Method for producing high-octane gasoline by hydrocatalysis of catalyst with stable initial activity

Abstract

The invention relates to the field of catalytic diesel conversion, and discloses a method for producing high-octane gasoline by hydrocatalysis of a catalyst with stable initial activity. The method comprises the following steps: carrying out activity stabilization on a hydrofining catalyst and a hydrocracking catalyst in the presence of active stable raw oil, hydrogen and C2-C6 alkane, filling the upstream of the hydrocracking catalyst with the hydrofining catalyst along a material flow direction, and filling n catalyst bed layers connected in series with the hydrofining catalyst and the hydrocracking catalyst, c2-C6 alkanes are injected into the n catalyst bed layers connected in series respectively, and n is a positive integer and is larger than or equal to 2; switching the active stable raw oil into catalytic diesel oil, and reducing the injection amount of C2-C6 alkanes until the injection amount is stopped; a mixed material containing catalytic diesel oil and C2-C6 alkane is sequentially contacted with a hydrofining catalyst and a hydrocracking catalyst which are stable in activity, and gasoline is obtained. The method greatly shortens the octane number standard period of gasoline products at the initial stage of operation.

Description

Technical Field

[0001] The present invention relates to the field of catalytic diesel conversion, and in particular to a method for producing high-octane gasoline by utilizing a catalyst with stable initial activity under hydrogen catalysis. Background Art

[0002] Crude oil prices have fluctuated more frequently in recent years, leading to a surge in demand for heavy crude oil. During high oil prices, heavy crude oil is widely used in the production of various fuels and chemical products, despite higher refining costs, to offset the rising raw material costs associated with high oil prices.

[0003] Catalytic cracking (FCC) technology occupies an important position in the refining and chemical industry in various countries around the world. It can convert high-boiling-point, high-molecular-weight hydrocarbon components in crude oil into more valuable gasoline, olefin gas and other products. In recent years, as the raw materials processed by catalytic cracking units have become increasingly heavy and inferior, the quality of catalytic cracking products, especially catalytic cracking light cycle oil (LCO), has further deteriorated. Catalytic cracking light cycle oil has high sulfur and aromatic content and poor engine ignition performance. It is mainly used abroad to blend fuel oil, non-automotive diesel and heating oil. How to convert aromatics in LCO into high-value-added high-octane fuel has been the direction of research for many years. At present, this field mainly adopts the method of moderate cracking and shallow hydrogenation to convert LCO into high-octane fuel, but there is a problem that the depth of hydrogenation affects the amount of octane loss during catalytic diesel hydrogenation conversion.

[0004] Reports indicate that hydrocracking processes can convert catalytically cracked light cycle oil (FCLO) into ultra-low sulfur diesel and high-octane gasoline blending components. Examples include the MAK-LCO technology developed by Mobil, Akzo Nobel, and MW Kellogg; and the LCO Unicracking™ technology and HC-190 specialized catalyst developed by UOP. These technologies can convert low-value FCLO components into high-octane gasoline and high-quality diesel blending components. Furthermore, the new LCO-X technology developed by UOP offers a novel approach to increasing aromatics production from FCLO through a hydroconversion-selective transalkylation route. This demonstrates that the utilization of relatively low-value FCLO has evolved from initial upgrading to the production of higher-value oil and chemical products.

[0005] Currently, in the early stages of operation of domestic catalytic diesel hydroconversion units, due to issues with the initial catalyst activity, the catalyst system exhibits strong hydrodesulfurization, denitrogenation, and dearomatization capabilities, while the hydrocracking catalyst also exhibits high hydrogenation performance. However, the initial mismatch between hydrogenation and cracking activity results in a low octane rating for the gasoline components, failing to meet production requirements. Producing high-octane gasoline requires a period of stable operation to meet the requirements.

[0006] CN111100696A discloses a method for catalytic diesel hydroconversion to produce high-octane gasoline. The method uses a single-stage cascade process, with catalysts of varying degrees of sulfidation (DFS) graded in the reaction zone. Highly aromatic catalytic diesel is mixed with recycled hydrogen and then introduced into the hydrorefining reaction zone for reaction. The DFS of the catalyst decreases along the flow direction, and the resulting oil then enters the hydrocracking reaction zone where it contacts at least two DFS catalysts for conversion. This method addresses the issue of activity matching and introduces catalysts with varying degrees of sulfidation. However, the preparation process for obtaining catalysts with varying degrees of sulfidation is complex, and there is also the problem of resulfurization of the active metals of the catalyst with low DFS with the hydrogen sulfide produced by the desulfurization reaction. The low-DFS catalyst is then converted into a high-DFS catalyst, further increasing its activity and causing excessive cracking.

[0007] CN111088073A discloses a method for the hydrocracking of catalytic diesel. This method employs a two-stage hydrocracking process, controlling the H2S and NH3 concentrations in the hydrofining reactor and the hydrocracking reactor in stages to shorten the initial stabilization period and extend the unit's operating cycle. However, it fails to consider the synergistic effects of the catalyst hydrogenation and cracking reactions.

[0008] CN104611050A discloses a method for converting catalytic cracking diesel. This method utilizes a single-stage cascade process, where the effluent from the hydrorefining reaction enters the hydrocracking reactor directly, where it reacts with the graded catalyst bed within the reactor. The hydrogenation activity of the hydrocracking catalyst decreases as the reactants flow. This method can reduce over-hydrogenation and secondary cracking of cracked naphtha while maintaining the diesel hydrocracking effect, lowering chemical hydrogen consumption and thereby increasing the naphtha's octane number and liquid yield. However, this method fails to address the issue of high catalyst activity during the initial stages of operation, preventing a reduction in setup time. Summary of the Invention

[0009] The present invention aims to overcome the problems of the prior art by providing a method for producing high-octane gasoline in the presence of hydrogen using a catalyst with stable initial activity. This method can significantly shorten the period of time required for the gasoline product to reach the octane standard during initial operation.

[0010] In order to achieve the above object, the present invention provides a method for producing gasoline by hydrogenation using a catalyst with stable initial activity, the method comprising the following steps:

[0011] (1) Under the action of active stable feedstock oil, hydrogen and C2-C6 alkanes, an activity stabilization process is performed on a hydrorefining catalyst and a hydrocracking catalyst, wherein, along the logistics direction, the hydrorefining catalyst is loaded upstream of the hydrocracking catalyst, the hydrorefining catalyst and the hydrocracking catalyst are loaded in n catalyst beds connected in series, and C2-C6 alkanes are respectively injected into the n catalyst beds connected in series, where n is a positive integer and n≥2;

[0012] (2) After the active stabilization process is completed, the active stabilization feedstock oil is switched to catalytic diesel, and the injection amount of C2-C6 alkanes is reduced until it stops;

[0013] (3) The mixture containing catalytic diesel and C2-C6 alkanes and hydrogen are contacted with an activity-stabilized hydrotreating catalyst and an activity-stabilized hydrocracking catalyst in sequence, and the hydrotreating reaction and the hydrocracking reaction are respectively carried out to obtain gasoline.

[0014] Through the above technical solution, the beneficial effects of the present invention include:

[0015] In the method provided by the present invention, a specific activity stabilization process is performed on the hydrorefining catalyst and the hydrocracking catalyst to regulate their initial activities, so that the activities of the hydrorefining catalyst and the hydrocracking catalyst are matched, thereby greatly shortening the time it takes for the octane number of the gasoline product to reach the octane number requirement of the national standard No. 92 gasoline in the initial stage of operation. DETAILED DESCRIPTION

[0016] 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.

[0017] In one aspect, the present invention provides a method for producing gasoline in the presence of hydrogen using a catalyst with stable initial activity, the method comprising the following steps:

[0018] (1) Under the action of active stable feedstock oil, hydrogen and C2-C6 alkanes, an activity stabilization process is performed on a hydrorefining catalyst and a hydrocracking catalyst, wherein, along the logistics direction, the hydrorefining catalyst is loaded upstream of the hydrocracking catalyst, the hydrorefining catalyst and the hydrocracking catalyst are loaded in n catalyst beds connected in series, and C2-C6 alkanes are respectively injected into the n catalyst beds connected in series, where n is a positive integer and n≥2;

[0019] (2) After the active stabilization process is completed, the active stabilization feedstock oil is switched to catalytic diesel, and the injection amount of C2-C6 alkanes is reduced until it stops;

[0020] (3) The mixture containing catalytic diesel and C2-C6 alkanes and hydrogen are contacted with an activity-stabilized hydrotreating catalyst and an activity-stabilized hydrocracking catalyst in sequence, and the hydrotreating reaction and the hydrocracking reaction are respectively carried out to obtain gasoline.

[0021] According to the present invention, preferably, the C2-C6 alkane is selected from at least one of ethane, propane, n-butane, isobutane, n-pentane, isopentane, neopentane, cyclopentane, n-hexane, cyclohexane, methylcyclopentane, 2-methylpentane, 3-methylpentane, 2,2-methylbutane and 2,3-dimethylbutane, preferably at least one of propane, n-butane, isobutane, n-pentane and isopentane. The use of the above-mentioned specific type of C2-C6 alkane, which is in the gas phase under certain conditions, regulates the reaction environment in the reaction zone, promotes moderate carbon deposition of the highly active catalyst, achieves the goal of activity regulation, and can better dissolve in the active and stable feedstock oil under low temperature conditions, which is more conducive to the step-by-step carbon deposition of the catalyst, better adjusts the initial activity of the catalyst, and ensures that the catalyst hydrogenation and conversion performance match.

[0022] In the present invention, the injection rate of C2-C6 alkanes into the n catalyst beds can be equal or gradually reduced. Preferably, along the flow direction, the injection rate of C2-C6 alkanes into the n catalyst beds gradually decreases relative to the volume of C2-C6 alkanes injected into the previous bed. This preferred embodiment facilitates modulation of the initial catalyst activity in each reaction zone, achieving activity matching between the hydrorefining catalyst and the hydroconversion catalyst, and shortening the time it takes for the gasoline product to reach a qualified octane number during the initial operation.

[0023] It should be noted that when the injection amount of C2-C6 alkanes in n catalyst beds gradually increases, there will be a problem that the passivation atmosphere of the catalyst system is concentrated in the cracking catalyst part, and the refined catalyst is insufficiently passivated, resulting in poor overall passivation effect of the catalyst system activity, and a longer passivation period is required.

[0024] Preferably, along the material flow direction, the injection volume of C2-C6 alkanes in each of the n catalyst beds is sequentially reduced by 10-55% by volume relative to the injection volume of C2-C6 alkanes in the previous bed. For example, the injection volume may be 10% by volume, 15% by volume, 20% by volume, 25% by volume, 30% by volume, 35% by volume, 40% by volume, 45% by volume, 50% by volume, 55% by volume, or any value within a range consisting of any two of these values. This preferred embodiment further facilitates the control of the initial catalyst activity in each zone.

[0025] It should be noted that the percentage of reduction of the above-mentioned C2-C6 alkanes is calculated as follows:

[0026] The percentage of reduction of C2-C6 alkanes in the current catalyst bed = (the injection volume of C2-C6 alkanes in the previous catalyst bed - the injection volume of C2-C6 alkanes in the current catalyst bed) / the injection volume of C2-C6 alkanes in the previous catalyst bed × 100%.

[0027] For example, the injection rate of C2-C6 alkanes in the second catalyst bed is 20% of the total injection rate of C2-C6 alkanes, and the injection rate of C2-C6 alkanes in the first catalyst bed is 25% of the total injection rate of C2-C6 alkanes, which is reduced by 20%.

[0028] In the present invention, for different catalyst beds, the reduction percentages of the C2-C6 alkanes may be equal or unequal.

[0029] According to the present invention, preferably, along the logistics direction, the injection amount of C2-C6 alkanes into the first catalyst bed is 20-40% by volume, based on the total volume of injected C2-C6 alkanes. For example, it can be 20% by volume, 23% by volume, 25% by volume, 28% by volume, 30% by volume, 33% by volume, 35% by volume, 38% by volume, 40% by volume, or any value within a range consisting of any two of these values. This preferred embodiment is beneficial for further enhancing the initial activity stabilization effect.

[0030] According to the present invention, preferably, along the logistics direction, the active stable feedstock oil and hydrogen flow from the first catalyst bed to the nth catalyst bed.

[0031] It is understandable that the active stable feedstock oil and hydrogen are introduced in a different manner from the C2-C6 alkanes.

[0032] In the present invention, the hydrorefining catalyst and hydrocracking catalyst can be separately loaded into two active and stable reactors connected in series, or they can be loaded into a single active and stable reactor. To reduce equipment modification investment, minimize process flow changes, and lower engineering costs, the hydrorefining catalyst and hydrocracking catalyst are preferably loaded into a single active and stable reactor.

[0033] The present invention does not particularly limit the type of activity stabilization reactor to be used, and various activity stabilization reactors that can achieve the activity stabilization process described in the present invention can be used. The present invention preferably uses a hydrogenation reactor as the activity stabilization reactor. Preferably, the activity stabilization process is carried out in a hydrogenation reactor, more preferably in a trickle bed reactor. This preferred embodiment is directly compatible with existing hydrogenation reaction process equipment, achieving process continuity, further simplifying and shortening the period for the octane number of gasoline products to meet qualified requirements in the initial stage of operation, and reducing enterprise transformation investment, saving cost investment, and obtaining higher economic returns.

[0034] According to the present invention, preferably, n is 2-18, preferably 4-11, and n is a positive integer. This preferred embodiment can effectively solve the problems of uneven activity modulation and poor activity synchronization synergy during the initial activity stabilization of the catalyst, achieving excellent results.

[0035] The present invention has no particular limitation on the specific number of loading beds of the hydrorefining catalyst and the specific number of loading beds of the hydrocracking catalyst, which can be appropriately selected according to different application scenarios and actual needs.

[0036] According to the present invention, preferably, the number of packed beds of the hydrotreating catalyst is 1-8, preferably 2-5.

[0037] According to the present invention, preferably, the number of packed beds of the hydrocracking catalyst is 2-10, preferably 2-6.

[0038] According to the present invention, preferably, the hydrotreating catalyst and the hydrocracking catalyst in step (1) are both catalysts that have been sulfided.

[0039] The present invention does not particularly limit the method of the vulcanization treatment, which may be a pre-vulcanization method conventionally used in the art, or a true vulcanization method or a full vulcanization method conventionally used in the art, and the present invention does not particularly limit this.

[0040] According to the present invention, preferably, the active stable feedstock oil has an initial distillation point of 150-300°C, preferably 150-250°C; and a final distillation point of 320-450°C, preferably 370-420°C.

[0041] According to the present invention, preferably, the density of the active stable feedstock oil is 0.84-0.89 g / cm 3 .

[0042] According to the present invention, preferably, the sulfur content of the active stable feedstock oil is not less than 6000 mg / g, preferably 9000-15000 mg / g.

[0043] According to the present invention, preferably, the organic nitrogen content in the active stable feedstock oil is not less than 50 μg / g, preferably 80-200 μg / g.

[0044] The use of active and stable feedstock oil with the above characteristics is more conducive to achieving initial activity modulation of the catalyst.

[0045] The present invention provides a wide range of options for the specific type of active, stable feedstock oil; any feedstock oil meeting the aforementioned characteristic requirements can be used. Preferably, the active, stable feedstock oil is a heavy diesel fraction, preferably selected from at least one of conventional second-tier oil, conventional third-tier oil, conventional fourth-tier oil, top-cut oil, first-tier oil, and straight-run diesel, and more preferably selected from at least one of conventional fourth-tier oil, top-cut oil, and first-tier oil. This preferred embodiment further facilitates fully utilizing the active, stable feedstock oil's high proportion of heavier components and low coke-prone materials, fully maximizing the catalyst activity and achieving appropriate control of catalyst activity.

[0046] According to the present invention, preferably, the conditions of the active stabilization process in step (1) include: a temperature of 310-400°C, preferably 320-370°C; a volume space velocity of 0.5-7h -1 , preferably 0.8-6h -1 ; The hydrogen-oil volume ratio is 300-2000:1, preferably 600-1500:1; the total pressure is 2.5-20MPa, preferably 6-15MPa; the time is 25-145h, preferably 40-90h.

[0047] The adoption of this preferred embodiment is more conducive to the catalyst achieving a process of stabilizing the synergistic initial activity of the hydrorefining catalyst and the hydroconversion catalyst according to the actual reaction conditions in the reaction zone.

[0048] The present invention does not specifically limit the loading ratio of the hydrorefining catalyst to the hydrocracking catalyst during the initial activity stabilization process. The loading ratio should be appropriately selected based on the needs of the subsequent hydrogenation reaction and can be determined by reference to conventional methods in the art. Generally, the loading ratio is adjusted based on the amount of catalyst required for the subsequent reaction and the amount of catalyst required for the initial activity stabilization. Preferably, the loading volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 7:3 to 3:7, and more preferably 6:4 to 4:6.

[0049] According to the present invention, preferably, the activity stabilization process in step (1) includes a first activity stabilization process and a second activity stabilization process.

[0050] According to the present invention, preferably, during the first activity stabilization process, the volume of C2-C6 alkanes is 10-18% of the volume of hydrogen, preferably 10-16%.

[0051] According to the present invention, preferably, the time of the first active stabilization process is 20-72 hours, preferably 24-50 hours.

[0052] According to the present invention, preferably, during the second active stabilization process, the volume of C2-C6 alkanes is 20-35% of the volume of hydrogen, preferably 20-30%.

[0053] According to the present invention, preferably, the time of the second active stabilization process is 16-60 hours, preferably 20-35 hours.

[0054] The use of the above-mentioned two specific initial activity stabilization periods is more conducive to achieving catalyst activity stability, more conducive to regulating hydrogenation activity, and achieving matching of the activities of the hydrorefining catalyst and the hydrocracking catalyst, thereby achieving the purpose of shortening the time it takes for the gasoline product octane number to reach the qualified level in the initial operation.

[0055] In step (2) of the present invention, the active stabilized feedstock oil is switched to catalytic diesel. In order to ensure the safety and controllability of the initial active stabilization process, including further controlling the impact of raw material degradation on the catalyst carbon deposition process after processing the catalytic diesel, the present invention preferably switches the active stabilized feedstock oil to catalytic diesel in batches, while simultaneously reducing the amount of C2-C6 alkanes in batches. Specifically, based on the total weight of the active stabilized feedstock oil and catalytic diesel, the increase in the proportion of catalytic diesel in each switch is 10-35% by weight.

[0056] It should be noted that the above-mentioned increase ratio is the difference between the percentage of catalytic diesel fuel in the current switch and the percentage of catalytic diesel fuel in the previous switch. The percentage refers to the mass content of catalytic diesel fuel based on the total weight of active stabilized feedstock oil and catalytic diesel fuel.

[0057] For example: based on the total weight of active stabilized crude oil and catalytic diesel, the injection amount of catalytic diesel during the second switching is 20%, and the injection amount of catalytic diesel during the first switching is 15%, and the increase ratio is 5%.

[0058] Preferably, based on the volume of hydrogen, the reduction ratio of C2-C6 alkanes in each switch is 2-8 volume %.

[0059] It should be noted that the above reduction ratio is the difference between the C2-C6 alkane injection percentage of the previous switch and the C2-C6 alkane injection percentage of the current switch. The percentage refers to the percentage of the injected volume of C2-C6 alkane to the volume of hydrogen.

[0060] For example, based on the volume of hydrogen, the injection amount of C2-C6 alkanes during the second switching is 20%, and the injection amount of C2-C6 alkanes during the first switching is 25%, and the reduction ratio is 5%.

[0061] Preferably, the time interval between two adjacent switches of catalytic diesel is 5-15 hours.

[0062] The present invention has no particular limitation on the number of switching times, as long as 100% switching to catalytic diesel is achieved and the C2-C6 alkane content meets the following requirements.

[0063] When switching 100% to catalytic diesel, the C2-C6 alkane content must be controlled within the specified levels described below. This helps further adjust the initial catalyst activity, avoiding excessive initial activity, oversaturation, and cracking, thereby rapidly adjusting catalyst activity to match the feedstock. Preferably, the volume of C2-C6 alkanes in the mixture is 10-20% of the volume of hydrogen.

[0064] It is understood that the volume of C2-C6 alkanes herein refers to the volume of C2-C6 alkanes when the fuel is switched to 100% catalytic diesel.

[0065] Preferably, the active stable feedstock is switched to 100% catalytic diesel and then the injection of C2-C6 alkanes is stopped.

[0066] The present invention does not particularly limit the type of the hydrotreating catalyst, and any conventional catalyst in the art may be used. Preferably, the hydrotreating catalyst comprises a first support and a first metal component supported on the first support, wherein the first support comprises a heat-resistant inorganic oxide, and the first metal component comprises a Group VIB metal component and a Group VIII metal component.

[0067] According to the present invention, preferably, the heat-resistant inorganic oxide is at least one selected from silicon oxide, aluminum oxide, amorphous silicon aluminum, zirconium oxide and titanium oxide.

[0068] According to the present invention, preferably, based on the total weight of the hydrotreating catalyst, the content of the first metal component in the hydrotreating catalyst, calculated as oxide, is 15-50 wt%, preferably 18-45 wt%.

[0069] According to the present invention, preferably, based on the total weight of the hydrotreating catalyst, the content of the Group VIII metal component in the hydrotreating catalyst, calculated as oxide, is 2-8 wt%, preferably 2.5-6 wt%.

[0070] According to the present invention, preferably, the Group VIB metal component is Mo and / or W.

[0071] According to the present invention, preferably, the Group VIII metal component is Ni and / or Co.

[0072] The source of the hydrotreating catalyst of the present invention is not particularly limited and can be obtained commercially or prepared by conventional methods.

[0073] The present invention does not particularly limit the type of the hydrocracking catalyst, and any conventional choice in the art may be used. Preferably, the hydrocracking catalyst comprises a second support and a second metal component supported on the second support, wherein the second support comprises a Y-type molecular sieve, and the second metal component comprises a Group VIB metal component and a Group VIII metal component.

[0074] In the present invention, the Y-type molecular sieve may be a pure Y molecular sieve or a modified Y molecular sieve, which may be selected conventionally in the art.

[0075] According to the present invention, preferably, based on the total weight of the hydrocracking catalyst, the content of the second metal component in the hydrocracking catalyst, calculated as oxide, is 3-50 wt%, preferably 18-40 wt%.

[0076] According to the present invention, preferably, based on the total weight of the hydrocracking catalyst, the content of the Group VIII metal component in the hydrocracking catalyst, calculated as oxide, is 1.5-8 wt%, preferably 2-6 wt%.

[0077] According to the present invention, preferably, the Group VIB metal component is Mo and / or W.

[0078] According to the present invention, preferably, the Group VIII metal component is Ni and / or Co.

[0079] The source of the hydrocracking catalyst of the present invention is not particularly limited and can be obtained commercially or prepared by conventional methods.

[0080] Preferably, the initial boiling point of the catalytic diesel is 65-290°C, preferably 100-200°C; and the final boiling point is 330-450°C.

[0081] Preferably, the aromatics content in the catalytic diesel is ≥50 wt%, preferably 50-75 wt%.

[0082] Further preferably, the polycyclic aromatic hydrocarbons content in the catalytic diesel is ≥35 wt%, preferably 40-50 wt%.

[0083] The present invention has no particular limitation on the specific type of the catalytic diesel, which can be a product of various catalytic cracking processes obtained by processing any base oil, as long as it meets the above characteristics.

[0084] The conditions of the hydrofining reaction of the present invention can be carried out according to conventional methods in the art. Preferably, the conditions of the hydrofining reaction include: a reaction temperature of 260-420°C, preferably 310-390°C; a volume space velocity of 0.4-10h -1 , preferably 0.6-5h -1 ; The hydrogen-oil volume ratio is 200-2000:1, preferably 500-1500:1; the reaction pressure is 2.5-20MPa, preferably 6-15MPa.

[0085] The conditions of the hydrocracking reaction of the present invention can be carried out according to conventional methods in the art. Further, the conditions of the hydrocracking reaction include: a reaction temperature of 350-450°C, preferably 355-420°C; a volume space velocity of 0.4-16h -1 , preferably 0.6-6h -1 , the hydrogen-oil volume ratio is 200-2000:1, preferably 500-1500:1; the reaction pressure is 2.5-20MPa, preferably 6-15MPa.

[0086] The method of the present invention can produce a product that meets the octane number requirements of the national standard No. 92 gasoline. Preferably, the sulfur content of the gasoline product is ≯10 μg / g, and the research octane number is ≮92.

[0087] The present invention will be described in detail below through examples.

[0088] In the following examples, the hydrorefining catalyst used was FHUDS-8, produced by the Fushun Branch of Sinopec Catalyst Company; and the hydrocracking catalyst used was FC-24, produced by the Fushun Branch of Sinopec Catalyst Company. The FHUDS-8 catalyst uses alumina as a support and Mo-Ni as the active component; the FC-24 catalyst uses a modified Y-type molecular sieve as a support and W-Ni as the active metal component.

[0089] Example 1

[0090] A conventional trickle-bed hydrogenation reactor was used, loaded with sulfided FHUDS-8 and FC-24 catalysts. The hydrorefining catalyst was loaded upstream of the hydrocracking catalyst in the direction of flow. The hydrorefining and hydrocracking catalysts were combined in five catalyst beds connected in series: two beds of hydrorefining catalyst and three beds of hydrocracking catalyst. The volume ratio of FHUDS-8 to FC-24 catalyst was 5:5. During the first stage of initial activation and stabilization, the active and stabilized feedstock oil-1 (minus first-line oil) and hydrogen in Table 1 were introduced into the reactor and flowed from the first catalyst bed to the fifth catalyst bed along the logistics direction. The volume of n-butane accounted for 15% of the volume of hydrogen. Along the logistics direction, the n-butane injection volume ratio in the five beds was 6:5:4:3:2, and the n-butane was injected into the inlet of each bed respectively. The reaction temperature was adjusted to 345°C for the hydrorefining catalyst bed and 355°C for the hydrocracking catalyst bed, and the volumetric space velocity was controlled to 0.9h-1. -1 The hydrogen-to-oil volume ratio was 1000:1, the total pressure was 8.0 MPa, and the temperature was kept constant for 24 hours. Then, the second initial activation stabilization process was carried out, with the volume of n-butane accounting for 25% of the volume of hydrogen. Other conditions remained unchanged and the temperature was kept constant for 26 hours.

[0091] After the initial activation and stabilization of the catalyst, catalytic diesel (composition see Table 1) was switched in batches. Based on the total weight of the active and stable feedstock oil and catalytic diesel, 25% catalytic diesel was switched in, the ratio of n-butane volume to hydrogen volume was reduced to 22%, and the temperature was kept constant for 6 hours; the catalytic diesel ratio was increased to 50%, and the ratio of n-butane volume to hydrogen volume was reduced to 19%, and the temperature was kept constant for 6 hours; the catalytic diesel ratio was increased to 75%, and the ratio of n-butane volume to hydrogen volume was reduced to 16%, and the temperature was kept constant for 6 hours; the catalytic diesel ratio was increased to 100%, and the ratio of n-butane volume to hydrogen volume was reduced to 13%, and the temperature was kept constant for 6 hours. After this process was completed, the addition of n-butane was stopped, and subsequent hydrofining and hydrocracking reactions were carried out. The reaction conditions of the hydrofining reaction and the hydrocracking reaction are shown in Table 2. The results are shown in Table 2.

[0092] Example 2

[0093] A conventional trickle-bed hydrogenation reactor was used, loaded with sulfided FHUDS-8 and FC-24 catalysts. The hydrorefining catalyst was loaded upstream of the hydrocracking catalyst along the flow direction. The hydrorefining and hydrocracking catalysts were loaded into a total of six catalyst beds connected in series, two for the hydrorefining catalyst and four for the hydrocracking catalyst. The volume ratio of FHUDS-8 to FC-24 catalyst was 4:6. The first stage of initial activation and stabilization was carried out. The active and stabilized feedstock oil-1 in Table 1 and hydrogen entered the hydrogenation reactor. Along the logistics direction, they flowed from the first catalyst bed to the sixth catalyst bed. The volume of n-butane accounted for 12% of the volume of hydrogen. The injection volume ratio of n-butane in the six beds was 6:5:4:3:2:1. The n-butane was injected into the inlet of each bed respectively. The reaction temperature was adjusted to 343°C in the hydrorefining catalyst loading zone and 353°C in the hydrocracking catalyst loading zone. The volumetric space velocity was controlled to 0.9h -1 The hydrogen-to-oil volume ratio is 1000:1, the total pressure is 8.0 MPa, and the temperature is kept constant for 48 hours. Then the second initial activation stabilization process is carried out, with the volume of n-butane accounting for 30% of the volume of hydrogen. Other conditions remain unchanged and the temperature is kept constant for another 24 hours.

[0094] After the initial activation and stabilization of the catalyst, the catalyst was switched to catalytic diesel (composition see Table 1) in batches. Based on the total weight of the active, stabilized feedstock oil and catalytic diesel, the catalyst was switched to 20% catalytic diesel, and the ratio of n-butane to hydrogen by volume was reduced to 26% and held at a constant temperature for 10 hours. The catalytic diesel ratio was increased to 40%, and the ratio of n-butane to hydrogen by volume was reduced to 22% and held at a constant temperature for 10 hours. The catalytic diesel ratio was increased to 60%, and the ratio of n-butane to hydrogen by volume was reduced to 18% and held at a constant temperature for 10 hours. The catalytic diesel ratio was increased to 80%, and the ratio of n-butane to hydrogen by volume was reduced to 14% and held at a constant temperature for 10 hours. The catalytic diesel ratio was increased to 100%, and the ratio of n-butane to hydrogen by volume was reduced to 10% and held at a constant temperature for 10 hours. After this process was completed, n-butane addition was stopped, and subsequent hydrofinishing and hydrocracking reactions were carried out. The reaction conditions for the hydrofinishing and hydrocracking reactions are shown in Table 2. The results are shown in Table 2.

[0095] Example 3

[0096] The method of Example 1 is followed, except that isobutane is used instead of n-butane as the C2-C6 alkane.

[0097] The results are shown in Table 2.

[0098] Example 4

[0099] The method of Example 1 was followed, except that the active stabilized feedstock oil was changed to conventional third-line oil. The specific composition is shown in Active Stabilized Feedstock Oil-2 in Table 1.

[0100] The results are shown in Table 2.

[0101] Example 5

[0102] The method of Example 1 was followed, except that the injection volumes of n-butane in the five catalyst beds were equal (ie, n-butane was equally divided).

[0103] The results are shown in Table 2.

[0104] Comparative Example 1

[0105] The method of Example 1 was followed, except that C2-C6 alkane (n-butane) was not used. Specifically:

[0106] Active stable feedstock oil and hydrogen enter the reactor, and the reaction temperature is adjusted to 345℃ for the hydrorefining catalyst bed and 355℃ for the hydrocracking catalyst bed. The volumetric space velocity is controlled to 0.9h -1 , the volume ratio of hydrogen to oil is 1000:1, and the constant temperature operation is 40h.

[0107] After the initial activation and stabilization of the catalyst, the catalytic diesel fuel was switched in batches. Based on the total weight of the active, stabilized feedstock oil and catalytic diesel fuel, the fuel was switched to 25% catalytic diesel fuel and held at a constant temperature for 6 hours. The catalytic diesel fuel ratio was then increased to 50% and held at a constant temperature for 6 hours. The catalytic diesel fuel ratio was increased to 75% and held at a constant temperature for 6 hours. The catalytic diesel fuel ratio was increased to 100% and held at a constant temperature for 6 hours. After this process was completed, the subsequent hydrofinishing and hydrocracking reactions were carried out. The reaction conditions for the hydrofinishing and hydrocracking reactions are shown in Table 2. The results are shown in Table 2.

[0108] Table 1

[0109]

[0110] Table 2

[0111]

[0112]

[0113] As can be seen from the results in Table 2, the method of the present invention can significantly shorten the gasoline product qualification period in the early stage of operation by performing a specific activity stabilization process on the hydrotreating catalyst and the hydrocracking catalyst, thereby regulating the initial activities of the two catalysts and matching their activities.

[0114] 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 producing gasoline by hydrogenation using a catalyst with stable initial activity, characterized in that: The method comprises the following steps: (1) Under the action of active stable feedstock oil, hydrogen and C2-C6 alkanes, an activity stabilization process is performed on a hydrorefining catalyst and a hydrocracking catalyst, wherein, along the logistics direction, the hydrorefining catalyst is loaded upstream of the hydrocracking catalyst, the hydrorefining catalyst and the hydrocracking catalyst are loaded in n catalyst beds connected in series, and C2-C6 alkanes are respectively injected into the n catalyst beds connected in series, where n is a positive integer and n≥2; (2) After the active stabilization process is completed, the active stabilization feedstock oil is switched to catalytic diesel, and the injection amount of C2-C6 alkanes is reduced until it stops; (3) The mixture containing catalytic diesel and C2-C6 alkanes and hydrogen are contacted with an activity-stabilized hydrotreating catalyst and an activity-stabilized hydrocracking catalyst in sequence, and the hydrotreating reaction and the hydrocracking reaction are respectively carried out to obtain gasoline.

2. The method according to claim 1, wherein The C2-C6 alkane is selected from at least one of ethane, propane, n-butane, isobutane, n-pentane, isopentane, neopentane, cyclopentane, n-hexane, cyclohexane, methylcyclopentane, 2-methylpentane, 3-methylpentane, 2,2-methylbutane and 2,3-dimethylbutane, and is preferably selected from at least one of propane, n-butane, isobutane, n-pentane and isopentane.

3. The method according to claim 1, wherein Along the logistics direction, the injection volume of C2-C6 alkanes in n catalyst beds is gradually reduced relative to the injection volume of C2-C6 alkanes in the previous bed, preferably reduced by 10-55 volume% in sequence; Preferably, along the logistics direction, based on the total volume of C2-C6 alkanes injected, the injection amount of C2-C6 alkanes in the first catalyst bed is 20-40% by volume; Preferably, along the logistics direction, the active stable feedstock oil and hydrogen flow from the first catalyst bed to the nth catalyst bed.

4. The method according to any one of claims 1 to 3, wherein: n is 2-18, preferably 4-11, and n is a positive integer; Preferably, the number of packed beds of the hydrotreating catalyst is 1-8, preferably 2-5; Preferably, the number of packed beds of the hydrocracking catalyst is 2-10, preferably 2-6; Preferably, the hydrotreating catalyst and the hydrocracking catalyst in step (1) are both catalysts that have been sulfided.

5. The method according to any one of claims 1 to 4, wherein: The active stable feedstock oil has an initial boiling point of 150-300°C, preferably 150-250°C; and a final boiling point of 320-450°C, preferably 370-420°C; Preferably, the density of the active stable feedstock oil is 0.84-0.89 g / cm 3 ; Preferably, the sulfur content of the active stable feedstock oil is not less than 6000 mg / g, preferably 9000-15000 mg / g; Preferably, the organic nitrogen content in the active stable feedstock oil is not less than 50 μg / g, preferably 80-200 μg / g; Preferably, the active stable feedstock oil is a heavy diesel fraction, preferably selected from at least one of conventional second-line oil, conventional third-line oil, conventional fourth-line oil, top-reduced oil, first-line oil and straight-run diesel, more preferably selected from at least one of conventional fourth-line oil, top-reduced oil and first-line oil.

6. The method according to any one of claims 1 to 5, wherein: The conditions of the active stabilization process in step (1) include: a temperature of 310-400°C, preferably 320-370°C; a volume space velocity of 0.5-7h -1 , preferably 0.8-6h -1 ; The hydrogen-oil volume ratio is 300-2000:1, preferably 600-1500:1; the total pressure is 2.5-20MPa, preferably 6-15MPa; the time is 25-145h, preferably 40-90h.

7. The method according to claim 6, wherein: The activity stabilization process in step (1) includes a first activity stabilization process and a second activity stabilization process; Preferably, during the first active stabilization process, the volume of C2-C6 alkanes is 10-18% of the volume of hydrogen, preferably 10-16%; Preferably, the time of the first active stabilization process is 20-72 hours, preferably 24-50 hours; Preferably, during the second active stabilization process, the volume of C2-C6 alkanes is 20-35% of the volume of hydrogen, preferably 20-30%; Preferably, the second active stabilization process lasts for 16-60 hours, preferably 20-35 hours.

8. The method according to any one of claims 1 to 7, wherein: In the mixed material, the volume of C2-C6 alkane is 10-20% of the volume of hydrogen.

9. The method according to any one of claims 1 to 8, wherein: The hydrorefining catalyst includes a first carrier and a first metal component supported on the first carrier, wherein the first carrier contains a heat-resistant inorganic oxide, and the first metal component includes a Group VIB metal component and a Group VIII metal component; Preferably, the heat-resistant inorganic oxide is selected from at least one of silicon oxide, aluminum oxide, amorphous silicon aluminum, zirconium oxide and titanium oxide; Preferably, based on the total weight of the hydrorefining catalyst, the content of the first metal component in the hydrorefining catalyst, calculated as oxide, is 15-50 wt%, preferably 18-45 wt%; Preferably, based on the total weight of the hydrotreating catalyst, the content of the Group VIII metal component in the hydrotreating catalyst, calculated as oxide, is 2-8 wt%, preferably 2.5-6 wt%.

10. The method according to any one of claims 1 to 9, wherein: The hydrocracking catalyst comprises a second carrier and a second metal component supported on the second carrier, wherein the second carrier comprises a Y-type molecular sieve, and the second metal component comprises a Group VIB metal component and a Group VIII metal component; Preferably, based on the total weight of the hydrocracking catalyst, the content of the second metal component in the hydrocracking catalyst, calculated as oxide, is 3-50 wt%, preferably 18-40 wt%; Preferably, based on the total weight of the hydrocracking catalyst, the content of the Group VIII metal component in the hydrocracking catalyst, calculated as oxide, is 1.5-8 wt%, preferably 2-6 wt%.

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