WO2022011866A1 - 一种柴油加氢的反应系统及方法 - Google Patents

一种柴油加氢的反应系统及方法 Download PDF

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WO2022011866A1
WO2022011866A1 PCT/CN2020/122722 CN2020122722W WO2022011866A1 WO 2022011866 A1 WO2022011866 A1 WO 2022011866A1 CN 2020122722 W CN2020122722 W CN 2020122722W WO 2022011866 A1 WO2022011866 A1 WO 2022011866A1
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Prior art keywords
hydrogenation
reaction
micro
reactor
hydrogen
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PCT/CN2020/122722
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English (en)
French (fr)
Inventor
张志炳
周政
张锋
李磊
孟为民
王宝荣
杨高东
罗华勋
杨国强
田洪舟
曹宇
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南京延长反应技术研究院有限公司
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Publication of WO2022011866A1 publication Critical patent/WO2022011866A1/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
    • 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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/008—Details of the reactor or of the particulate material; Processes to increase or to retard the rate of reaction
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/02—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds
    • B01J8/04—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds
    • B01J8/0446—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds the flow within the beds being predominantly vertical
    • 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
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/04—Diesel oil

Definitions

  • the invention relates to the field of diesel hydrogenation, in particular to a reaction system and method for diesel hydrogenation.
  • the first object of the present invention is to provide a reaction system for diesel hydrogenation, which reduces energy consumption, lowers reaction temperature, and improves reaction yield by combining a hydrogenation reactor and a micro-interface generator. , improve the utilization rate of raw materials, especially the utilization rate of hydrogen, and effectively improve the production capacity, thereby improving the quality and yield of products, in addition to saving equipment costs and saving equipment area.
  • the second object of the present invention is to provide a method for hydrodesulfurization of diesel using the above reaction system, the desulfurized diesel obtained by the reaction is environmentally friendly and clean, has sufficient power as a fuel, is widely used, and improves the applicability of the desulfurized diesel.
  • the invention provides a reaction system for diesel hydrogenation, comprising: a hydrogenation reactor, wherein a plurality of fixed beds are arranged in sequence from top to bottom in the hydrogenation reactor, and a catalyst is filled in the fixed beds; A micro-interface generator is arranged between the adjacent fixed beds, and a hydrogen inlet is arranged on the micro-interface generator; a product outlet is arranged on the side wall of the hydrogenation reactor, and the product outlet is connected to the hydrorefining reaction
  • the reaction product from the bottom of the hydrorefining reactor is passed into the gas-liquid separation tank to separate the gas phase containing hydrogen sulfide and hydrogen from the product, and the final product desulfurized diesel oil is obtained.
  • the diesel hydrogenation reaction system of the present invention is provided with a micro-interface generator inside the hydrogenation reactor to disperse and break the incoming hydrogen into micro-bubbles, thereby improving the mass transfer effect.
  • the main function of straight-run diesel is to cooperate with the dispersion and crushing of the gas, which is equivalent to the role of the medium.
  • the micro-interface generator is arranged inside the hydrogenation reactor, preferably in a straight line along the vertical direction and evenly arranged between the adjacent fixed beds. , to ensure better catalytic hydrogenation effect of diesel in the gap between the two fixed beds, which is equivalent to simultaneous dispersing and crushing and catalytic reaction, making the operation of dispersing and crushing more closely connected with the reaction.
  • the catalyst added to the catalytic hydrogenation reaction It is a solid, so it is filled into the reactor by filling, and the dispersion and crushing are mainly to improve the catalytic reaction effect and reduce energy consumption. Therefore, since the catalyst is arranged inside the reactor, the micro-interface generator is the most suitable. The best way is to set it inside the reactor.
  • the setting position of the micro-interface generator is also obtained through practical design. , which needs to be designed according to the different characteristics of different reactions.
  • the fixed bed has 3 stages and the micro-interface generators are 3, wherein two micro-interface generators are arranged between two fixed beds, and another micro-interface generator is arranged in the hydrogenation process bottom of the reactor.
  • the number of 3 micro-interface generators can already ensure the effect of dispersion and crushing.
  • micro-interface generator arranged inside the hydrogenation reactor is of a pneumatic type, and the mass transfer effect is improved by passing hydrogen into the micro-interface generator and then directly contacting with straight-run diesel oil and then breaking into micro-bubbles.
  • the micro-interface generator can also be set up in the subsequent hydrotreating reactor to reduce energy consumption and improve the utilization rate of raw materials.
  • the micro-interface generator is arranged in the hydrogenation reactor and the hydrotreating reactor, which can improve the efficiency of the reaction and reduce the energy consumption of the reaction.
  • micro-interface generator used in the present invention has been embodied in the inventor's prior patents, such as application numbers CN201610641119. Patents of CN205833127U and CN207581700U. In the previous patent CN201610641119.6, the specific product structure and working principle of the micro-bubble generator (that is, the micro-interface generator) were introduced in detail.
  • the body is provided with an inlet communicating with the cavity, the opposite first and second ends of the cavity are open, wherein the cross-sectional area of the cavity is from the middle of the cavity to the first and second ends of the cavity.
  • the second end is reduced; the secondary crushing piece is arranged at at least one of the first end and the second end of the cavity, a part of the secondary crushing piece is arranged in the cavity, and both ends of the secondary crushing piece and the cavity are open An annular channel is formed between the through holes of the micro-bubble generator.
  • the micro-bubble generator also includes an air inlet pipe and a liquid inlet pipe.” From the specific structure disclosed in the application document, we can know that its specific working principle is: the liquid enters the micron tangentially through the liquid inlet pipe.
  • the micro-bubble generator in this patent belongs to the pneumatic micro-interface generation. device.
  • the previous patent 201610641251.7 records that the primary bubble breaker has a circulating liquid inlet, a circulating gas inlet and a gas-liquid mixture outlet, and the secondary bubble breaker communicates the feed port with the gas-liquid mixture outlet, indicating that the bubble breaker is both It needs to be mixed with gas and liquid.
  • the primary bubble breaker mainly uses circulating liquid as power, so in fact, the primary bubble breaker belongs to the hydraulic micro-interface generator, and the secondary bubble breaker is a gas-liquid breaker. The mixture is simultaneously fed into the elliptical rotating ball for rotation, so that the bubbles are broken during the rotation, so the secondary bubble breaker is actually a gas-liquid linkage type micro-interface generator.
  • both hydraulic micro-interface generators and gas-liquid linkage micro-interface generators belong to a specific form of micro-interface generators.
  • the micro-interface generators used in the present invention are not limited to the above-mentioned forms.
  • the specific structure of the bubble breaker described in the prior patent is only one of the forms that the micro-interface generator of the present invention can take.
  • the previous patent 201710766435.0 records that "the principle of the bubble breaker is to achieve high-speed jets to achieve gas collision", and it is also stated that it can be used in micro-interface enhanced reactors to verify the relationship between the bubble breaker and the micro-interface generator.
  • the prior patent CN106187660 also has related records for the specific structure of the bubble breaker, see the specific description in paragraphs [0031]-[0041], and the accompanying drawings, which are related to the bubble breaker S-2 The specific working principle of the bubble breaker is explained in detail.
  • the top of the bubble breaker is the liquid phase inlet, and the side is the gas phase inlet.
  • the liquid phase entering from the top provides the entrainment power, so as to achieve the effect of crushing into ultra-fine bubbles, which can also be seen in the accompanying drawings.
  • the bubble breaker has a conical structure, and the diameter of the upper part is larger than that of the lower part, so that the liquid phase can provide better entrainment power.
  • micro-interface generator Since the micro-interface generator was just developed in the early stage of the previous patent application, it was named as micro-bubble generator (CN201610641119.6), bubble breaker (201710766435.0), etc., and later changed its name to micro-interface generator with continuous technological improvement.
  • the micro-interface generator in the present invention is equivalent to the previous micro-bubble generator, bubble breaker, etc., but the names are different.
  • the micro-interface generator of the present invention belongs to the prior art, although some bubble breakers belong to the type of pneumatic bubble breakers, some belong to the type of hydraulic bubble breakers, and some belong to the type of gas bubble breakers.
  • the type of liquid-linked bubble breaker but the difference between the types is mainly selected according to the specific working conditions.
  • the connection between the micro-interface generator and the reactor and other equipment, including the connection structure and connection position depends on the micro-interface generator. It depends on the structure of the interface generator, which is not limited.
  • the type of the hydrogenation reactor for the hydrogenation reaction is a fixed bed reactor, the catalyst in the fixed bed reactor is fixed on the bed layer, and the catalyst for the hydrogenation reaction is generally a nickel-based catalyst, preferably the catalyst can be a supported catalyst Nickel-based catalysts of type 1, or nickel-based catalysts modified with alkaline earth metal oxides or rare earth metal oxides are more preferable.
  • the type of the hydrotreating reactor is also a fixed-bed reactor, and the type of catalyst selected can be the same as or not the same as the type of catalyst selected by the hydrotreating reactor.
  • the amount of catalyst in the hydrotreating reactor and the hydrotreating reactor can be specifically determined according to the properties of the feedstock oil, product quality requirements and catalyst properties. Generally, the amount of catalyst in the hydrotreating reactor is smaller than that in the hydrotreating reactor.
  • the reason for the amount of catalyst used is that what is carried out in the previous hydrogenation reactor is a preliminary reaction, and what is carried out in the subsequent hydrotreating reactor is a deep reaction, and after adding a micro-interface generator in the hydrogenation reactor, it is It can improve the reaction efficiency, so it is also possible to appropriately reduce the amount of the catalyst, which does not affect the normal progress of the hydrogenation reaction.
  • the volume ratio of the catalyst dosage in the hydrogenation reactor and the hydrotreating reactor can be (1:12)-(1:15).
  • a gas-phase outlet is provided on the top of the gas-liquid separation tank, and the gas-phase outlet is connected to a desulfurization tower to realize the absorption of hydrogen sulfide in the gas phase.
  • the top of the desulfurization tower is provided with a desulfurization agent spray port, and the desulfurization tower is provided with multi-layer packing to increase the mass transfer effect.
  • the reaction product after the hydrogenation reaction is passed into the gas-liquid separation tank, the hydrogen sulfide and hydrogen go out from the top of the gas-liquid separation tank, and go to the desulfurization tower for hydrogen sulfide removal, and the liquid phase is discharged from the bottom of the gas-liquid separation tank. Collect and store in finished product tanks.
  • a desulfurizing agent circulating tank is arranged beside the desulfurization tower, the desulfurizing agent circulating tank is connected with the lower part of the side wall of the desulfurizing tower through a pipeline, and the top of the desulfurizing agent circulating tank is connected with the desulfurizing agent through a pipeline.
  • Sprinkler connection is arranged beside the desulfurization tower, the desulfurizing agent circulating tank is connected with the lower part of the side wall of the desulfurizing tower through a pipeline, and the top of the desulfurizing agent circulating tank is connected with the desulfurizing agent through a pipeline.
  • the desulfurizing agent is generally lye, which is circulated into the desulfurization tower and is in reverse contact with the gas phase entering from the bottom of the desulfurization tower, sprayed and absorbed to achieve desulfurization, and the lye after desulfurization is returned to the desulfurizer circulation tank. In the case of serious pollution during the circulating use of the lye in the circulating tank, stop for replacement.
  • each of the hydrogen inlets is collected and connected to a hydrogen main pipeline, the hydrogen main pipeline is connected to a hydrogen storage tank, and the hydrogen recovered from the top of the desulfurization tower is merged with the hydrogen main pipeline through a pipeline.
  • a heater is provided between the hydrogenation reactor and the hydrofinishing reactor to heat the product from the product outlet. Because the hydrogenation reaction generally has a relatively high temperature, setting a heater between the two reactors can improve the reaction efficiency and correspondingly reduce the energy consumption of the reactor itself.
  • the reaction product from the bottom of the hydrotreating reactor is cooled by a heat exchanger before entering the gas-liquid separation tank.
  • the bottom of the hydrogenation reactor is provided with a straight-run diesel inlet, and the straight-run diesel inlet is connected to the heat exchanger through a pipeline for heating the straight-run diesel before entering the hydrogenation reactor .
  • a heat exchanger is provided to achieve heat exchange between the reaction product after hydrogenation and the incoming straight-run diesel.
  • a pump body can be provided on the pipelines connected between the corresponding devices according to actual needs.
  • the present invention also provides a reaction method for diesel hydrogenation, comprising the following steps:
  • the pressure of the hydrogenation reaction is 2-8MPa, and the temperature of the hydrogenation reaction is 250-320°C. Most preferably, the pressure of the hydrogenation reaction is 3.5 MPa, and the temperature of the hydrogenation reaction is 300°C.
  • the pressure of the refining hydrogenation reaction is 2-8 MPa, and the temperature of the refining hydrogenation reaction is 250-320°C.
  • the pressure of the refining hydrogenation reaction is 3.5MPa, and the temperature of the refining hydrogenation reaction is 300°C.
  • reaction pressure and reaction temperature will be slightly higher than the hydrogenation reaction temperature and pressure during actual operation.
  • the desulfurized diesel product obtained by adopting the diesel hydrogenation reaction of the invention has good quality and high yield, and the desulfurization rate can reach 99.95%.
  • the reaction method for diesel hydrogenation of the invention has low reaction temperature, greatly reduced pressure and high liquid hourly space velocity, which is equivalent to increasing the production capacity, and the final desulfurization rate is close to 100%, which is nearly 1 percentage point higher than before.
  • reaction system of diesel hydrogenation of the present invention reduces the energy consumption, reduces the reaction temperature, improves the reaction yield, and improves the utilization rate of the raw material by combining the hydrogenation reactor and the micro-interface generator;
  • reaction system of diesel hydrogenation of the present invention is most favorable for improving mass transfer effect by setting the micro-interface generator at a specific position;
  • the reaction system of diesel hydrogenation of the present invention is to realize the heat exchange between the reaction product after the hydrogenation and the straight-run diesel oil that enters by setting a heat exchanger;
  • the reaction method for diesel hydrogenation of the present invention is low in reaction temperature, greatly reduced in pressure, and high in liquid hourly space velocity, which is equivalent to increasing the production capacity, and the final desulfurization rate is close to 100%, which is nearly 1 percentage point higher than before.
  • FIG. 1 is a schematic structural diagram of a reaction system for diesel hydrogenation provided in an embodiment of the present invention.
  • 90-Desulfurization tower 901-Desulfurizer spray port;
  • the terms “installed”, “connected” and “connected” should be understood in a broad sense, unless otherwise expressly specified and limited, for example, it may be a fixed connection or a detachable connection Connection, or integral connection; can be mechanical connection, can also be electrical connection; can be directly connected, can also be indirectly connected through an intermediate medium, can be internal communication between two elements.
  • installed should be understood in a broad sense, unless otherwise expressly specified and limited, for example, it may be a fixed connection or a detachable connection Connection, or integral connection; can be mechanical connection, can also be electrical connection; can be directly connected, can also be indirectly connected through an intermediate medium, can be internal communication between two elements.
  • FIG. 1 it is a reaction system for diesel hydrogenation according to an embodiment of the present invention, which includes a hydrogenation reactor 10.
  • a product outlet 105 is provided on the side wall of the hydrogenation reactor 10, and the product outlet 105 is connected to the hydrogenation reactor.
  • the refining reactor 20 is a hydrorefining reactor 20 for deep hydrogenation reaction.
  • a heater 30 is provided between the hydrogenation reactor 10 and the hydrorefining reactor 20 to heat the product from the product outlet 105 .
  • the reaction product from the bottom of the hydrotreating reactor 20 is cooled by the heat exchanger 40 before entering the gas-liquid separation tank 70 .
  • the bottom of the hydrogenation reactor 10 is provided with a straight-run diesel oil inlet 104, and the straight-run diesel oil inlet 104 is connected to the heat exchanger 40 through a pipeline for heating the straight-run diesel oil before entering the hydrogenation reactor 10. .
  • the purpose of the heat exchanger 40 is to exchange heat between the raw diesel and the obtained finished diesel, so as to ensure that the raw diesel enters the hydrogenation reactor 10 for hydrogenation reaction after being heated.
  • a plurality of fixed beds 102 are arranged in sequence from top to bottom, and the fixed beds 102 are filled with catalysts; a micro-interface generator 101 is arranged between the adjacent fixed beds 102,
  • the fixed bed 102 is preferably three sections, and the micro-interface generators 101 are preferably three, wherein two micro-interface generators 101 are arranged between the two fixed beds 102, and another micro-interface generator 101 is arranged in the
  • the bottom of the hydrogenation reactor 10 basically needs three micro-interface generators 101 in order to satisfy the effect of dispersion and crushing.
  • the micro-interface generator 101 is provided with a hydrogen inlet 103, and the hydrogen inlet 103 is collected and connected to the hydrogen main pipeline. 50.
  • the hydrogen main pipeline 50 is connected to the hydrogen storage tank 60, and the raw hydrogen enters the micro-interface generator 101 after being branched into each branch pipeline through the hydrogen main pipeline 50 after entering from the hydrogen storage tank 60.
  • the product after coming out of the hydrotreating reactor 20 first passes through the gas-liquid separation tank 70 to separate the gas phase containing hydrogen sulfide and hydrogen from the product, and the separated product goes to the finished product tank 80 from the bottom of the gas-liquid separation tank 70,
  • the finished product is desulfurized diesel, of course, the desulfurized diesel can also be further purified.
  • the top of the gas-liquid separation tank 70 is provided with a gas phase outlet 701, and the gas phase containing hydrogen and hydrogen sulfide coming out of the gas phase outlet 701 goes to the desulfurization tower 90 to realize the absorption of hydrogen sulfide in the gas phase, and the top in the desulfurization tower 90 is provided with a desulfurizing agent
  • the desulfurization tower 90 is provided with multiple layers of packing to increase the mass transfer effect.
  • a desulfurizing agent circulating tank 902 is arranged on the side of the desulfurization tower 90.
  • the desulfurizing agent circulating tank 902 is connected with the lower part of the side wall of the desulfurizing tower 90 through a pipeline.
  • the top of the desulfurizing agent circulating tank 902 is connected with the desulfurizing agent spray port 901 through a pipeline.
  • the lye is circulated into the desulfurization tower 90 and is in reverse contact with the gas phase entering from the bottom of the desulfurization tower 90, sprayed and absorbed to achieve desulfurization, and the lye after desulfurization is returned to the desulfurizer circulation tank 902, when the desulfurizer When the lye in the circulation tank 902 is seriously polluted, it is stopped for replacement.
  • micro-interface generator 101 it is not limited to include three micro-interface generators 101.
  • additional micro-interface generators 101 can also be added.
  • the installation position is actually not limited. It can also be built-in, and it can also be installed on the side wall of the kettle when it is built in, so that the micro-bubbles from the outlet of the micro-interface generator 101 can be hedged. The way of the built-in micro interface generator 101.
  • the type of the hydrogenation reactor 10 can be not only a fixed-bed reactor, but also other types such as a ebullated-bed reactor.
  • the way of feeding and discharging materials is also not limited, and can be fed from below.
  • the top discharging method can also be used for the top feeding and the bottom discharging method, but the method of side feeding and top discharging is more preferable.
  • Hydrogen and straight-run diesel are firstly hydrogenated in the hydrogenation reactor 10. Before the hydrogenation reaction, hydrogen is passed into the hydrogenation reactor 10. In the micro-interface generator 101, the gas forms micro-bubbles after being dispersed and broken, which is more conducive to the efficient reaction. After the hydrogenation reaction, the reaction product is sent to the hydrofinishing reactor 20 through a pipeline for deep hydrogenation reaction.
  • the pressure of the above hydrogenation reaction is 2-8MPa, and the temperature of the hydrogenation reaction is 250-320°C. Most preferably, the pressure of the hydrogenation reaction is 3.5 MPa, and the temperature of the hydrogenation reaction is 300°C.
  • the pressure of the refining hydrogenation reaction is 2-8MPa, and the temperature of the refining hydrogenation reaction is 250-320°C. Most preferably, the pressure of the refining hydrogenation reaction is 3.5MPa, and the temperature of the refining hydrogenation reaction is 300°C.
  • the catalyst used in both the hydrogenation reaction and the refining hydrogenation reaction is a Mo-Ni type catalyst, and the volume ratio of the catalyst dosage in the hydrogenation reactor 10 and the hydrogenation refining reactor 20 can be (1:12)-(1:15) , preferably 1:13.
  • the gas phase at the top goes to the desulfurization tank for desulfurization treatment, and the product separated at the bottom goes to the finished product tank 80 for storage.
  • the removal rate of desulfurization can reach 99.95%, which is increased by nearly 1 percentage point compared with the previous hydrogenation reaction process.
  • the pressure and temperature of the hydrogenation reactor are reduced, and the energy consumption is sufficiently reduced.
  • the reaction system for diesel hydrogenation of the present invention has fewer equipment components, small footprint, low energy consumption, low cost, high safety, controllable reaction, and raw materials.
  • the conversion rate is high, which is equivalent to providing a reaction system with stronger operability for the field of diesel hydrogenation, which is worthy of widespread application.

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  • Oil, Petroleum & Natural Gas (AREA)
  • Chemical Kinetics & Catalysis (AREA)
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Abstract

一种柴油加氢的反应系统及方法，所述反应系统包括：加氢反应器（10），所述加氢反应器（10）内从上至下依次设置有多段固定床层（102），所述固定床层（102）内装填有催化剂；相邻所述固定床层（102）之间设置有微界面发生器（101），所述微界面发生器（101）上设置有氢气进口（103）；所述加氢反应器（10）的侧壁设置有产物出口（105），所述产物出口（105）连接加氢精制反应器（20）进行深度加氢反应；所述加氢精制反应器（20）底部出来的反应产物通入气液分离罐（70）以将含硫化氢、氢气的气相与产品分离后，得到最终产品脱硫柴油。该反应系统通过与微界面发生器（101）进行组合后，降低了能耗，降低了反应温度，提高了反应产率，提高了原料的利用率，尤其是提高氢气的利用率，同时有效的提高了产能。

Description

一种柴油加氢的反应系统及方法 技术领域
本发明涉及柴油加氢领域，具体而言，涉及一种柴油加氢的反应系统及方法。
背景技术
随着人们环保意识的提高以及环保法规的日益严格，生产和使用清洁车用燃料越来越成为一种发展趋势，柴油深度加氢脱硫技术的开发则成为了目前研究的热点。目前，大多柴油加氢精制装置操作流程是原料油通过同精制柴油、反应产物换热，并经加热炉加热至反应要求温度后进行加氢，这种加氢过程虽然操作方便，容易产业化，但是能耗高，加氢反应器压力高、温度高，产能也比较低。
有鉴于此，特提出本发明。
发明内容
本发明的第一目的在于提供一种柴油加氢的反应系统，该反应系统通过将加氢反应器与微界面发生器进行组合后，降低了能耗，降低了反应温度，提高了反应产率，提高了原料的利用率，尤其是提高氢气的利用率，同时有效的提高了产能，进而提高了产品的品质以及收率，此外也起到了节省设备成本，节约设备占地面积的作用。
本发明的第二目的在于提供一种采用上述反应系统进行柴油加氢脱硫的方法，反应得到的脱硫柴油环保、清洁，作为燃料动力足，应用广泛，提高了脱硫柴油的适用面。
为了实现本发明的上述目的，特采用以下技术方案：
本发明提供了一种柴油加氢的反应系统，包括：加氢反应器，所述加氢反应器内从上至下依次设置有多段固定床层，所述固定床层内装填有催化剂；相邻所述固定床层之间设置有微界面发生器，所述微界面发生器上设置有氢气进口；所述加氢反应器的侧壁设置有产物出口，所述产物出口连接加氢精制反应器进行深度加氢反应；所述加氢精制反应器底部出来的反应产物通入气液分离罐以将含硫化氢、氢气的气相与产品分离后，得到最终产品脱硫柴油。
本发明的柴油加氢的反应系统，通过在加氢反应器内部设置有微界面发生器，将进入的氢气进行分散破碎成微气泡，从而提高传质效果，在微界面发生器内部通入的直馏柴油主要作用是配合气体的分散破碎，相当于介质的作用。
并且，该微界面发生器设置在加氢反应器的内部，最好沿垂直方向呈一条直线依次均匀设置在相邻所述固定床层之间，这样的设置方式可以在加氢脱硫反应的同时，以保证两个固定床层间隙的柴油催化加氢效果更佳，相当于分散破碎与催化反应同时进行，使得分散破碎操作与反应的进行联系更加紧密，另外该催化加氢反应所添加的催化剂为固体，因此是采用装填的方式装填到反应器内部的，而分散破碎主要就是为了配合提高催化反应效果的，降低能耗的，所以鉴于催化剂设置在反应器的内部，那么微界面发生器最优的方式也是设置在反应器的内部，可见将微界面发生器设置在加氢反应器的内部对于提高传质效果是最为有利的，因此微界面发生器的设置位置也是经过实践设计所得到的，需要根据不同反应的不同特点进行特定的设计。
优选地，所述固定床层为3段，所述微界面发生器为3个，其中两个微界面发生器设置两段固定床层之间，另外一个微界面发生器设置在所述加氢反应器内的底部。3个微界面发生器的数量已经可以保证分散破碎的效果。
上述设置在加氢反应器内部的微界面发生器为气动式，通过将氢气通入微界面发生器后与直馏柴油直接触后破碎形成微气泡的方式，提高传质效果。
当然，除了在加氢反应器内部设置微界面发生器，在后续的加氢精制反应器内部也可以相应的设置微界面发生器以起到相应的降低能耗、提高原料利用 率的效果，同时在加氢反应器内以及加氢精制反应器内设置微界面发生器，更能提高反应的效率，降低反应的能耗。
本领域所属技术人员可以理解的是，本发明所采用的微界面发生器在本发明人在先专利中已有体现，如申请号CN201610641119.6、201610641251.7、CN201710766435.0、CN106187660、CN105903425A、CN109437390A、CN205833127U及CN207581700U的专利。在先专利CN201610641119.6中详细介绍了微米气泡发生器(即微界面发生器)的具体产品结构和工作原理，该申请文件中记载了“微米气泡发生器包括本体和二次破碎件、本体内具有空腔，本体上设有与空腔连通的进口，空腔的相对的第一端和第二端均敞开，其中空腔的横截面积从空腔的中部向空腔的第一端和第二端减小；二次破碎件设在空腔的第一端和第二端中的至少一个处，二次破碎件的一部分设在空腔内，二次破碎件与空腔两端敞开的通孔之间形成一个环形通道。微米气泡发生器还包括进气管和进液管。”从该申请文件中公开的具体结构可以知晓其具体工作原理为：液体通过进液管切向进入微米气泡发生器内，超高速旋转并切割气体，使气体气泡破碎成微米级别的微气泡，从而提高液相与气相之间的传质面积，而且该专利中的微米气泡发生器属于气动式微界面发生器。
另外，在先专利201610641251.7中有记载一次气泡破碎器具有循环液进口、循环气进口和气液混合物出口，二次气泡破碎器则是将进料口与气液混合物出口连通，说明气泡破碎器都是需要气液混合进入，另外从后面的附图中可知，一次气泡破碎器主要是利用循环液作为动力，所以其实一次气泡破碎器属于液动式微界面发生器，二次气泡破碎器是将气液混合物同时通入到椭圆形的旋转球中进行旋转，从而在旋转的过程中实现气泡破碎，所以二次气泡破碎器实际上是属于气液联动式微界面发生器。其实，无论是液动式微界面发生器，还是气液联动式微界面发生器，都属于微界面发生器的一种具体形式，然而本发明所采用的微界面发生器并不局限于上述几种形式，在先专利中所记载的气泡破碎器的具体结构只是本发明微界面发生器可采用的其中一种形式而已。
此外，在先专利201710766435.0中记载到“气泡破碎器的原理就是高速射流以达到气体相互碰撞”，并且也阐述了其可以用于微界面强化反应器，验证本身气泡破碎器与微界面发生器之间的关联性；而且在先专利CN106187660中对于气泡破碎器的具体结构也有相关的记载，具体见说明书中第[0031]-[0041]段，以及附图部分，其对气泡破碎器S-2的具体工作原理有详细的阐述，气泡破碎器顶部是液相进口，侧面是气相进口，通过从顶部进来的液相提供卷吸动力，从而达到粉碎成超细气泡的效果，附图中也可见气泡破碎器呈锥形的结构，上部的直径比下部的直径要大，也是为了液相能够更好的提供卷吸动力。
由于在先专利申请的初期，微界面发生器才刚研发出来，所以早期命名为微米气泡发生器(CN201610641119.6)、气泡破碎器(201710766435.0)等，随着不断技术改进，后期更名为微界面发生器，现在本发明中的微界面发生器相当于之前的微米气泡发生器、气泡破碎器等，只是名称不一样。
综上所述，本发明的微界面发生器属于现有技术，虽然有的气泡破碎器属于气动式气泡破碎器类型，有的气泡破碎器属于液动式气泡破碎器类型，还有的属于气液联动式气泡破碎器类型，但是类型之间的差别主要是根据具体工况的不同进行选择，另外关于微界面发生器与反应器、以及其他设备的连接，包括连接结构、连接位置，根据微界面发生器的结构而定，此不作限定。
优选地，进行加氢反应的加氢反应器的类型为固定床反应釜，固定床反应釜内催化剂固定在床层上，加氢反应的催化剂一般采用的镍基催化剂，优选地催化剂可以为负载型的镍基催化剂，或者采用碱土金属氧化物或稀土金属氧化物改性过的镍基催化剂更优。
优选地，所述加氢精制反应器的类型也为固定床反应釜，其选择的催化剂类型可以与加氢反应器所选择的催化剂类型一致，也不可以不一致。
另外，加氢反应器与加氢精制反应器内的催化剂用量可以根据原料油的性质和产品质量要求以及催化剂性质具体来确定，一般加氢反应器内的催化剂用量要小于加氢精制反应器内所用催化剂的用量，原因在于前面的加氢反应器内 所进行的为初步反应，后续加氢精制反应器内进行的为深度反应，并且在加氢反应器内添加了微界面发生器后，是可以提高反应效率的，因此适当的降低催化剂的用量也是可以的，并不影响加氢反应的正常进行。加氢反应器与加氢精制反应器内的催化剂用量体积比可以为(1:12)-(1:15)。
优选地，所述气液分离罐的顶部设置有气相出口，所述气相出口连接脱硫塔以实现对气相中硫化氢的吸收。所述脱硫塔内的顶部设置有脱硫剂喷洒口，所述脱硫塔内布设有多层填料以增加传质效果。
进行了加氢反应之后的反应产物通入气液分离罐，硫化氢以及氢气从气液分离罐的顶部出去，去往脱硫塔进行硫化氢脱除，液相则从气液分离罐的底部出去进行收集，储存在成品罐中。
优选地，所述脱硫塔旁侧设置有脱硫剂循环罐，所述脱硫剂循环罐与所述脱硫塔的侧壁下部通过管道连接，所述脱硫剂循环罐的顶部通过管道与所述脱硫剂喷洒口连接。
脱硫剂一般选择的是碱液，碱液循环进入到脱硫塔与从脱硫塔底部进入的气相逆向接触，喷洒吸收以实现脱硫，脱硫之后的碱液再返回到脱硫剂循环罐中，当脱硫剂循环罐中的碱液循环使用过程中污染比较严重的情况下，停车进行更换。
优选地，每个所述氢气进口汇集后连接氢气总管道，所述氢气总管道连接氢气储罐，所述脱硫塔顶部回收的氢气通过管道与所述氢气总管道汇合。
优选地，所述加氢反应器与所述加氢精制反应器之间设置有加热器以实现从所述产物出口出来的产物进行加热。因为加氢反应一般温度比较高，因此在两个反应器之间设置加热器是可以提高反应效率的，也相应的降低了反应器本身的能耗。
优选地，所述加氢精制反应器底部出来的反应产物先经过换热器冷却后再进入所述气液分离罐。
优选地，所述加氢反应器的底部设置有直馏柴油进口，所述直馏柴油进口 通过管道与所述换热器连接以用于将直馏柴油加热后再进入所述加氢反应器。通过设置换热器以实现加氢出来之后的反应产物与进入的直馏柴油之间的换热。
本发明的反应系统中可根据实际需要在相应的设备之间连接的管道上设置泵体。
本发明还提供了一种柴油加氢的反应方法，包括如下步骤：
将直馏柴油与氢气混合微界面分散破碎后，进行加氢反应，再进行精制加氢后，气液分离以实现将含硫化氢、氢气的气相与产品分离。
优选地，所述加氢反应的压力2-8MPa，所述加氢反应的温度为250-320℃。最优选地，加氢反应的压力为3.5MPa，加氢反应的温度为300℃。
优选地，所述精制加氢反应的压力2-8MPa，所述精制加氢反应的温度为250-320℃。最优选地，精制加氢反应的压力为3.5MPa，精制加氢反应的温度为300℃。
上述反应方法中，由于精制加氢反应之前并没有像加氢反应一样预先对原料进行分散破碎，所以实际操作时，其反应压力、反应温度较加氢反应的温度、压力会稍微高一些。
采用本发明柴油加氢反应得到的脱硫柴油品品质好、收率高，脱硫率可以达到99.95％。
本发明的柴油加氢的反应方法反应温度低、压力大幅度下降，液时空速高，相当于提高了产能，最终的脱硫率接近100％，较以往提高了近1个百分点。
与现有技术相比，本发明的有益效果在于：
(1)本发明的柴油加氢的反应系统通过将加氢反应器与微界面发生器进行组合后，降低了能耗，降低了反应温度，提高了反应产率，提高了原料的利用率；
(2)本发明的柴油加氢的反应系统通过将微界面发生器设置在特定的位 置，从而对于提高传质效果是最为有利的；
(3)本发明的柴油加氢的反应系统通过设置换热器以实现加氢出来之后的反应产物与进入的直馏柴油之间的换热；
(4)本发明的柴油加氢的反应方法反应温度低、压力大幅度下降，液时空速高，相当于提高了产能，最终的脱硫率接近100％，较以往提高了近1个百分点。
附图说明
通过阅读下文优选实施方式的详细描述，各种其他的优点和益处对于本领域普通技术人员将变得清楚明了。附图仅用于示出优选实施方式的目的，而并不认为是对本发明的限制。而且在整个附图中，用相同的参考符号表示相同的部件。在附图中：
图1为本发明实施例提供的柴油加氢的反应系统的结构示意图。
附图说明：
10-加氢反应器；                 101-微界面发生器；
102-固定床层；                  103-氢气进口；
104-直馏柴油进口；              105-产物出口；
20-加氢精制反应器；             30-加热器；
40-换热器；                     50-氢气总管道；
60-氢气储罐；                   70-气液分离罐；
701-气相出口；                  80-成品罐；
90-脱硫塔；                     901-脱硫剂喷洒口；
902-脱硫剂循环罐。
具体实施方式
下面将结合附图和具体实施方式对本发明的技术方案进行清楚、完整地描述，但是本领域技术人员将会理解，下列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，仅用于说明本发明，而不应视为限制本发明的范围。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。实施例中未注明具体条件者，按照常规条件或制造商建议的条件进行。所用试剂或仪器未注明生产厂商者，均为可以通过市售购买获得的常规产品。
在本发明的描述中，需要说明的是，术语“中心”、“上”、“下”、“左”、“右”、“竖直”、“水平”、“内”、“外”等指示的方位或位置关系为基于附图所示的方位或位置关系，仅是为了便于描述本发明和简化描述，而不是指示或暗示所指的装置或元件必须具有特定的方位、以特定的方位构造和操作，因此不能理解为对本发明的限制。此外，术语“第一”、“第二”、“第三”仅用于描述目的，而不能理解为指示或暗示相对重要性。
在本发明的描述中，需要说明的是，除非另有明确的规定和限定，术语“安装”、“相连”、“连接”应做广义理解，例如，可以是固定连接，也可以是可拆卸连接，或一体地连接；可以是机械连接，也可以是电连接；可以是直接相连，也可以通过中间媒介间接相连，可以是两个元件内部的连通。对于本领域的普通技术人员而言，可以具体情况理解上述术语在本发明中的具体含义。
为了更加清晰的对本发明中的技术方案进行阐述，下面以具体实施例的形式进行说明。
实施例
参阅图1所示，为本发明实施例的柴油加氢的反应系统，其包括了加氢反应器10，加氢反应器10的侧壁设置有产物出口105，所述产物出口105连接加氢精制反应器20进行深度加氢反应的加氢精制反应器20，加氢反应器10 与加氢精制反应器20之间设置有加热器30以实现对从产物出口105出来的产物进行加热。加氢精制反应器20底部出来的反应产物先经过换热器40冷却后再进入所述气液分离罐70。加氢反应器10的底部设置有直馏柴油进口104，所述直馏柴油进口104通过管道与所述换热器40连接以用于将直馏柴油加热后再进入所述加氢反应器10。换热器40设置的目的是为了将原料柴油与得到的成品柴油进行换热，从而保证了原料柴油加热之后进入加氢反应器10进行加氢反应。
在加氢反应器10的内部从上至下依次设置有多段固定床层102，所述固定床层102内装填有催化剂；相邻所述固定床层102之间设置有微界面发生器101，固定床层102最好为3段，所述微界面发生器101最好为3个，其中两个微界面发生器101设置两段固定床层102之间，另外一个微界面发生器101设置在所述加氢反应器10内的底部，为了满足分散破碎的效果基本上需要3个微界面发生器101，在微界面发生器101上设置有氢气进口103，氢气进口103汇集后连接氢气总管道50，所述氢气总管道50连接氢气储罐60，原料氢气从氢气储罐60进来后通过氢气总管道50分支到各个分支管道后，进入微界面发生器101中。
从加氢精制反应器20出来之后的产物先经过气液分离罐70以将含硫化氢、氢气的气相与产品分离后，分离出的产品从气液分离罐70的底部去往成品罐80，成品为脱硫柴油，当然脱硫柴油也可以后续进行进一步的纯化。
气液分离罐70的顶部设置有气相出口701，气相出口701出来的含有氢气、硫化氢的气相去往脱硫塔90以实现对气相中硫化氢的吸收，脱硫塔90内的顶部设置有脱硫剂喷洒口901，所述脱硫塔90内布设有多层填料以增加传质效果，填料一般为2层，均布在脱硫塔90的内部。
脱硫塔90旁侧设置有脱硫剂循环罐902，脱硫剂循环罐902与脱硫塔90的侧壁下部通过管道连接，脱硫剂循环罐902的顶部通过管道与脱硫剂喷洒口901连接，脱硫剂一般选择的是碱液，碱液循环进入到脱硫塔90与从脱硫塔 90底部进入的气相逆向接触，喷洒吸收以实现脱硫，脱硫之后的碱液再返回到脱硫剂循环罐902中，当脱硫剂循环罐902中的碱液污染比较严重的情况下，停车进行更换。
在上述实施例中，并不局限于包括3个微界面发生器101，为了增加分散、传质效果，也可以多增设额外的微界面发生器101，安装位置其实也是不限的，可以外置也可以内置，内置时还可以采用安装在釜内的侧壁上相对设置，以实现从微界面发生器101的出口出来的微气泡发生对冲，当然对于本发明的方案来说最优的是采用内置微界面发生器101的方式。
在上述实施例中，加氢反应器10的类型除了可以是固定床反应釜以外，还可以是沸腾床反应釜等其他类型，除此之外进出料的方式也不限，可以从下方进料上方出料，也可以采用上方进料下方出料的方式，但是比较优选地是侧方进料，上方出料的方式。
在上述实施例中，泵体的个数并没有具体要求，可根据需要在相应的位置设置。
以下简要说明本发明的柴油加氢反应系统的工作过程和原理：
氮气吹扫反应系统中的各个设备，然后开车进行操作，氢气与直馏柴油先在加氢反应器10内进行加氢反应，进行加氢反应之前，将氢气通入位于加氢反应器10内部的微界面发生器101中，进行分散破碎后使得气体形成微气泡，更有利于反应高效的进行，加氢反应后，反应产物通过管道送入加氢精制反应器20中进行深度加氢反应。
其中，上述加氢反应的压力2-8MPa，所述加氢反应的温度为250-320℃。最优选地，加氢反应的压力为3.5MPa，加氢反应的温度为300℃。
精制加氢反应的压力2-8MPa，所述精制加氢反应的温度为250-320℃。最优选地，精制加氢反应的压力为3.5MPa，精制加氢反应的温度为300℃。
加氢反应与精制加氢反应均采用的催化剂为Mo-Ni型催化剂，加氢反应器10与加氢精制反应器20内的催化剂用量体积比可以为(1:12)-(1:15)，较优 地为1:13。
从精制加氢反应器10出来的物质经过气液分离罐70的分离后，顶部气相去往脱硫罐进行脱硫处理，底部分离出的产品去往成品罐80储存。
以上各个工艺步骤循环往复，以使整个合成系统平稳的运行。
通过采用本发明的加氢反应工艺，脱硫的脱除率可以达到99.95％，较以往的加氢反应工艺，其脱除率提高了近1个百分点。
此外，通过铺设微界面发生器101降低了加氢反应釜的压力以及温度，充分降低了能耗。
总之，与现有技术的柴油加氢的反应系统相比，本发明的柴油加氢的反应系统设备组件少、占地面积小、能耗低、成本低、安全性高、反应可控，原料转化率高，相当于为柴油加氢领域提供了一种操作性更强的反应系统，值得广泛推广应用。
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，而非对其限制；尽管参照前述各实施例对本发明进行了详细的说明，本领域的普通技术人员应当理解：其依然可以对前述各实施例所记载的技术方案进行修改，或者对其中部分或者全部技术特征进行等同替换；而这些修改或者替换，并不使相应技术方案的本质脱离本发明各实施例技术方案的范围。

Claims (10)

  1. 一种柴油加氢的反应系统，其特征在于，包括：加氢反应器，所述加氢反应器内从上至下依次设置有多段固定床层，所述固定床层内装填有催化剂；相邻所述固定床层之间设置有微界面发生器，所述微界面发生器上设置有氢气进口；
    所述加氢反应器的侧壁设置有产物出口，所述产物出口连接加氢精制反应器进行深度加氢反应；
    所述加氢精制反应器底部出来的反应产物通入气液分离罐以将含硫化氢、氢气的气相与产品分离后，得到最终产品脱硫柴油。
  2. 根据权利要求1所述的反应系统，其特征在于，所述气液分离罐的顶部设置有气相出口，所述气相出口连接脱硫塔以实现对气相中硫化氢的吸收；
    所述脱硫塔内的顶部设置有脱硫剂喷洒口，所述脱硫塔内布设有多层填料以增加传质效果。
  3. 根据权利要求2所述的反应系统，其特征在于，所述脱硫塔旁侧设置有脱硫剂循环罐，所述脱硫剂循环罐与所述脱硫塔的侧壁下部通过管道连接，所述脱硫剂循环罐的顶部通过管道与所述脱硫剂喷洒口连接。
  4. 根据权利要求1所述的反应系统，其特征在于，所述固定床层为3段，所述微界面发生器为3个，其中两个微界面发生器设置两段固定床层之间，另外一个微界面发生器设置在所述加氢反应器内的底部。
  5. 根据权利要求1所述的反应系统，其特征在于，每个所述氢气进口汇集后连接氢气总管道，所述氢气总管道连接氢气储罐，所述脱硫塔顶部回收的氢气通过管道与所述氢气总管道汇合。
  6. 根据权利要求1所述的反应系统，其特征在于，所述加氢反应器与所述加氢精制反应器之间设置有加热器以实现从所述产物出口出来的产物进行加热。
  7. 根据权利要求1所述的反应系统，其特征在于，所述加氢精制反应器 底部出来的反应产物先经过换热器冷却后再进入所述气液分离罐。
  8. 根据权利要求7所述的反应系统，其特征在于，所述加氢反应器的底部设置有直馏柴油进口，所述直馏柴油进口通过管道与所述换热器连接以用于将直馏柴油加热后再进入所述加氢反应器。
  9. 采用权利要求1-8任一项所述的柴油加氢反应系统的反应方法，其特征在于，包括：
    将直馏柴油与氢气混合微界面分散破碎后，进行加氢反应，再进行精制加氢后，气液分离以实现将含硫化氢、氢气的气相与产品分离。
  10. 根据权利要求9所述的反应方法，其特征在于，所述加氢反应的压力2-8MPa，所述加氢反应的温度为250-320℃；
    优选地，所述精制加氢反应的压力2-8MPa，所述精制加氢反应的温度为250-320℃。
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