Method and system for hydrogenating hydrocarbon oil

Abstract

A method and a system for hydrogenating hydrocarbon oil. The method comprises the following steps: 1, in the presence of hydrogen, subjecting a hydrocarbon oil raw material to a gas phase hydrogenation reaction in a gas phase hydrogenation reactor (2), so as to obtain a gas phase hydrogenation reaction product; 2, subjecting the gas phase hydrogenation reaction product to gas-liquid separation to obtain a gas phase component and a liquid phase component; and 3, subjecting the liquid phase component to pressurization, and optionally, hydrogen mixing, and then subjecting same to a liquid phase hydrogenation reaction in a liquid phase hydrogenation reactor (9) to obtain a liquid phase hydrogenation reaction product. By means of the method and system for hydrogenating hydrocarbon oil, the effect of deep hydrogenation can be achieved, and the process flow is simplified as a whole, the reaction severity is reduced, and the consumption of energy and hydrogen is reduced.

Description

Hydrogenation method and hydrogenation system for hydrocarbon oil Technical Field

[0001] The present invention relates to the fields of oil refining and chemical industry, and in particular to a hydrocarbon oil hydrogenation method and a hydrogenation system. Background Art

[0002] Conventional fixed-bed hydrogenation technology is widely used and is currently applied in many fields such as diesel, jet fuel, naphtha, and reformate. However, with increasingly stringent environmental regulations and the introduction of technical requirements such as energy conservation and consumption reduction and hydrogen consumption reduction, existing fixed-bed hydrogenation technology has highlighted some shortcomings. For example, for diesel hydrogenation, from the perspective of the reaction mechanisms of deep desulfurization and deep dearomatization, there are significant differences in the requirements for the reaction environment between the two. For desulfurization reactions, the removal of small molecular sulfur mainly follows the direct desulfurization route, that is, hydrogenolysis. Large molecular sulfides with lower reaction activity follow the hydrodesulfurization reaction route, that is, first hydrogenation of the aromatic ring and then hydrogenolysis. In the upper part of the reactor, under a relatively low temperature and high hydrogen partial pressure environment, the main processes are hydrogenolysis of small molecular sulfides (endothermic), hydrogenation of nitrides, and conversion of some dicyclic aromatic hydrocarbons (largely exothermic). In the lower part of the reactor, the accumulation of heat and hydrogen sulfide creates a high temperature and low hydrogen partial pressure reaction environment. This environment, severely thermodynamically limited by aromatic hydrogenation, is unfavorable for both the hydrodesulfurization pathway of large sulfides and the saturation of polycyclic aromatic hydrocarbons. Especially towards the end of the reaction, catalyst activity declines. Raising the temperature can compensate for the loss of desulfurization activity, but this further compromises dearomatization. Besides the different reaction environment requirements, competitive adsorption of aromatics on the catalyst surface inhibits deep desulfurization, making it difficult for conventional hydrogenation technologies to achieve both ultra-deep desulfurization and efficient aromatic saturation. Furthermore, existing hydrogenation technologies also face issues with unnecessary hydrogenation of reactants and over-hydrogenation. While there are no restrictions on nitrogen compounds in diesel fuel, during deep desulfurization and dearomatization, nitrogen compounds are more readily adsorbed on the catalyst surface than sulfides and aromatics. Consequently, they not only significantly inhibit desulfurization and dearomatization reactions but also waste hydrogen resources due to their preferential hydrogenation. This problem is particularly pronounced when hydrogenating diesel fuel with high nitrogen content, such as the diesel fraction from waste tire pyrolysis oil, to produce automotive diesel. Furthermore, diesel fuel only limits the content of polycyclic aromatic hydrocarbons (PAHs), but during the hydrogenation reaction, some monocyclic aromatic hydrocarbons undergo hydrogenation saturation, further leading to hydrogen waste.

[0003] For existing technologies, deep dearomatization on top of deep desulfurization can be achieved by using existing catalyst systems and reducing the reaction space velocity, that is, reducing the processing volume or adding reactors. However, this is not economically feasible. Alternatively, a two-stage process technology can be used. After conventional hydrogenation, the generated oil is stripped to remove hydrogen sulfide before entering a precious metal hydrogenation reactor. This significantly increases catalyst costs and process complexity, and does not address the hydrogen waste caused by excessive hydrogenation. Alternatively, the feedstock can be divided into light and heavy components. The light component is shallowly hydrogenated to minimize saturation of monocyclic aromatics, while the heavy component is deeply hydrodesulfurized and dearomatized. However, shallow hydrogenation generally uses medium-pressure hydrogenation, which increases the energy consumption of the entire technology and has not yet been industrialized.

[0004] In terms of jet fuel hydrogenation, although there are no clear restrictions on nitrogen content in the quality standards for jet fuel, indicators such as jet fuel color stability and stability in use are significantly correlated with nitrogen content. Therefore, when processing high-nitrogen raw materials, a reaction pressure of 4 MPa or above is generally required to achieve effective removal of nitrides. However, increasing the reaction pressure not only increases the energy consumption of the device, but also causes excessive removal of sulfides, which deteriorates the lubricity of the refined jet fuel and requires the reverse addition of anti-wear agents, which increases factory costs. At the same time, under higher reaction pressures, thiophene sulfides containing aromatic rings follow both the hydrogenation reaction pathway and the hydrogenolysis reaction pathway, and the hydrogenation pathway has a higher hydrogen consumption. This results in high energy consumption and high hydrogen consumption when removing nitrides by increasing the device pressure.

[0005] Some companies are also increasing jet fuel yield by cutting the jet fuel fraction to a heavier level. This increases the final distillation point of the jet fuel and incorporates more diesel components into the jet fuel, effectively addressing the current surge in jet fuel demand and declining diesel demand. However, as the jet fuel fraction becomes heavier, some of the aromatics in the diesel are recycled into the jet fuel fraction. The increased aromatic content in the jet fuel can cause the jet fuel product's smoke point to fail. Existing jet fuel plants have a relatively low pressure design rating, typically 3-4 MPa. Aromatic hydrogenation in a fixed-bed reaction environment requires a higher reaction pressure, making it difficult to remove aromatics using pressurization methods in existing plants, and thus failing to significantly improve smoke point performance. Therefore, to produce high-smoke-point jet fuel products by cutting the jet fuel to a heavier level, the plant pressure must be increased, resulting in wasted energy.

[0006] However, judging from the desulfurization and dearomatization reaction mechanisms in jet fuel fractions, the two have significantly different requirements for the reaction environment. For desulfurization reactions, the removal of small-molecule mercaptans and sulfides primarily follows a direct desulfurization route, namely, hydrogenolysis, which does not require high reaction pressures. Aromatic hydrocarbons, on the other hand, are primarily concentrated in the heavy components of the cut, and pressure-dearomatization of these heavy components is sufficient. If high-hydrogenation-activity catalysts are used under high-pressure reaction conditions while desulfurization and dearomatization reactions occur simultaneously, the desulfurization depth will be increased, resulting in poor lubricity of the refined jet fuel. At the same time, some thiophene sulfides will be removed through the hydrogenation pathway under high pressure, increasing chemical hydrogen consumption. Furthermore, the hydrogen sulfide produced during the desulfurization process will affect hydrogenolysis and dearomatization, reducing the hydrogen partial pressure. Furthermore, the low concentration of aromatic hydrocarbons in the mixed components is detrimental to efficient dearomatization.

[0007] In the hydrogenation of reformate oil, unreasonable hydrogen consumption also occurs. In existing technologies, catalyst activity is high in the initial stages of the hydrogenation reactor. Furthermore, since some large olefins (C6 and C7) are difficult to remove, the reaction temperature or pressure must be increased. This results in losses in aromatics hydrogenation, particularly for key chemicals like benzene, toluene, and xylene. Saturation in hydrogenation can reduce profits for manufacturers.

[0008] CN108085058B discloses a method for deep dearomatization of hydrocarbon oils. This method uses relatively mild temperature and pressure conditions to pass the crude oil and hydrogen over a highly dispersed Pt-Pd / Al2O3 catalyst. The post-reaction gas phase is compressed and recycled, and the liquid phase is a low-aromatic product. However, the Pt-Pd / Al2O3 hydrogenation catalyst is primarily suitable for dearomatization of low-sulfur diesel feedstocks and is not ideal for desulfurization and dearomatization of low-quality diesel.

[0009] CN109926067A discloses a platinum-palladium-cobalt ternary metal hydrodearomatization catalyst and its preparation method. This method uses a step-by-step impregnation method with active metal precursors, resulting in a high utilization rate of the platinum and palladium precious metals in the prepared catalyst, strong synergistic catalytic ability, and enhanced aromatics hydrogenation performance of the non-precious metal cobalt due to the introduction of platinum and palladium. However, this catalyst is still designed for hydrocarbon oils after the sulfur compounds have been substantially removed, to avoid the effects of hydrogen sulfide on the deep dearomatization of the precious metals, and is unable to achieve desulfurization and dearomatization of low-quality diesel.

[0010] Summary of the Invention

[0011] In response to the deficiencies of the prior art, the present invention provides a method and system for hydrogenating hydrocarbon oil. The method and system provided by the present invention can achieve the effect of deep hydrogenation, while also simplifying the process flow as a whole, reducing the severity of the reaction, and reducing energy consumption and hydrogen consumption.

[0012] A first aspect of the present invention provides a method for hydrogenating hydrocarbon oil, the method comprising the following steps:

[0013] (1) in the presence of hydrogen, subjecting a hydrocarbon oil feedstock to a gas phase hydrogenation reaction in a gas phase hydrogenation reactor to obtain a gas phase hydrogenation reaction product;

[0014] (2) performing gas-liquid separation on the gas-phase hydrogenation reaction product to obtain a gas phase component and a liquid phase component;

[0015] (3) pressurizing the liquid phase components and optionally mixing them with hydrogen, and then performing a liquid phase hydrogenation reaction in a liquid phase hydrogenation reactor to obtain a liquid phase hydrogenation reaction product.

[0016] Preferably, the conditions of the gas phase hydrogenation reaction include: pressure of 0.1 to 2.8 MPa, preferably 0.3 to 2 MPa, temperature of 80 to 400°C, preferably 100 to 380°C, hydrogen to oil volume ratio of 5 to 500, preferably 5 to 300, volume space velocity of 0.5 to 15 h -1 , preferably 0.5 to 13 hours -1 .

[0017] Preferably, the conditions of the liquid phase hydrogenation reaction include: pressure of 1 to 10 MPa, preferably 1.5 to 7 MPa, temperature of 50 to 400°C, preferably 80 to 360°C, volume space velocity of 0.1 to 10 h -1 , preferably 0.1 to 9 hours -1 .

[0018] A second aspect of the present invention provides a hydrocarbon oil hydrogenation system, comprising a gas-phase hydrogenation reactor filled with a first hydrogenation catalyst, a gas-liquid separation unit, a pressurizing unit, and a liquid-phase hydrogenation reactor filled with a second hydrogenation catalyst, which are connected in sequence; the liquid phase component outlet of the gas-liquid separation unit is connected to the inlet of the pressurizing unit;

[0019] The gas phase hydrogenation reactor is used to carry out a gas phase hydrogenation reaction on the hydrocarbon oil feedstock and hydrogen to obtain a gas phase hydrogenation reaction product;

[0020] The pressurizing unit is used to pressurize the liquid phase component obtained by the gas-liquid separation unit to obtain pressurized logistics;

[0021] The liquid phase hydrogenation reactor is used to contact the pressurized logistics with the second hydrogenation catalyst to carry out a liquid phase hydrogenation reaction to obtain a liquid phase hydrogenation reaction product.

[0022] Preferably, the system further comprises a hydrogen mixing unit disposed between the pressurizing unit and the liquid phase hydrogenation reactor, wherein the hydrogen mixing unit is used to mix the pressurized logistics with hydrogen.

[0023] Compared with the prior art, the hydrogenation method and system of the present invention have the following advantages:

[0024] (1) The hydrogenation method and system of the present invention can only carry out low-hydrogen consumption hydrogenation reaction under low pressure and low hydrogen-to-oil ratio conditions in the gas-phase hydrogenation reactor to remove small molecular impurities, such as desulfurization and deolefination; the liquid phase component obtained by gas-liquid separation is pressurized and optionally mixed with hydrogen, and then subjected to liquid-phase hydrogenation reaction to remove large molecular impurities, such as removal of polycyclic aromatic hydrocarbons and removal of large molecular nitrogen compounds, thereby avoiding the defects of the prior art that the removal of small molecular impurities and large molecular impurities is placed in the same reaction system and the reaction conditions are difficult to be compatible. The hydrogenation method and system of the present invention can only carry out hydrogenation removal reaction of small molecular sulfides or small molecular olefins by controlling the reaction conditions of the gas-phase hydrogenation reactor, thereby realizing the separation of the removal reactions of small molecular sulfides, small molecular olefins and aromatic hydrocarbons, nitrides, and large molecular olefins, respectively optimizing the reaction conditions and improving the reaction efficiency of each.

[0025] (2) The present invention can achieve hydrorefining of hydrocarbon oils under milder operating conditions and a more simplified process flow. The entire reaction system does not require a hydrogen compressor as in a fixed-bed reaction system and a circulating oil pump as in a liquid-phase hydrogenation reaction system, thereby reducing investment costs, simplifying the process flow, improving reaction efficiency, and reducing reaction severity. BRIEF DESCRIPTION OF THE DRAWINGS

[0026] FIG1 is a schematic diagram of a method and system for hydrogenating hydrocarbon oil according to a specific embodiment of the present invention.

[0027] Description of Reference Numerals

[0028] In Figure 1: 1- Diesel feedstock and hydrogen 2- Gas phase hydrogenation reactor 3- Gas phase hydrogenation reaction product 4- Heat exchange unit 5- Gas-liquid separation unit 6- Liquid phase component 7- Gas phase component 8- Pressurization unit 9- Liquid phase hydrogenation reactor 10- Hydrogenated heavy component 11- Stripping unit 12- Refined diesel product DETAILED DESCRIPTION

[0029] The endpoints of the ranges and any values ​​disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0030] The description of the exemplary embodiments is intended to be read in conjunction with the accompanying drawings, which are to be considered a part of the entire written description. In the specification, relative terms such as "lower," "upper," "horizontal," "vertical," "above," "below," "upward," "downward," "top," and "bottom," and their derivatives (e.g., "horizontally," "downwardly," "upwardly," etc.) are to be interpreted as referring to the directions shown in the accompanying drawings being described at the time. These relative terms are for convenience of description and do not require that the device be constructed or operated in a particular orientation. Unless otherwise specified, "connected" as used herein refers to a relationship whereby structures are affixed or connected to each other, directly or indirectly, through intermediate structures.

[0031] A first aspect of the present invention provides a method for hydrogenating hydrocarbon oil, the method comprising the following steps:

[0032] (1) in the presence of hydrogen, subjecting a hydrocarbon oil feedstock to a gas phase hydrogenation reaction in a gas phase hydrogenation reactor to obtain a gas phase hydrogenation reaction product;

[0033] (2) performing gas-liquid separation on the gas-phase hydrogenation reaction product to obtain a gas phase component and a liquid phase component;

[0034] (3) pressurizing the liquid phase components and optionally mixing them with hydrogen, and then performing a liquid phase hydrogenation reaction in a liquid phase hydrogenation reactor to obtain a liquid phase hydrogenation reaction product.

[0035] The hydrogen in step (1) of the present invention can be any hydrogen-containing gas that can provide hydrogen, and can be fresh hydrogen, recycled hydrogen, or hydrogen-rich gas. Those skilled in the art will clearly understand the hydrogen-containing gas described in the present invention after understanding the technical solution of the present invention.

[0036] The present invention allows for a wide range of hydrocarbon oil feedstocks. The method provided by the present invention is applicable to any hydrocarbon oil requiring simultaneous removal of small molecule impurities and other macromolecular impurities, including but not limited to at least one of diesel, waste tire pyrolysis oil, jet fuel, and reformate. The method provided by the present invention primarily removes small molecule impurities (e.g., sulfides, small molecule nitrides, and small molecule olefins) in a vapor-phase hydrogenation reactor. The liquid phase component obtained by gas-liquid separation of the vapor-phase hydrogenation reaction product is pressurized and optionally mixed with hydrogen before entering a liquid-phase hydrogenation reactor for deep hydrorefining.

[0037] The present invention has a wide range of properties for hydrocarbon oil feedstock. Preferably, the diesel has an initial boiling point of 140-200°C, an end boiling point of 300-400°C, an S content of ≯20,000 μg / g, preferably ≯15,000 μg / g, a N content of ≯1,000 μg / g, preferably ≯800 μg / g, and a polycyclic aromatic hydrocarbon content of ≯50 wt%. For example, the S content of the diesel may be, but is not limited to, 8,000 μg / g, 9,000 μg / g, 10,000 μg / g, 11,500 μg / g, etc., the N content may be, but is not limited to, 500 μg / g, 600 μg / g, 700 μg / g, etc., and the polycyclic aromatic hydrocarbon content may be, but is not limited to, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt%, etc.

[0038] The present invention does not specifically limit the source of the diesel fuel, and includes, but is not limited to, one or more of straight-run diesel fuel, catalytic diesel fuel, coker diesel fuel, and ebullated-bed residue hydrotreated diesel fuel. The straight-run diesel fuel, catalytic diesel fuel, coker diesel fuel, and ebullated-bed residue hydrotreated diesel fuel described in the present invention have common definitions in the art and are not specifically limited in the present invention.

[0039] The method provided by the present invention, when the hydrocarbon oil feedstock is diesel, can achieve a deep dearomatization effect in addition to desulfurization. Specifically, a gas-phase desulfurization reaction is carried out in a gas-phase hydrogenation reactor, which primarily removes small molecular sulfides. The gas-phase desulfurization reaction product is subjected to gas-liquid separation to obtain a liquid-phase component, which, after pressurization and hydrogenation, enters a liquid-phase hydrogenation reactor for a polycyclic aromatic hydrocarbon hydrogenation saturation reaction.

[0040] In the present invention, the waste tire pyrolysis oil refers to the diesel fraction cut from the pyrolysis oil components obtained after the waste tires are pyrolyzed. The present invention does not specifically limit the specific operating methods and means of the pyrolysis process, and those skilled in the art can select the appropriate method based on actual needs.

[0041] During their research, the inventors discovered that while waste tire pyrolysis oil has a relatively high nitrogen content, its other properties are generally satisfactory for use as automotive diesel. Using the technical solution of the present invention to treat this oil, the effects of nitrogen compounds on desulfurization can be avoided by utilizing gas-phase conditions, while also minimizing hydrogen losses associated with hydrogenation of the nitrogen compounds.

[0042] In the present invention, a wide range of properties can be selected for the waste tire pyrolysis oil. Preferably, the waste tire pyrolysis oil raw material has an initial boiling point of 130-180°C, a final boiling point of 320-350°C, a sulfur content of ≤15,000 μg / g, preferably ≤11,000 μg / g, a nitrogen content of 5,000-10,000 μg / g, and a polycyclic aromatic hydrocarbon content of ≤50 wt%. For example, the S content of the waste tire pyrolysis oil can be, but is not limited to, 8000 μg / g, 9000 μg / g, 10000 μg / g, 11000 μg / g, etc.; the N content can be, but is not limited to, 5000 μg / g, 6000 μg / g, 7000 μg / g, 8000 μg / g, 9000 μg / g, 10000 μg / g, etc.; and the polycyclic aromatic hydrocarbon content can be, but is not limited to, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt%, etc.

[0043] According to the method provided by the present invention, when the hydrocarbon oil raw material is waste tire pyrolysis oil, the waste tire pyrolysis oil raw material undergoes a gas-phase desulfurization reaction in a gas-phase hydrogenation reactor. The gas-phase desulfurization reaction mainly removes small molecular sulfides and does not remove nitrogen compounds. The gas-phase desulfurization reaction product is subjected to gas-liquid separation to obtain a liquid phase component, which, after pressurization and hydrogen mixing, enters a liquid phase hydrogenation reactor to undergo a deep hydrogenation dearomatization reaction.

[0044] The aviation kerosene described herein can be any conventionally processed aviation kerosene in the art, for example, heavy-cut aviation kerosene. Preferably, the aviation kerosene has an initial boiling point of 80-200°C, a final boiling point of 220-320°C, and a sulfur content of 5000 μg / g or less, preferably 3000 μg / g or less. Preferably, the nitrogen content of the aviation kerosene is 100 μg / g or less, more preferably 60 μg / g or less. Preferably, the smoke point of the aviation kerosene is 18-25 mm, preferably 20-24 mm. For example, the sulfur content of the aviation kerosene can be, but is not limited to, 1000 μg / g, 2000 μg / g, 2500 μg / g, etc., the nitrogen content can be, but is not limited to, 10 μg / g, 20 μg / g, 30 μg / g, etc., and the smoke point can be, but is not limited to, 19 mm, 21 mm, 23 mm, etc.

[0045] The method provided by the present invention is not only applicable to the processing of high-nitrogen aviation kerosene raw materials, but also applicable to the production of high-smoke point aviation kerosene products. According to the method provided by the present invention, when the hydrocarbon oil raw material is a high-nitrogen aviation kerosene raw material (for example, a method with an N content of 11-30), the high-nitrogen aviation kerosene raw material undergoes only a low-hydrogen-consumption gas-phase hydrodesulfurization reaction under low-pressure conditions in a gas-phase hydrogenation reactor, and the liquid phase component obtained by gas-liquid separation of the gas-phase desulfurization reaction product enters a liquid-phase hydrogenation reactor after pressurization to undergo a nitride hydrogenation saturation reaction. According to the method provided by the present invention, when the hydrocarbon oil raw material is an aviation kerosene raw material, in order to obtain a high-smoke point aviation kerosene product, the aviation kerosene raw material undergoes only a low-hydrogen-consumption gas-phase hydrodesulfurization reaction under low-pressure conditions in a gas-phase hydrogenation reactor, and the liquid phase component obtained by gas-liquid separation of the gas-phase desulfurization reaction product enters a liquid-phase hydrogenation reactor after pressurization to undergo an aromatics hydrogenation saturation reaction.

[0046] The reformed oil of the present invention can be any reformed oil that needs to be processed conventionally in the art. The present invention has no particular limitation on this. Preferably, the distillation range of the reformed oil is 40 to 100° C., and the bromine index is ≯8000 μg / g. According to a preferred embodiment of the present invention, when the hydrocarbon oil is a reformed oil, the method further comprises dechlorinating and / or re-contacting the reformed oil, and then performing a gas phase hydrogenation reaction. The method of the present invention is used to treat the reformed oil, which can remove the C6 and C7 olefins that are the most difficult to remove in the reformed oil, and the aromatics loss rate is low. Specifically, a hydrogenation saturation reaction of small molecule olefins occurs in the gas phase hydrogenation reactor, which can remove C6 and below olefins, and then benzene is allowed to enter the gas phase without further hydrogenation saturation by gas-liquid separation. The liquid phase components (mainly including C7 olefins and toluene, ethylbenzene, xylene, etc.) undergo a selective hydrogenation deolefination reaction in the liquid phase hydrogenation reactor, further reducing aromatics losses.

[0047] In the present invention, the symbol “≯” indicates not greater than, and the symbol “≮” indicates not less than.

[0048] According to a preferred embodiment of the present invention, the gas-phase hydrogenation reaction is carried out in the presence of a first hydrogenation catalyst, and the first hydrogenation catalyst is a catalyst having a hydrodesulfurization function and / or a hydrodeolefination function.

[0049] The present invention provides a wide range of catalysts for hydrodesulfurization, which can be selected from conventional options in the art. For example, the hydrodesulfurization catalyst comprises a support and a hydrogenation-active metal, wherein the support is an inorganic refractory oxide, typically selected from one or more of alumina, amorphous silica-alumina, silica, and titanium oxide; and the hydrogenation-active metal comprises a Group VIB metal component and / or a Group VIII metal component. In the hydrodesulfurization catalyst, the Group VIB metal component is preferably selected from tungsten and / or molybdenum, with an oxide content of 5% to 30%, preferably 15% to 30%, in the catalyst. The Group VIII metal component is preferably selected from nickel and / or cobalt, with an oxide content of 1% to 6%, preferably 2% to 6%, in the catalyst. Preferably, the first hydrogenation catalyst is a Mo-Co hydrodesulfurization catalyst. The hydrodesulfurization catalyst may also contain an additive component, such as at least one of phosphorus, silicon, boron, magnesium, and fluorine, with the mass content of the additive in the hydrodesulfurization catalyst generally being less than 6% by weight.

[0050] The catalyst having the hydrodesulfurization function can be prepared by any method or can be a commercial catalyst, for example, it can be the FHUDS-5 catalyst and / or FHUDS-7 catalyst developed by Sinopec Fushun Research Institute of Petrochemical (FRIPP).

[0051] The catalyst having the function of hydrogenating and deolefination can be any catalyst capable of deolefination conventionally used in the art, preferably a precious metal catalyst. Specifically, the precious metal catalyst includes a carrier and a precious metal component supported on the carrier. The present invention does not particularly limit the type of carrier, for example, it can be a carrier conventionally defined in the art, such as alumina. There is no particular limitation on the type of precious metal in the precious metal catalyst, and it can be a precious metal conventionally defined in the art. Preferably, the precious metal component is Pt and / or Pd. In the present invention, there is no particular limitation on the source of the catalyst having the function of hydrogenating and deolefination, for example, it can be prepared by any method, or it can be obtained commercially, for example, it can be the HDO-18 catalyst developed by Sinopec Fushun Research Institute of Petrochemical Industry (FRIPP).

[0052] According to the method provided by the present invention, preferably, the conditions of the gas phase hydrogenation reaction include: a pressure of 0.1 to 2.8 MPa, preferably 0.3 to 2 MPa, more preferably 0.5 to 1.5 MPa, a temperature of 80 to 400°C, preferably 100 to 380°C, more preferably 150 to 380°C, a hydrogen to oil volume ratio of 5 to 500, preferably 5 to 300, more preferably 20 to 220, and a volume space velocity of 0.5 to 15 h -1 , preferably 0.5 to 13 hours -1In the method provided by the present invention, the gas-phase hydrogenation reaction can be carried out at a lower pressure and a lower hydrogen-to-oil volume ratio, thereby greatly saving energy consumption.

[0053] It should be noted that during the gas phase hydrogenation reaction in step (1), not only the sulfur and / or olefins in the hydrocarbon oil feedstock are removed, but also some small molecular nitrogen compounds (such as fatty amines) therein are removed.

[0054] Under the above disclosure, those skilled in the art can match catalysts of different functions in gas phase hydrogenation reactor according to specific hydrocarbon oil types. For example, when the hydrocarbon oil is diesel, jet fuel or waste tire pyrolysis oil, according to the method provided by the present invention, the catalyst with hydrodesulfurization function is loaded in gas phase hydrogenation reactor, by matching the use of high hydrogenation desulfurization path catalyst in gas phase hydrogenation reactor, controlling desulfurization reaction to follow the reaction path of low hydrogen consumption more, reducing the chemical hydrogen consumption of desulfurization reaction, while also reducing the hydrogen-to-oil ratio of gas phase hydrogenation reactor, reducing reaction severity. Simultaneously by controlling the reaction conditions of preferred gas phase hydrogenation reactor, not only for gas phase desulfurization, but more mainly in order to better cooperate with liquid phase hydrogenation reactor, so as to be more conducive to subsequent depth hydrorefining (such as depth dearomatization, deolefination, denitrogenation). And the preferred gas phase hydrogenation reaction conditions of the present invention avoid the hydrogenation saturation of macromolecular nitrides, reduce hydrogen consumption. Most large molecular nitrides have structures containing aromatic rings and nitrogen heterocycles, and the CN bond can only be broken after the aromatic rings and nitrogen heterocycles are completely hydrogenated and saturated. Under the reaction conditions of the gas-phase hydrogenation reactor, large molecular nitrides are difficult to undergo hydrogenation reactions, and a large amount of them enter the gas phase components through gas-liquid separation without deep hydrogenation, thereby reducing the hydrogen consumption of the chemical reaction.

[0055] And when described hydrocarbon oil is reformate oil, preferably in gas phase hydrogenation reactor, load the catalyst with hydrogenation deolefination function.The olefin that is most difficult to remove in reformate oil is C6, C7 olefin, need improve reaction pressure for removing this part olefin, cause aromatic hydrocarbon loss, the present invention is by arranging gas phase hydrogenation reactor (low pressure), with C6 and following olefin (accounting for 70 weight % of total olefin content) all saturated, then retain hydrogenation light component by gas-liquid separation, make benzene enter in gas phase and do not carry out further hydrogenation saturation.C7 and above olefin and toluene, ethylbenzene, dimethylbenzene etc. are as liquid phase components, and selective hydrogenation deolefination reaction takes place in liquid phase hydrogenation reactor, further reduce aromatic hydrocarbon loss.Under this kind of embodiment, although be low pressure reaction in gas phase hydrogenation reactor, by the setting of larger hydrogen-to-oil ratio, promote olefin hydrogenation reaction, and can retain benzene to greatest extent. After passing through the gas phase hydrogenation reactor, a large amount of hydrogen is consumed, and the hydrogen-dissolved liquid phase components are pressurized to undergo hydrogenation reaction of the heavy components. At this time, since the liquid phase components are in a slightly "hydrogen-poor" state, that is, a small amount of dissolved hydrogen is only used to meet the hydrogenation of C7 and above olefins, thereby avoiding the hydrogenation loss of toluene, ethylbenzene, and xylene.

[0056] Those skilled in the art can match different gas phase hydrogenation reaction conditions in the gas phase hydrogenation reactor according to the specific hydrocarbon oil type. For example, when the hydrocarbon oil is diesel or waste tire pyrolysis oil, the gas phase hydrogenation reaction conditions preferably include: pressure of 0.5-2 MPa, temperature of 330-380°C, hydrogen-to-oil volume ratio of 50-220, volume space velocity of 0.5-2 h -1 For example, when the hydrocarbon oil is jet fuel, the gas phase hydrogenation reaction conditions preferably include: a pressure of 0.5-2 MPa, a temperature of 300-350°C, a hydrogen-to-oil volume ratio of 20-100, and a volume space velocity of 4-13 h -1 For example, when the hydrocarbon oil is reformed oil, the gas phase hydrogenation reaction conditions preferably include: pressure of 0.5-2 MPa, temperature of 150-200°C, hydrogen-to-oil volume ratio of 20-50, and volume space velocity of 4-8 h -1 .

[0057] According to the method of the present invention, when step (1) involves a gas-phase hydrodesulfurization reaction, it is preferred to follow a low hydrogen consumption hydrogenation reaction path. In this preferred embodiment, the sulfur content in the gas-phase hydrogenation reaction product obtained in step (1) preferably accounts for at least 1wt% of the sulfur content in the hydrocarbon oil feedstock, more preferably 5-60wt%, such as 5wt%, 10wt%, 15wt%, 20wt%, 25wt%, 30wt%, 35wt%, 40wt%, 45wt%, 50wt%, 55wt%, 60wt% and values ​​between any two groups. Under low pressure conditions, the aromatic rings of macromolecular sulfides cannot be hydrogenated, and only small molecular sulfur is removed. Therefore, this preferred embodiment follows the hydrogenation desulfurization path more, reducing the hydrogen consumption of the chemical reaction. For macromolecular sulfides such as 4,6-DMDBT with steric hindrance, they can be deeply removed in the subsequent hydrogenation reaction.

[0058] The present invention has no particular limitation on the type of the gas-phase hydrogenation reactor. Preferably, the gas-phase hydrogenation reactor is a fixed-bed reactor.

[0059] In step (2) of the present invention, the gas-phase hydrogenation reaction product is subjected to gas-liquid separation. The method preferably further comprises subjecting the gas-phase hydrogenation reaction product to heat exchange and cooling before the gas-liquid separation is performed. The present invention has no particular limitation on the heat exchange conditions, and preferably can ensure that the macromolecular heavy substances in the gas-phase hydrogenation reaction product are liquefied. The present invention has no particular limitation on the specific operation of the heat exchange, and can be carried out using any heat exchange medium conventionally used in the art. From the perspective of energy conservation, it is preferred that the hydrocarbon oil feedstock of step (1) be subjected to heat exchange with the gas-phase hydrogenation reaction product. The heat exchange can be carried out in a heat exchange unit such as a heat exchanger. The specific type of the heat exchanger in the present invention is not particularly limited, for example, it can be a commercial heat exchanger conventionally defined in the art, or it can be a tubular heat exchanger.

[0060] The present invention uses a small amount of heat exchange under high temperature conditions for the gas phase desulfurization reaction product, generally cooling it by 8 to 50°C, preferably by 15 to 50°C, thereby retaining a large amount of reaction heat while ensuring the liquefaction of macromolecules, meeting the feeding conditions of the liquid phase hydrogenation reactor, and saving energy consumption.

[0061] The present invention has no particular limitation on the gas-liquid separation in step (2), which can be carried out in a gas-liquid separator. The gas-liquid separation is used to establish a gas-liquid equilibrium of the effluent after heat exchange, so that the gas phase component flows out upward and the liquid phase component flows out downward. The present invention has a wide range of conditions for selecting the gas-liquid separation. Those skilled in the art can make appropriate selections based on the composition of the hydrocarbon oil feedstock, so as to ensure that the components that need further hydrorefining are retained in the liquid phase component. Preferably, based on the total amount of the gas phase hydrogenation reaction product, the mass percentage of the liquid phase component is 1% to 80%, preferably 10% to 75%, for example, 10%, 15%, 20%, 25%, 30%, 35%, 40%, 45%, 50%, 55%, 60%, 65%, 70%, and values ​​between any two groups. Depending on the selection of different hydrocarbon oils, the preferred range of the mass percentage of the liquid phase component may vary.

[0062] Preferably, the hydrocarbon oil feedstock is diesel, and the mass percentage of the liquid phase component is 1 to 50%, more preferably 10 to 30%, based on the total amount of the gas phase hydrogenation reaction product.

[0063] Preferably, the hydrocarbon oil feedstock is aviation kerosene, and the mass percentage of the liquid phase component is 10-35% based on the total amount of the gas phase hydrogenation reaction product.

[0064] Preferably, the hydrocarbon oil raw material is waste tire pyrolysis oil, and the mass percentage of the liquid phase component is 20-40% based on the total amount of the gas phase hydrogenation reaction product.

[0065] Preferably, the hydrocarbon oil feedstock is reformed oil, and the mass percentage of the liquid phase component is 20-80% based on the total amount of the gas phase hydrogenation reaction product, more preferably 40-75%, and even more preferably 60-75%.

[0066] According to the present invention, in step (3), the specific operation and condition selection range for pressurizing the liquid phase component and optionally mixing hydrogen are relatively wide. The hydrogen mixing is used to supplement the hydrogen required for the reaction on the basis of pressurizing the liquid phase component, and preferably pressurizing and optionally mixing hydrogen to meet the reaction conditions and hydrogen consumption requirements in the liquid phase hydrogenation reactor. The method provided by the present invention can perform hydrogen mixing after pressurizing the liquid phase component, or can directly perform liquid phase hydrogenation reaction without hydrogen mixing. When liquid phase hydrogenation requires additional hydrogen, the method can mix hydrogen with the pressurized logistics and then send it into the liquid phase hydrogenation reactor. The present invention does not specifically limit the specific operation of pressurization and hydrogen mixing. The specific process can be that the liquid phase component is first pressurized by the pressurization unit, and then enters the hydrogen mixing unit, and the hydrogen is pressurized by the compressor to a pressure slightly higher than the liquid phase component (higher than the liquid phase component by 0.05 to 0.1 MPa), and also enters the hydrogen mixing unit. The material flowing out of the hydrogen mixing unit is a liquid phase component dissolved with hydrogen.

[0067] In the present invention, the pressurization and hydrogen mixing in step (3) can be performed in a pressurization unit and a hydrogen mixing unit, respectively. There is no particular limitation on the selection of the pressurization unit and the hydrogen mixing unit, as long as the hydrogen mixing and pressurization of the liquid phase components can be achieved. According to a preferred embodiment of the present invention, in step (3), the liquid phase components are pressurized and optionally hydrogenated, preferably to a pressure suitable for a liquid phase hydrogenation reaction. The range of the pressurization, whether hydrogen mixing is performed, and the amount of hydrogen mixing may vary depending on different hydrocarbon feedstocks and different product requirements.

[0068] Preferably, the hydrocarbon oil feedstock is diesel. In this preferred embodiment, the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst, which is a hydrodearomatization catalyst. In this preferred embodiment, the method includes pressurizing and hydrogenating the liquid phase component. Further preferably, during the hydrogenation process, relative to 1m 3 The liquid phase component, the hydrogen replenishment amount is 10~40Nm 3 .

[0069] According to another preferred embodiment of the present invention, the hydrocarbon oil feedstock is waste tire pyrolysis oil. In this preferred embodiment, the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst, which is a hydrodearomatization catalyst. In this preferred embodiment, the method includes pressurizing and hydrogenating the liquid phase component. Further preferably, during the hydrogenation process, relative to 1m 3 The liquid phase component, the hydrogen replenishment amount is 10~40Nm3 , preferably 20 to 30 Nm 3 .

[0070] According to another preferred embodiment of the present invention, the hydrocarbon oil feedstock is aviation kerosene. In this preferred embodiment, the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst, and the second hydrogenation catalyst is a hydrogenation dearomatization catalyst. When the hydrocarbon oil feedstock is aviation kerosene, the method of the present invention can effectively remove aromatics from aviation kerosene and improve the smoke point of aviation kerosene. In particular, when the aviation kerosene fraction is cut to increase the production of aviation kerosene and the aromatic content in the raw material is increased, a reaction environment suitable for the removal of aromatics can be more preferably selected to achieve the production of high smoke point aviation kerosene. In this preferred embodiment, the method includes pressurizing the liquid phase component. When the amount of hydrogen introduced at the inlet of the gas phase hydrogenation reactor in step (1) can meet the hydrogen demand of the liquid phase hydrogenation, hydrogen mixing may not be required after the pressurization in step (3).

[0071] According to another preferred embodiment of the present invention, the hydrocarbon oil feedstock is high-nitrogen aviation kerosene. In this preferred embodiment, the liquid-phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst, and the second hydrogenation catalyst is a hydrodenitrogenation catalyst. When the hydrocarbon oil feedstock is high-nitrogen aviation kerosene, the method provided by the present invention can achieve desulfurization and denitrification using high-nitrogen aviation kerosene as a raw material under more moderate operating conditions and a more simplified process flow. In the gas-phase hydrogenation reactor, only low-hydrogen consumption desulfurization reaction occurs under low-pressure conditions, reducing the hydrogen consumption of the reaction. The liquid phase component separated by gas and liquid undergoes a nitride hydrogenation saturation reaction after pressurization. In this preferred embodiment, the method includes pressurizing the liquid phase component. When the amount of hydrogen introduced at the inlet of the gas-phase hydrogenation reactor in step (1) can meet the hydrogen demand of the liquid phase hydrogenation, no hydrogen mixing is required after the pressurization in step (3).

[0072] According to another preferred embodiment of the present invention, the hydrocarbon oil feedstock is reformate. In this preferred embodiment, the liquid-phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst, which is a hydrodeolefination catalyst. In this preferred embodiment, the method includes pressurizing the liquid-phase component. When the amount of hydrogen introduced into the inlet of the gas-phase hydrogenation reactor in step (1) can meet the hydrogen demand of the liquid-phase hydrogenation, hydrogen mixing after the pressurization in step (3) may not be required.

[0073] The present invention has a wide range of selection for the type of the above-mentioned second hydrogenation catalyst, as long as it can achieve the above-mentioned purpose, for example, it can be a hydrodearomatization catalyst with a hydrodearomatization function, a hydrodenitrogenation catalyst with a hydrodenitrification function, or a hydrodeolefination catalyst with a hydrodeolefination function.

[0074] According to a preferred embodiment of the present invention, the hydrodearomatization catalyst can be a non-precious metal catalyst or a precious metal catalyst. The non-precious metal catalyst can include a support and a hydrogenation-active metal, wherein the support is an inorganic refractory oxide, generally selected from one or more of alumina, amorphous silica-alumina, silica, and titania, preferably alumina; and the hydrogenation-active metal includes a Group VIB metal component and / or a Group VIII metal component. In the hydrodearomatization catalyst, the Group VIB metal component is preferably selected from tungsten and / or molybdenum, and its content in the catalyst is 5% to 30% by weight, preferably 15% to 30% by weight of the oxide. The Group VIII metal component is preferably selected from nickel and / or cobalt, and its content in the catalyst is 1% to 6% by weight, preferably 2% to 5% by weight of the oxide. The hydrodearomatization catalyst may also contain an additive component, such as at least one of phosphorus, silicon, boron, magnesium, and fluorine, with its content in the catalyst having hydrodenitrogenation function generally being less than 6% by weight. The hydrodearomatization catalyst is preferably a Mo-Ni catalyst.

[0075] The hydrodearomatization catalyst may be prepared by any method or may be a commercial catalyst, for example, at least one of FHUDS-6, FHUDS-8 and FHUDS-10 developed by Fushun Research Institute of Petrochemical Industry (FRIPP) of Sinopec.

[0076] The noble metal catalyst preferably uses Pt, Pd or the like as an active metal, which can also be prepared by any method or a commercial catalyst, such as the FHDA-10 catalyst developed by Sinopec Fushun Research Institute of Petrochemicals (FRIPP).

[0077] According to a preferred embodiment of the present invention, the hydrodenitrogenation catalyst can be any catalyst capable of achieving hydrodenitrogenation, preferably a non-precious metal catalyst or a precious metal catalyst. The non-precious metal catalyst can include a support and a hydrogenation-active metal, wherein the support is an inorganic refractory oxide, generally selected from one or more of alumina, amorphous silica-alumina, silica, and titanium oxide, preferably alumina; and the hydrogenation-active metal includes a Group VIB metal component and / or a Group VIII metal component. In the hydrodearomatization catalyst, the Group VIB metal component is preferably selected from tungsten and / or molybdenum, and its content in the catalyst is 10% to 20% by weight of the oxide. The Group VIII metal component is preferably selected from nickel and / or cobalt, and its content in the catalyst is 1% to 9% by weight of the oxide. The hydrodenitrogenation catalyst may also contain an auxiliary component, such as at least one of phosphorus, silicon, boron, magnesium, and fluorine, whose mass content in the catalyst having the hydrodenitrogenation function is generally less than 6wt%. The hydrodenitrogenation catalyst can be prepared by any method or can be a commercially available catalyst, for example, the FH-40 series catalysts developed by Sinopec Fushun Research Institute of Petrochemical Industry (FRIPP), such as FH-40A and FH-40B. The selection range of the precious metal catalyst can be the same as that of the hydrodearomatization catalyst described above.

[0078] According to a preferred embodiment of the present invention, the hydrodeolefination catalyst can be any catalyst capable of removing olefins. The catalyst is preferably a precious metal catalyst comprising Pt and Pd as metal components and supported by AlO. The catalyst can also be prepared by any method or a commercially available catalyst, such as the HDO-18 catalyst developed by Sinopec Fushun Research Institute of Petrochemicals (FRIPP).

[0079] It should be noted that the catalyst with a certain function described in the present invention does not mean that it can only play this function, but mainly plays this function. For example, the hydrodesulfurization catalyst does not mean that it can only play a hydrodesulfurization role, but mainly plays a hydrodesulfurization role in its application environment.

[0080] According to a preferred embodiment of the present invention, the pressure of the liquid-phase hydrogenation reaction is at least 0.5 MPa higher than the pressure of the gas-phase hydrodesulfurization, preferably 0.5 to 6 MPa higher, for example, 0.5 MPa, 1 MPa, 1.5 MPa, 2 MPa, 2.5 MPa, 3 MPa, 3.5 MPa, 4 MPa, 4.5 MPa, 5 MPa, 5.5 MPa, 6 MPa, and values ​​between any two groups. The method provided by the present invention can be carried out at medium to low pressures, greatly reducing the reaction severity and saving energy consumption.

[0081] According to a preferred embodiment of the present invention, the hydrocarbon oil is diesel or waste tire pyrolysis oil, and the pressure of the liquid phase hydrogenation reaction is at least 1 MPa higher than the pressure of the gas phase hydrogenation reaction, preferably 2 to 6 MPa higher.

[0082] According to a preferred embodiment of the present invention, the hydrocarbon oil is jet fuel, and the pressure of the liquid phase hydrogenation reaction is at least 1.5 MPa higher than the pressure of the gas phase hydrogenation reaction, preferably 1.5 to 5 MPa higher.

[0083] According to a preferred embodiment of the present invention, the hydrocarbon oil is reformed oil, and the pressure of the liquid-phase hydrogenation reaction is at least 0.5 MPa higher than the pressure of the gas-phase hydrogenation reaction, preferably 0.5-2 MPa higher.

[0084] According to a preferred embodiment of the present invention, the temperature of the liquid-phase hydrogenation reaction is at least 5°C lower than that of the gas-phase hydrogenation reaction, preferably 10-60°C lower. This preferred embodiment is more conducive to liquefying macromolecules that require further reaction while retaining reaction heat.

[0085] According to a preferred embodiment of the present invention, the hydrocarbon oil is diesel or waste tire pyrolysis oil, and the temperature of the liquid phase hydrogenation reaction is at least 15°C lower than the temperature of the gas phase hydrogenation reaction, preferably 15-40°C lower.

[0086] According to a preferred embodiment of the present invention, the hydrocarbon oil is jet fuel, and the temperature of the liquid phase hydrogenation reaction is at least 10°C lower than the temperature of the gas phase hydrogenation reaction, preferably 10-50°C lower.

[0087] According to a preferred embodiment of the present invention, the hydrocarbon oil is reformed oil, and the temperature of the liquid phase hydrogenation reaction is at least 10° C. lower than that of the gas phase hydrogenation reaction, preferably 30-50° C. lower.

[0088] According to the method provided by the present invention, preferably, the conditions of the liquid phase hydrogenation reaction include: pressure of 1 to 10 MPa, preferably 1.5 to 7 MPa, temperature of 50 to 400°C, preferably 80 to 360°C, volume space velocity of 0.1 to 10 h -1 , preferably 0.1 to 9 hours -1 .

[0089] Those skilled in the art can match different gas phase hydrogenation reaction conditions in the liquid phase hydrogenation reactor according to the specific hydrocarbon oil type. For example, when the hydrocarbon oil is diesel or waste tire pyrolysis oil, the liquid phase hydrogenation reaction conditions preferably include: pressure of 2-10 MPa, preferably 3-7 MPa, temperature of 200-400°C, preferably 260-360°C, volume space velocity of 0.1-5 h -1 , preferably 0.1 to 1h-1 For example, when the hydrocarbon oil is jet fuel, the liquid phase hydrogenation reaction conditions preferably include: a pressure of 1 to 6 MPa, preferably 2 to 5 MPa, a temperature of 150 to 300°C, preferably 220 to 300°C, and a volume space velocity of 0.1 to 4 h -1 , preferably 0.5 to 1.5 hours -1 For example, when the hydrocarbon oil is reformed oil, the liquid phase hydrogenation reaction conditions preferably include: pressure of 3-6 MPa, temperature of 260-345°C, volume space velocity of 0.2-1 h -1 .

[0090] The present invention has no particular limitation on the reactor type of the liquid phase hydrogenation reactor. Preferably, the liquid phase hydrogenation reactor is a fixed bed reactor.

[0091] According to a specific embodiment of the present invention, the liquid-phase hydrogenation reaction product is discharged downward from the bottom of the liquid-phase hydrogenation reactor.

[0092] The present invention has no particular limitation on the application of the gas phase component (also referred to as hydrogenation light component) and the liquid phase hydrogenation reaction product obtained by gas-liquid separation, and can be adaptively treated according to the specific needs of the product. Preferably, the method further comprises subjecting the gas phase component and the liquid phase hydrogenation reaction product to dehydrogensulfurization treatment. The gas phase component and the liquid phase hydrogenation reaction product can be treated separately or together, preferably together. According to a preferred embodiment of the present invention, the method further comprises mixing the gas phase component and the liquid phase hydrogenation reaction product, and then subjecting them to dehydrogensulfurization treatment. The present invention has no particular limitation on the specific operation and conditions of the dehydrogensulfurization treatment, and can be carried out according to conventional technical means in the art, preferably by steam stripping. The steam stripping can be carried out in a stripping unit such as a stripping tower.

[0093] A second aspect of the present invention provides a hydrocarbon oil hydrogenation system, comprising a gas-phase hydrogenation reactor filled with a first hydrogenation catalyst, a gas-liquid separation unit, a pressurizing unit, and a liquid-phase hydrogenation reactor filled with a second hydrogenation catalyst, which are connected in sequence; the liquid phase component outlet of the gas-liquid separation unit is connected to the inlet of the pressurizing unit;

[0094] The gas phase hydrogenation reactor is used to carry out a gas phase hydrogenation reaction on the hydrocarbon oil feedstock and hydrogen to obtain a gas phase hydrogenation reaction product;

[0095] The pressurizing unit is used to pressurize the liquid phase component obtained by the gas-liquid separation unit to obtain pressurized logistics;

[0096] The liquid phase hydrogenation reactor is used to contact the pressurized logistics with the second hydrogenation catalyst to carry out a liquid phase hydrogenation reaction to obtain a liquid phase hydrogenation reaction product.

[0097] Preferably, the system further comprises a hydrogen mixing unit disposed between the pressurizing unit and the liquid phase hydrogenation reactor, wherein the hydrogen mixing unit is used to mix the pressurized logistics with hydrogen.

[0098] In the system provided by the present invention, the types of the first hydrogenation catalyst and the second hydrogenation catalyst can be matched according to the type of hydrocarbon oil feedstock. The specific criteria can be selected based on the content described in the first aspect and will not be repeated here.

[0099] Preferably, the hydrocarbon oil feedstock is diesel or waste tire pyrolysis oil, the first hydrogenation catalyst is a hydrodesulfurization catalyst, and the second hydrogenation catalyst is a hydrodearomatization catalyst; or

[0100] The hydrocarbon oil feedstock is heavy jet fuel, the first hydrogenation catalyst is a hydrodesulfurization catalyst, and the second hydrogenation catalyst is a hydrodearomatization catalyst; or

[0101] The hydrocarbon oil feedstock is high-nitrogen jet fuel, the first hydrogenation catalyst is a hydrodesulfurization catalyst, and the second hydrogenation catalyst is a hydrodenitrogenation catalyst; or

[0102] The hydrocarbon oil feedstock is reformed oil, the first hydrogenation catalyst is a hydrodeolefination catalyst, and the second hydrogenation catalyst is a hydrodeolefination catalyst.

[0103] The specific selection of the first hydrogenation catalyst and the second hydrogenation catalyst can be the same as described in the first aspect, and will not be repeated here.

[0104] According to the system provided by the present invention, preferably, the gas phase hydrogenation reactor is a fixed bed reactor.

[0105] According to a specific embodiment of the present invention, the gas-liquid separation unit is provided with a gas phase component outlet at the top and a liquid phase component outlet at the bottom.

[0106] According to a specific embodiment of the present invention, a liquid-phase hydrogenation reaction product outlet is provided at the bottom of the liquid-phase hydrogenation reactor.

[0107] According to the system provided by the present invention, preferably, the liquid phase hydrogenation reactor is a fixed bed reactor.

[0108] According to a preferred embodiment of the present invention, the system further comprises a stripping unit, the inlet of which is connected to the outlet of the liquid-phase hydrogenation reaction product of the liquid-phase hydrogenation reactor and / or the outlet of the gas-phase component of the gas-liquid separation unit. The stripping unit is used to remove hydrogen sulfide from the liquid-phase hydrogenation reaction product and / or the gas-phase component obtained by the gas-liquid separation unit.

[0109] Preferably, the liquid phase hydrogenation reaction product outlet of the liquid phase hydrogenation reactor is connected to the gas phase component outlet of the gas-liquid separation unit for mixing the liquid phase hydrogenation reaction product and the gas phase component obtained by the gas-liquid separation unit and then performing hydrogen sulfide removal treatment.

[0110] According to a preferred embodiment of the present invention, the system further comprises a heat exchange unit disposed between the gas phase hydrogenation reactor and the gas-liquid separation unit, wherein the heat exchange unit is used to heat and cool the gas phase hydrogenation reaction product before it enters the gas-liquid separation unit.

[0111] The present invention has no particular limitation on the heat exchange unit, and it can be a conventional choice in the art, such as a heat exchanger.

[0112] The present invention has no particular limitation on the gas-liquid separation unit, which may be a conventional choice in the art, for example, a gas-liquid separator.

[0113] Preferably, the pressurizing unit is used to increase the pressure of the liquid phase component, preferably to a pressure that can meet the reaction condition requirements in the liquid phase hydrogenation reactor.

[0114] Preferably, the hydrogen mixing unit is used to supplement the hydrogen required for the reaction on the basis of pressurizing the liquid phase components, and the hydrogen mixed in is preferably sufficient to meet the hydrogen consumption requirement of the reaction in the liquid phase hydrogenation reactor.

[0115] The system provided by the present invention has no particular limitation on the selection of the pressurizing unit and the hydrogen mixing unit, as long as pressurization and hydrogen mixing can be achieved.

[0116] Preferably, the boosting unit is a gas phase booster.

[0117] Preferably, the hydrogen mixing unit is selected from at least one of a membrane-tube hydrogen dissolving assembly, a hydrogen mixer (e.g., a high-efficiency hydrogen mixer), a microbubble generator, and a bubble fractalizer. For example, a hydrogen mixer with strong convection backmixing or a hydrogen mixer with micropores generating microbubbles is used to increase solubility. A support plate with evenly distributed pores and downcomers is preferably used to improve gas phase dispersion.

[0118] The scheme and effects of the present invention are further illustrated by the following examples.

[0119] In the present invention, unless otherwise explicitly stated, percentages and contents are all based on mass.

[0120] The method and system of the present invention are described in detail below with reference to FIG1 . An exemplary explanation is given by taking diesel as an example of a hydrocarbon oil raw material. The diesel raw material and hydrogen 1 enter the gas-phase hydrogenation reactor 2 (also referred to as the first hydrogenation reactor), where a gas-phase hydrogenation reaction occurs to obtain a gas-phase hydrogenation reaction product 3; the gas enters the heat exchange unit 4 (heat exchanger), and after cooling by the heat exchanger, enters the gas-liquid separation unit 5 (gas-liquid separator) to obtain a gas phase component 7 and a liquid phase component 6. The gas phase component 7 is discharged upward from the gas-liquid separator, and the liquid phase component 6 is discharged downward from the gas-liquid separator, and then passes through the boosting unit 8 and the hydrogen mixing unit (hydrogen mixer, not shown) to enter the liquid phase hydrogenation reactor 9 for liquid phase hydrogenation reaction to obtain a hydrogenated heavy component 10. The gas phase component 7 is mixed with the hydrogenated heavy component 10 and enters the stripping unit 11 (stripping tower) to finally obtain a refined diesel product 12.

[0121] In the following examples and comparative examples, chemical hydrogen consumption refers to the mass percentage of hydrogen consumed per unit mass of raw material.

[0122] The calculation formula for chemical hydrogen consumption is:

[0123] The wear scar diameter of aviation kerosene products is measured according to the petrochemical standard SH / T 0687-2017.

[0124] Examples A1-A3

[0125] The process was carried out using the flow diagram shown in Figure 1. Two 100mL fixed-bed hydrogenation reactors were connected in series, namely a gas-phase hydrogenation reactor and a liquid-phase hydrogenation reactor. A heat exchanger, a gas-liquid separator, a booster pump, and a hydrogen mixer were set between the reactors. The gas-phase hydrogenation reactor was loaded with 50mL of Mo-Co type diesel hydrogenation catalyst A, and the liquid-phase hydrogenation reactor was loaded with 50mL of Mo-Ni type diesel hydrogenation catalyst B. A gas-phase component outlet was set above the gas-liquid separator, which was connected to the bottom effluent of the liquid-phase hydrogenation reactor and then entered the stripping tower together to obtain a refined diesel product.

[0126] A mixture of straight wood, coke, and catalytic wood was used as the raw material. Catalyst properties are shown in Table A1, raw oil properties are shown in Table A2, and reaction process conditions and results are shown in Table A3.

[0127] Example A4

[0128] The same process flow as in Examples A1-A3 was used, except that the hydrogen-to-oil volume ratio at the gas phase reactor inlet was increased. Catalyst properties are shown in Table A1, feed oil properties are shown in Table A2, and reaction conditions and results are shown in Table A3.

[0129] Example A5

[0130] The same process flow as in Examples A1-A3 was used, except that the separation conditions for gas-liquid separation were changed. Catalyst properties are shown in Table A1, feed oil properties are shown in Table A2, and reaction process conditions and results are shown in Table A3.

[0131] Example A6

[0132] The hydrogenation process flow was the same as in Examples A1-A3, differing only in that the hydrogen-to-oil volume ratio at the inlet of the gas-phase hydrogenation reactor was increased, and the hydrogen mixer before the liquid-phase hydrogenation reactor was eliminated in favor of a liquid-phase booster pump. The raw materials and catalyst properties were the same as in Examples A1-A3. The reaction conditions and results are shown in Table A3.

[0133] Comparative Example A1

[0134] A conventional fixed-bed diesel hydrogenation process was employed, with two hydrogenation reactors connected in series: the first and second reactors. A stripping tower was placed between the reactors to remove hydrogen sulfide. The first reactor was loaded with 50 mL of Mo-Co diesel hydrogenation catalyst A, and the second reactor was loaded with 50 mL of Mo-Ni diesel hydrogenation catalyst B. Following the reactors, conventional processes such as high-pressure separation, low-pressure separation, and stripping were performed to produce the diesel product. After hydrogen sulfide removal, the hydrogen was pressurized and recycled using a circulating hydrogen compressor. The raw materials and catalyst properties were the same as in Examples A1-A3. The reaction conditions and results are shown in Table A3.

[0135] Table A1 Catalyst Physicochemical Properties

[0136] Table A2 Raw oil properties

[0137] Table A3 Hydrogenation process conditions and results

[0138] As can be seen from Table A3, in Example A4, due to the excessively large hydrogen-to-oil volume ratio at the inlet of the gas phase reactor, the oil partial pressure was low and liquefaction was difficult during the heat exchange process, the liquefaction ratio (liquid phase component content) was reduced, and a portion of the polycyclic aromatic hydrocarbons and macromolecular sulfides entered the gas phase components and could not undergo further hydrogenation reaction, resulting in a poor saturated desulfurization effect of polycyclic aromatic hydrocarbons.

[0139] In Example A5, the liquefaction ratio (liquid phase content) after heat exchange is high, and the liquid phase temperature does not meet the feed reaction temperature, requiring reheating, which increases energy consumption. In addition, unnecessary nitrogen compounds and monocyclic aromatic hydrocarbons enter the liquid phase, increasing hydrogen consumption in the chemical reaction.

[0140] In Example A6, no hydrogen mixer was placed before the liquid-phase hydrogenation reactor. A sufficient hydrogen source was provided by increasing the hydrogen-to-oil volume ratio at the gas-phase hydrogenation reactor inlet. However, this high hydrogen-to-oil volume ratio at the gas-phase hydrogenation reactor inlet resulted in a low oil-phase partial pressure within the separator, resulting in insufficient liquefaction. This meant that it was difficult for all PAHs to enter the liquid-phase components for the liquid-phase hydrogenation reaction, resulting in a high PAH content in the refined diesel product.

[0141] Comparative Example A1 uses conventional fixed-bed hydrogenation technology and a traditional catalyst loading system. The pressure of the first hydrogenation reactor is relatively high, resulting in simultaneous desulfurization, denitrogenation, and dearomatization, which affect each other. The reaction conditions of the two reactors cannot be optimized for a certain reaction, resulting in poor reaction effect and high hydrogen consumption. Hydrogen is recycled and energy consumption is high. However, under the premise that the pressure of the first hydrogenation reactor is significantly reduced, the reaction system of the present invention divides desulfurization and dearomatization and denitrogenation into two reaction systems, which can optimize the reaction conditions in a targeted manner, obtain a better dearomatization effect, and better retain the nitrides that do not need hydrogenation, avoiding the restriction of nitrides on the hydrogenation reaction and improving the reaction efficiency; sulfides are removed through the hydrogenolysis path, and nitrides are hydrogenated in small amounts, reducing hydrogen consumption; hydrogen does not need to be recycled, reducing energy consumption.

[0142] Examples B1-B3

[0143] The process is carried out using the flow diagram shown in Figure 1. Two 100mL fixed-bed hydrogenation reactors are connected in series, namely a gas-phase hydrogenation reactor (gas-phase hydrodesulfurization reactor) and a liquid-phase hydrogenation reactor (liquid-phase hydrodearomatization reactor). A heat exchanger, a gas-liquid separator and a booster pump are set between the reactors. The gas-phase hydrogenation reactor is filled with 30mL Mo-Co type hydrogenation catalyst A, and the liquid-phase hydrogenation reactor is filled with 70mL Mo-Ni type hydrogenation catalyst B. A gas-phase component outlet is set above the gas-liquid separator. After being connected to the bottom effluent of the liquid-phase hydrogenation reactor, the gas-phase component and the liquid-phase dearomatization product enter the stripping tower together to obtain a refined jet fuel product.

[0144] Heavy jet fuel was used as the raw material. Catalyst properties are shown in Table B1, feedstock oil properties are shown in Table B2, and reaction process conditions and results are shown in Table B3.

[0145] Example B4

[0146] The same process flow as in Examples B1-B3 was used, except that the hydrogen-to-oil volume ratio at the inlet of the gas-phase hydrogenation reactor was increased. Catalyst properties are shown in Table B1, feed oil properties are shown in Table B2, and reaction process conditions and results are shown in Table B3.

[0147] Example B5

[0148] The same process flow as in Examples B1-B3 was used, except that the separation conditions for gas-liquid separation were changed. Catalyst properties are shown in Table B1, feed oil properties are shown in Table B2, and reaction process conditions and results are shown in Table B3.

[0149] Example B6

[0150] The same process as in Example B1-B3 was used, except that the hydrogen-to-oil volume ratio at the inlet of the gas phase hydrogenation reactor was reduced, and the pressurization system of the hydrodearomatization reactor was changed to a pressurization and hydrogen mixing system. 3 The liquid phase component, the hydrogen replenishment amount is 20Nm 3 The gas-phase hydrogenation reactor was loaded with 30 mL of Mo-Co hydrogenation catalyst A, and the liquid-phase hydrogenation reactor was loaded with 70 mL of Mo-Ni hydrogenation catalyst B. Catalyst properties are shown in Table B1, feedstock properties are shown in Table B2, and reaction process conditions and results are shown in Table B3.

[0151] Comparative Example B1

[0152] A conventional fixed-bed jet fuel hydrogenation process was used, with a single hydrogenation reactor installed. The reactor was loaded with 100 mL of Mo-Co light distillate hydrogenation catalyst A. Conventional processes such as low-pressure fractionation and stripping were set up after the reactor to produce the jet fuel product. After hydrogen sulfide removal, the hydrogen was compressed and recycled using a circulating hydrogen compressor. The raw materials and catalyst properties were the same as in Examples B1-B3. The reaction conditions and results are shown in Table B3.

[0153] Table B1 Catalyst physicochemical properties

[0154] Table B2 Raw oil properties

[0155] Table B3 Hydrogenation process conditions and results

[0156] As can be seen from Table B3, the hydrogen-to-oil volume ratio at the inlet of the gas-phase hydrogenation reactor in Example B4 is too large, resulting in difficulty in liquefaction during the gas-liquid separation process. A portion of the aromatic hydrocarbons enter the gas phase component without hydrogenation reaction, and the product smoke point is low. The heat exchange during the gas-liquid separation process of Example B5 is large, and the liquefaction ratio is high. However, reheating is required to meet the reaction conditions during the liquid phase reaction, resulting in high energy consumption. Example B6 reduces the hydrogen-to-oil volume ratio due to the low hydrogen consumption of the gas-phase hydrogenation reactor. However, due to the reduction in hydrogen partial pressure, the difficulty of gasifying the raw material becomes greater, requiring a higher reaction temperature, resulting in excessive removal of sulfides and increased energy and hydrogen consumption.

[0157] Comparative Example 1 uses conventional fixed-bed hydrogenation technology and a traditional catalyst loading system. The reactor pressure is high, resulting in simultaneous desulfurization and dearomatization, which affect each other. The reaction conditions cannot be optimized for a specific reaction. Sulfide removal is deep under high pressure, affecting product lubricity and increasing hydrogen consumption. The hydrogen is recycled, resulting in high energy consumption. However, the reaction system of the present invention significantly reduces the pressure in the hydrodesulfurization reactor. By separating desulfurization and dearomatization into two reaction systems, the reaction conditions are optimized in a targeted manner. This allows for better retention of jet fuel sulfides, improved lubricity, and a better dearomatization effect. Sulfides are removed via hydrogenolysis, reducing hydrogen consumption. The hydrogen does not need to be recycled, reducing energy consumption.

[0158] Examples C1-C3

[0159] The process is carried out using the flow diagram shown in Figure 1. Two 100mL fixed-bed hydrogenation reactors are connected in series, namely a gas-phase hydrogenation reactor and a liquid-phase hydrogenation reactor. A heat exchanger, a gas-liquid separator and a booster pump are set between the reactors. The gas-phase hydrogenation reactor is filled with 20mL Mo-Co type hydrogenation catalyst A, and the liquid-phase hydrogenation reactor is filled with 80mL Mo-Ni type hydrogenation catalyst B. A gas phase component outlet is set above the gas-liquid separator, which is connected to the bottom effluent of the liquid-phase hydrogenation reactor (liquid-phase hydrogenation denitrogenation reaction product) and then enters the stripping tower together to obtain a refined jet fuel product. High-nitrogen jet fuel is used as raw material. The catalyst properties are shown in Table B1, the crude oil properties are shown in Table C2, and the reaction process conditions and results are shown in Table C3.

[0160] Example C4

[0161] The same process flow as in Example 1-3 was used, except that the hydrogen-to-oil volume ratio at the gas phase reactor inlet was increased. The reaction process conditions and results are shown in Table C3.

[0162] Comparative Example C1

[0163] A conventional fixed-bed jet fuel hydrogenation process was used, with a hydrogenation reactor installed. The reactor was loaded with 100 mL of Mo-Co hydrogenation catalyst A. Conventional processes such as low-pressure separation and stripping were followed after the reactor to produce the jet fuel product. After hydrogen sulfide removal, the hydrogen was compressed and recycled using a circulating hydrogen compressor. The raw materials and catalyst properties were the same as in Examples C1-C3. The reaction conditions and results are shown in Table C3.

[0164] Table C2 Raw oil properties

[0165] Table C3 Hydrogenation Process Conditions and Results

[0166] As shown in Table C3, the excessively high hydrogen-to-oil volume ratio at the inlet of the first hydrogenation reactor in Example C4 resulted in difficulty in liquefaction during the gas-liquid separation process. Most of the nitrides entered the gas phase without undergoing hydrogenation, resulting in a high nitrogen content in the product. Comparative Example 1 employed conventional fixed-bed hydrogenation technology and a traditional catalyst loading system. The high pressure in the first hydrogenation reactor resulted in simultaneous desulfurization and denitrification, which affected each other. The reaction conditions could not be optimized for either reaction. Sulfide removal was deep at high pressure, affecting product lubricity and increasing hydrogen consumption. The hydrogen was recycled, resulting in high energy consumption.

[0167] Examples D1-D3

[0168] The process was carried out using the flow diagram shown in Figure 1. Two 100mL fixed-bed hydrogenation reactors were connected in series: a gas-phase hydrogenation reactor and a liquid-phase hydrogenation reactor. A heat exchanger, a gas-liquid separator, and a booster pump were installed between the reactors. The gas-phase hydrogenation reactor was loaded with 60mL of Pt-Pd type hydrogenation catalyst A, and the liquid-phase hydrogenation reactor was loaded with 40mL of Pt-Pd type hydrogenation catalyst A. A gas phase component outlet was located above the gas-liquid separator and connected to the bottom effluent of the liquid-phase hydrogenation reactor. The gas phase component then entered the stripping tower together to obtain the reformed oil product.

[0169] Reforming oil was used as the raw material. Catalyst properties are shown in Table D1, raw oil properties are shown in Table D2, and reaction process conditions and results are shown in Table D3.

[0170] Example D4

[0171] The process flow is the same as that of Example 1-3, except that the liquid phase reaction pressure is increased. The reaction process conditions and results are shown in Table 3.

[0172] Example D5

[0173] The process flow is the same as that of Example 1-3, except that the liquid phase reaction temperature is increased. The reaction process conditions and results are shown in Table 3.

[0174] Comparative Example D1

[0175] A conventional fixed-bed hydrogenation process was used, with a gas-phase hydrogenation reactor installed. The reactor was loaded with 100 mL of Pt-Pd hydrogenation catalyst A. After hydrogen sulfide removal, the hydrogen was compressed and recycled using a circulating hydrogen compressor. The raw materials and catalyst properties were the same as in Examples D1-D3. The reaction conditions and results are shown in Table D3.

[0176] Comparative Example D2

[0177] A conventional fixed-bed hydrogenation process was used, with a liquid-phase hydrogenation reactor installed. The reactor was loaded with 100 mL of Pt-Pd hydrogenation catalyst A. Hydrogen was mixed with the feedstock oil in a hydrogen mixer before entering the reactor. The feedstock and catalyst properties were the same as in Examples D1-D3. The reaction conditions and results are shown in Table D3.

[0178] Table D1 Catalyst Physicochemical Properties

[0179] Table D2 Raw oil properties

[0180] Table D3 Hydrogenation process conditions and results

[0181] As shown in Table D3, Comparative Example 1 employed conventional fixed-bed hydrogenation technology, requiring a circulating hydrogen compressor for hydrogen recycle, resulting in high costs and significant aromatics losses. Comparative Example 2 employed liquid-phase hydrogenation technology, which required a single hydrogen dissolution to meet the reaction's hydrogen consumption, leading to higher pressures and increased aromatics losses.

[0182] Examples E1-E3

[0183] The process flow diagram shown in Figure 1 was used. Two 100mL fixed-bed hydrogenation reactors were connected in series, one for gas phase hydrogenation and the other for liquid phase hydrogenation. A heat exchanger, gas-liquid separator, booster pump, and hydrogen mixer were installed between the reactors. The first hydrogenation reactor was loaded with 50mL of Mo-Co type hydrogenation catalyst A, and the second hydrogenation reactor was loaded with 50mL of Mo-Ni type hydrogenation catalyst B. A gas phase outlet was set above the gas-liquid separator, which was connected to the bottom effluent of the second hydrogenation reactor and then entered the stripping tower together to obtain the waste tire pyrolysis oil product.

[0184] Middle distillate waste tire pyrolysis oil was used as the raw material. Catalyst properties are shown in Table A1, raw oil properties are shown in Table E2, and reaction process conditions and results are shown in Table E3.

[0185] Example E4

[0186] The same process flow as in Examples E1-E3 was used, except that the separation conditions for gas-liquid separation were changed. The reaction process conditions and results are shown in Table E3.

[0187] Comparative Example E1

[0188] The same process flow as in Examples E1-E3 was used, with the only difference being an increase in the hydrogen-to-oil volume ratio at the inlet of the gas-phase hydrogenation reactor, the elimination of the hydrogen mixer before the liquid-phase hydrogenation reactor, and the use of only a liquid-phase booster pump. The reaction process conditions and results are shown in Table E3.

[0189] Table E2 Crude Oil Properties

[0190] Table E3 Hydrogenation process conditions and results

[0191] As shown in Table E3, Comparative Example E1 does not incorporate a hydrogen mixer before the liquid-phase hydrogenation reactor, but instead provides a sufficient hydrogen source by increasing the hydrogen-to-oil volume ratio at the gas-phase hydrogenation reactor inlet. However, the high hydrogen-to-oil volume ratio in the gas-phase hydrogenation reactor results in a low oil-phase partial pressure within the gas-liquid separator, resulting in insufficient liquefaction. This means that it is difficult for all PAHs to enter the liquid phase for the liquid-phase hydrogenation reaction, resulting in a relatively high PAH content in the refined oil. Furthermore, the hydrogen circulation requires the use of a compressor, resulting in higher energy consumption than in the Examples.

[0192] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. A method for hydrogenating hydrocarbon oil, characterized in that: The method comprises the following steps: (1) subjecting a hydrocarbon oil feedstock to a gas phase hydrogenation reaction in a gas phase hydrogenation reactor in the presence of hydrogen to obtain a gas phase hydrogenation reaction product; (2) separating the gas-phase hydrogenation reaction product into gas and liquid to obtain a gas phase component and a liquid phase component; (3) pressurizing the liquid phase components and optionally mixing hydrogen, and then performing a liquid phase hydrogenation reaction in a liquid phase hydrogenation reactor to obtain a liquid phase hydrogenation reaction product.

2. The method according to claim 1, wherein: The hydrocarbon oil raw material is selected from at least one of diesel, waste tire pyrolysis oil, jet fuel and reforming oil; Preferably, the diesel has an initial boiling point of 140-200°C, a final boiling point of 300-400°C, an S content of ≯20000 μg / g, preferably ≯15000 μg / g, an N content of ≯1000 μg / g, preferably ≯800 μg / g, and a polycyclic aromatic hydrocarbon content of ≯50 wt%; Preferably, the initial boiling point of the waste tire pyrolysis oil is 130-180°C, the final boiling point is 320-350°C, the S content is ≯15000 μg / g, preferably ≯11000 μg / g, the N content is 5000-10000 μg / g, and the polycyclic aromatic hydrocarbon content is ≯50wt%; Preferably, the initial boiling point of the aviation kerosene is 80-200°C, the final boiling point is 220-320°C, the S content is ≯5000μg / g, preferably ≯3000μg / g, and / or the N content of the aviation kerosene is ≯100μg / g, and / or the smoke point of the aviation kerosene is 18-25mm; Preferably, the distillation range of the reformed oil is 40-180° C., the bromine index is ≯8000 mg Br / 100 g, and the aromatic content is 60-85 wt %.

3. The method according to claim 1 or 2, wherein: The gas phase hydrogenation reaction is carried out in the presence of a first hydrogenation catalyst, wherein the first hydrogenation catalyst is a catalyst having a hydrodesulfurization function and / or a hydrodeolefination function; Preferably, the conditions of the gas phase hydrogenation reaction include: pressure of 0.1-2.8 MPa, preferably 0.3-2 MPa, temperature of 80-400°C, preferably 100-380°C, hydrogen-to-oil volume ratio of 5-500, preferably 5-300, volume space velocity of 0.5-15h -1 , preferably 0.5 to 13 hours -1 ; Preferably, the gas phase hydrogenation reactor is a fixed bed reactor.

4. The method according to any one of claims 1 to 3, wherein: The sulfur content of the gas phase hydrogenation reaction product obtained in step (1) accounts for at least 1 wt % of the sulfur content in the hydrocarbon oil feedstock, preferably 5-60 wt %.

5. The method according to any one of claims 1 to 4, wherein: Based on the total amount of the gas phase hydrogenation reaction product, the mass percentage of the liquid phase component is 1% to 80%, preferably 10% to 75%; Preferably, the method further comprises, before performing the gas-liquid separation, subjecting the gas-phase hydrogenation reaction product to heat exchange and cooling; Preferably, the heat exchange causes the gas phase hydrogenation reaction product to drop in temperature by 8 to 50°C, preferably by 15 to 50°C.

6. The method according to any one of claims 1 to 5, wherein: The conditions of the liquid phase hydrogenation reaction include: pressure of 1 to 10 MPa, preferably 1.5 to 7 MPa, temperature of 50 to 400°C, preferably 80 to 360°C, volume space velocity of 0.1 to 10 h -1 , preferably 0.1~9h -1 ; Preferably, the liquid phase hydrogenation reactor is a fixed bed reactor.

7. The method according to any one of claims 1 to 6, wherein: The pressure of the liquid phase hydrogenation reaction is at least 0.5 MPa higher than the pressure of the gas phase hydrodesulfurization, preferably 0.5 to 6 MPa higher; And / or, the temperature of the liquid phase hydrogenation reaction is at least 5°C lower than the temperature of the gas phase hydrodesulfurization reaction, preferably 10-60°C lower.

8. The method according to any one of claims 1 to 7, wherein: The hydrocarbon oil feedstock is diesel or waste tire pyrolysis oil, the gas phase hydrogenation reaction is carried out in the presence of a first hydrogenation catalyst, and the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst; The first hydrogenation catalyst is a catalyst having a hydrodesulfurization function; the second hydrogenation catalyst is a hydrodearomatization catalyst; Preferably, step (3) comprises pressurizing and mixing the liquid phase components with hydrogen, relative to 1m 3 The liquid phase component, the hydrogen replenishment amount is 10~40Nm 3 ; Preferably, the conditions of the gas phase hydrogenation reaction include: pressure of 0.5-2 MPa, temperature of 330-380°C, hydrogen-to-oil volume ratio of 50-220, volume space velocity of 0.5-2 h -1 ; Based on the total amount of the gas phase hydrogenation reaction product, the mass percentage of the liquid phase component is 10 to 40%; the conditions of the liquid phase hydrogenation reaction include: pressure of 3 to 7 MPa, temperature of 260 to 360 ° C, volume space velocity of 0.1 to 1 h -1 .

9. The method according to any one of claims 1 to 7, wherein: The hydrocarbon oil raw material is aviation kerosene, The gas phase hydrogenation reaction is carried out in the presence of a first hydrogenation catalyst, and the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst; The first hydrogenation catalyst is a catalyst having a hydrodesulfurization function; The second hydrogenation catalyst is a hydrodearomatization catalyst and / or a hydrodenitrogenation catalyst; Preferably, the conditions of the gas phase hydrogenation reaction include: pressure of 0.5-2 MPa, temperature of 300-350°C, hydrogen-to-oil volume ratio of 20-100, and volume space velocity of 4-13 h -1 ; Based on the total amount of the gas phase hydrogenation reaction product, the mass percentage of the liquid phase component is 10 to 35%; the conditions of the liquid phase hydrogenation reaction include: pressure of 2 to 5 MPa, temperature of 220 to 300 ° C, volume space velocity of 0.5 to 1.5 h -1 .

10. The method according to any one of claims 1 to 7, wherein: The hydrocarbon oil feedstock is reformed oil, the gas phase hydrogenation reaction is carried out in the presence of a first hydrogenation catalyst, and the liquid phase hydrogenation reaction is carried out in the presence of a second hydrogenation catalyst; The first hydrogenation catalyst and the second hydrogenation catalyst are each independently a hydrodeolefination catalyst; Preferably, the conditions of the gas phase hydrogenation reaction include: pressure of 0.5-2 MPa, temperature of 150-200°C, hydrogen-to-oil volume ratio of 20-50, volume space velocity of 4-8 h -1 ; Based on the total amount of the gas phase hydrogenation reaction product, the mass percentage of the liquid phase component is 40-75%; the conditions of the liquid phase hydrogenation reaction include: pressure of 3-6MPa, temperature of 260-345°C, volume space velocity of 0.2-1h -1 .

11. A hydrocarbon oil hydrogenation system, characterized in that: The system comprises a gas phase hydrogenation reactor filled with a first hydrogenation catalyst, a gas-liquid separation unit, a pressurizing unit and a liquid phase hydrogenation reactor filled with a second hydrogenation catalyst, which are connected in sequence; the liquid phase component outlet of the gas-liquid separation unit is connected to the inlet of the pressurizing unit; The gas phase hydrogenation reactor is used to carry out a gas phase hydrogenation reaction on the hydrocarbon oil feedstock and hydrogen to obtain a gas phase hydrogenation reaction product; The pressurizing unit is used to pressurize the liquid phase component obtained by the gas-liquid separation unit to obtain a pressurized logistics; The liquid phase hydrogenation reactor is used to contact the pressurized logistics with the second hydrogenation catalyst to perform a liquid phase hydrogenation reaction to obtain a liquid phase hydrogenation reaction product; Optionally, the system further comprises a hydrogen mixing unit disposed between the pressurizing unit and the liquid phase hydrogenation reactor, wherein the hydrogen mixing unit is used to mix the pressurized logistics with hydrogen.

12. The hydrogenation system according to claim 11, wherein: The gas phase hydrogenation reactor and the liquid phase hydrogenation reactor are fixed bed reactors.

13. The hydrogenation system according to claim 11 or 12, wherein: The system further comprises a stripping unit, the inlet of which is connected to the outlet of the liquid phase hydrogenation reaction product of the liquid phase hydrogenation reactor and / or to the outlet of the gas phase component of the gas-liquid separation unit.

14. The hydrogenation system according to any one of claims 11 to 13, wherein: The hydrocarbon oil feedstock is diesel or waste tire pyrolysis oil, the first hydrogenation catalyst is a catalyst having a hydrodesulfurization function, and the second hydrogenation catalyst is a hydrodearomatization catalyst; or The hydrocarbon oil feedstock is heavy jet fuel, the first hydrogenation catalyst is a catalyst having a hydrodesulfurization function, and the second hydrogenation catalyst is a hydrodearomatization catalyst; or The hydrocarbon oil feedstock is high-nitrogen jet fuel, the first hydrogenation catalyst is a catalyst having a hydrodesulfurization function, and the second hydrogenation catalyst is a hydrodenitrogenation catalyst; or The hydrocarbon oil feedstock is reformed oil, the first hydrogenation catalyst is a hydrogenation deolefination catalyst, and the second hydrogenation catalyst is a hydrogenation deolefination catalyst.

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