力7.5MPa)所获得的加氢柴油产物中硫含量和氮含量较为接近，表明与常规催化加氢相比，对2号混合柴油进行超临界流体强化加氢，反应压力可以降低3.0MPa，因此能够明显降低氢气压缩带来的反应能耗。
实施例8～9
实施例8和实施例9采用图1所示的工艺流程进行1号蜡油(性质见表1)的超/亚临界流体强化加氢反应，其中所用的溶剂为正庚烷：环己烷：甲苯质量比为2：2：1的混合物，所用的加氢催化剂为工业加氢精制催化剂(性质见表2)，并将此加氢催化剂装填于固定床反应器中形成一个催化剂床层。
具体的，将一定比例的氢气、溶剂与新鲜1号蜡油混合并加热到适宜温度后进入固定床反应器，混合物料与工业加氢精制催化剂接触并在超临界或亚临界流体条件下发生催化加氢反应，获得催化加氢反应流出物；对催化加氢反应流出物进行分离，获得含氢气体、含溶剂轻馏分和加氢油品。实施例8和实施例9所采用的超/亚临界流体强化加氢工艺条件见表8，蜡油加氢评价结果见表9。
比较例8～9
比较例8和比较例9采用常规固定床加氢反应流程：将1号蜡油(性质 见表1)与氢气混合后加热到一定温度，送入固定床反应器(与实施例8～9相同)进行催化加氢(所用加氢催化剂与实施例8～9相同，具体性质见表2)。对催化加氢反应流出物进行分离，获得含氢气体和加氢油品。比较例8～9所用的加氢反应条件见表8，混合蜡油加氢评价结果见表9。
表8实施例8～9、比较例8～9中加氢反应操作条件
  实施例8 比较例8 实施例9 比较例9
反应温度，℃ 370 370 360 360
反应压力，MPa 10.0 10.0 7.0 7.0
氢油体积比 200 200 200 400
液时空速，h -1 1.5 1.5 1.5 1.5
剂油比 2 / 1 /
表9混合蜡油加氢处理评价结果
  实施例8 比较例8 实施例9 比较例9
柴油收率，m％ 8.7 8.6 8.5 8.4
柴油硫含量，μg/g 1.8 2.3 4.2 4.8
柴油氮含量，μg/g 1.6 2.2 2.6 3.7
蜡油收率，m％ 88.2 88.2 88.3 88.3
蜡油脱硫率，％ 98.5 91.9 96.1 95.9
蜡油脱氮率，％ 96.4 89.7 89.0 88.6
蜡油脱金属率，％ 88.1 77.2 79.2 79.2
由表9中实施例8和比较例8的加氢反应结果对比可以看出，在相同的氢气和蜡油原料进料条件以及相同的反应条件下，对1号蜡油进行加氢处理，柴油和蜡油的收率基本相当。但是采用实施例8的工艺条件，所获得的加氢蜡油产物中，脱硫率、脱氮率和脱金属率均明显高于对应的比较例8。由此可证实，采用本发明的方法对蜡油进行超/亚临界流体强化加氢处理，可以在更加缓和的反应条件下，实现蜡油中硫、氮和金属等杂质的有效脱除，从而为后续催化裂化或加氢裂化提供优质原料。
由表9中实施例9和比较例9的加氢反应结果对比可以看出，在相同的蜡油原料进料条件和相同的反应条件下，对1号蜡油进行加氢处理，柴油和 蜡油的收率基本相当，实施例9(氢油体积比200)和比较例9(氢油体积比400)所获得的加氢蜡油产物中硫含量、氮含量和金属含量较为接近，表明2号混合柴油的超临界流体强化加氢工艺对比常规加氢可以降低一半的氢油体积比，从而明显降低大量氢气循环带来的反应能耗。
实施例10～12
实施例10～12参考图1所示的工艺流程进行1号蜡油(性质见表1)的超/亚临界流体强化加氢反应。其中，所用的溶剂为正庚烷：环己烷：甲苯质量比为2：2：1的混合物，所采用的加氢催化剂为工业加氢精制催化剂和工业加氢裂化催化剂(性质均见表2)，并将此两种加氢催化剂按照体积比1：1装填于一个固定床反应器中各形成一个催化剂床层；对催化加氢反应流出物进行分离，获得含氢气体、含溶剂轻馏分和加氢油品。所采用的超/亚临界流体强化加氢工艺条件见表10，蜡油加氢评价结果见表12。
表10实施例10～12中超/亚临界流体强化加氢工艺条件
  实施例10 实施例11 实施例12
精制反应段温度，℃ 390 390 375
裂化反应段温度，℃ 395 397 385
反应压力，MPa 15.0 15.0 15.0
氢油体积比 600 450 600
液时空速，h -1 1.0 1.0 1.0
剂油比 2 2 2
表11比较例10～12的加氢反应操作条件
  比较例10 比较例11 比较例12
精制反应段温度，℃ 390 390 375
裂化反应段温度，℃ 395 397 385
反应压力，MPa 15.0 15.0 15.0
氢油体积比 600 450 600
液时空速，h -1 1.0 1.0 1.0
比较例10～12
采用常规固定床加氢反应流程：将1号蜡油(性质见表1)与氢气混合 后加热到一定温度，然后进入固定床反应器(与实施例10～12相同)进行催化加氢(所用加氢催化剂与实施例10～12相同，具体性质见表2)，获得催化加氢反应流出物；对催化加氢反应流出物进行分离，获得含氢气体和加氢油品。比较例10～12中的加氢反应条件见表11，蜡油加氢评价结果见表12。
表12混合蜡油加氢裂化评价结果
Figure PCTCN2021104430-appb-000002
注：“—”代表没有尾油产品。
由表12中实施例10～12和比较例10～12的加氢反应结果对比可以看出，在相同的氢气和蜡油进料条件以及相同的反应条件下，对1号蜡油进行加氢裂化，汽油、柴油和尾油的收率基本相当。但是，采用实施例10～12的工艺方法和条件，所获得的加氢蜡油产物中硫含量和氮含量均明显低于对应的比较例。由此可证实，采用本发明的方法对蜡油进行超/亚临界流体强化加氢处理，可以在较低的氢油比条件下，实现蜡油中硫、氮等杂质的有效脱除。
此外，对比实施例10～12的结果可知，适当提高反应温度，混合蜡油加氢产品中轻馏分含量增加，且各个馏分的硫含量明显降低。
实施例13
实施例13参考图1所示的工艺流程进行1号生物油(即棕榈油，性质见表1)的超/亚临界流体强化加氢反应。其中，所用的溶剂为沸点低于150℃的轻石脑油，所采用的加氢催化剂为工业加氢精制催化剂和工业加氢裂化催 化剂(性质均见表2)，并将此两种加氢催化剂按照体积比1：1装填于一个固定床反应器中各形成一个催化剂床层(分别对应为“精制反应段”和“裂化反应段”)；对催化加氢反应流出物进行分离，获得含氢气体、含溶剂轻馏分和加氢油品。所采用的超/亚临界流体强化加氢工艺条件见表13，棕榈油加氢评价结果见表14。
比较例13
采用常规固定床加氢反应流程：将1号生物油(即棕榈油，性质见表1)与氢气混合后加热到一定温度，然后进入固定床反应器(与实施例13相同)进行催化加氢(所用加氢催化剂与实施例13相同，具体性质见表2)，获得催化加氢反应流出物；对催化加氢反应流出物进行分离，获得含氢气体和加氢油品。比较例13所用的加氢反应条件见表13，棕榈油加氢评价结果见表14。
表13实施例13和对比例13中加氢反应操作条件
  实施例13 比较例13
反应温度，℃ 360 360
反应压力，MPa 10.0 10.0
氢油体积比 800 800
液时空速，h -1 0.6 0.6
剂油比 2 /
表14生物油加氢裂化评价结果
  实施例13 比较例13
脱氮率，％ 97.9 93.5
脱氧率，％ 98.5 92.6
由表14中实施例13和比较例13的加氢反应结果对比可以看出，在相同的进料条件和相同的反应条件下，对1号生物油进行加氢精制，实施例13所获得的加氢产物中氧含量和氮含量均明显低于对应的比较例。这表明，采用超/亚临界流体强化加氢工艺，能够实现生物油中氧化物和氮化物的有效脱除，脱氧率和脱氮率均明显优于常规固定床加氢工艺的结果，从而获得品质更高的加氢生物油。
综合上述结果说明：在适宜的条件下，超/亚临界流体的引入极大地促进 了氢气在馏分油加氢反应体系中的溶解性能，大幅度改善了有效氢的供给，从而促进了馏分油中硫化物、氮化物和氧化物的高效加氢转化。此外，与常规的催化加氢工艺相比，超/亚临界流体的引入还降低了反应能耗、延长了装置的运转周期。
最后应说明的是：以上各实施例仅用以说明本发明的技术方案，但是，本发明并不限于上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变型，或者对部分技术特征进行替换，或者通过任何合适的方式进行组合，而这些修改只要不违背本发明的思想，其同样应当视为本发明所公开的内容和范围。

Claims (10)

  1. 一种馏分油超/亚临界流体强化加氢方法，其特征在于，至少包括：
    将馏分油原料、溶剂和氢气预热后通入第一段固定床反应器中，在超临界或亚临界流体状态下进行第一段催化加氢，获得催化加氢反应流出物；
    对所述催化加氢反应流出物进行分离，分出其中的含氢气体并获得加氢油品以及含溶剂轻馏分；将所述含溶剂轻馏分返回所述第一段固定床反应器中替代溶剂循环使用；
    其中，所述溶剂选自如下材料中的至少一种：碳原子数为4～10的链烷烃；碳原子数为5～10的环烷烃；碳原子数为6～8的芳烃；汽油、汽油馏分、石脑油、石脑油馏分。
  2. 根据权利要求1所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述馏分油原料选自如下材料中的至少一种：
    煤油、柴油、蜡油、润滑油、催化裂化回炼油、催化裂化油浆和生物油。
  3. 根据权利要求1或2所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述第一段催化加氢过程中，平均反应温度为240～460℃，优选260～430℃；反应压力为3～30MPa，优选4～25MPa；剂油质量比为0.1～6：1，优选0.1～4.0：1；氢油体积比为50～1200：1，优选80～1000：1；液时体积空速为0.2～12.0h -1，优选0.3～6.0h -1。
  4. 根据权利要求3所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述馏分油原料为煤油和/或柴油，所述第一段催化加氢的平均反应温度为240～440℃，优选260～420℃，进一步优选为280～420℃；反应压力为3～25MPa，优选4～20MPa；剂油质量比为0.1～5.0：1，优选0.1～3.0：1；氢油体积比为50～1000：1，优选80～800：1，进一步优选为80～600：1；液时体积空速为0.5～12.0h -1，优选0.5～6.0h -1。
  5. 根据权利要求3所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述馏分油原料为蜡油、润滑油、催化裂化回炼油、催化裂化油浆和生物油中的至少一种，所述第一段催化加氢的平均反应温度为300～460℃，优选320～430℃，进一步优选为340～420℃；反应压力为4～30MPa，优选5～25MPa，进一步优选为6～20MPa；剂油质量比为0.1～6.0：1，优选0.1～5.0：1，进一步优选为0.2～4.0：1；氢油体积比为50～1200：1，优选80～ 1000：1，进一步优选为100～800：1；液时体积空速为0.2～8.0h -1，优选0.3～5.0h -1，进一步优选为0.5～4.0h -1。
  6. 根据权利要求1-5任一项所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述第一段催化加氢包括加氢脱硫、加氢脱氮、加氢脱氧、烯烃加氢饱和、芳烃加氢饱和、加氢脱金属和加氢裂化中的至少一种。
  7. 根据权利要求1-6任一项所述的馏分油超/亚临界流体强化加氢方法，其特征在于，所述第一段催化加氢过程中所用的加氢催化剂包括加氢精制催化剂、加氢处理催化剂、加氢改质催化剂和加氢裂化催化剂中的至少一种；所述加氢催化剂包括载体以及负载在载体上的活性金属成分，其中，所述活性金属成分选自ⅥB族和Ⅷ族金属中的至少一种。
  8. 根据权利要求1-7任一项所述的馏分油超/亚临界流体强化加氢方法，其特征在于，还包括：将经过所述第一段催化加氢的反应流出物通入第二段固定床反应器中进行第二段催化加氢，获得第二段催化加氢反应流出物；对所述第二段催化加氢反应流出物进行所述分离。
  9. 根据权利要求1-7任一项所述的馏分油超/亚临界流体强化加氢方法，其特征在于，还包括：将所述加氢油品和氢气通入第二段固定床反应器中进行第二段催化加氢，得到第二段催化加氢反应流出物；
    对所述第二段催化加氢反应流出物进行分离，分出其中的含氢气体并获得第二段加氢油品。
  10. 根据权利要求1所述的馏分油超/亚临界流体强化加氢方法，其特征在于，将所述含氢气体排出，或者将所述含氢气体提纯后循环利用。
PCT/CN2021/104430 2020-07-10 2021-07-05 一种馏分油超/亚临界流体强化加氢方法 WO2022007736A1 (zh)

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