A processing method for coking gasoline and diesel

Abstract

The present invention discloses a processing method for coking gasoline and diesel oil, including the following: The coking gasoline and diesel oil raw materials are cut into heavy diesel fraction, jet fuel fraction and gasoline fraction; The first, second and third hydrogenation reaction zones are arranged in the fixed-bed reactor; The jet fuel fraction enters from below the third hydrogenation reaction zone, the gasoline fraction and hydrogen enter from above the first hydrogenation reaction zone, the heavy diesel fraction enters in liquid phase from above the second hydrogenation reaction zone, the effluent from the top of the second hydrogenation reaction zone is mixed with the jet fuel fraction and enters the third hydrogenation reaction zone, and the reaction effluent from the third hydrogenation reaction zone flows out from the top; The effluent from the bottom of the second hydrogenation reaction zone dissolved with hydrogen enters the first hydrogenation reaction zone for deep hydrodesulfurization reaction. After the effluent from the bottom of the first hydrogenation reaction zone and the reaction effluent from the third hydrogenation reaction zone are subjected to gas-liquid separation, ethylene cracking stock, jet fuel and diesel oil are obtained through stripping and fractionation. The present invention avoids excessive hydrodesulfurization reaction, reduces hydrogen consumption and improves liquid yield.

Description

Technical Field

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

[0002] Delayed coking is one of the important means for refining enterprises to process inferior heavy oils. However, the products of delayed coking generally need to be further purified to remove impurities. For example, coker gasoline and diesel oil contain high levels of impurities such as sulfur, nitrogen, olefins, gums, and silicon-containing compounds. The gasoline fraction is used as a raw material for units such as ethylene cracking, fertilizers, and reforming. The aviation kerosene and diesel fractions can be used to produce high-quality fuel oils, but they must all undergo hydrofining and desiliconization treatments. The silicon-containing compounds come from the silicon-containing defoaming agent added during the coking process of delayed coking. The silicon in the silicon-containing defoaming agent enters the reactor with the feedstock oil and deposits on the catalyst, resulting in permanent deactivation of the catalyst due to silicon poisoning and also causing a short single-start cycle of the unit. Multiple literature reports such as "Direct Determination of Silicon Content in Coker Distillate Oil by ICP-OES" show that the main product formed after the thermal cracking of polydimethylsiloxane as a defoaming agent is cyclic siloxane, and most of it exists in the naphtha fraction. The optimal reaction temperature for the silicon-trapping catalyst used for desiliconization is above 270°C, and it is usually used in a graded manner with the hydrodesulfurization catalyst.

[0003] Coker gasoline and diesel oil themselves contain high levels of impurities such as silicon, sulfur, nitrogen, olefins, gums, and silicon, and need to be used after hydrotreating. The article "Research on the Production of Jet Fuel by Hydrotreating Coker Gasoline and Diesel Oil" on pages 173-176 of the 3rd issue of Petrochemical Technology & Application in 2004 introduced the method of mixing coker gasoline and diesel oil for deep hydrofining and then fractionating to obtain gasoline, kerosene, and diesel, etc. This method only deeply refines the entire coker fraction and then separates the kerosene fraction from it, resulting in over-hydrogenation. The article "Hydrofining of Coker Gasoline and Diesel Oil under Low Pressure with FH-5 Catalyst" published by Sun Guang in the 5th issue of Refining Technology and Engineering in 1994 used a set of hydrofining devices to hydrofine coker gasoline and coker diesel. Specifically, a switching-feed 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, continuous switching is required, and the operating conditions need to be frequently changed, which brings inconvenience to the operation and has a greater impact on the performance of the catalyst.

[0004] CN112852480A discloses a coker gasoline hydrotreating catalyst grading and hydrotreating method, including: sequentially arranging a replaceable zone and a main reaction zone in series along the material flow direction. Among them, the replaceable zone sequentially fills a diolefin saturation catalyst in the first reactor and a silicon-trapping agent in the second reactor along the material flow direction, and the second reactor is a parallel double reactor that can be switched online; the main reaction zone sequentially fills an arsenic-removing agent and a gasoline hydrotreating catalyst along the material flow direction.

[0005] CN101003751A discloses a processing method for coking whole fraction oil, including: separating the coking whole fraction oil into coking light fraction oil and coking heavy fraction oil, wherein the light fraction oil contains part of light diesel fraction, and the heavy fraction oil is heavy diesel fraction with a higher dry 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. This method has problems such as over-hydrogenation of light components and over-desulfurization.

[0006] CN111321005A discloses a hydroprocessing method for producing diesel oil 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 the first hydrogenation reaction zone, and then reacts with the light fraction diesel raw material oil in the second hydrogenation reaction zone under the action of a hydrodesulfurization catalyst.

[0007] 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 the first hydrogenation reaction zone and contacted with a hydrofining catalyst, the heavy fraction is sent to the second hydrogenation reaction zone and contacted with a hydrofining catalyst, and the liquid phase products of the two reaction zones are mixed into a refined fraction oil. This process requires the raw material to be fractionated into light and heavy components, and the process is relatively complex.

[0008] The processes selected in the above patents are all conventional gas-phase circulating trickle-bed hydrogenation processes. However, they all fail to consider the matching of reaction temperature rise, desiliconization catalyst and reaction fraction desulfurization requirements. Some products have the situation of over-hydrodesulfurization, which is likely to cause problems such as excessive product quality, reduced liquid yield, excessive hydrogen consumption and energy consumption. Summary of the Invention

[0009] Aiming at the deficiencies of the prior art, the purpose of the present invention is to provide a processing method for coking gasoline and diesel, which deeply couples the reaction fraction, reaction sequence and reaction temperature, realizes reactions such as desulfurization, desiliconization and dealkylation and mercaptan desulfurization, avoids over-hydrodesulfurization reaction, reduces hydrogen consumption, and improves liquid yield; at the same time, reduces the reaction severity, can better play the role of each reaction zone, and better utilizes the heat balance, reduces the energy consumption of the device, and produces high-quality diesel oil, jet fuel and naphtha for ethylene cracking.

[0010] The processing method of coking gasoline and diesel oil of the present invention includes the following: The coking gasoline and diesel oil raw materials are cut into three fractions: heavy diesel fraction, aviation kerosene fraction and gasoline fraction; A first hydrogenation reaction zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone are arranged from bottom to top in the fixed-bed reactor; A hydrodesulfurization catalyst is loaded in the first hydrogenation reaction zone, a silicon capture agent and a hydrodearomatization catalyst are loaded in the second hydrogenation reaction zone, and a hydrodesulfurization mercaptan catalyst is loaded in the third hydrogenation reaction zone; The aviation kerosene fraction enters from below the third hydrogenation reaction zone, the gasoline fraction and hydrogen enter from above the first hydrogenation reaction zone in a gaseous state, and the heavy diesel fraction enters from above the second hydrogenation reaction zone in a liquid state. The gas-liquid countercurrent contact is carried out for desiliconization, dearomatization and desulfurization reactions. The effluent from the top of the second hydrogenation reaction zone is mixed with the aviation kerosene fraction and enters the third hydrogenation reaction zone for hydrodesulfurization mercaptan and desulfurization reactions. The reaction effluent from the third hydrogenation reaction zone flows out from the top; The effluent from the bottom of the second hydrogenation reaction zone dissolved with hydrogen enters the first hydrogenation reaction zone for deep hydrodesulfurization reaction. After the effluent from the bottom of the first hydrogenation reaction zone and the reaction effluent from the third hydrogenation reaction zone are subjected to gas-liquid separation, ethylene cracking stock, aviation kerosene and diesel oil are obtained through stripping and fractionation.

[0011] In the method of the present invention, the initial boiling point of the coking gasoline and diesel oil raw materials is 30°C to 100°C, and the final boiling point temperature is 350°C to 400°C. The proportion of coking gasoline and diesel oil in the coking gasoline and diesel oil raw materials is more than 50%, and one or more of straight-run diesel oil, residue-added diesel oil, wax-added diesel oil, straight-run aviation kerosene, catalytic gasoline, ethylene cracking gasoline, straight-run naphtha and other oil products can be mixed for processing; The coking gasoline and diesel oil raw materials contain 2 to 500 μg / g of silicon -1 , the olefin content is not higher than 35%, the sulfur content is not higher than 15,000 μg / g -1 , and the nitrogen content is not higher than 1500 μg / g -1 . In the method of the present invention, the cutting points for cutting the coking gasoline and diesel oil raw materials into heavy diesel fraction, aviation kerosene fraction and gasoline fraction are 220°C to 280°C and 150°C to 200°C.

[0012] In the method of the present invention, the first, second and third hydrogenation reaction zones can also be three reactors used in series; Several catalyst beds can be set in each hydrogenation reaction zone according to needs, and preferably 1 to 3 are set.

[0013] In the method of the present invention, the reaction conditions of the first hydrogenation reaction zone are: hydrogen partial pressure is 4.0 MPa to 16.0 MPa, preferably 8.0 MPa to 12.0 MPa, and the volume space velocity is 0.3 h -1 ~10.0 h -1 , and the average reaction temperature is 330°C to 420°C, preferably 340°C to 410°C.

[0014] In the method of the present invention, the reaction conditions in the second hydrogenation reaction zone are as follows: the hydrogen partial pressure is 1.0 MPa to 10.0 MPa, preferably 2.0 MPa to 8.0 MPa, and the volume hourly space velocity is 0.8 h -1 to 10.0 h -1 , and the hydrogen-oil volume ratio is 100:1 to 2500:1; the average reaction temperature in the second hydrogenation reaction zone is 60 °C to 130 °C lower than the average reaction temperature in the first hydrogenation reaction zone.

[0015] In the method of the present invention, the reaction temperature at the gas-phase inlet at the lower part of the second hydrogenation reaction zone is 265 °C to 330 °C, preferably 275 °C to 310 °C.

[0016] In the method of the present invention, the pressure in the second hydrogenation reaction zone is reduced by 2.0 MPa to 8.0 MPa compared with the first hydrogenation reaction zone.

[0017] In the method of the present invention, the hydrogen sulfide concentration in the effluent at the top of the second hydrogenation reaction zone is 4000 ppm to 16000 ppm, preferably 5000 ppm to 12000 ppm.

[0018] In the method of the present invention, the reaction conditions in the third hydrogenation reaction zone are as follows: the hydrogen partial pressure is 0.5 MPa to 8.0 MPa, preferably 1.5 MPa to 6.0 MPa, and the volume hourly space velocity is 1.6 h -1 to 10.0 h -1 , and the hydrogen-oil volume ratio is 80:1 to 2500:1, preferably 150:1 to 2000:1.

[0019] In the method of the present invention, the reaction temperature at the inlet of the third hydrogenation reaction zone is 240 °C to 330 °C, preferably 260 °C to 310 °C.

[0020] In the method of the present invention, the pressure in the third hydrogenation reaction zone is reduced by 0.5 MPa to 4.0 MPa compared with the second hydrogenation reaction zone.

[0021] In the method of the present invention, the jet fuel fraction is mixed with the gas-phase product in the second hydrogenation reaction zone and enters the third hydrogenation reaction zone. Alternatively, the jet fuel fraction can be mixed with hydrogen first and then mixed with the gas-phase product in the second hydrogenation reaction zone and enter the third hydrogenation reaction zone.

[0022] In the method of the present invention, the hydrofining catalyst generally uses a refractory porous oxide as the carrier, such as one or more of alumina, silica, amorphous silica-alumina, and titanium oxide. The hydrogenation active combination of the catalyst is W-Ni, Mo-Ni or W-Mo-Ni. The content of the hydrogenation active component based on oxides is generally 15 wt% to 60 wt%, preferably 16 wt% to 46 wt%. Among them, the nickel oxide content is 1.8 wt% to 8.5 wt%, preferably 2.0 wt% to 5.5 wt%. The shape is spherical or bar-shaped. The spherical diameter is 0.04 to 12 mm, preferably 0.04 to 5 mm. The bar-shaped one has a length of 2 to 15 mm, preferably 2 to 8 mm, and a diameter of 1 to 6 mm, preferably 1.5 to 3.5 mm. The specific surface area is 150 to 650 m 2 / g, preferably 200 to 450 m 2 / g, and the pore volume is 0.2 to 1.5 mL / g, preferably 0.30 to 0.85 mL / g. The hydrofining catalyst can select 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.

[0023] In the method of the present invention, the silicon capture agent is well-known to those skilled in the art. As one of the specific embodiments, the active metals of the silicon capture agent are oxides of Group VIB metals and oxides of Group VIII metals. Based on the total weight of the silicon capture agent, the Group VIB metal is 2-25% in terms of oxide, preferably 5-15%, and the Group VIII metal is 2-20% in terms of oxide, preferably 4-15%. The pore volume of the silicon capture agent is 0.5 to 1.0 mL / g, preferably 0.7 to 1.0 mL / g. The specific surface area is 260 to 550 m 2 / g, preferably 350-500 m 2 / g. The amount of B acid is 0.05 to 0.35 mmol / g, preferably 0.07 to 0.21 mmol / g. The preparation method of the silicon capture agent is well-known in the art and can adopt one or more of impregnation method, co-extrusion method, and co-precipitation method.

[0024] In the method of the present invention, the hydrodearomatization catalyst is well-known to those skilled in the art. As one of the specific embodiments, the active metal of the hydrodearomatization catalyst is a Group VIB metal oxide or / and a Group VIII metal oxide. The carrier is alumina or alumina modified with an auxiliary agent. Based on the weight of the hydrodearomatization catalyst, the content of the Group VIB metal oxide is 5% - 30%, preferably 5% - 15%, and the content of the Group VIII metal oxide is 1% - 15%, preferably 1.5% - 6%. The shape is spherical or cylindrical. The spherical diameter is 0.2 - 20 mm, preferably 0.3 - 5 mm; the cylindrical shape has a length of 2 - 15 mm, preferably 3 - 8 mm, and a diameter of 1.5 - 6 mm, preferably 2.0 - 5.5 mm. The specific surface area is 100 - 550 m 2 / g, preferably 150 - 420 m 2 / g, and the pore volume is 0.3 - 1.5 mL / g, preferably 0.4 - 0.9 mL / g. The preparation method of the catalyst is well-known in the art and can be one or more of the impregnation method, co-extrusion method, and co-precipitation method. Preferably, it is a regenerated catalyst after the deactivated catalyst is regenerated.

[0025] In the method of the present invention, the hydrodesulfurization catalyst is well-known to those skilled in the art. As one of the specific embodiments, the active metal of the hydrodesulfurization catalyst is a Group VIB metal oxide or / and a Group VIII metal oxide. The carrier is alumina or alumina modified with an auxiliary agent. Based on the weight of the hydrodesulfurization catalyst, the content of the Group VIB metal oxide is 5% - 30%, preferably 5% - 15%, and the content of the Group VIII metal oxide is 1% - 15%, preferably 1.5% - 6%; the specific surface area is 80 - 500 m 2 / g, preferably 100 - 380 m 2 / g, and the pore volume is 0.3 - 1.5 mL / g, preferably 0.4 - 0.9 mL / g. The preparation method of the catalyst is well-known in the art and can be one or more of the impregnation method, co-extrusion method, and co-precipitation method, or it can also be a regenerated catalyst after the deactivated catalyst is regenerated.

[0026] In the method of the present invention, the volume ratio of the catalysts in the first hydrogenation reaction zone, the second hydrogenation reaction zone, and the third hydrogenation reaction zone is 70:20:10 - 40:30:30, preferably 60:25:15 - 50:25:25.

[0027] In the method of the present invention, the volume ratio of the silicon capture agent and the hydrodearomatization catalyst in the second hydrogenation reaction zone is 20:80 - 80:20, preferably 35:65 - 65:35.

[0028] In the process of the present invention, the fractionation system is a system that can accurately fractionate gasoline, jet fuel, and diesel.

[0029] A coking gasoline and diesel treatment method of the present invention has the following beneficial effects:

[0030] 1. The coking gasoline and diesel treatment method provided by the present invention divides the feedstock oil into different fractions, optimizes the catalyst grading according to different product requirements, deeply couples the reaction fractions, reaction sequence, reaction pressure, hydrogen sulfide concentration and various reactions, achieves precise control of the reaction depth, balances the reaction heat, fully exerts the performance of the catalyst, achieves deep desulfurization of the diesel fraction, controls the strong exothermic reaction of olefins, avoids passive desulfurization of the jet fuel fraction and gasoline fraction, and realizes the purpose of progressive controllability of the reaction process.

[0031] 2. The coking gasoline and diesel treatment method provided by the present invention is based on the fact that the best reaction temperature for the hydrodesiliconization reaction needs to be above 270 °C. The coking gasoline is gasified and undergoes a countercurrent reaction with the diesel fraction in the hydrodesiliconization and deolefination reaction zone. Due to the countercurrent reaction, the gasified coking gasoline enters the reaction zone at the highest reaction temperature point of this process. As the gas phase moves upward, the reaction temperature decreases under the action of the circulating oil, making it easier to control the temperature of the desiliconization reaction zone, realizing the matching of the activity of the silicon capture agent and the reactants, effectively controlling the reaction temperature of the silicon capture catalyst bed, and well solving the problem of reaction heat release in the original process and the mismatch of the reaction temperature at the initial stage of the operation of the silicon capture catalyst.

[0032] 3. The first, second and third hydrogenation reaction zones provided by the present invention are respectively equipped with hydrofining catalysts, silicon capture catalysts, hydrodeolefination catalysts and hydrodesulfurization mercaptan catalysts, and the corresponding reaction temperatures and reaction pressures are adjusted. Following the principle that the reactants become easier from difficult, and the reaction temperature and reaction pressure decrease from high to low, it well exerts the high-temperature hydrodesulfurization, desiliconization, hydrodesulfurization mercaptan activity and deolefination performance of the catalyst, avoiding the passive over-desulfurization of sulfur in the jet fuel fraction and gasoline fraction in the conventional coking gasoline and diesel hydrogenation process to ensure ultra-deep removal of thiophene sulfur, especially for the coking gasoline fraction. Under the countercurrent reaction, the desiliconization reaction temperature is effectively controlled to avoid over-desulfurization. This can reduce the sulfur injection amount of the downstream ethylene plant, improve the total liquid yield while increasing the yields of gasoline and jet fuel components, and reduce the hydrogen consumption.

[0033] 4. The diesel fraction provided by the present invention enters the second hydrogenation reaction zone and undergoes a countercurrent reaction with the circulating hydrogen and the gas-phase gasoline fraction to carry out a simple hydrodesulfurization reaction; the reaction product of the diesel fraction in the second hydrogenation reaction zone first reacts with the gas-phase gasoline fraction, and then countercurrently with the circulating hydrogen at the upper part of the first hydrogenation reaction zone. While high-pressure mixed hydrogen is introduced, hydrogen sulfide in the liquid phase is stripped out, so that the diesel fraction with high dissolved hydrogen and low hydrogen sulfide enters the first hydrogenation reaction zone as a liquid phase, making it easier to carry out liquid-phase hydrofining reaction.

[0034] 5. The method of the present invention gradually controls the reaction heat release, effectively matches the heat of fusion and the reaction heat balance, effectively controls the reaction temperature rise, reduces the probability of thermal cracking, and ensures high liquid yield, low hydrogen consumption and energy consumption. BRIEF DESCRIPTION OF THE DRAWINGS

[0035] Figure 1 It is a schematic diagram of the process adopted in the embodiment of the present invention. DETAILED DESCRIPTION OF THE INVENTION

[0036] The method disclosed in the present invention will be described in more detail below in conjunction with the drawings and specific embodiments. Figure 1 Many necessary devices are omitted, such as oil pumps, valves, heating furnaces and other devices.

[0037] As Figure 1 shown, the diesel fraction of coking gasoline and diesel oil enters the upper part of the second hydrogenation reaction zone of the reactor 5 through pipeline 1 and flows downward, forming a countercurrent reaction with the gasoline fraction in pipeline 3 and the recycled hydrogen in pipeline 4, and undergoes hydrodesiliconization, hydrodearomatization and hydrodesulfurization reactions. The diesel fraction flows downward into the first hydrogenation reaction zone for hydrodesulfurization reaction, flows out from the bottom of the reactor, and enters the cold high-pressure separator 8 through pipeline 6; the gas-phase gasoline and recycled hydrogen mixture pass through the second hydrogenation reaction zone, are mixed with the jet fuel fraction in pipeline 2 and enter the third hydrogenation reaction zone for hydrodesulfurization and hydrodesulfurization reactions, and the mixed material flows out from the top of the reactor; enters the cold high-pressure separator 8 through pipeline 7, the separated liquid phase enters the cold low-pressure separator 10 through pipeline 9, and then enters the fractionation unit through pipeline 11; the recycled hydrogen separated by the cold high-pressure separator 8 enters the recycled hydrogen desulfurization system through pipeline 12, and then returns to pipeline 4 through the recycled hydrogen compressor 13.

[0038] The present invention will be further described below in conjunction with preferred embodiments. It should be understood that the preferred embodiments described herein are only used to illustrate and explain the present invention and are not used to limit the present invention.

[0039] In the examples and comparative examples of the present invention, the hydrofining catalyst in the first hydrogenation reaction zone is FHUDS-10 developed by Fushun Research Institute of Petroleum and Petrochemicals and produced by Fushun Branch of Sinopec Catalyst Company; the silicon-capturing agent in the second hydrogenation reaction zone is FHRS-2 developed by Fushun Research Institute of Petroleum and Petrochemicals and produced by Fushun Branch of Sinopec Catalyst Company; the olefin-removing catalyst in the second hydrogenation reaction zone is the regenerated FH-40C developed by Fushun Research Institute of Petroleum and Petrochemicals and produced by Fushun Branch of Sinopec Catalyst Company; the low-temperature mercaptan-removing catalyst in the third hydrogenation reaction zone is the regenerated FH-40A developed by Fushun Research Institute of Petroleum and Petrochemicals and produced by Fushun Branch of Sinopec Catalyst Company. The FHUDS-10 catalyst uses modified alumina as the carrier and Mo-Ni as the active metal components; the FH-40C catalyst uses alumina as the carrier and W-Mo-Ni as the active metal components; the FHRS-2 catalyst uses expanded-pore alumina as the carrier for acidic modification and Mo-Ni as the active metal components; the FH-40A catalyst uses alumina as the carrier for acidic modification and Mo-Ni as the active metal components. Among them, ICP is used to analyze the composition content in the feedstock oil and products, and the silicon types in the feedstock oil and products are analyzed by nuclear magnetic resonance 29Si MAS NMR to be siloxane compounds such as dimethylsiloxane, trimethylsiloxane, and tetramethylsiloxane.

[0040] In Comparative Example 1 of the present invention, the traditional coker gasoline and diesel hydrofining process is adopted. The reactor is a trickle-bed reactor, and the coker gasoline and diesel are mixed with hydrogen and enter the reactor for hydrodealkylation, desiliconization, and desulfurization reactions.

[0041] In Comparative Example 2 of the present invention, the traditional coker gasoline and diesel hydrofining process is adopted. The reactor is a trickle-bed reactor. The reactor is divided into the first, second, and third hydrogenation reaction zones from bottom to top. The first hydrogenation reaction zone is filled with a mercaptan-removing catalyst, the second hydrogenation reaction zone is filled with an olefin-removing catalyst and a silicon-capturing catalyst in sequence, and the third hydrogenation reaction zone is filled with a hydrofining catalyst. The coker diesel fraction is mixed with hydrogen and enters the third hydrogenation reaction zone, the coker gasoline fraction is mixed with the product of the third hydrogenation reaction zone and enters the second hydrogenation reaction zone, and the coker jet fuel fraction is mixed with the product of the second hydrogenation reaction zone and enters the first hydrogenation reaction zone. The feed is divided into regions, and each fraction enters the reactor for hydrodealkylation, desiliconization, and desulfurization reactions.

[0042] The examples of the present invention adopt Figure 1 the hydrogenation process flow. The reactor is a fixed-bed reactor, which is divided into the first, second, and third hydrogenation reaction zones from bottom to top. The first hydrogenation reaction zone is filled with a hydrofining catalyst, the second hydrogenation reaction zone is filled with a silicon-capturing catalyst and an olefin-removing catalyst in sequence, and the third hydrogenation reaction zone is filled with a mercaptan-removing catalyst. The feedstock oils used in the examples and comparative examples are shown in Table 1. The main operating process conditions and gasoline product properties in the examples and comparative examples of the present invention are shown in Table 2.

[0043] Table 1 Properties of feedstock oil

[0044]

[0045] Table 2 Main operating process conditions and gasoline product properties at different times

[0046]

[0047] Table 2 (continued) Main operating process conditions and gasoline product properties at different times

[0048]

[0049] It can be seen from the data results of the examples and comparative examples that the most prominent feature of the method of the present invention is that by simply modifying a conventional coking gasoline and diesel hydrogenation refining unit, adopting the process of a trickle bed reactor, optimizing the catalyst grading method, through fractionation, the reaction phase state, reaction sequence, reaction form, and reaction process conditions of the reactants are deeply coupled, the reaction depth is controlled, the reaction heat is balanced, the performance of the catalyst in each reaction zone is maximally exerted, and the purpose of reducing the hydrogen consumption and energy consumption of the unit is achieved. The liquid yield is increased, and the operation cycle of the coking gasoline hydrogenation unit is extended. It brings considerable economic and social benefits to enterprises and has great practical application advantages.

Claims

1. A processing method for coking gasoline and diesel, characterized in that It includes the following: The coking gasoline and diesel raw materials are cut into three fractions: heavy diesel fraction, jet fuel fraction and gasoline fraction; A first hydrogenation reaction zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone are arranged from bottom to top in the fixed-bed reactor; A hydrofining catalyst is loaded in the first hydrogenation reaction zone, a silicon-capturing agent and a hydrodearomatization catalyst are sequentially loaded in the second hydrogenation reaction zone, and a hydrodesulfurization mercaptan catalyst is loaded in the third hydrogenation reaction zone; The jet fuel fraction enters from below the third hydrogenation reaction zone, the gasoline fraction and hydrogen enter from above the first hydrogenation reaction zone in a gas phase, and the heavy diesel fraction enters from above the second hydrogenation reaction zone in a liquid phase, so that they are in countercurrent gas-liquid contact for desiliconization, dearomatization and desulfurization reactions. The effluent from the top of the second hydrogenation reaction zone is mixed with the jet fuel fraction and then enters the third hydrogenation reaction zone for hydrodesulfurization mercaptan and desulfurization reactions. The reaction effluent from the third hydrogenation reaction zone flows out from the top; The effluent from the bottom of the second hydrogenation reaction zone dissolved with hydrogen enters the first hydrogenation reaction zone for deep hydrodesulfurization reaction. After the effluent from the bottom of the first hydrogenation reaction zone and the reaction effluent from the third hydrogenation reaction zone are subjected to gas-liquid separation, ethylene cracking stock, jet fuel and diesel are obtained through stripping and fractionation; The reaction conditions of the first hydrogenation reaction zone are: the hydrogen partial pressure is 4.0 MPa to 16.0 MPa, and the average reaction temperature is 330 °C to 420 °C; The average reaction temperature of the second hydrogenation reaction zone is 60 °C to 130 °C lower than that of the first hydrogenation reaction zone; The reaction temperature at the lower gas-phase inlet of the second hydrogenation reaction zone is 265 °C to 330 °C; The pressure of the second hydrogenation reaction zone is 2.0 MPa to 8.0 MPa lower than that of the first hydrogenation reaction zone; The reaction temperature at the inlet of the third hydrogenation reaction zone is 240 °C to 330 °C, and the pressure of the third hydrogenation reaction zone is 0.5 MPa to 4.0 MPa lower than that of the second hydrogenation reaction zone.

2. The processing method according to claim 1, characterized in that: The initial boiling point of the coking gasoline and diesel raw material is 30°C to 100°C, and the final boiling point temperature is 350°C to 400°C; the proportion of coking gasoline and diesel in the coking gasoline and diesel raw material is more than 50 wt%, and the silicon content of the coking gasoline and diesel raw material is 2 to 500 μg / g -1 , the olefin content is not higher than 35%, and the sulfur content is not higher than 15,000 μg / g -1 , the nitrogen content is not higher than 1,500 μg / g -1 .

3. The processing method according to claim 1, characterized in that: The coking gasoline and diesel raw materials contain one or more of straight-run diesel, straight-run jet fuel, catalytic gasoline, ethylene cracking gasoline and straight-run naphtha.

4. The processing method according to claim 1, characterized in that: Several catalyst beds are arranged in the first, second and third hydrogenation reaction zones as required.

5. The processing method according to claim 1, characterized in that: 1 to 3 catalyst beds are respectively arranged in the first, second and third hydrogenation reaction zones.

6. The processing method according to claim 1, characterized in that: The reaction conditions of the first hydrogenation reaction zone are as follows: the hydrogen partial pressure is 8.0 MPa to 12.0 MPa, the volume space velocity is 0.3 h -1 to 10.0 h -1 , and the average reaction temperature is 340 °C to 410 °C.

7. The processing method according to claim 1, characterized in that: The reaction conditions in the second hydrogenation reaction zone are as follows: the hydrogen partial pressure is 1.0 MPa to 10.0 MPa, the volume hourly space velocity is 0.8 h -1 to 10.0 h -1 , and the hydrogen-to-oil volume ratio is 100:1 to 2500:

1.

8. The processing method according to claim 1, wherein: The reaction temperature at the lower gas-phase inlet of the second hydrogenation reaction zone is 275 °C to 310 °C.

9. The processing method according to claim 1, characterized in that: The hydrogen sulfide concentration in the effluent from the top of the second hydrogenation reaction zone is 4000 ppm to 16000 ppm.

10. The processing method according to claim 1, wherein: The reaction conditions of the third hydrogenation reaction zone described are as follows: the hydrogen partial pressure is 0.5 MPa to 8.0 MPa, the volume space velocity is 1.6 h -1 to 10.0 h -1 , and the hydrogen-oil volume ratio is 80:1 to 2500:

1.

11. The processing method according to claim 1, characterized in that: The reaction temperature at the inlet of the third hydrogenation reaction zone is 260 °C to 310 °C.

12. The processing method according to claim 1, characterized in that: The jet fuel fraction is mixed with hydrogen, then mixed with the effluent from the top of the second hydrogenation reaction zone, and enters the third hydrogenation reaction zone.

13. The processing method according to claim 1, wherein: The described hydrofining catalyst uses a refractory porous oxide as the carrier, and the refractory porous oxide is selected from one or more of alumina, silica, amorphous silica-alumina, and titanium oxide. The hydrofining activity combination of the hydrofining catalyst is W-Ni, Mo-Ni, or W-Mo-Ni. The content of the hydrofining active component is 15 wt% to 60 wt% based on oxides, and the nickel oxide content is 1.8 wt% to 8.5 wt%. The shape of the hydrofining catalyst is spherical or bar-shaped. The spherical diameter is 0.04 to 12 mm, the bar-shaped length is 2 to 15 mm, the bar-shaped diameter is 1 to 6 mm, the specific surface area is 150 to 650 m 2 / g, and the pore volume is 0.2 to 1.5 mL / g.

14. The processing method according to claim 1, characterized in that: The active metals of the silicon-capturing agent are oxides of Group VIB metals and oxides of Group VIII metals, and the carrier is alumina or alumina modified with additives. Based on the total weight of the silicon-capturing agent, the Group VIB metal is 2-25% in terms of oxide, and the Group VIII metal is 2-20% in terms of oxide; the pore volume of the silicon-capturing agent is 0.5-1.0 mL / g, the specific surface area is 260-550 m 2 / g, and the amount of B acid is 0.05-0.35 mmol / g.

15. The processing method according to claim 1, characterized in that: The active metals of the hydrodearomatization catalyst are metal oxides of Group VIB or / and metal oxides of Group VIII, and the carrier is alumina or alumina modified with additives. Based on the weight of the hydrodearomatization catalyst, the content of the metal oxide of Group VIB is 5% - 30%, and the content of the metal oxide of Group VIII is 1% - 15%. The shape of the hydrodearomatization catalyst is spherical or bar-shaped. The diameter of the spherical catalyst is 0.2 - 20 mm, the length of the bar-shaped catalyst is 2 - 15 mm, the diameter of the bar is 1.5 - 6 mm, the specific surface area is 100 - 550 m 2 / g, and the pore volume is 0.3 - 1.5 mL / g.

16. The processing method according to claim 1, characterized in that: The active metals of the hydrodesulfurization mercaptan catalyst are metal oxides of Group VIB and / or Group VIII metals, and the carrier is alumina or alumina modified with additives. Based on the weight of the hydrodesulfurization mercaptan catalyst, the content of the metal oxide of Group VIB is 5% to 30%, the content of the metal oxide of Group VIII is 1% to 15%, the specific surface area is 80 to 500 m 2 / g, and the pore volume is 0.3 to 1.5 mL / g.

17. The processing method according to claim 1, wherein: The volume ratio of the catalysts in the first hydrogenation reaction zone, the second hydrogenation reaction zone and the third hydrogenation reaction zone is 70:20:10 to 40:30:

30.

18. The processing method according to claim 1, wherein: The volume ratio of the silicon-capturing agent and the hydrodearomatization catalyst in the second hydrogenation reaction zone is 20:80 to 80:20.

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