Method for hydrogenating coking gasoline and diesel oil

Abstract

The present invention relates to the technical field of oil refining, and discloses a method for hydrogenating coker gasoline and diesel. The method comprises the following steps: (1) subjecting the coker gasoline and diesel feedstock to a first fractionation to obtain a light fraction of the feedstock and a heavy fraction of the feedstock; (2) in the presence of a first hydrogenation catalyst, subjecting the heavy fraction of the feedstock obtained in step (1) to a first hydrogenation reaction with hydrogen to obtain a first hydrogenation reaction product; (3) in the presence of a second hydrogenation catalyst, subjecting the first hydrogenation reaction product obtained in step (2) to a second hydrogenation reaction with the light fraction of the feedstock obtained in step (1) to obtain a second hydrogenation reaction product; (4) subjecting the second hydrogenation reaction product obtained in step (3) to gas-liquid separation to obtain a gas phase and a liquid phase, and subjecting the liquid phase to stripping and a second fractionation to obtain a hydrogenated product. By controlling the temperature of the reaction zone, this method avoids excessive hydrodesulfurization.

Description

Technical Field

[0001] The present invention relates to the field of petroleum refining, and relates to a method for hydrotreating coker gasoline and diesel. Background Art

[0002] In today's world, the demand for clean oil products is increasing continuously, but the growth rate of crude oil production is slow and the crude oil is becoming increasingly heavy, and at the same time, the product quality requirements are becoming more and more stringent. In order to make better use of petroleum resources, higher requirements are also put forward for the secondary processing capacity of crude oil in petroleum refining enterprises. The secondary processing units of petroleum refining enterprises mainly include coking, fluid catalytic cracking, hydrogenation and catalytic reforming units, etc. In particular, delayed coking has become an important means for heavy oil conversion.

[0003] Delayed coking is one of the important means for refining enterprises to process inferior heavy oil, but the products obtained therefrom need to be further purified to remove impurities. In particular, coker gasoline contains high contents of impurities such as sulfur and olefins, and has poor stability, and is not suitable as vehicle gasoline. However, coker naphtha can be used as a raw material for units such as ethylene cracking, fertilizer and reforming after hydrotreating. Since coker naphtha itself contains high contents of impurities such as sulfur, nitrogen, olefins and gum, in order to be used as a raw material for units such as ethylene cracking, fertilizer and reforming, it must first be hydrofined. Coker diesel also contains high contents of impurities such as sulfur, nitrogen and olefins. Coker diesel can be used to produce vehicle diesel after hydrofining according to its product characteristics. The article "Research on Producing Jet Fuel by Hydrotreating Coker Gasoline and Diesel" on pages 173-176 of the 3rd issue of Petrochemical Technology & Application in 2004 introduced that coker gasoline and diesel were mixed for deep hydrofining and then gasoline, kerosene and diesel were fractionated out. This method only deeply refined the whole coker fraction and then separated the kerosene fraction therefrom, resulting in over-hydrogenation. The article "Hydrofining Coker Gasoline and Diesel under Low Pressure with FH-5 Catalyst" published by Sun Guang on pages 39-42 of the 5th issue of Petroleum Refinery Engineering in 1994 used a set of hydrofining devices to hydrofine coker gasoline and coker diesel. Specifically, a switching feeding method was adopted, that is, after processing a batch of coker gasoline, the raw material was switched to coker diesel. This method is only applicable to small-scale coking units. During the production process, it is necessary to continuously switch, and the operating conditions need to be frequently changed, which brings inconvenience to the operation and has a great impact on the performance of the catalyst.

[0004] CN101003751A discloses a processing method for coking whole-fraction oil, comprising: separating coking whole-fraction oil into coking light-fraction oil and coking heavy-fraction oil, wherein the light-fraction oil contains part of the light diesel fraction, and the heavy-fraction oil is a heavy diesel fraction with a higher final boiling point. The coking heavy-fraction oil is subjected to hydrocracking, and the cracking products are mixed with the coking light-fraction oil and then subjected to hydrofining treatment. CN111321005A discloses a hydroprocessing method for producing diesel with low energy consumption and long cycle. First, the diesel raw material is fractionated into a light fraction and a heavy fraction, the heavy fraction is subjected to hydrodesulfurization reaction in a first reaction zone, and then reacts with the light-fraction diesel raw material oil in a second hydroprocessing reaction zone under the action of a hydrodesulfurization catalyst. CN200410050729.6 discloses a hydrotreating method for diesel fraction. First, the raw material oil is fractionated into a light fraction and a heavy fraction, the light fraction is sent to a first reaction zone and contacted with a hydrofining catalyst, the heavy fraction is sent to a second reaction zone and contacted with a hydrofining catalyst, and the liquid-phase products of the two reaction zones are mixed to form a refined fraction oil. The raw material of this process needs to be fractionated into light and heavy components, and the process is relatively complex.

[0005] The processes selected in the above patents are all conventional gas-phase circulating trickle-bed hydroprocessing processes. However, they all have problems such as excessive hydrodesulfurization reaction, which is likely to cause over-quality of products, reduction of liquid yield, and excessive hydrogen consumption and energy consumption. Summary of the Invention

[0006] The object of the present invention is to overcome the problems of excessive hydrodesulfurization and the like existing in the prior art, and provide a method for hydroprocessing coking gasoline and diesel. This method controls the temperature of the reaction zone to avoid excessive hydrodesulfurization reaction and improve the liquid yield; at the same time, it reduces the reaction severity, can better play the role of each reaction zone, makes better use of the heat balance, reduces the energy consumption and hydrogen consumption of the device, and produces high-quality diesel and naphtha for ethylene cracking.

[0007] To achieve the above object, on the one hand, the present invention provides a method for hydroprocessing coking gasoline and diesel, which comprises the following steps:

[0008] (1) Subjecting the coking gasoline and diesel raw material oil to a first fractionation to obtain a raw material oil light fraction and a raw material oil heavy fraction;

[0009] (2) In the presence of a first hydroprocessing catalyst, subjecting the raw material oil heavy fraction obtained in step (1) to a first hydroprocessing reaction with hydrogen to obtain a first hydroprocessing reaction product;

[0010] (3) In the presence of a second hydroprocessing catalyst, subjecting the first hydroprocessing reaction product obtained in step (2) to a second hydroprocessing reaction with the raw material oil light fraction obtained in step (1) to obtain a second hydroprocessing reaction product;

[0011] (4) Carry out gas-liquid separation on the second hydrogenation reaction product obtained in step (3) to obtain a gas phase and a liquid phase, and subject the liquid phase to stripping and second fractionation to obtain a hydrogenation product;

[0012] Among them, in step (2), the average reaction temperature of the first hydrogenation reaction is 340 - 420 °C;

[0013] In step (3), the average reaction temperature of the second hydrogenation reaction is 255 - 285 °C.

[0014] Preferably, in step (2), the conditions of the first hydrogenation reaction include: hydrogen partial pressure is 2 - 16 MPa, volume space velocity is 0.3 - 10 h 1 , and the hydrogen-oil volume ratio is 200:1 to 2500:1.

[0015] Preferably, the organic sulfur content in the reaction effluent of the first hydrogenation reaction is 5 - 10 μg / g.

[0016] Preferably, the hydrogen sulfide concentration in the reaction effluent in the first reaction zone is 5000 - 15000 ppm, preferably 6000 - 10000 ppm.

[0017] Preferably, in step (2), the first hydrogenation catalyst includes a first carrier and a first hydrogenation active component supported on the first carrier, and the first hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni.

[0018] Preferably, the first carrier is alumina, silica, amorphous silica-alumina, titanium oxide, and their composite oxides or mixtures.

[0019] Preferably, based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 15 - 50 wt%, preferably 18 - 45 wt%.

[0020] Preferably, the specific surface area of the first hydrogenation catalyst is 100 - 600 m 2 / g, preferably 200 - 350 m 2 / g, and the pore volume is 0.2 - 1.5 mL / g, preferably 0.35 - 0.7 mL / g.

[0021] Preferably, in step (3), the conditions of the second hydrogenation reaction include: hydrogen partial pressure is 0.5 - 10 MPa, preferably 1.5 - 8 MPa, volume space velocity is 1 - 10 h 1 , and the hydrogen-oil volume ratio is 80:1 to 2500:1.

[0022] Preferably, the organic sulfur content in the reaction effluent of the second hydrogenation reaction is 50 - 2000 μg / g.

[0023] Preferably, in step (3), the second hydrogenation catalyst comprises a second support and a second hydrogenation active component supported on the second support, and the second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII.

[0024] Preferably, based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5%-30%, preferably 5%-15%, and the content of the metal oxide of Group VIII is 1%-15%, preferably 2%-6%.

[0025] Preferably, the specific surface area of the second hydrogenation catalyst is 100-500 m 2 / g, preferably 300-500 m 2 / g, and the pore volume is 0.3-1.2 mL / g, preferably 0.4-0.8 mL / g.

[0026] Preferably, in step (2), the average reaction temperature of the first hydrogenation reaction is 340-370 °C; in step (3), the average reaction temperature of the second hydrogenation reaction is 260-275 °C.

[0027] Preferably, the average reaction temperature of the second hydrogenation reaction is 55-145 °C lower than that of the first hydrogenation reaction, preferably 65-85 °C lower.

[0028] Preferably, the hydrogen partial pressure of the first hydrogenation reaction is 4-16 MPa, and the hydrogen partial pressure of the second hydrogenation reaction is 1.5-8 MPa.

[0029] Preferably, the hydrogen partial pressure of the second hydrogenation reaction is 1-6 MPa lower than that of the first hydrogenation reaction, preferably 2-6 MPa lower.

[0030] Preferably, the first hydrogenation reaction and the second hydrogenation reaction are sequentially carried out in different reaction zones in the same reactor.

[0031] Preferably, the volume ratio of the first hydrogenation catalyst to the second hydrogenation catalyst is 70:30 to 50:50, preferably 65:35 to 50:50.

[0032] Preferably, in step (1), the initial boiling point temperature of the coking gasoline and diesel raw material oil is 30-100 °C, and the final boiling point temperature is 350-400 °C.

[0033] Preferably, the cut-off point between the light fraction and the heavy fraction of the raw material oil is 170-280 °C.

[0034] Preferably, the method further comprises: after the hydrogen sulfide in the gas phase obtained in step (4) is removed, returning it to step (1) to be mixed with the coking gasoline and diesel raw material oil.

[0035] Preferably, the hydrogen sulfide concentration in the material after gas-phase hydrogen sulfide removal obtained in step (4) is 1500 ppm or less, preferably 50 - 1000 ppm.

[0036] Preferably, the hydrogenation products are gasoline and diesel.

[0037] Preferably, the sulfur content in the gasoline is 50 - 1500 ppm.

[0038] Through the above technical solution, the method for hydrogenating coking gasoline and diesel of the present invention has the following beneficial effects:

[0039] 1. The coking gasoline and diesel hydrogenation process method provided by the present invention divides the feedstock oil into light and heavy fractions, optimizes the catalyst grading according to different reaction requirements, deeply couples the reaction fractions, reaction sequence, reaction pressure, hydrogen sulfide concentration and various reactions, controls the reaction depth, balances the reaction heat, maximally exerts the efficiency of the catalyst in the reaction, avoids passive desulfurization of the gasoline fraction, and achieves the purpose of controllable reaction.

[0040] 2. In the present invention, the heavy fraction first enters the first reaction zone for hydrofining desulfurization reaction, avoiding the strong exotherm of a large amount of olefins in the light fraction, controlling the reaction temperature rise, and avoiding the situation where the high point temperature of the reaction is too high. At the same time, the product of the first reaction zone is used as a heat-carrying medium to be mixed with the light fraction and enter the second reaction zone, increasing the heat capacity, and avoiding the excessive passive rise of the temperature after the strong exothermic reaction of the light component olefins. The effluent from the first reaction zone is mixed with the light fraction, which can further reduce the reaction temperature, achieve a low reaction temperature matching, and the reaction process is a gentle reaction, which can better reduce the thermal cracking caused by too high temperature and reduce the influence on the yield of the target product.

[0041] 3. In the present invention, a first hydrogenation catalyst and a low-temperature second hydrogenation catalyst are respectively arranged in the first and second reaction zones, and the corresponding reaction temperature and reaction pressure are adjusted, which well exerts the high-temperature hydrodesulfurization, low-temperature hydrodesulfurization alcohol activity and olefin removal performance of the catalyst. Furthermore, by performing a pressure reduction operation in the second reaction zone, it is possible to avoid the passive excessive desulfurization of sulfur in the gasoline fraction in the conventional coking gasoline and diesel hydrogenation process to ensure ultra-deep removal of thiophene sulfur. This can reduce the sulfur injection amount of the downstream ethylene plant, reduce the operating cost, and while increasing the total liquid yield, increase the yield of gasoline components.

[0042] 4. In the present invention, the heavy fraction enters the first reaction zone and undergoes a deep desulfurization reaction with high-purity hydrogen under the action of a catalyst, avoiding the influence of excessive hydrogen sulfide concentration on the reaction, obtaining high-quality diesel components, regulating the hydrogen sulfide concentration in the product of the first reaction zone, and then mixing and hydrogenating with the light fraction, which can achieve the purpose of inhibiting the desulfurization depth, reducing the sulfur removal rate of the gasoline fraction that is too high, reducing the sulfur injection amount of the downstream ethylene plant, and realizing the reduction of the operation cost of the ethylene plant.

[0043] 5. The method provided by the present invention gradually controls the reaction heat release, effectively matches the heat capacity and reaction heat balance, effectively controls the reaction temperature rise, reduces the probability of thermal cracking, ensures high liquid yield, and can operate under reduced pressure in the second reactor, reducing the operation safety risk of the device, and at the same time reducing the hydrogen consumption and energy consumption of the device.

[0044] 6. The method provided by the present invention controls the reaction process, controls the desulfurization depth, reduces the high point temperature of the reaction, and reduces the reaction hydrogen consumption and the energy consumption of the device.

[0045] 7. As the reaction device used in the method of the present invention, an existing coking gasoline and diesel hydrogenation or other similar hydrogenation devices can be used for minor modifications, reducing the transformation cost. At the same time, the operation steps change little, also reducing the operation risk.

[0046] 8. The coking gasoline and diesel hydrogenation process of the present invention has a simple process, convenient operation, good safety, environmental friendliness, high product added value, and a long operation cycle. BRIEF DESCRIPTION OF THE DRAWINGS

[0047] Figure 1 is a schematic diagram of the process adopted in the embodiment of the present invention.

[0048] DESCRIPTION OF THE REFERENCE NUMERALS

[0049] 1. Heavy fraction introduction pipeline 2. Fresh hydrogen introduction pipeline 3. Second hydrogen circulation pipeline

[0050] 4. Light fraction introduction pipeline 5. Hydrofining reactor 6. Hydrogenation product export pipeline

[0051] 7. Cold high-pressure separator 8. Liquid phase pipeline 9. Cold low-pressure separator

[0052] 10. Export pipeline 11. First hydrogen circulation pipeline DETAILED DESCRIPTION OF THE INVENTION

[0053] The endpoints and any values in the ranges disclosed herein are not limited to the exact ranges or values. These ranges or values should be understood to include values close to these ranges or values. For numerical ranges, the values between the endpoints of each range, between the endpoints of each range and individual point values, and between individual point values can be combined with each other to obtain one or more new numerical ranges, and these numerical ranges should be considered to be specifically disclosed herein.

[0054] The method for hydrogenation of coking gasoline and diesel provided by the present invention comprises the following steps:

[0055] (1) Subjecting the coking gasoline and diesel feedstock to a first fractionation to obtain a light fraction of the feedstock and a heavy fraction of the feedstock;

[0056] (2) In the presence of a first hydrogenation catalyst, subjecting the heavy fraction of the feedstock obtained in step (1) to a first hydrogenation reaction with hydrogen to obtain a first hydrogenation reaction product;

[0057] (3) In the presence of a second hydrogenation catalyst, subjecting the first hydrogenation reaction product obtained in step (2) to a second hydrogenation reaction with the light fraction of the feedstock obtained in step (1) to obtain a second hydrogenation reaction product;

[0058] (4) Separating the second hydrogenation reaction product obtained in step (3) into a gas phase and a liquid phase by gas-liquid separation, and subjecting the liquid phase to stripping and a second fractionation to obtain a hydrogenated product;

[0059] Wherein, in step (2), the average reaction temperature of the first hydrogenation reaction is 340 - 420 °C;

[0060] In step (3), the average reaction temperature of the second hydrogenation reaction is 255 - 285 °C.

[0061] In the present invention, the first hydrogenation reaction in step (2) is mainly used for hydrodesulfurization, and the second hydrogenation reaction in step (3) is mainly used for hydrodesulfurization of mercaptans and hydrodearomatization reaction.

[0062] In the present invention, as the coking gasoline and diesel feedstock, it may contain various hydrocarbons of gasoline and / or diesel fractions, and specifically may contain a mixture of one or more of C4 - C25 hydrocarbons. The initial boiling point temperature of the coking gasoline and diesel feedstock may be 30 - 100 °C, preferably 60 - 90 °C, and the final boiling point temperature may be 340 - 400 °C, preferably 350 - 380 °C.

[0063] According to the present invention, in step (1), the coking gasoline and diesel raw material oil is separated into a raw material oil light fraction and a raw material oil heavy fraction by the first fractionation. As the cut-off point for the raw material oil light fraction and the raw material oil heavy fraction, it is preferably 170 - 280 °C, more preferably 190 - 230 °C. By fractionating under the above conditions, the yield of gasoline obtained can be increased, and the economic benefits can be improved.

[0064] According to the present invention, in step (2), the conditions of the first hydrogenation reaction may include: the hydrogen partial pressure is 2 - 16 MPa, preferably 4 - 16 MPa, and the volume space velocity is 0.3 - 10 h 1 , preferably 0.6 - 2 h 1 , and the hydrogen-oil volume ratio is 200:1 - 2500:1, preferably 300:1 - 1000:1.

[0065] According to the present invention, preferably, the content of organic sulfur in the reaction effluent in the first reaction zone is 5 - 10 μg / g.

[0066] According to the present invention, preferably, the concentration of hydrogen sulfide in the reaction effluent in the first reaction zone is 5000 - 15000 ppm, preferably 6000 - 10000 ppm.

[0067] According to the present invention, in step (2), the first hydrogenation catalyst can be selected as a suitable commercial catalyst according to the needs of the process flow, can also be prepared by a conventional method, or can be a regenerated catalyst after the deactivated catalyst is regenerated. The first hydrogenation catalyst includes a first carrier and a first hydrogenation active component supported on the first carrier. The first hydrogenation active component can be W-Ni, Mo-Ni, or W-Mo-Ni. The first carrier is generally a refractory porous oxide, such as alumina, silica, amorphous silica-alumina, titanium oxide, and their composite oxides or mixed oxide carriers, etc. Among them, materials without acidity or with weak acidity are generally used as the carrier. Preferably, based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 15 - 50 wt%, preferably 18 - 45 wt%, and the content of nickel oxide is 2 - 8 wt%, preferably 2.3 - 6 wt%.

[0068] The shape of the first hydrogenation catalyst is not particularly limited. For example, it can be spherical or bar-shaped. The spherical diameter is 0.04 - 15 mm, preferably 0.04 - 5 mm; the bar shape has a length of 2 - 15 mm, preferably 2 - 8 mm, and a diameter of 1 - 6 mm, preferably 1.5 - 3.5 mm.

[0069] Preferably, the specific surface area of the first hydrogenation catalyst is 100 - 600 m 2 / g, preferably 200 - 350 m 2 / g, the pore volume is 0.2 - 1.5 mL / g, preferably 0.35 - 0.7 mL / g.

[0070] According to the present invention, in step (3), the conditions of the second hydrogenation reaction may include: the hydrogen partial pressure is 0.5 - 10 MPa, preferably 1.5 - 8 MPa, the volume space velocity is 1 - 10 h 1 , the hydrogen-oil volume ratio is 80:1 - 2500:1, preferably 200:1 - 1000:1. Additionally, the hydrogen for the second hydrogenation reaction can come from the product of the first hydrogenation reaction, or can be additionally added as needed.

[0071] According to the present invention, preferably, the content of organic sulfur in the reaction effluent in the second reaction zone is 50 - 2000 μg / g, preferably 150 - 600 μg / g. When the content of organic sulfur in the reaction effluent in the second reaction zone of the present invention is within the above range, it can prevent the coking gasoline and diesel raw oil from being over-hydrogenated, which is more beneficial for the product to be used as a raw material for ethylene cracking. Downstream ethylene cracking units need to inject sulfur-containing media (such as dimethyl sulfide (DMDS)) when using raw materials such as hydrotreated naphtha, hydrocracked light naphtha, hydrocracked oil, and ethane. If the ethylene cracking raw material contains a sufficient amount of sulfur, sulfur injection is not required. The ethylene cracking unit needs to maintain the sulfur concentration in the raw material above 100 ppm, and the sulfur injection amount is calculated according to the feed rate of the cracking furnace.

[0072] According to the present invention, in step (3), the second hydrogenation catalyst can be selected as a suitable commercial catalyst according to the needs of the process flow, can also be prepared by existing methods, or can be a regenerated catalyst after the deactivated catalyst is regenerated. The second hydrogenation catalyst includes a second carrier and a second hydrogenation active component supported on the second carrier. The second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII. For example, it can be Mo, W, Co, etc. The second carrier can be alumina or alumina modified with additives. Preferably, based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5 - 30 wt%, preferably 5 - 15 wt%, and the content of the metal oxide of Group VIII is 1 - 15 wt%, preferably 2 - 6 wt%.

[0073] Preferably, the specific surface area of the second hydrogenation catalyst is 100 - 500 m 2 / g, preferably 300 - 500 m 2 / g, the pore volume is 0.3 - 1.2 mL / g, preferably 0.4 - 0.8 mL / g.

[0074] According to the present invention, the preparation methods of the first hydrogenation catalyst and the second hydrogenation catalyst are well-known in the art. One or more of the impregnation method, co-extrusion method, and co-precipitation method can be used, or they can also be regenerated catalysts after the deactivated catalysts are regenerated. In order to obtain better hydrogenation catalytic effects, preferably, the volume ratio of the first hydrogenation catalyst to the second hydrogenation catalyst is 70:30 to 50:50, preferably 65:35 to 50:50.

[0075] According to the present invention, the average reaction temperature of the first hydrogenation reaction can be 420°C or lower, 410°C or lower, 400°C or lower, 390°C or lower, 380°C or lower, 370°C or lower, 360°C or lower, or 350°C or lower, and 340°C or higher, 350°C or higher, 360°C or higher, 370°C or higher, 380°C or higher, 390°C or higher, or 400°C or higher, preferably 340 - 370°C; the average reaction temperature of the second hydrogenation reaction can be 285°C or lower, 280°C or lower, 270°C or lower, or 260°C or lower, and 255°C or higher, 260°C or higher, or 270°C or higher, preferably 260 - 275°C. Further preferably, the average reaction temperature of the second hydrogenation reaction is 55°C or higher, 60°C or higher, or 65°C or higher lower than the average reaction temperature of the first hydrogenation reaction, and 145°C or lower, 135°C or lower, 125°C or lower, 115°C or lower, 105°C or lower, 95°C or lower, or 85°C or lower.

[0076] According to the present invention, the hydrogen partial pressure of the first hydrogenation reaction can be, for example, 4MPa, 5MPa, 6MPa, 7MPa, 8MPa, 9MPa, 10MPa, 11MPa, 12MPa, 13MPa, 14MPa, 15MPa, or 16MPa; the hydrogen partial pressure of the second hydrogenation reaction is 1.5MPa, 2MPa, 3MPa, 4MPa, 5MPa, 6MPa, 7MPa, or 8MPa. Preferably, the hydrogen partial pressure of the second hydrogenation reaction is 1 - 6MPa lower than that of the first hydrogenation reaction, preferably 2 - 6MPa, more preferably 4 - 6MPa. Specifically, the hydrogen partial pressure of the second hydrogenation reaction can be, for example, 1MPa, 1.5MPa, 2MPa, 2.5MPa, 3MPa, 3.5MPa, 4MPa, 4.5MPa, 5MPa, 5.5MPa, or 6MPa lower than that of the first hydrogenation reaction. By adopting the above reaction conditions, the sulfur content of the obtained hydrogenation product can be further regulated.

[0077] As the apparatus for carrying out the above first hydrogenation reaction and second hydrogenation reaction, any apparatus used in the art for hydrogenating coking gasoline and diesel raw materials can be employed, and a trickle bed reactor is preferably used. In the above method, the first reaction zone for the first hydrogenation reaction or the second reaction zone for the second hydrogenation reaction can be a reactor or a part of a reactor, or can be composed of multiple reactors. If multiple reactors are selected, preferably 2 - 4 reactors. Inside the first reaction zone or the second reaction zone, a catalyst bed layer can be independently provided respectively, or multiple catalyst bed layers can be provided. If multiple catalyst bed layers are selected, preferably 2 - 5 catalyst bed layers.

[0078] According to the present invention, the first hydrogenation reaction and the second hydrogenation reaction are sequentially carried out in different reaction zones in the same reactor. For example, the first hydrogenation reaction is carried out in the first reaction zone, and the second hydrogenation reaction is carried out in the second reaction zone, and the first reaction zone and the second reaction zone are provided in the same reactor. For example, the process as shown in Figure 1 can be adopted, Figure 1 in which equipment such as oil pumps, valves, heating furnaces, etc. are omitted. The reaction apparatus includes a hydrofining reactor 5, which includes a first reaction zone located in the upper part and a second reaction zone located in the lower part. A first hydrogenation catalyst is provided in the first reaction zone, and a second hydrogenation catalyst is provided in the second reaction zone. At this time, when the catalyst is fully filled, the volume ratio of the first hydrogenation catalyst to the second hydrogenation catalyst can be represented by the volume ratio of the first reaction zone to the second reaction zone.

[0079] According to the present invention, the method may further include: after removing hydrogen sulfide from the gas phase obtained in step (4), returning it to step (1) to be mixed with the coking gasoline and diesel raw materials. Specifically, the alcohol amine adsorption method can be used for desulfurization. Preferably, the hydrogen sulfide concentration in the material (i.e., the recycle hydrogen at the inlet of the first reaction zone) after removing hydrogen sulfide from the gas phase obtained in step (4) is 1500 ppm or less, preferably 50 - 1000 ppm.

[0080] According to the present invention, the hydrogenation product obtained by the hydrogenation method of the present invention can be gasoline and diesel. In step (4), the fractionation system used in the second fractionation is a system that can accurately fractionate gasoline and diesel.

[0081] By using the hydrogenation method of the present invention to hydrogenate the coker gasoline and diesel raw materials, diesel with a sulfur content of less than 10 ppm, preferably 3 - 8 ppm, and gasoline with a sulfur content of 50 - 1500 ppm can be obtained. Among them, the sulfur content of the gasoline is preferably 200 - 1000 ppm, and more preferably 200 - 500 ppm. In addition, the olefin content of the gasoline can reach 1.5% by volume or less, preferably 1% by volume or less. It can be seen from this that the hydrogenation method of the present invention can avoid over-hydrogenation of the coker gasoline and diesel raw materials, which is more conducive to the application of the product as an ethylene cracking raw material and avoids the need to increase the sulfur content in subsequent steps.

[0082] According to a preferred embodiment of the present invention, the method further includes returning the gas phase obtained in step (4) as recycle hydrogen to steps (2) and (3). Specifically, the gas phase obtained in step (4) can be desulfurized by a desulfurization tower and then returned to steps (2) and (3).

[0083] According to a specific embodiment of the present invention, when hydrogenating coker gasoline and diesel, as Figure 1 shown, the following process can be adopted: The heavy fraction of coker gasoline and diesel is introduced through the heavy fraction pipeline 1, mixed with the fresh hydrogen in the fresh hydrogen pipeline 2 and the recycle hydrogen in the second hydrogen recycle pipeline 3, and enters from the top of the hydrofining reactor 5. The material passes through the first reaction zone and undergoes a hydrodesulfurization reaction in the presence of the first hydrogenation catalyst. The light fraction of coker gasoline and diesel is introduced through the light fraction pipeline 4 and mixed with the recycle hydrogen in the hydrogen recycle pipeline 3 and enters the middle of the hydrofining reactor 5, where it is mixed with the reaction effluent from the upper first reaction zone and enters the second reaction zone, and undergoes hydrodesulfurization alcohol and hydrodearomatization in the presence of the second hydrogenation catalyst. The mixed material flows out from the bottom of the hydrofining reactor 5; it enters the cold high-pressure separator 7 through the hydrogenation product export pipeline 6, the obtained liquid phase passes through the liquid phase pipeline 8, enters the cold low-pressure separator 9, and then enters the fractionation unit through the export pipeline 10; the recycle hydrogen separated by the cold high-pressure separator 7 enters the recycle hydrogen desulfurization system for desulfurization through the first hydrogen recycle pipeline 11, and then returns to the hydrogen recycle pipeline 3 through the recycle hydrogen compressor 12.

[0084] The present invention will be described in detail below through examples. In the following examples, FHUDS-8 is the FHUDS-8 hydrofining catalyst (the active component is Mo-Ni) produced by the Fushun Branch of Sinopec Catalyst Company, and FH-40C is the FH-40C hydrofining catalyst (the active component is W-Mo-Ni) produced by the Fushun Branch of Sinopec Catalyst Company.

[0085] Examples 1 - 8 and Comparative Examples 1 - 4

[0086] Adopt as Figure 1The device shown is used for the hydrotreating of coker gasoline and diesel. The hydrofining reactor 5 is a trickle bed reactor, which includes a first reaction zone located in the upper part and a second reaction zone located in the lower part. A first hydrotreating catalyst is arranged in the first reaction zone, and a second hydrotreating catalyst is arranged in the second reaction zone.

[0087] The coker gasoline and diesel feedstock shown in Table 1 is fractionated at the splitting point of 210 °C to obtain a light fraction of the feedstock and a heavy fraction of the feedstock. The heavy fraction of the coker gasoline and diesel is mixed with the fresh hydrogen in the fresh hydrogen introduction pipeline 2 and the recycle hydrogen in the second hydrogen recycle pipeline 3 through the heavy fraction introduction pipeline 1, and enters from the top of the hydrofining reactor 5. The material passes through the first reaction zone and undergoes a hydrodesulfurization reaction in the presence of the first hydrotreating catalyst. The light fraction of the coker gasoline and diesel is mixed with the recycle hydrogen in the hydrogen recycle pipeline 3 through the light fraction introduction pipeline 4 and enters the middle part of the hydrofining reactor 5, where it is mixed with the reaction effluent from the first reaction zone in the upper part and enters the second reaction zone, where it undergoes hydrodesulfurization alcohol and hydrodearomatization in the presence of the second hydrotreating catalyst. The mixed material flows out from the bottom of the hydrofining reactor 5; it enters the cold high-pressure separator 7 through the hydrogenation product export pipeline 6. The obtained liquid phase passes through the liquid phase pipeline 8, enters the cold low-pressure separator 9, and then enters the fractionation unit through the export pipeline 10; the recycle hydrogen separated by the cold high-pressure separator 7 enters the recycle hydrogen desulfurization system through the first hydrogen recycle pipeline 11, is desulfurized in the desulfurization tower, and then returns to the hydrogen recycle pipeline 3 through the recycle hydrogen compressor 12.

[0088] The feedstock oils used in the examples and comparative examples of the present invention are shown in Table 1. The main operating process conditions and product properties in the examples and comparative examples of the present invention are shown in Tables 2 - 4.

[0089] Table 1 Properties of the feedstock oil

[0090] Property Coked gasoline and diesel <![CDATA[Density (20 °C), g / cm 3 > 0.8501 Distillation range, °C (ASTM D86) Initial boiling point / 10% 45 / 95 50％ / 90％ 256 / 355 95% / Dry point (final boiling point) 369 / 375 Sulfur, μg / g 9655 Nitrogen, μg / g 1016 Olefins, v% 25

[0091] Table 2

[0092] Item Example 1 Example 2 Example 3 Example 4 First hydrogenation catalyst FHUDS-8 FHUDS-8 FHUDS-8 FHUDS-8 Second hydrogenation catalyst FH-40C FH-40C FH-40C FH-40C Hydrogen partial pressure in the first and second reaction zones, MPa 8.0 / 3.0 8.0 / 3.0 8.0 / 3.0 8.0 / 3.0 Hydrogen-oil volume ratio at the reactor inlet 500:1 500:1 500:1 500:1 <![CDATA[Overall volumetric space velocity, h -1 > 1.3 1.3 1.3 1.3 Volume ratio of the first reaction zone to the second reaction zone, v% 55：45 55：45 60：40 60：40 Average reaction temperature in the first reaction zone, °C 342 342 340 340 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 6500 7600 6000 7200 Average reaction temperature in the second reaction zone, °C 260 270 272 272 Liquid yield, % 98.5 98.6 98.4 98.5 Diesel sulfur content, ppm 7 7 8 8 Gasoline sulfur content, ppm 226 235 229 235 Gasoline olefin content, v% <1 <1 <1 <1

[0093] Table 3

[0094] Item Comparative Example 1 Comparative Example 2 Comparative Example 3 Comparative Example 4 First hydrogenation catalyst FHUDS-8 FHUDS-8 FHUDS-8 FHUDS-8 Second hydrogenation catalyst FHUDS-8 FH-40C FH-40C FH-40C Hydrogen partial pressure in the first and second reaction zones, MPa 8.0 / 7.5 8.0 / 3.0 8.0 / 3.0 8.0 / 3.0 Hydrogen-oil volume ratio at the reactor inlet 500:1 500:1 500:1 500:1 <![CDATA[Total volume space velocity, h -1 > 1.2 1.3 1.3 1.3 Volume ratio of the first reaction zone to the second reaction zone, v% 40：60 55：45 55：45 55：45 Average reaction temperature in the first reaction zone, °C 325 325 405 380 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 3500 3500 8000 6500 Average reaction temperature in the second reaction zone, °C 345 315 220 360 Liquid yield, % 97.2 98.6 97.5 96.5 Diesel sulfur content, ppm 9 155 <2 3 Gasoline sulfur content, ppm <1 <1 1675 <1 Gasoline olefin content, v% <1 <1 15.6 <1

[0095] Table 4

[0096] Item Example 5 Example 6 Example 7 Example 8 First hydrogenation catalyst FHUDS-8 FHUDS-8 FHUDS-8 FHUDS-8 Second hydrogenation catalyst FH-40C FH-40C FH-40C FH-40C Hydrogen partial pressure in the first and second reaction zones, MPa 8.0 / 3.0 10.0 / 5.0 8.0 / 7.0 8.0 / 2.5 Hydrogen-oil volume ratio at the reactor inlet 500:1 500:1 500:1 500:1 <![CDATA[Overall volumetric space velocity, h -1 > 1.3 1.3 1.3 1.3 Volume ratio of the first reaction zone to the second reaction zone, v% 55：45 55：45 55：45 55：45 Average reaction temperature in the first reaction zone, °C 395 342 342 342 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 8500 6500 6500 6500 Average reaction temperature in the second reaction zone, °C 280 260 260 260 Liquid yield, % 97.2 98.4 97.9 98.2 Diesel sulfur content, ppm <2 5 7 7 Gasoline sulfur content, ppm 50 195 95 560 Gasoline olefin content, v% <1 <1 <1 1.5

[0097] According to the results of the above examples and comparative examples, it can be seen from the comparison between Example 1 and Example 5 that by making the average reaction temperatures of the first and second hydrotreating reactions be in the ranges of 340 - 370 °C and 260 - 275 °C respectively, the liquid yield can be further increased and the sulfur content of the obtained gasoline and diesel can be better controlled.

[0098] It can be seen from the comparison between Examples 1, 6, 8 and Example 7 that by controlling the hydrogen partial pressure of the second hydrogenation reaction to be 2-6 MPa lower than that of the first hydrogenation reaction, the liquid yield can be further increased and the sulfur content of the obtained gasoline can be better controlled.

[0099] As described above, the biggest feature of the hydrogenation method of the present invention is that it adopts a simple transformation of a conventional coking gasoline and diesel hydrogenation refining unit, the trickle bed reactor process, optimizes the catalyst grading method, and through fractionation cutting, deeply combines the reaction sequence and reaction process conditions of the reactants, controls the reaction depth, balances the reaction heat, maximally exerts the performance of the catalysts in each reaction zone, and achieves the purpose of reducing the hydrogen consumption and energy consumption of the unit. The hydrogenation method of the present invention can increase the liquid yield, extend the operation cycle of the coking gasoline and diesel hydrogenation unit, bring considerable economic and social benefits to the enterprise, and has great practical application advantages.

[0100] The preferred embodiments of the present invention have been described in detail above. However, the present invention is not limited thereto. Within the technical concept scope of the present invention, various simple modifications can be made to the technical solutions of the present invention, including any other suitable combination of each technical feature. These simple modifications and combinations should also be regarded as the content disclosed by the present invention and fall within the protection scope of the present invention.

Claims

1. A method for hydrotreating coking gasoline and diesel, characterized in that, The method comprises the following steps: (1) subjecting the coking gasoline and diesel raw material oil to a first fractionation to obtain a light fraction of the raw material oil and a heavy fraction of the raw material oil; (2) in the presence of a first hydrogenation catalyst, subjecting the heavy fraction of the raw material oil obtained in step (1) to a first hydrogenation reaction with hydrogen to obtain a first hydrogenation reaction product; (3) in the presence of a second hydrogenation catalyst, subjecting the first hydrogenation reaction product obtained in step (2) to a second hydrogenation reaction with the light fraction of the raw material oil obtained in step (1) to obtain a second hydrogenation reaction product; (4) subjecting the second hydrogenation reaction product obtained in step (3) to gas-liquid separation to obtain a gas phase and a liquid phase, and subjecting the liquid phase to stripping and a second fractionation to obtain a hydrogenation product; wherein, in step (2), the average reaction temperature of the first hydrogenation reaction is 340-420 °C; in step (3), the average reaction temperature of the second hydrogenation reaction is 255-285 °C, the average reaction temperature of the second hydrogenation reaction is 55-145 °C lower than the average reaction temperature of the first hydrogenation reaction, the hydrogen partial pressure of the first hydrogenation reaction is 4-16 MPa, the hydrogen partial pressure of the second hydrogenation reaction is 1.5-8 MPa, and the hydrogen partial pressure of the second hydrogenation reaction is 1-6 MPa lower than that of the first hydrogenation reaction, the first hydrogenation reaction in step (2) is mainly used for hydrodesulfurization, and the second hydrogenation reaction in step (3) is mainly used for hydrodesulfurization of mercaptans and hydrodearomatization reaction, the hydrogen sulfide concentration in the first hydrogenation reaction product is 5000-15000 ppm.

2. The method according to claim 1, wherein In step (2), the conditions of the first hydrogenation reaction include: hydrogen partial pressure is 2 - 16 MPa, space velocity is 0.3 - 10 h -1 , and hydrogen-to-oil volume ratio is 200:1 - 2500:

1.

3. The method according to claim 1, wherein In step (2), the first hydrogenation catalyst comprises a first carrier and a first hydrogenation active component supported on the first carrier, and the first hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni.

4. The method according to claim 3, wherein The first carrier is one or more of alumina, silica, amorphous silica-alumina and titanium oxide.

5. The method according to claim 3, wherein Based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 15-50 wt%.

6. The method according to claim 5, wherein, Based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 18-45 wt%.

7. The method according to claim 3, wherein The specific surface area of the first hydrogenation catalyst is 100 - 600 m 2 / g, and the pore volume is 0.2 - 1.5 mL / g.

8. The method according to claim 7, wherein The specific surface area of the first hydrogenation catalyst is 200 - 350 m 2 / g, and the pore volume is 0.35 - 0.7 mL / g.

9. The method according to claim 1, wherein In step (3), the conditions for the second hydrogenation reaction include: a hydrogen partial pressure of 0.5 - 10 MPa, a space velocity of 1 - 10 h -1 , and a hydrogen-to-oil volume ratio of 80:1 to 2500:

1.

10. The method according to claim 1, wherein, In step (3), the second hydrogenation catalyst comprises a second carrier and a second hydrogenation active component supported on the second carrier, and the second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII.

11. The method according to claim 10, wherein, Based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5-30 wt%, and the content of the metal oxide of Group VIII is 1-15 wt%.

12. The method according to claim 11, wherein, Based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5-15 wt%, and the content of the metal oxide of Group VIII is 2-6 wt%.

13. The method according to claim 10, wherein The specific surface area of the second hydrogenation catalyst is 100 - 500 m 2 / g, and the pore volume is 0.3 - 1.2 mL / g.

14. The method according to any one of claims 1-13, wherein, In step (2), the average reaction temperature of the first hydrogenation reaction is 340-370 °C; in step (3), the average reaction temperature of the second hydrogenation reaction is 260-275 °C.

15. The method according to any one of claims 1-13, wherein, The average reaction temperature of the second hydrogenation reaction is 65-85 °C lower than the average reaction temperature of the first hydrogenation reaction.

16. The method according to any one of claims 1-13, wherein, The hydrogen partial pressure of the second hydrogenation reaction is 2 - 6 MPa lower than that of the first hydrogenation reaction.

17. The method according to any one of claims 1-13, wherein, The first hydrogenation reaction and the second hydrogenation reaction are sequentially carried out in different reaction zones in the same reactor.

18. The method according to claim 17, wherein, The volume ratio of the first hydrogenation catalyst to the second hydrogenation catalyst is 70:30 to 50:

50.

19. The method according to claim 18, wherein The volume ratio of the first hydrogenation catalyst to the second hydrogenation catalyst is 65:35 to 50:

50.

20. The method according to any one of claims 1-13, wherein In step (1), the initial boiling point temperature of the coker gasoline and diesel feedstock is 30 - 100 °C, and the final boiling point temperature is 350 - 400 °C.

21. The method according to claim 20, wherein, The cut-off point between the light fraction and the heavy fraction of the feedstock is 170 - 280 °C.

22. The method according to any one of claims 1-13, wherein, This method further includes: after removing hydrogen sulfide from the gas phase obtained in step (4), returning it to step (1) to be mixed with the coker gasoline and diesel feedstock.

23. The method according to claim 22, wherein, The hydrogen sulfide concentration in the material of the gas phase obtained in step (4) after removing hydrogen sulfide is below 1500 ppm.

24. The method according to claim 23, wherein The hydrogen sulfide concentration in the material of the gas phase obtained in step (4) after removing hydrogen sulfide is 50 - 1000 ppm.

25. The method according to any one of claims 1-24, wherein, The hydrogenation product is gasoline and diesel.

26. The method according to claim 25, wherein, The sulfur content in the gasoline is 50 - 1500 ppm.

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