Method for producing gasoline by catalytic hydrogenation of diesel

Abstract

The present invention discloses a method for producing gasoline by catalytic diesel hydrogenation. The method comprises: (1) passivating raw oil and C2-C6 alkanes, mixing them with hydrogen and entering a hydrogenation reactor, and sequentially contacting with a hydrofining catalyst and a hydrocracking catalyst for a passivation reaction; (2) after the passivation reaction is completed, switching the passivated raw oil to catalytic diesel in batches until the final feed is a mixture of catalytic diesel and C2-C6 alkanes; (3) the mixture of catalytic diesel and C2-C6 alkanes and hydrogen sequentially contact with the passivated hydrofining catalyst and the hydrocracking catalyst, respectively, and performing a hydrofining reaction and a hydrocracking reaction, respectively, to obtain gasoline. The present invention can adopt an existing trickle bed hydrogenation reaction device, can shorten the adjustment time for the octane number of gasoline products to reach a qualified level in the initial stage of start-up, and can better reduce investment.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining and relates to a method for producing gasoline by catalytic diesel hydrogenation. Background Art

[0002] One of the main technological means to lighten heavy oil is catalytic cracking technology. In recent years, as the quality of raw materials processed by catalytic cracking units has become increasingly poor, the quality of catalytic cracking products, especially catalytic cracking light cycle oil, has deteriorated further. Catalytic cracking light cycle oil has high sulfur and aromatic content and poor engine ignition performance. It is mainly used for blending fuel oil, non-automotive diesel and heating oil. The depth of hydrogenation during catalytic diesel hydrogenation conversion determines the loss of octane number.

[0003] At present, the catalytic diesel hydroconversion units in China all have the problem that the initial activity of the catalyst is high at the beginning of the operation, the hydrodesulfurization, denitrogenation and dearomatization capabilities are strong, the hydrocracking catalyst has high hydrogenation performance, and the initial matching of hydrogenation activity and cracking activity is poor. This shows that the octane number of the gasoline component is not high at the beginning of the operation, which cannot meet the production requirements. It needs to be stabilized for a long time before the octane number of the produced gasoline can meet the requirement of not less than 90.

[0004] US4971680 discloses a hydrocracking process. Its hydrocracking catalyst can selectively crack aromatics, especially bicyclic aromatics, in diesel into gasoline components. While this method cracks and produces gasoline, due to the presence of a large amount of hydrogen, a hydrogenation reaction occurs at the same time, resulting in a low octane number of the gasoline product, which still needs further treatment.

[0005] CN104611050A discloses a catalytic cracking diesel conversion method. The method adopts a single-stage series process, and the hydrofining reaction effluent directly enters the hydrocracking reactor and contacts and reacts with the graded catalyst bed in the cracking reactor; according to the flow direction of the reaction material, the hydrogenation activity of the hydrocracking catalyst tends to decrease. The inventive method can reduce the over-hydrogenation and secondary cracking of cracked naphtha while ensuring the diesel hydrocracking effect, reduce chemical hydrogen consumption, and thus improve the octane number and liquid yield of naphtha. However, the problem of high catalyst activity in the initial stage of the device is not considered, and the device adjustment time cannot be shortened. Summary of the invention

[0006] In view of the shortcomings of the prior art, the purpose of the present invention is to provide a method for catalytic diesel hydrogenation to produce gasoline, especially by using the existing trickle bed hydrogenation reaction device, which can shorten the adjustment time for the octane number of the gasoline product to reach a qualified level in the initial stage of start-up, and can better reduce investment.

[0007] The present invention provides a method for producing gasoline by catalytic hydrogenation of diesel, comprising the following steps:

[0008] (1) Passivating the feedstock oil and C2-C6 alkanes, mixing them with hydrogen, and entering the hydrogenation reactor, where they sequentially contact with the hydrorefining catalyst and the hydrocracking catalyst for passivation reaction;

[0009] (2) After the passivation reaction is completed, the passivation feedstock oil is switched to catalytic diesel in batches until the final feedstock is a mixture of catalytic diesel and C2-C6 alkanes;

[0010] (3) Catalytically treating a mixture of diesel and C2-C6 alkanes and hydrogen in turn with a passivated hydrotreating catalyst and a hydrocracking catalyst to perform a hydrotreating reaction and a hydrocracking reaction, respectively, to obtain gasoline.

[0011] Furthermore, the density of the passivated feedstock oil is 0.84-0.89 g / cm -3 The sulfur content is not less than 6000μg / g, preferably 9000μg / g~15000μg / g, the nitrogen content is not less than 50μg / g, preferably 80μg / g~300μg / g, the initial distillation point is 150℃~300℃, preferably 150℃~250℃, and the final distillation point is 370℃~420℃.

[0012] Furthermore, the passivated feedstock oil is a heavy diesel fraction, such as one or more of normal second-line, normal third-line, normal fourth-line, reduced top oil, reduced first-line, straight-run diesel, etc., preferably one or more of normal fourth-line, reduced top oil and reduced first-line.

[0013] Furthermore, the C2-C6 alkane is one or more of ethane, propane, n-butane, isobutane, n-pentane, isopentane, neopentane, cyclopentane, n-hexane, cyclohexane, methylcyclopentane, 2-methylpentane, 3-methylpentane, 2,2-methylbutane, 2,3-dimethylbutane, etc., preferably one or more of propane, n-butane, isobutane, n-pentane, and isopentane.

[0014] Furthermore, the hydrogenation reactor is a trickle bed reactor.

[0015] Furthermore, the hydrotreating catalyst and the hydrocracking catalyst have been subjected to conventional sulfurization or activation treatment operations before the passivation reaction.

[0016] Furthermore, the conditions of the passivation reaction are: reaction temperature is 310°C to 380°C, preferably 320°C to 350°C, total pressure is 2.5 to 20.0 MPa, preferably 6.0 to 15.0 MPa, volume space velocity is 0.5 h -1 ~7.0h -1 , preferably 0.8h -1 ~6h -1The volume ratio of hydrogen to oil is 300:1 to 2000:1, preferably 600:1 to 1500:1, and the total passivation time is 40 to 100 hours.

[0017] Furthermore, the passivation reaction includes two stages. In the first stage, C2-C6 alkanes account for 10%-18% of the total volume of hydrogen, preferably 12%-16%, and the operation time is 20-72 hours, preferably 24-50 hours; in the second stage, C2-C6 alkanes account for 20%-30% of the total volume of hydrogen, preferably 20%-25%, and the operation time is 16-60 hours, preferably 20-35 hours.

[0018] Further, the hydrorefining catalyst includes a carrier and a hydrogenation active component, wherein the carrier is generally a non-acidic or weakly acidic refractory porous oxide, such as at least one of alumina, silicon oxide, amorphous silicon aluminum, titanium oxide, and a composite oxide or mixed oxide carrier of several elements. The hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni. Based on the weight of the hydrorefining catalyst, the content of the hydrogenation active component in terms of oxide is generally 15wt% to 50wt%, preferably 18wt% to 45wt%. Among them, the content of nickel oxide is further preferably 2wt% to 8wt%, more preferably 2.6wt% to 6wt%. The hydrorefining catalyst can select a suitable commercial catalyst according to the needs of the process flow, can also be prepared according to the existing method, or can be a regenerated catalyst after the deactivated catalyst is regenerated. For example, it can be FHUDS-8 produced by Fushun Branch of Sinopec Catalyst Company.

[0019] Furthermore, the hydrocracking catalyst comprises a carrier and a hydrogenation active component, wherein the carrier is an acidic material, such as Y-type and / or USY molecular sieve. Based on the weight of the hydrocracking catalyst, the content of the carrier is 20wt% to 70wt%, preferably 30wt% to 60wt%; the hydrogenation active component is one or more of W, Mo, Ni and Co, and based on the weight of the hydrocracking catalyst, the content of the hydrogenation active component in terms of oxide is generally 3wt% to 50wt%. The hydrocracking catalyst can be selected from suitable commercial catalysts according to the needs of the process flow, or can be prepared according to existing methods, or can be a regenerated catalyst after a deactivated catalyst is regenerated. For example, it can be FC-24 produced by Fushun Branch of Sinopec Catalyst Company.

[0020] Furthermore, the volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 3:7 to 7:3, preferably 4:6 to 6:4.

[0021] Furthermore, the initial distillation point of the catalytic diesel is 65°C to 290°C, preferably 100°C to 200°C, the final distillation point is 330°C to 450°C, the aromatic content of the catalytic diesel is ≥50wt%, preferably 50wt% to 75wt%, and the polycyclic aromatic hydrocarbon content is ≥35wt%, preferably 40wt% to 50wt%.

[0022] Furthermore, the catalytic diesel can be a product of various catalytic cracking processes obtained by processing any base oil, such as one or more of DCC diesel, MIP diesel, RTC diesel, LTAG diesel, FDFCC diesel, CPP diesel and ARGG diesel. The catalytic diesel can also be one or more of mixed aromatics extraction unit raffinate, coal liquefaction diesel, coal tar diesel and shale oil diesel fraction.

[0023] Furthermore, the passivated stock oil is switched to catalytic diesel in batches, specifically, based on the total weight of the passivated stock oil and the catalytic diesel, the proportion of the catalytic diesel is increased by 10% to 35% each time.

[0024] Furthermore, each time the proportion of catalytic diesel is increased, the proportion of C2-C6 alkanes in the total volume of hydrogen is reduced by 1% to 4%.

[0025] Furthermore, when the final feed is a mixture of catalytic diesel and a passivation aid, the proportion of C2-C6 alkanes in the total volume of hydrogen is 10%-15%.

[0026] Furthermore, the interval between each switching of catalytic diesel is 5 to 15 hours.

[0027] Furthermore, the process conditions of the hydrotreating reactor are as follows: reaction temperature is 260°C to 420°C, preferably 310°C to 390°C; volume space velocity is 0.4h -1 ~10.0h -1 , preferably 0.6h -1 ~5h -1 ; The volume ratio of hydrogen to oil is 200:1 to 2000:1, preferably 500:1 to 1500:1, and the total pressure is 2.5 to 20.0 MPa, preferably 6.0 to 15.0 MPa.

[0028] Furthermore, the conditions of the hydrocracking reaction include: a reaction temperature of 350°C to 450°C, preferably 360°C to 420°C; a volume space velocity of 0.4h -1 ~16.0h -1 , preferably 0.6h -1 ~6h -1The volume ratio of hydrogen to oil is 200:1 to 2000:1, preferably 500:1 to 1500:1, and the total pressure is 2.5 to 20.0 MPa, preferably 6.0 to 15.0 MPa.

[0029] Furthermore, the reaction products obtained through the hydrofining reaction and the hydrocracking reaction are separated into gas and liquid, and then enter the fractionation zone to obtain gasoline through stripping and fractionation.

[0030] Furthermore, the sulfur content of the obtained gasoline product is ≯10 μg / g, and the research octane number is ≮92.

[0031] Compared with the prior art, the method of the present invention has the following advantages:

[0032] 1. The present invention adopts the existing catalytic diesel hydroconversion process, strengthens the catalyst passivation stage at the initial stage of operation, selects the passivated raw oil and C 2～ C6 alkanes can be used to reduce the activity of the catalyst and achieve a controllable carbon deposition process of the catalyst.

[0033] 2. The present invention introduces C at the beginning of the catalyst passivation process. 2～ C6 alkanes act as passivation aids and exist in gaseous form in a high-temperature atmosphere, achieving the effect of changing the hydrogen partial pressure of the reaction system at different times, realizing partial carbon deposition on the catalyst, avoiding excessive carbon deposition on the catalyst, matching the hydrogenation activity of the catalyst system with the cracking performance, and also avoiding the situation where the operating pressure of the circulating hydrogen compressor is too low and the compressor cannot work.

[0034] 3. The method of the present invention comprises adding C 2～ C6 alkanes can be separated through oil separation and distillation systems to achieve controllable operation, avoid excessive enrichment of circulating hydrogen, and quickly switch to normal production conditions without affecting unit operation and other product quality.

[0035] 4. The method of the present invention introduces the recycling of liquefied gas and light gasoline, which can further reduce the hydrogen partial pressure, especially for the hydrogenation conversion catalyst, which is more conducive to exerting the cracking activity and reducing its hydrogenation performance.

[0036] 5. The present invention strengthens the initial passivation process, optimizes the passivation medium, fully matches the activity, and controls the carbon deposition rate of the catalyst, thereby shortening the period in which the octane number of the gasoline product meets the octane number requirements of the 92-grade national standard gasoline. In addition, the method has few operating steps, simple cutting process, convenient operation, and good safety. DETAILED DESCRIPTION

[0037] The following non-limiting embodiments may enable those skilled in the art to more fully understand the present invention, but are not intended to limit the present invention in any way.

[0038] The hydrorefining catalyst used in the embodiments and comparative examples of the present invention is FHUDS-8 developed by Fushun Petrochemical Research Institute and produced by Fushun Branch of Sinopec Catalyst Company; the hydrocracking catalyst is FC-24 developed by Fushun Petrochemical Research Institute and produced by Fushun Branch of Sinopec Catalyst Company. The FHUDS-8 catalyst uses alumina as a carrier and Mo-Ni as an active component; the FC-24 catalyst uses a modified Y-type molecular sieve as a conversion component and W-Ni as an active metal component.

[0039] The conventional one-stage series hydrocracking process used in the comparative example of the present invention has a trickle bed reactor as the reactor.

[0040] Example 1

[0041] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 1:1. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, propane and hydrogen entered the hydrogenation reactor, where the propane injection amount accounted for 15% of the total volume of hydrogen, and the reaction temperature was adjusted to 346°C in the hydrorefining zone and 356°C in the hydrocracking zone, and the volume space velocity was controlled to 0.5h -1 After the constant temperature operation, the second stage of catalyst passivation was started. The ratio of propane injection volume to total hydrogen volume was increased to 25%. Other conditions remained unchanged and the constant temperature was continued for 27 hours.

[0042] After the passivation constant temperature is completed, based on the total weight of the passivated raw oil and catalytic diesel, switch to 25% catalytic diesel, and reduce the volume ratio of propane to the total hydrogen to 21%, and keep the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 50% again, and reduce the volume ratio of propane to the total hydrogen to 18%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 75% again, and reduce the volume ratio of propane to the total hydrogen to 15%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 100%, and reduce the volume ratio of propane to the total hydrogen to 13%, and continue the constant temperature for 12 hours. After this process is completed, stop adding propane. Carry out subsequent hydrofining and hydrocracking reactions.

[0043] Example 2

[0044] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 1:1. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, pentane and hydrogen entered the hydrogenation reactor. The pentane injection amount accounted for 16% of the total volume of hydrogen, and the reaction temperature was adjusted to 346°C in the hydrorefining zone and 356°C in the hydrocracking zone. The volume space velocity was controlled to 0.5h -1After the constant temperature operation, the second stage of catalyst passivation was started. The ratio of pentane injection volume to total hydrogen volume was increased to 30%. Other conditions remained unchanged and the constant temperature was continued for 35 hours.

[0045] After the passivation constant temperature is completed, based on the total weight of the passivated raw oil and catalytic diesel, switch to 20% catalytic diesel, and reduce the total volume ratio of pentane to hydrogen to 26%, and keep the constant temperature interval for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 40% again, and reduce the total volume ratio of pentane to hydrogen to 22%, and continue the constant temperature interval for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 60% again, and reduce the total volume ratio of pentane to hydrogen to 18%, and continue the constant temperature interval for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 80%, and reduce the total volume ratio of pentane to hydrogen to 14%, and continue the constant temperature interval for 12 hours. After the constant temperature is completed, increase the proportion of catalytic diesel to 100%, and reduce the total volume ratio of pentane to hydrogen to 10%, and continue the constant temperature interval for 12 hours. After this process is completed, stop adding propane. Carry out subsequent hydrofining and hydrocracking reactions.

[0046] Example 3

[0047] Compared with Example 1, the only difference is that the passivation auxiliary agent is changed to isobutane.

[0048] Comparative Example 1

[0049] The sulfurized FHUDS-8 and FC-24 catalysts are loaded into the hydrogenation reactor in a volume ratio of 1:1. The passivated feedstock oil and hydrogen enter the hydrogenation reactor, and the reaction temperature is adjusted to 346°C in the hydrotreatment zone and 356°C in the hydroconversion zone. The air velocity is controlled to 0.50 and the constant temperature is operated for 75 hours. After the passivation constant temperature is completed, the catalytic diesel is switched four times. Based on the total weight of the passivated feedstock oil and catalytic diesel, the proportion of catalyst diesel is increased by 25% each time. After each increase in the proportion of catalyst diesel, the constant temperature interval is 12 hours. Finally, switch to 100% catalytic diesel and continue the constant temperature interval for 12 hours. After this process is completed, stop adding the passivator.

[0050] Comparative Example 2

[0051] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 1:1. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, propane and hydrogen entered the hydrogenation reactor, where the propane injection amount accounted for 15% of the total volume of hydrogen, and the reaction temperature was adjusted to 346°C in the hydrorefining zone and 356°C in the hydrocracking zone, and the volume space velocity was controlled to 0.5h -1 After the constant temperature operation, the second stage of catalyst passivation was started. The ratio of propane injection volume to total hydrogen volume was increased to 25%. Other conditions remained unchanged and the constant temperature was continued for 27 hours.

[0052] After the passivation is finished, all the passivated raw oil is directly switched to catalytic diesel, and the ratio of the injection amount of the passivator to the total volume of hydrogen is reduced to 13%, and subsequent hydrofining and hydrocracking reactions are carried out.

[0053] The properties of the raw oil grade passivated raw oil used in Examples 1-3 and Comparative Examples 1-2 are shown in Table 1. The main operating process conditions and gasoline product properties in Examples 1-3 and Comparative Examples 1-2 of the present invention are shown in Table 2.

[0054] Table 1 Properties of passivated feedstock oil and feedstock oil

[0055]

[0056]

[0057] Table 2 Main operating process conditions and gasoline product properties

[0058]

[0059] It can be seen from the data results of Examples 1-3 and Comparative Examples 1-2 that the catalyst passivation method for the catalytic diesel hydroconversion process of the present invention, by optimizing the passivation raw materials, formulating the passivation medium, and the gradient passivation steps, effectively controls the catalyst carbon deposition rate, realizes the passivation process of the enhanced catalyst, and greatly shortens the gasoline product qualification cycle.

Claims

1. A method for producing gasoline by catalytic hydrogenation of diesel, characterized in that: The following steps are involved: (1) The deactivated feedstock oil and C2-C6 alkanes are mixed with hydrogen and enter the hydrogenation reactor, where they are contacted with the hydrorefining catalyst and the hydrocracking catalyst in sequence for deactivation reaction; (2) After the passivation reaction is completed, the passivation feedstock oil is switched to catalytic diesel in batches until the final feedstock is a mixture of catalytic diesel and C2-C6 alkanes; (3) Catalyzing the mixture of diesel and C2-C6 alkanes and hydrogen in turn in contact with a passivated hydrofining catalyst and a hydrocracking catalyst to perform a hydrofining reaction and a hydrocracking reaction, respectively, to obtain gasoline; The passivation reaction includes two stages. In the first stage, C2-C6 alkanes account for 10%-18% of the total volume of hydrogen, and the operation time is 20-72 hours. In the second stage, C2-C6 alkanes account for 20%-30% of the total volume of hydrogen, and the operation time is 16-60 hours. The passivated feedstock oil is switched to catalytic diesel in batches, specifically, the proportion of catalytic diesel is increased by 10% to 35% each time based on the total weight of the passivated feedstock oil and the catalytic diesel; Each time the proportion of catalytic diesel is increased, the proportion of C2~C6 alkanes in the total volume of hydrogen is reduced by 1%~4%. When the final feed is a mixture of catalytic diesel and C2~C6 alkanes, the proportion of C2~C6 alkanes in the total volume of hydrogen is 10%~15%.

2. The method according to claim 1, characterized in that The density of the passivated raw oil is 0.84-0.89 g / cm -3 , the sulfur content is not less than 6000 mg / g, the nitrogen content is not less than 50 mg / g, the initial distillation point is 150℃~300℃, and the final distillation point is 370℃~420℃.

3. The method according to claim 2, characterized in that The sulfur content of the passivated raw oil is 9000mg / g-15000mg / g, the nitrogen content is 80mg / g-300mg / g, and the initial distillation point is 150°C-250°C.

4. The method according to claim 1, characterized in that: The C2-C6 alkane is one or more 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.

5. The method according to claim 1, characterized in that The conditions of the passivation reaction are: reaction temperature of 310°C to 380°C, total pressure of 2.5 to 20.0 MPa, volume space velocity of 0.5 h -1 ~7.0h -1 The volume ratio of hydrogen to oil is 300:1~2000:1, and the total passivation time is 40~100h.

6. The method according to claim 5, characterized in that The conditions of the passivation reaction are: reaction temperature of 320°C to 350°C, total pressure of 6.0 to 15.0 MPa, volume space velocity of 0.8 h -1 ~6h -1 , the volume ratio of hydrogen to oil is 600:1~1500:

1.

7. The method according to claim 1, characterized in that The passivation reaction includes two stages. In the first stage, C2-C6 alkanes account for 12%-16% of the total volume of hydrogen, and the operation time is 24-50 hours. In the second stage, C2-C6 alkanes account for 20%-25% of the total volume of hydrogen, and the operation time is 20-35 hours.

8. The method according to claim 1, characterized in that The initial distillation point of the catalytic diesel is 65° C. to 290° C., the final distillation point is 330° C. to 450° C., the aromatic content of the catalytic diesel is ≥50wt%, and the polycyclic aromatic hydrocarbon content is ≥35wt%.

9. The method according to claim 8, characterized in that The initial distillation point of the catalytic diesel is 100-200° C., the aromatic content of the catalytic diesel is 50wt%-75wt%, and the polycyclic aromatic hydrocarbon content is 40wt%-50wt%.

10. The method according to claim 1, characterized in that The process conditions of the hydrofining reaction are: reaction temperature of 260℃~420℃, volume space velocity of 0.4h -1 ~10.0h -1 , the volume ratio of hydrogen to oil is 200:1~2000:1, and the total pressure is 2.5~20.0MPa; and / or, The conditions of the hydrocracking reaction include: a reaction temperature of 350°C to 450°C, a volume space velocity of 0.4h -1 ~16.0h -1 The volume ratio of hydrogen to oil is 200:1~2000:1, and the total pressure is 2.5~20.0MPa.

11. The method according to claim 10, characterized in that The process conditions of the hydrofining reaction are: reaction temperature of 310°C to 390°C; volume space velocity of 0.6h -1 ~5h -1 ; The volume ratio of hydrogen to oil is 500:1~1500:1, and the total pressure is 6.0~15.0MPa; and / or, The conditions of the hydrocracking reaction include: a reaction temperature of 360°C to 420°C; a volume space velocity of 0.6 h -1 ~6h -1 The volume ratio of hydrogen to oil is 500:1~1500:1, and the total pressure is 6.0~15.0MPa.

12. The method according to claim 1, characterized in that The sulfur content of the obtained gasoline product is ≯10mg / g, and the research octane number is ≮92.

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