Method for preparing benzene series from polycyclic aromatic hydrocarbon and application of benzene series

Abstract

The invention relates to the field of hydrogenation utilization of inferior oil products, in particular to a method for preparing benzene series from polycyclic aromatic hydrocarbon and application of the benzene series. The method comprises the following steps: carrying out a first hydrogenation reaction on a polycyclic aromatic hydrocarbon raw material in the presence of a first catalyst, and carrying out a second hydrogenation reaction on a product flow of the first hydrogenation reaction in the presence of a second catalyst; the inlet temperature of the first hydrogenation reaction is lower than 260 DEG C, and the inlet temperature of the second hydrogenation reaction is not lower than 260 DEG C; the first catalyst and the second catalyst respectively and independently comprise a sulfide of a hydrogenation active element and a carrier, and the sulfide of the hydrogenation active element is of a tower-shaped structure. According to the present invention, the polycyclic aromatic hydrocarbon raw material is subjected to two-step hydrogenation such as the low-temperature hydrogenation reaction and the high-temperature hydrogenation reaction by using the catalyst comprising the sulfide of the hydrogenation active element and the carrier, and the sulfide of the hydrogenation active element has the tower-shaped structure, such that the polycyclic aromatic hydrocarbon in the polycyclic aromatic hydrocarbon raw material is efficiently converted into the monocyclic aromatic hydrocarbon, and the deep denitrification is simultaneously performed.

Description

Technical Field

[0001] This invention relates to the hydrogenation utilization of inferior oil products, specifically to a method for preparing benzene compounds from polycyclic aromatic hydrocarbons and its application. Background Technology

[0002] Low-quality aromatic-rich distillate oils are mainly light cycle oil (LCO) from catalytic cracking and byproducts of ethylene cracking. LCO has a total aromatic content as high as 80%, with naphthalene-based bicyclic aromatics accounting for about 70%, monocyclic and tricyclic aromatics each accounting for about 15%, and the remainder being alkanes, cycloalkanes, and alkenes. LCO has a sulfur content of 0.2–1.5 wt%, a nitrogen content of 100–1000 ppm, a cetane number of only 15–25, and poor ignition performance. Ethylene cracking byproducts include cracked C9 and ethylene tar fractions. Cracked C9 mainly comes from the cracked gasoline C9 fraction separated after passing through the BTX tower, with an aromatic content exceeding 70%, accounting for 11%–22% of ethylene production. Domestically, the vast majority of cracked C9 is sold only as inexpensive primary fuel oil or after preliminary processing.

[0003] LCO exhibits poor stability, rapidly increasing gum content and darkening color during storage and transportation, severely impacting its usability. The primary reason for LCO's poor stability is its high content of basic and non-basic nitrogen compounds, thiols, and thiophenols, with nitrogen compounds having a particularly significant impact. Nitrogen compounds in catalytic diesel are classified as basic and non-basic nitrogen compounds. The NOx emitted after combustion is a major contributor to acid rain and smog. Furthermore, the presence of nitrogen compounds severely hinders the deep removal of sulfur. On one hand, during hydrodesulfurization, nitrogen compounds and sulfides compete for adsorption at catalyst active sites, with nitrogen compounds exhibiting stronger adsorption capacity, thus inhibiting the hydrodesulfurization reaction. On the other hand, basic nitrogen compounds can poison the hydrodesulfurization catalyst.

[0004] Producing high-quality diesel fuel with low sulfur, low aromatics, and high cetane number is gradually becoming the global trend in automotive diesel fuel production. The increasing deterioration and heavier composition of crude oil has resulted in higher aromatic and cycloalkanes content in diesel fuel produced by refineries. These aromatics and cycloalkanes contribute relatively little to the cetane number of diesel fuel, making it difficult for diesel products to meet the required cetane number standards. Therefore, reducing the aromatic and cycloalkanes content in diesel fuel has become a real challenge for refineries in upgrading diesel quality.

[0005] Currently, the most common processing methods for low-quality aromatic-rich distillate oils are hydrorefining and hydrocracking. Hydrorefining of low-quality aromatic-rich distillate oils mainly involves desulfurization, denitrogenation, and selective hydrogenation saturation reactions of aromatics, which can improve its color and stability, but the increase in cetane number is relatively small, far from meeting the cetane number requirements for automotive diesel. Hydrocracking processes, such as UOP's Unicracking process, can produce gasoline, jet fuel, and diesel products, but when the proportion of secondary processed oils such as low-quality aromatic-rich distillate oils is too high, qualified jet fuel products cannot be obtained.

[0006] Research on catalysts for the hydrotreating of heavy distillate oils has long been one of the most active areas of research. Alumina-supported Co(Ni)Mo(W) / Al₂O₃ catalysts were first applied in industrial hydrotreating reactions as early as 1943 and remain commonly used catalysts in many hydrorefining processes. Currently, the most common combination of active components in typical hydrotreating catalysts is Co-Mo, Ni-Mo, and Ni-W, which typically contain 1-4 wt% Co(Ni)O, 8-16 wt% MoO₃, or 12-25 wt% WO₃.

[0007] CN1064988C discloses a method for hydroconversion of diesel fractions. This method uses a hydroconversion catalyst containing molecular sieves and employs a single-stage, tandem single-stage, and two-stage hydroconversion process. It utilizes a supported hydroconversion catalyst containing molecular sieves to remove aromatics, desulfurize, and improve the cetane number of low-quality diesel, especially LCO. However, the diesel produced by this method has low specifications; its sulfur and nitrogen content, aromatics, and cetane number do not meet current national standards. A large amount of aromatic components in the catalytic cracking diesel feedstock are hydrosaturated and not effectively utilized.

[0008] CN1872959A discloses a hydrogenation catalyst using alumina as a support, nickel, molybdenum, and tungsten as active components, and fluorine as an auxiliary agent. Compared with traditional bimetallic hydrogenation catalysts, the trimetallic catalyst provided by this catalyst exhibits improved activity. However, due to limitations in support performance and other factors, the improvement is limited, and the activity remains relatively low. Furthermore, this catalyst also faces challenges such as fluorine loss during industrial operation, fluorine corrosion of equipment, and environmental pollution.

[0009] CN1040610A discloses a hydrorefining catalyst supported on γ-Al₂O₃ containing TiO₂. The supported catalyst γ-Al₂O₃ contains 5-30% titanium oxide, with W, Mo, and Ni as active components. The performance of the hydrorefining catalyst supported on TiO₂-modified Al₂O₃ is improved to some extent. However, the support has a lower acidity, especially fewer moderately strong acid centers, which is unfavorable for the ring-opening and breaking of nitrogen heterocycles, resulting in a less significant denitrification effect.

[0010] The petroleum hydrocarbon hydrotreating method proposed in US4880524A employs a highly active hydrogenation catalyst. This catalyst is of the Ni-Mo / Al₂O₃ type with a specific surface area greater than 300 m². 2 / g, with pore sizes smaller than 7nm exceeding 70%. This catalyst exhibits good hydrorefining activity for light distillate oils, but its hydrorefining effect on catalytic cracking light cycle oils is poor. Summary of the Invention

[0011] A comprehensive analysis of the above technologies reveals that existing technologies involve the production of gasoline and diesel fractions, petroleum coke, etc., from inferior aromatic oils, wasting valuable aromatic resources in these oils. Selective hydrorefining of aromatic oils, along with catalyst innovation and process optimization, promotes the conversion of polycyclic aromatic hydrocarbons (PAHs) in the oils, maximizing the selectivity of monocyclic aromatic hydrocarbons in the products. This provides high-quality feedstock for hydrocracking, maximizing the production of benzene, toluene, and xylene. This approach fully utilizes inferior aromatic oils, increases their added value, achieves chemical-grade utilization of refining byproducts, and enhances industry competitiveness.

[0012] The purpose of this invention is to overcome the problems of low conversion rate of polycyclic aromatic hydrocarbons and low selectivity of monocyclic aromatic hydrocarbons (benzene series compounds) in the prior art, and to provide a method for preparing benzene series compounds from polycyclic aromatic hydrocarbons and its application. This method has the characteristics of high conversion rate of polycyclic aromatic hydrocarbons and high selectivity of monocyclic aromatic hydrocarbons (benzene series compounds).

[0013] To achieve the above objectives, the first aspect of the present invention provides a method for preparing benzene compounds from polycyclic aromatic hydrocarbons, the method comprising: a polycyclic aromatic hydrocarbon feedstock undergoing a first hydrogenation reaction in the presence of a first catalyst, and the product stream of the first hydrogenation reaction undergoing a second hydrogenation reaction in the presence of a second catalyst; the inlet temperature of the first hydrogenation reaction is below 260°C, and the inlet temperature of the second hydrogenation reaction is not below 260°C; the first catalyst and the second catalyst each independently comprise a sulfide of a hydrogenation active element and a support, and the sulfide of the hydrogenation active element has a tower-shaped structure.

[0014] The second aspect of the present invention provides an application of the method described in the first aspect in the hydrogenation of inferior aromatic-rich distillate oil.

[0015] Through the above technical solution, the present invention has the following advantages:

[0016] This invention utilizes a catalyst containing a sulfide with hydrogenation active elements and a support, wherein the sulfide with hydrogenation active elements has a tower-shaped structure, to perform a two-step hydrogenation reaction on polycyclic aromatic hydrocarbon feedstock, namely a low-temperature hydrogenation reaction and a high-temperature hydrogenation reaction. This enables the polycyclic aromatic hydrocarbon feedstock to be efficiently converted into monocyclic aromatic hydrocarbons, while simultaneously achieving deep denitrification.

[0017] The method of this invention for hydrotreating inferior aromatic-rich distillate oils has a high conversion rate of polycyclic aromatic hydrocarbons such as acenaphthenes and acenaphthenes, and the products have a high aromatic hydrocarbon retention rate and monocyclic aromatic hydrocarbon selectivity, as well as extremely low nitrogen content, making them a high-quality feedstock for hydrocracking. Attached Figure Description

[0018] Figure 1 This is a TEM image of the catalyst in Example 1;

[0019] Figure 2 This is a TEM image of the catalyst in Example 11;

[0020] Figure 3 These are the H2-TPR diagrams of the catalysts in Examples 1 and 11;

[0021] Figure 4 This is a process flow diagram for preparing benzene compounds from polycyclic aromatic hydrocarbons according to a preferred embodiment of the present invention. Detailed Implementation

[0022] The endpoints and any values ​​of the ranges 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 endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0023] In this invention, benzene compounds refer to monocyclic aromatic hydrocarbons.

[0024] This invention provides a method for preparing benzene compounds from polycyclic aromatic hydrocarbons. The method includes: a first hydrogenation reaction of the polycyclic aromatic hydrocarbon feedstock in the presence of a first catalyst; and a second hydrogenation reaction of the product stream of the first hydrogenation reaction in the presence of a second catalyst. The inlet temperature of the first hydrogenation reaction is below 260°C, and the inlet temperature of the second hydrogenation reaction is not below 260°C. The first catalyst and the second catalyst each independently include a sulfide of a hydrogenation active element and a support, and the sulfide of the hydrogenation active element has a tower-shaped structure.

[0025] This invention utilizes a catalyst containing a sulfide with hydrogenation active elements and a support, wherein the sulfide with hydrogenation active elements has a tower-shaped structure, to perform a two-step hydrogenation reaction on polycyclic aromatic hydrocarbon feedstock, namely a low-temperature hydrogenation reaction below 260°C and a high-temperature hydrogenation reaction at a temperature not lower than 260°C. This enables the polycyclic aromatic hydrocarbon feedstock to be efficiently converted into monocyclic aromatic hydrocarbons, while simultaneously achieving deep denitrification.

[0026] According to a preferred embodiment of the present invention, the hydrogenation active element comprises Group VIB elements and / or Group VIII elements, as well as lanthanide metals and / or Group VA elements. By adopting the aforementioned preferred embodiment, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0027] According to a preferred embodiment of the present invention, the number of stacking layers of the tower-shaped structure is 3-8 layers, preferably 3-5 layers.

[0028] According to a preferred embodiment of the present invention, the stacking length of the tower-shaped structure is 3-15 nm, preferably 3-6 nm.

[0029] According to a preferred embodiment of the present invention, the interlayer spacing of the stacked layers of the tower-shaped structure is 0.2-3 nm, preferably 0.2-2 nm.

[0030] According to a preferred embodiment of the present invention, the hydrogen reduction peak temperature of the first catalyst and the second catalyst is 350-450°C, preferably 350-400°C.

[0031] In this invention, the support in the catalyst can be any conventional choice in the art as long as the purpose of this invention can be achieved. This is an illustrative example, but does not limit the scope of this invention. According to a preferred embodiment of this invention, the support in the catalyst is selected from at least one of Al2O3, TiO2, and SiO2.

[0032] In this invention, the catalyst support is selected from at least one of Al2O3, Al2O3-TiO2 (Al2O3 90-99%, TiO2 1-10%), Al2O3-TiO2-SiO2 (Al2O3 60-98%, TiO2 1-10%, SiO2 1-30%), and Al2O3-SiO2 (Al2O3 70-99%, SiO2 1-30%), preferably one or two of Al2O3, Al2O3-TiO2 (Al2O3 90-99%, TiO2 1-10%), and Al2O3-TiO2-SiO2 (Al2O3 60-98%, TiO2 1-10%, SiO2 1-30%). By adopting the aforementioned preferred scheme, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series (monocyclic aromatic hydrocarbons) can be further improved.

[0033] According to a preferred embodiment of the present invention, the hydrogenation active element comprises Group VIB elements, Group VIII elements, lanthanides, and / or Group VA elements, wherein the mass ratio of Group VIB elements to Group VIII elements, calculated as oxides, is 0.1-15, preferably 0.8-9. By adopting the aforementioned preferred embodiment, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0034] According to a preferred embodiment of the present invention, based on the total mass of the unsulfurized catalyst (pre-sulfurized composite metal oxide catalyst), the oxide content of the hydrogenation active element in each of the first and second catalysts is not less than 30 g / (L support), preferably 30-600 g / (L support), more preferably 40-560 g / (L support). By adopting the aforementioned preferred embodiment, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0035] In this invention, there are no special requirements for the type of Group VIB element in the first catalyst. According to a preferred embodiment of this invention, the Group VIB element in the first catalyst is selected from Mo.

[0036] In this invention, there are no special requirements for the type of Group VIB element in the second catalyst. According to a preferred embodiment of the invention, the Group VIB element in the second catalyst is selected from Mo and / or W, preferably Mo and W, and the mass ratio of Mo to W as oxides is 0.25-4, for example, 0.25, 0.5, 1, 2, 3, 4.

[0037] In this invention, there are no special requirements for the Group VIII elements in the first catalyst and the second catalyst. According to a preferred embodiment of this invention, the Group VIII elements in the first catalyst and the second catalyst are each independently selected from at least one of Ni, Co, Fe, Pt, and Pd.

[0038] In this invention, as long as the purpose of this invention can be achieved, the lanthanide metals in the first catalyst and the second catalyst can be conventional choices in the art. According to a preferred embodiment of this invention, the lanthanide metals in the first catalyst and the second catalyst are each independently selected from La and / or Ce.

[0039] In this invention, as long as the purpose of this application can be achieved, there are no special requirements for the Group VA elements in the first catalyst and the second catalyst. According to a preferred embodiment of this invention, the Group VA element in the first catalyst and the second catalyst is P.

[0040] According to a preferred embodiment of the present invention, the average pore size of the first catalyst and the second catalyst is 7-20 nm.

[0041] According to a preferred embodiment of the present invention, the pore volumes of the first catalyst and the second catalyst are each 0.5-0.98 cm³. 3 / g.

[0042] According to a preferred embodiment of the present invention, the specific surface area of ​​the first catalyst and the second catalyst are each 190-380 m². 2 / g.

[0043] Catalysts with the above-mentioned pore size, pore volume, and specific surface area characteristics can further improve the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series (monocyclic aromatic hydrocarbons).

[0044] As long as a catalyst with the aforementioned characteristics of the present invention can be prepared, there are no special requirements for the preparation method of the catalyst. This is an illustrative example, but it does not limit the scope of the present invention. According to a preferred embodiment of the present invention, the preparation methods of the first catalyst and the second catalyst in the present invention each independently include: mixing an active component source, a support, a chelating surfactant, and a water-soluble additive to obtain a mixture, impregnation, curing, drying, calcination, and sulfidation.

[0045] The impregnation solution (mixture) of this invention has a high metal concentration, weak interaction between the active component and the carrier, and low viscosity. After subsequent curing, drying, calcination, and sulfidation, the active metal of the catalyst has a tower-shaped structure, with active sites such as edges, corners, and sides fully exposed. It has good hydrogenation activity of polycyclic aromatic hydrocarbons, high desulfurization and denitrification rates, and can be used under relatively mild conditions with a long service life.

[0046] According to a preferred embodiment of the present invention, the chelating surfactant is selected from at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and ethylenediaminetetraacetic acid, preferably a combination of 1,2-cyclohexanediaminetetraacetic acid and / or ethylenediaminetetraacetic acid and citric acid, more preferably the mass ratio of 1,2-cyclohexanediaminetetraacetic acid and / or ethylenediaminetetraacetic acid and citric acid in the composition is 1-2:2-1.

[0047] According to a preferred embodiment of the present invention, the water-soluble additive is selected from one or more of ethanol, ethylene glycol, and acetone;

[0048] According to a preferred embodiment of the present invention, the chelated surfactant content in the mixture is 0.01-10% by mass, preferably 1.5-7.5%.

[0049] According to a preferred embodiment of the present invention, the water-soluble additive in the mixture has a mass content of 0.1-20%, preferably 5-12.5%.

[0050] The immersion conditions described in this invention have a wide range of options and are illustrative, but do not limit the scope of this invention. According to a preferred embodiment of this invention, the immersion conditions include: a contact temperature of 0-60°C, preferably 20-40°C, and an immersion time determined according to the immersion temperature, for example, a contact time of 5-60 min, preferably 8-40 min.

[0051] According to a preferred embodiment of the present invention, the conditions for conditioning include: a conditioning temperature of 0-60℃, preferably 20-40℃, and a conditioning time determined according to the immersion temperature, for example, a conditioning time of 2-96h, preferably 6-48h.

[0052] The drying conditions described in this invention have a wide range of options and are illustrative, but do not limit the scope of the invention. According to a preferred embodiment of the invention, the drying conditions include: a drying temperature of 50-130°C, preferably 60-120°C; a drying time determined according to the impregnation temperature, for example, a drying time of 3-46 hours, preferably 6-36 hours; and drying at an air volume hourly space velocity of 2-250 h⁻¹. -1 Preferably 20-150h -1 In the process.

[0053] The calcination conditions described in this invention have a wide range of selectable options and are illustrative, but do not limit the scope of this invention. According to a preferred embodiment of this invention, the calcination conditions include: a calcination temperature of 150-520℃, preferably 280-500℃; a calcination time determined based on the impregnation temperature, for example, a calcination time of 2-24 hours, preferably 3-12 hours; and calcination at an air volume space velocity of 2-250 h⁻¹. -1 Preferably 25-150h -1 In the process.

[0054] In this invention, the vulcanization conditions are vulcanization performed by programmed temperature increase, including: initial temperature of 25-150℃, heating rate of 5-15℃ / min, step temperature interval of 50-80℃, step temperature holding time of 4-10h, final temperature of 320-360℃, and final temperature treatment time of 12-35h.

[0055] According to a preferred embodiment of the present invention, the inlet temperature of the first hydrogenation reaction is at least 15°C lower than the inlet temperature of the second hydrogenation reaction, preferably 15-60°C, for example, 15°C, 20°C, 25°C, 30°C, 35°C, 50°C, or 60°C. By adopting the aforementioned preferred embodiment, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0056] According to a preferred embodiment of the present invention, the outlet temperature of the first hydrogenation reaction is 260-330°C, preferably 270-325°C.

[0057] According to a preferred embodiment of the present invention, the outlet temperature of the second hydrogenation reaction is 330-380°C, preferably 340-365°C.

[0058] According to a preferred embodiment of the present invention, the conditions for the first hydrogenation reaction include: the temperature of the pre-hydrogenation stream, i.e., the inlet temperature, is 200-260°C, preferably 220-250°C; the temperature of the post-hydrogenation stream is at most 85°C higher than the temperature of the pre-hydrogenation stream, preferably 50-85°C, more preferably 55-80°C, for example, 55°C, 65°C, 75°C, or 80°C. By adopting the aforementioned preferred scheme, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0059] In this invention, as long as the purpose of this invention can be achieved, the pressure of the first hydrogenation reaction can be a conventional choice in the art. This is an illustrative example, but does not limit the scope of this invention. According to a preferred embodiment of this invention, the pressure of the first hydrogenation reaction is 5.0-12.0 MPa.

[0060] In this invention, as long as the purpose of this invention can be achieved, the hydrogen-to-oil volume ratio of the first hydrogenation reaction can be a conventional choice in the art. This is an illustrative example, but does not limit the scope of this invention. According to a preferred embodiment of this invention, the hydrogen-to-oil volume ratio of the first hydrogenation reaction is 400-3000.

[0061] In this invention, as long as the objective of the invention can be achieved, the hydrogen-to-oil volume ratio in the first hydrogenation reaction can be a conventional choice in the art. This is illustrative but does not limit the scope of the invention. According to a preferred embodiment of the invention, the liquid hourly space velocity (LHSV) of the first hydrogenation reaction is 0.2-3 h⁻¹. -1 .

[0062] According to a preferred embodiment of the present invention, the conditions for the second hydrogenation reaction include: the temperature of the pre-hydrogenation stream, i.e., the inlet temperature, is 260-330°C, preferably 280-325°C; the temperature of the post-hydrogenation stream is at most 55°C higher than the temperature of the pre-hydrogenation stream, preferably 20-55°C, more preferably 25-55°C, for example, 25°C, 35°C, 45°C, or 55°C. By adopting the aforementioned preferred scheme, the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons) can be further improved.

[0063] In this invention, the pressure of the second hydrogenation reaction can be any conventional choice in the art as long as the purpose of this invention can be achieved. This is an illustrative example, but does not limit the scope of this invention. According to a preferred embodiment of this invention, the pressure of the second hydrogenation reaction is 5.0-12.0 MPa.

[0064] In this invention, as long as the purpose of this invention can be achieved, the hydrogen-to-oil volume ratio of the second hydrogenation reaction can be a conventional choice in the art. This is an illustrative example, but does not limit the scope of this invention. According to a preferred embodiment of this invention, the hydrogen-to-oil volume ratio of the second hydrogenation reaction is 400-3000.

[0065] In this invention, as long as the objective of the invention can be achieved, the hydrogen-to-oil volume ratio in the second hydrogenation reaction can be a conventional choice in the art. This is an illustrative example, but does not limit the scope of the invention. According to a preferred embodiment of the invention, the liquid hourly space velocity (LHSV) of the second hydrogenation reaction is 0.2-3 h⁻¹. -1 .

[0066] According to a preferred embodiment of the present invention, a portion of the product stream after the first hydrogenation reaction is recycled back to the first hydrogenation reaction process, and another portion is sent to the second hydrogenation reaction process. The mass of the product stream recycled back to the first hydrogenation reaction process accounts for 5-100% of the mass of the polycyclic aromatic hydrocarbon feedstock, preferably 10-50%. By adopting the aforementioned preferred embodiment, the product stream of the first hydrogenation reaction process can be partially recycled, thereby improving the utilization rate of the polycyclic aromatic hydrocarbon feedstock, thus saving resources and reducing costs.

[0067] According to a preferred embodiment of the present invention, the second hydrogenation reaction requires the addition of fresh hydrogen gas, preferably 5-50% of the volume of hydrogen gas in the first hydrogenation reaction, and more preferably 10-30%. By adopting the aforementioned preferred embodiment, the reaction effect of the second hydrogenation reaction can be further enhanced.

[0068] According to a preferred embodiment of the present invention, the polycyclic aromatic hydrocarbon feedstock and the hydrogen feedstock are preheated to 220-250°C before the first hydrogenation reaction.

[0069] According to a preferred embodiment of the present invention, the first hydrogenation reaction is carried out in a catalytic system containing a first catalyst and a protective agent, wherein the protective agent comprises a hydrogenation active component and an oxide support, and the content of the hydrogenation active component, calculated as oxide, is 5-80 g / L (support).

[0070] In this invention, there are no particular requirements for the volume ratio of the first catalyst to the protective agent in the catalytic system. This is an illustrative example, but it does not limit the scope of the invention. According to a preferred embodiment of the invention, the volume ratio of the catalyst to the protective agent in the catalytic system is 4-25, for example, it can be 4, 10, 15, 20, or 25.

[0071] According to a preferred embodiment of the present invention, the oxide content of the hydrogenation active element in the first catalyst is at least 400 g / L (support) higher than the oxide content of the hydrogenation active component in the protective agent, preferably 400-550 g / L (support), for example, 400 g / L (support), 450 g / L (support), 500 g / L (support), or 550 g / L (support). By adopting the aforementioned preferred embodiment, the reaction effect of the first hydrogenation reaction can be further enhanced.

[0072] According to a preferred embodiment of the present invention, the hydrogenation active component in the protective agent comprises a Group VIB element and / or a Group VIII element, and the oxide support is Al2O3.

[0073] According to a preferred embodiment of the present invention, the average pore size of the protective agent is 15-30 nm.

[0074] According to a preferred embodiment of the present invention, the pore volume of the protective agent is 0.7-0.98 cm³. 3 / g.

[0075] According to a preferred embodiment of the present invention, the specific surface area of ​​the protective agent is 120-250 m². 2 / g.

[0076] Protective agents with the above-mentioned pore size, pore volume, and specific surface area characteristics can further improve the conversion rate of polycyclic aromatic hydrocarbons and the selectivity of benzene series compounds (monocyclic aromatic hydrocarbons).

[0077] In this invention, as long as the purpose of this application can be achieved, there are no special requirements for the group VIB elements in the protective agent. According to a preferred embodiment of the present invention, the group VIB elements in the protective agent are selected from Mo and / or W.

[0078] In this invention, as long as the purpose of this application can be achieved, there are no special requirements for the group VIII element in the protective agent. According to a preferred embodiment of the present invention, the group VIII element in the protective agent is selected from at least one of Ni, Co, Fe, Pt, and Pd.

[0079] In this invention, as long as the purpose of this application can be achieved, there are no special requirements for the content of polycyclic aromatic hydrocarbons in the polycyclic aromatic hydrocarbon raw material. According to a preferred embodiment of the present invention, the content of polycyclic aromatic hydrocarbons in the polycyclic aromatic hydrocarbon raw material is not less than 65 wt%, preferably 75-92 wt%.

[0080] In this invention, the method of this invention can be used to prepare benzene compounds from any polycyclic aromatic hydrocarbon. According to a preferred embodiment of this invention, the polycyclic aromatic hydrocarbon includes at least one of naphthalene compounds, acenaphthene compounds, anthracene compounds, and phenanthrene compounds, more preferably acenaphthene compounds.

[0081] This invention provides an application of the method in the hydrogenation of inferior aromatic-rich distillate oil.

[0082] The method of this invention is used to hydrotreat inferior aromatic-rich distillate oils, and the resulting product has a high conversion rate of polycyclic aromatic hydrocarbons such as acenaphthene, a high content of monocyclic aromatic hydrocarbons, and a low nitrogen content, making it a high-quality feedstock for hydrocracking.

[0083] In this invention, the purification method can be a conventional choice in the art, such as purification by a high-precision separator or a phase separator.

[0084] The present invention will be described in detail below through embodiments. The following embodiments:

[0085] TEM images were obtained using a G2F30 transmission electron microscope (TEM) manufactured by FEI Corporation, USA.

[0086] The H2-TPR spectrum was measured using an AutoChem 2920 dynamic adsorption analyzer from Micron Instruments, Inc., USA. The specific parameters were: reducing gas was a 10% H2-Ar mixture, sample mass was 50 mg, gas flow rate was 50 ml / min, and the temperature was increased from room temperature to 800℃ at a rate of 10℃ / min.

[0087] The elemental composition of the catalyst was determined using a ZSX-100e 4580 X-ray fluorescence spectrometer manufactured by Rigaku Corporation, Japan.

[0088] Nitrogen testing was performed on an Antek 900 sulfur-nitrogen analyzer, and the nitrogen content testing method referenced standard SH / T0657-2007 (chemiluminescence method).

[0089] The pore structure parameters of the carrier were measured using a Micrometrics Tristar3000 surface area analyzer at a test temperature of -196℃. Before the test, the sample was vacuum activated at 300℃ for 6 hours.

[0090] Test methods for product composition:

[0091] The composition of hydrogenation products from the preparation of benzene series compounds from polycyclic aromatic hydrocarbons was analyzed by gas chromatography-mass spectrometry (GC-MS) using an Agilent 6890GC-5973MSD instrument and the standard method SH / T 0606-2005.

[0092] The naphthalene conversion rate and tetrahydronaphthalene selectivity in the product are calculated using the following formula:

[0093]

[0094] Unless otherwise specified, all raw materials are commercially available products, and all pressures are gauge pressures.

[0095] Example 1

[0096] 1. Preparation of hydrogenation catalysts

[0097] 1 L of Al₂O₃ support (pore size: 12 nm) was mixed with 0.85 L of a mixed solution containing nickel acetate, lanthanum nitrate, ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid, citric acid, phosphoric acid, ethylene glycol, and acetone (containing 83 g NiO, 440 g MoO₃, 5 g La₂O₃, and 10 g P₂O₅). The amounts of 1,2-cyclohexanediaminetetraacetic acid, citric acid, ethylene glycol, and acetone were 5.5% and 2.5% of the total mass of the mixed solution, respectively. After impregnation, the mixture was cured at 25 °C for 24 hours, dried at 110 °C for 6 hours, and calcined at 450 °C for 4 hours to prepare catalyst A1. Catalyst A1 has a pore size of 11 nm and a pore volume of 0.71 cm³. 3 / g, specific surface area 253m² 2 / g.

[0098] The A1 catalyst contains 83 g / L NiO, 440 g / L MoO3, 5 g / L La2O3, and 10 g / L P2O5.

[0099] 2. Preparation of Protective Agent

[0100] 1 L of Al₂O₃ support (pore size: 18 nm) was mixed with 0.75 L of a mixed solution of ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid (1,2-cyclohexanediaminetetraacetic acid), and phosphoric acid (containing 13 g NiO, 56 g MoO₃, 5 g La₂O₃, 25 g 1,2-cyclohexanediaminetetraacetic acid, and 5 g P₂O₅), wherein the amount of 1,2-cyclohexanediaminetetraacetic acid was 1.8% of the mass of the mixed solution. The mixture was dried at 110 °C for 6 hours and calcined at 450 °C for 4 hours to prepare protective agent B1. Protective agent B1 has a pore size of 18 nm and a pore volume of 0.94 cm⁻¹. 3 / g, specific surface area 213m² 2 / g.

[0101] B1 protective agent contains 13 g / L NiO, 56 g / L MoO3, 5 g / L La2O3, and 5 g / L P2O5.

[0102] 3. Loading of hydrogenation catalyst and protective agent

[0103] B1 and A1 are loaded into hydrogenation reactor 1 in a volume ratio of 2:8, while hydrogenation reactor 2 is filled entirely with A1 catalyst.

[0104] 4. Sulfation of hydrogenation catalyst and protective agent

[0105] After the airtightness is verified, sulfidation begins. The system pressure is 5.5 MPa, and the hydrogen volume hourly space velocity is 800 h⁻¹. -1 Hydrogen gas was introduced, and the temperature was increased from room temperature to 150°C at a rate of 30°C / h. A sulfurizing agent containing 0.5 wt% sulfur was then introduced, and the temperature was increased to 180°C at a rate of 5°C / h for 6 hours, held at this temperature for 5 hours, then increased to 230°C at a rate of 5°C / h for 10 hours, held at this temperature for 4 hours, then increased to 280°C at a rate of 10°C / h for 5 hours, held at this temperature for 4 hours, then increased to 320°C at a rate of 10°C / h for 4 hours, and held at this temperature for 2 hours. The sulfidation of the catalyst was then complete. The sulfidated A1 catalyst was tested using a G2F30 transmission electron microscope (TEM) from FEI Corporation, USA. Figure 1 As can be seen from the figure, the sulfides of the hydrogenation active elements in the sulfide-type catalyst exhibit a tower-like structure with 3-5 stacked layers; the spacing between the stacked layers is 0.2-2 nm; and the stack length is 3-6 nm. Figure 3 The reduction peak temperature of catalyst A1 was measured to be 380℃ by H2-TPR.

[0106] 5. Catalyst Evaluation

[0107] The product streams from LCO feedstocks of low-quality aromatic distillate fractions are shown in Table 1a, such as... Figure 4 It is mixed with fresh hydrogen and circulating hydrogen, preheated in the heater, and then fed into the hydrogenation reactor 1 via feed thermometer 1. After passing through the discharge thermometer 1, part of the material is circulated back to the inlet of the hydrogenation reactor 1 via the heat exchanger. The circulation ratio (product circulation mass of hydrogenation reactor 1 / feed mass of inferior aromatic oil) is 0.3. Part of the product is mixed with cold hydrogen (the amount used is 20% of the volume of fresh hydrogen in hydrogenation reactor 1) and then fed into the hydrogenation reactor 2 via feed thermometer 2. The hydrogenated product passes through the discharge thermometer 2 and then enters the high-precision tank and oil-water separator in sequence before entering the hydrocracking unit.

[0108] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 320℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0109] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 305℃, an outlet temperature of 355℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 ,

[0110] The evaluation results (the effluent stream from hydrogenation reactor 2, hereinafter the same) are listed in Table 2.

[0111] Example 2

[0112] Same as Example 1, except that A1 in hydrogenation reactor 2 is replaced with A2, and the preparation of hydrogenation catalyst A2 is as follows:

[0113] 1 L of Al₂O₃ support (pore size: 12 nm) was mixed with 0.85 L of a mixed solution containing nickel acetate, ammonium metatungstate, lanthanum nitrate, ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid, citric acid, phosphoric acid, ethylene glycol, and acetone (containing 83 g NiO, 340 g WO₃, 100 g MoO₃, 5 g La₂O₃, and 10 g P₂O₅). The amounts of 1,2-cyclohexanediaminetetraacetic acid, citric acid, ethylene glycol, and acetone were 5.5% and 2.5% of the total mass of the mixed solution, respectively. After impregnation, the mixture was cured at 25 °C for 24 hours, dried at 110 °C for 6 hours, and calcined at 450 °C for 4 hours to prepare catalyst A₂. Catalyst A₂ has a pore size of 12 nm and a pore volume of 0.73 cm³. 3 / g, specific surface area 231m² 2 / g. The TEM image obtained after sulfidation is similar to that of A1. The sulfides of hydrogenated active elements have a tower-like structure with 3-5 stacked layers; the stacking layer spacing is 0.2-2nm; the stacking length is 3-6nm. The H2-TPR image is similar to that of A1, and the reduction peak temperature is 385℃.

[0114] The A2 catalyst contains 83 g / L NiO, 100 g / L MoO3, 340 g / L WO3, 5 g / L La2O3, and 10 g / L P2O5.

[0115] The evaluation results are listed in Table 2.

[0116] Example 3

[0117] Same as Example 1, except that the inferior aromatic-rich distillate feedstock comes from LCO, as shown in Table 1b. Figure 4 It is mixed with fresh hydrogen and recycled hydrogen, preheated in the heater, and then fed into the hydrogenation reactor 1 via feed thermometer 1. After passing through the discharge thermometer 1, part of the material is circulated back to the inlet of the hydrogenation reactor 1 via the heat exchanger. The circulation ratio (recycled mass of product from hydrogenation reactor 1 / feed mass of inferior aromatic oil) is 0.5. Part of the product is mixed with cold hydrogen (20% of the volume of fresh hydrogen in hydrogenation reactor 1) and fed into the hydrogenation reactor 2 via feed thermometer 2. The hydrogenated product passes through the discharge thermometer 2 and enters the high-precision tank and oil-water separator in sequence before entering the hydrocracking unit.

[0118] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 325℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0119] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 310℃, an outlet temperature of 365℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1500, and a liquid hourly space velocity of 0.8 h⁻¹. -1 .

[0120] The evaluation results are listed in Table 2.

[0121] Example 4

[0122] Same as Example 1, except that A1 is replaced with A3. Preparation of hydrogenation catalyst A3:

[0123] 1 L of Al₂O₃ support (pore size: 12 nm) was mixed with 0.85 L of a mixed solution containing nickel acetate, lanthanum nitrate, ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid, citric acid, phosphoric acid, ethylene glycol, and acetone (containing 83 g NiO, 470 g MoO₃, 5 g La₂O₃, and 10 g P₂O₅). The amounts of 1,2-cyclohexanediaminetetraacetic acid, citric acid, ethylene glycol, and acetone were 5.5% and 2.5% of the total mass of the mixed solution, respectively. After impregnation, the mixture was cured at 25 °C for 24 hours, dried at 110 °C for 6 hours, and calcined at 450 °C for 4 hours to prepare catalyst A3. Catalyst A3 has a pore size of 11 nm and a pore volume of 0.67 cm⁻¹. 3 / g, specific surface area 261m² 2 / g.

[0124] The A3 catalyst contains 83 g / L NiO, 470 g / L MoO3, 5 g / L La2O3, and 10 g / L P2O5. After sulfidation, the TEM image is similar to that of A1. The sulfides of the hydrogenation active elements exhibit a tower-like structure with 3-5 stacked layers; the interlayer spacing is 0.2-2 nm; the stack length is 3-6 nm; the H2-TPR pattern is similar to A1, and the reduction peak temperature is 382℃.

[0125] The evaluation results are listed in Table 2.

[0126] Example 5

[0127] Same as Example 1, except that:

[0128] The pre-hydrogenation stream temperature in hydrogenation reactor 1 is 250℃, the outlet temperature is 355℃, the pressure is 6.0 MPa, the hydrogen-to-oil volume ratio is 1000, the recycle ratio of the first hydrogenation reactor (recycled product mass of hydrogenation reactor 2 / fresh oil feed mass) is 0.1, and the fresh oil liquid hourly space velocity is 0.8 h⁻¹. -1 .

[0129] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 263℃, an outlet temperature of 290℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 .

[0130] The evaluation results are listed in Table 2.

[0131] Example 6

[0132] Same as Example 1, except that:

[0133] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 260℃, an outlet temperature of 330℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0134] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 315℃, an outlet temperature of 365℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 .

[0135] The evaluation results are listed in Table 2.

[0136] Example 7

[0137] Same as Example 1, except that:

[0138] Replace A1 with A2.

[0139] The evaluation results are listed in Table 2.

[0140] Example 8

[0141] Same as Example 1, except that the product from the first hydrogenation reactor is not recycled.

[0142] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 360℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0143] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 340℃, an outlet temperature of 359℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 .

[0144] The evaluation results are listed in Table 2.

[0145] Example 9

[0146] Same as Example 1, except that B1 is replaced with B2. The preparation of protective agent B2 is as follows:

[0147] 1 L of Al₂O₃ support (pore size: 18 nm) was mixed with 0.75 L of a mixed solution of ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid (CPA), and phosphoric acid (containing 33 g NiO, 168 g MoO₃, 5 g La₂O₃, 25 g CPA, and 5 g P₂O₅), wherein the amount of CPA was 1.8% of the mass of the mixed solution. The mixture was dried at 110 °C for 6 hours and calcined at 450 °C for 4 hours to prepare protective agent B₂. Protective agent B₂ has a pore size of 17 nm and a pore volume of 0.93 cm⁻¹. 3 / g, specific surface area 203m² 2 / g.

[0148] B2 protective agent contains 33 g / L NiO, 168 g / L MoO3, 5 g / L La2O3, and 5 g / L P2O5.

[0149] The evaluation results are listed in Table 2.

[0150] Example 10

[0151] Same as Example 1, except that no protective agent is added to the first reactor.

[0152] The evaluation results are listed in Table 2.

[0153] Example 11

[0154] Same as Example 1, except that A1 is replaced with A4. Preparation of hydrogenation catalyst A4:

[0155] 1 L of Al₂O₃ support (pore size: 12 nm) was mixed with 0.85 L of a mixed solution of nickel acetate, lanthanum nitrate, ammonium molybdate, citric acid, phosphoric acid, and ethylene glycol (containing 83 g NiO, 440 g MoO₃, 5 g La₂O₃, and 10 g P₂O₅). The amount of citric acid and ethylene glycol was 2.5% of the mass of the mixed solution. After impregnation, the mixture was cured at 25 °C for 24 hours, dried at 110 °C for 6 hours, and calcined at 450 °C for 4 hours to prepare catalyst A4. Catalyst A4 has a pore size of 11 nm and a pore volume of 0.71 cm³. 3 / g, specific surface area 253m² 2 / g.

[0156] The A4 catalyst contains 83 g / L NiO, 440 g / L MoO3, 5 g / L La2O3, and 10 g / L P2O5.

[0157] The sulfurized A1 catalyst was tested using a G2F30 transmission electron microscope (TEM) from FEI Corporation, USA. Figure 2As can be seen from the figure, the number of sulfide stacking layers of the hydrogenation active element in the sulfide-type catalyst is 3-8 layers; the stacking layer spacing is 0.2-3 nm; the stacking length is 3-15 nm, and the hydrogenation active phase is unevenly distributed, exhibiting obvious metal sintering characteristics, such as... Figure 3 The reduction peak temperature measured by H2-TPR was 425℃.

[0158] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 320℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0159] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 305℃, an outlet temperature of 355℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 ,

[0160] The evaluation results are listed in Table 2.

[0161] Comparative Example 1

[0162] Same as Example 1, except that:

[0163] The hydrogenation reactor 1 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 275℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1000, and a fresh oil liquid phase volume hourly space velocity of 0.8 h⁻¹. -1 .

[0164] The hydrogenation reactor 2 has a pre-hydrogenation stream temperature of 250℃, an outlet temperature of 275℃, a pressure of 6.0 MPa, a hydrogen-to-oil volume ratio of 1200, and a liquid hourly space velocity of 0.8 h⁻¹. -1 ,

[0165] The evaluation results are listed in Table 2.

[0166] Comparative Example 2

[0167] Same as Example 1, except that A1 is replaced with A5. Preparation of hydrogenation catalyst A5:

[0168] 1.45 L of Al₂O₃ powder (pore size: 12 nm) was mixed with 83 g NiO, 440 g MoO₃, and 5 g La₂O₃. This mixture was then combined with a mixed solution containing P₂O₅, 1,2-cyclohexanediaminetetraacetic acid (1,2-cyclohexanediaminetetraacetic acid), citric acid, phosphoric acid, ethylene glycol, acetone, and nitric acid. The mixture was kneaded, shaped into strips, and extruded. The amounts of 1,2-cyclohexanediaminetetraacetic acid, citric acid, ethylene glycol, and acetone were 5.5% and 2.5% of the mass of the mixed solution, respectively. After extrusion, the strips were cured at 25 °C for 24 hours, dried at 110 °C for 6 hours, and calcined at 450 °C for 4 hours to prepare catalyst A5. Catalyst A5 has a pore size of 8 nm and a pore volume of 0.61 cm³. 3 / g, specific surface area 223m² 2 / g. The sulfide of the hydrogenation active element in the sulfidated catalyst after sulfidation has no tower-like structure.

[0169] The A5 catalyst contains 83 g / L NiO, 440 g / L MoO3, 5 g / L La2O3, and 10 g / L P2O5.

[0170] The evaluation results are listed in Table 2.

[0171]

[0172] Table 2

[0173]

[0174]

[0175] The preferred embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A process for the preparation of benzene homologues from fused ring aromatic hydrocarbons, characterized in that, The method comprises: subjecting a condensed ring aromatic hydrocarbon feedstock to a first hydrogenation reaction in the presence of a first catalyst, and subjecting a product stream of the first hydrogenation reaction to a second hydrogenation reaction in the presence of a second catalyst; An inlet temperature of the first hydrogenation reaction is lower than 260°C, and an inlet temperature of the second hydrogenation reaction is not lower than 260°C; The first catalyst and the second catalyst each independently comprise sulfides of hydrogenation active elements and a carrier, and the sulfides of hydrogenation active elements are in a tower structure.

2. The method according to claim 1, wherein The hydrogenation active elements comprise Group VI B elements and / or Group VIII elements, and lanthanide metals and / or Group VA elements; and / or The tower structure has a stack layer number of 3-8, preferably 3-5, and / or a stack length of 3-15 nm, preferably 3-6 nm, and / or an interlayer spacing of 0.2-3 nm, preferably 0.2-2 nm; and / or The first catalyst and the second catalyst each have a hydrogen reduction peak temperature of 350-450°C, preferably 350-400°C.

3. The method according to claim 1 or 2, wherein The carrier is selected from at least one of Al2O3, TiO2, and SiO2; and / or The hydrogenation active elements comprise Group VI B elements, Group VIII elements, lanthanide metals, and / or Group VA elements, wherein a mass ratio of the Group VI B elements and the Group VIII elements in terms of oxides is 0.1-15, preferably 0.8-9; and / or An oxide content of the hydrogenation active elements in the first catalyst and the second catalyst each is not less than 30 g / L, preferably 30-600 g / L, more preferably 40-560 g / L, based on a total mass of the un-sulfided catalyst; Preferably, The Group VI B elements in the first catalyst are selected from Mo; and / or the Group VI B elements in the second catalyst are selected from Mo and / or W, preferably Mo and W, and a mass ratio of Mo and W in terms of oxides is 0.25-4; and / or The Group VIII elements in the first catalyst and the second catalyst each independently are selected from at least one of Ni, Co, Fe, Pt, and Pd; and / or The lanthanide metals in the first catalyst and the second catalyst each independently are selected from La and / or Ce; and / or The Group VA elements in the first catalyst and the second catalyst are P.

4. The method according to any one of claims 1-3, wherein The first catalyst and the second catalyst each have an average pore diameter of 7-20 nm; and / or The first and second catalysts each have a pore volume of 0.5 to 0.98 cm 3 / g; and / or The specific surface area of the first catalyst and the second catalyst is each 190-380 m 2 / g.

5. The method of any of claims 1-4, wherein, An inlet temperature of the first hydrogenation reaction is at least 15°C, preferably 15-60°C, lower than an inlet temperature of the second hydrogenation reaction; more preferably, An outlet temperature of the first hydrogenation reaction is 260-330°C, preferably 270-325°C; An outlet temperature of the second hydrogenation reaction is 330-380°C, preferably 340-365°C.

6. The method according to any one of claims 1-5, wherein The conditions of the first hydrogenation reaction comprise: the temperature of the stream before hydrogenation is 200-260°C, preferably 220-250°C; the temperature of the stream after hydrogenation is higher than the temperature of the stream before hydrogenation by at most 85°C, preferably 50-85°C, more preferably 55-80°C; and / or the pressure is 5.0-12.0 MPa; and / or the hydrogen / oil volume ratio is 400-3000; and / or Liquid hourly space velocity is 0.2-3 h -1 ; and / or the conditions of the second hydrogenation reaction include: the temperature of the stream before hydrogenation is 260-330°C, preferably 280-325°C; the temperature of the stream after hydrogenation is higher than the temperature of the stream before hydrogenation by at most 55°C, preferably 20-55°C, more preferably 25-55°C; and / or the pressure is 5.0-12.0 MPa; and / or the hydrogen / oil volume ratio is 400-3000; and / or Liquid hourly space velocity is 0.2-3 h -1 .

7. The method of any of claims 1-6, wherein, a part of the product stream after the first hydrogenation reaction is recycled back to the first hydrogenation reaction process, and another part is sent to the second hydrogenation reaction process, wherein the ratio of the mass of the product stream recycled back to the first hydrogenation reaction process to the mass of the polycyclic aromatic hydrocarbon feedstock is 5-100%, preferably 10-50%.

8. The method of any one of claims 1-7, wherein, The second hydrogenation reaction requires fresh hydrogen, preferably the volume of the fresh hydrogen supplied is 5-50% of the volume of the hydrogen in the first hydrogenation reaction, preferably 10-30%.

9. The method according to any one of claims 1-8, wherein the polycyclic aromatic hydrocarbon feedstock and the hydrogen feedstock before the first hydrogenation reaction are preheated to 220-250°C; and / or the first hydrogenation reaction is carried out in a catalytic system containing a first catalyst and a protective agent, the protective agent comprising a hydrogenation active component and an oxide carrier, the content of the hydrogenation active component in terms of oxide is 5-80 g / L; preferably, the volume ratio of the first catalyst to the protective agent in the catalytic system is 4-25; more preferably, the oxide content of the hydrogenation active element in the first catalyst is at least 400 g / L higher than the oxide content of the hydrogenation active component in the protective agent, preferably 400-550 g / L.

10. The method according to claim 9, wherein the hydrogenation active component in the protective agent comprises a Group VIB element and / or a Group VIII element, and the oxide carrier is Al2O3; and / or the average pore size of the protective agent is 15-30 nm; and / or The pore volume of the protective agent is 0.7-0.98 cm 3 / g; and / or The specific surface area of the protective agent is 120-250 m 2 / g; preferably, the Group VIB element in the protective agent is selected from Mo and / or W; and / or the Group VIII element in the protective agent is selected from at least one of Ni, Co, Fe, Pt and Pd.

11. The method of any of claims 1-10, wherein, The content of the polycyclic aromatic hydrocarbons in the polycyclic aromatic hydrocarbon feedstock is not less than 65 wt%, preferably 75-92 wt%; preferably, the polycyclic aromatic hydrocarbons include at least one of naphthalene compounds, acenaphthene compounds, anthracene compounds and phenanthrene compounds, more preferably acenaphthene compounds.

12. Use of the method according to any one of claims 1-11 in the hydrogenation of poor quality aromatic-rich distillate oil.

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