A combined process for liquefied petroleum gas-catalytic gasoline hydrogenation

Abstract

This invention discloses a combined process for hydrogenating liquefied petroleum gas (LPG) and catalytic gasoline, comprising: (1) a first and a second fixed-bed reactor connected in series for simultaneous sulfidation; after sulfidation, maintaining a CO2 concentration of ≥3000 μg / g in the reaction system, introducing catalytic gasoline feedstock for a period of time, and then switching to hydrogen; (2) the mixture enters the first fixed-bed reactor and reacts with a mercaptan transfer catalyst to obtain a first reaction stream; (3) the first reaction stream undergoes gas-liquid separation to obtain gaseous and liquid products, the gaseous product is desulfurized to obtain LPG, and the liquid product is fractionated to obtain light and heavy fractions; (4) the heavy fraction and hydrogen enter the second fixed-bed reactor and react with a selective hydrogenation catalyst to obtain a second reaction stream, which is then separated into heavy catalytic gasoline products. This process can simultaneously perform efficient selective hydrogenation of LPG and catalytic gasoline, reducing equipment investment costs, lowering hydrogen consumption, and improving production efficiency.

Description

Technical Field

[0001] This invention belongs to the field of clean refining technology, and specifically relates to a combined process of liquefied petroleum gas-catalytic gasoline hydrogenation. Background Technology

[0002] Liquefied petroleum gas (LPG) mainly comes from catalytic cracking and delayed coking units in refineries. It is a mixture of gaseous hydrocarbons that can be used as a clean fuel or, after separation and purification, as a chemical feedstock for producing high-value-added products. However, LPG contains certain amounts of sulfide impurities, especially methanethiol and ethanethiol, which produce harmful SO₂ during combustion. x Emissions can lead to catalyst poisoning and deactivation in downstream processing and promote the oxidation of reactive hydrocarbons in liquefied petroleum gas (LPG) to form gum. Hydrodesulfurization technology can remove mercaptans from LPG, but it easily causes olefin saturation, thus reducing the utilization value of LPG products.

[0003] Catalytic cracking gasoline has a high sulfur and olefin content. To upgrade the quality of gasoline, traditional desulfurization technology saturates high-octane olefin components to generate low-octane alkanes, resulting in a significant decrease in octane number. In order to achieve the dual goals of deep desulfurization of FCC gasoline and minimal loss of octane number, different processes and catalysts are adopted at home and abroad to achieve this goal.

[0004] IFP company Prime-G + The technology consists of a selective hydrotreating unit (SHU), a fractionation tower (to separate LCN from MCN or HCN), and a dual-catalyst hydrotreating unit for processing both MCN and HCN. The process uses FCC gasoline full-fraction feedstock, undergoing hydrotreating pretreatment in the SHU unit to hydrogenate and saturate dienes and isomerize double bonds; small-molecule thiols and sulfides are converted into large-molecule sulfides. Then, fractionation separates the gasoline into olefin-rich light gasoline and sulfur-rich heavy gasoline; the sulfur-rich heavy gasoline is then subjected to selective deep hydrodesulfurization using a dual-catalyst, minimizing olefin saturation and thus minimizing octane number loss.

[0005] CN102311783B discloses a combined process for hydrogenating liquefied petroleum gas (LPG) and coking gasoline: coking gasoline undergoes hydrogenation in a coking gasoline hydrogenation reaction section, and the reaction effluent is mixed with LPG feedstock and enters the LPG hydrogenation reaction section. After separation, hydrogenated LPG and hydrogenated coking gasoline are obtained. This method facilitates the diffusion of reaction heat, lowers the reaction temperature, reduces the influence of temperature on the reaction equilibrium, and can simultaneously obtain high-quality hydrogenated coking gasoline and hydrogenated LPG. CN109777506B discloses a combined processing method for refinery gas hydrogenation. In this method, jet fuel feedstock and circulating oil are mixed with hydrogen in a hydrogen dissolving device, and then fed into a hydrogenation catalyst bed in a jet fuel hydrogenation reactor for reaction under liquid-phase hydrogenation conditions. The resulting reaction stream is mixed with refinery gas and hydrogen in a gas dissolving device and then fed into a supplementary hydrogenation reactor for reaction under liquid-phase hydrogenation conditions. The hydrogenation reaction effluent is separated into gas and liquid phases. The separated gas phase is desulfurized and further separated to obtain hydrogen, refinery gas, naphtha, and jet fuel products. All of the above hydrogenation methods involve first hydrogenating the liquid feedstock and then mixing it with liquefied petroleum gas (LPG) for hydrogenation. While this removes mercaptans from the LPG, it also produces a large amount of olefin saturation, reducing the utilization value of the LPG products. CN103820149A discloses a method for reducing the sulfur content in liquefied petroleum gas (LPG), comprising: a mixing process of hydrogenated or unhydrogenated LPG with other petroleum fractions having a boiling point higher than that of LPG; a hydrogenation reaction process, including a sulfideation reaction, carried out in a fixed-bed reactor on LPG or a mixed fraction containing LPG; and a process for fractionating the mixed fraction to obtain a low-sulfur LPG fraction. This invention can significantly reduce the sulfur content in LPG, and the resulting low-sulfur LPG fraction can be used to produce MTBE or alkylated gasoline (isooctane), but the process is limited to obtaining a low-sulfur LPG fraction. Summary of the Invention

[0006] To address the shortcomings of existing technologies, this invention provides a combined process for hydrogenating liquefied petroleum gas (LPG) and catalytic gasoline. This process can simultaneously and efficiently selectively hydrogenate both LPG and catalytic gasoline, enabling the simultaneous production of high-value LPG products and gasoline products with minimal octane number loss. This reduces equipment investment costs, lowers hydrogen consumption, improves production efficiency, and effectively solves the problem of high temperature rise during LPG hydrogenation.

[0007] The liquefied petroleum gas-catalytic gasoline hydrogenation combined process of the present invention includes the following:

[0008] (1) The first fixed-bed reactor and the second fixed-bed reactor are connected in series. The first fixed-bed reactor is filled with a mercaptan transfer catalyst and the second fixed-bed reactor is filled with a selective hydrodesulfurization catalyst. The two reactors are subjected to sulfidation treatment simultaneously. After sulfidation, CO2 is introduced into the two reactors to maintain the CO2 concentration in the reaction system at ≥3000μg / g. Catalytic gasoline feedstock is introduced to react for a period of time and then switched to a hydrogen atmosphere. (2) Catalytic gasoline feedstock, liquefied petroleum gas and hydrogen are mixed in the first dissolved gas device. The mixture enters the first fixed-bed reactor and reacts with the mercaptan transfer catalyst to obtain the first reaction stream. (3) The first reaction stream is separated into gas phase product and liquid phase product. The gas phase product is further desulfurized to obtain liquefied petroleum gas product. The liquid phase product is fractionated to obtain light fraction and heavy fraction. (4) The heavy fraction is mixed with hydrogen in the second dissolved gas device. The mixture enters the second fixed-bed reactor and reacts with the selective hydrodesulfurization catalyst to obtain the second reaction stream. The second reaction stream is separated into gas and liquid to obtain the heavy catalytic gasoline product.

[0009] In the process of this invention, before the sulfidation treatment described in step (1), the first and second fixed-bed reactors have completed the conventional operation process before the sulfidation treatment, such as nitrogen gas sealing, catalyst drying, hydrogen replacement, hydrogen gas sealing and establishing hydrogen circulation.

[0010] In the process of this invention, the sulfidation treatment conditions in step (1) are as follows: the amount of sulfiding agent introduced is 90% to 150% of the theoretical sulfur required by the catalyst; the sulfidation process adopts programmed temperature rise, the temperature is raised to 200 to 230℃ and held at 260 to 290℃ for 4 to 16 hours; the sulfiding agent is generally one or more of carbon disulfide, dimethyl disulfide, methyl sulfide, and n-butyl sulfide.

[0011] In the process of this invention, the amount of CO2 introduced in step (1) is ≥3000μg / g, preferably ≥6000μg / g, based on its concentration in hydrogen.

[0012] In the process of this invention, the catalytic gasoline feedstock introduced in step (1) is fed at a volume hourly space velocity (VHSV) of 2 to 10 h⁻¹. -1 The preferred volume hourly space velocity is 4-8 h⁻¹. -1 The reaction time is 48 to 600 hours, preferably 72 to 480 hours. The inlet temperature of the first fixed-bed reactor is 180 to 220°C, preferably 190 to 210°C. The inlet temperature of the second fixed-bed reactor is 280 to 320°C, preferably 290 to 310°C.

[0013] In the process of this invention, after the sulfidation in step (2) is completed, the concentration of hydrogen sulfide in the hydrogen gas is generally replaced by hydrogen circulation, which is ≤300mg / m³. 3 Preferred concentration ≤100mg / m³ 3 .

[0014] In the process of this invention, the CO2 content in the hydrogen atmosphere described in step (1) is generally ≤20μg / g, preferably ≤10μg / g.

[0015] In the process of this invention, the catalytic gasoline feedstock in step (2) includes fresh catalytic gasoline feedstock and recycled oil, wherein the properties of the fresh catalytic gasoline feedstock are: sulfur content 100~2000μg / g, olefin content 20v%~60v%.

[0016] In the process of this invention, the liquefied petroleum gas mentioned in step (2) comes from the catalytic cracking unit and delayed coking unit of the refinery, with a sulfur content of 10~300μg / g and an olefin content of 20v%~80v.

[0017] In the process of this invention, the first dissolved gas device in step (2) is any mixing device that can premix the catalytic gasoline feedstock, liquefied petroleum gas and hydrogen. The mixing device can be one or more of the following: static mixer, dissolved gas pump, mechanical stirring device, colloid mill, microporous plate nano / micro hydrogen dispersion component, microbubble generator, ceramic membrane nano / micro hydrogen dispersion component, jet mixer, and microchannel mixer.

[0018] In the process of this invention, the thiol transfer conditions in step (2) are as follows: reaction pressure of 1.0~4.0 MPa, preferably 1.6~3.2 MPa; reaction temperature of 90~200℃, preferably 110~180℃; and volume hourly space velocity of the catalytic gasoline feedstock and liquefied petroleum gas feedstock of 1.0~10.0 h⁻¹. -1 Preferably 2.0~8.0h -1 The volume ratio of hydrogen to catalytic gasoline feedstock and liquefied petroleum gas feedstock is 100:1 to 1:1, preferably 50:1 to 1:1, and the volume ratio of catalytic gasoline feedstock to liquefied petroleum gas is 0.5:1 to 20:1, preferably 2:1 to 10:1.

[0019] In the process of this invention, the thiol transfer catalyst mentioned in step (2) can be a well-known catalyst with thiol transfer function. It can be a commercially available product or prepared according to existing technology, such as the ExxonMobil HR-845 catalyst. The thiol transfer catalyst generally includes a hydrogenation active component and a catalyst support. The hydrogenation active component includes molybdenum and nickel, with a weight content of 3% to 40% based on oxides. The catalyst support is alumina, amorphous aluminum silicate, silicon dioxide, titanium dioxide, etc., and may also contain other additives such as P, Si, B, Ti, Zr, Ca, etc.

[0020] In the process of this invention, after the gaseous product obtained in step (3) is desulfurized by a gas separator, hydrogen is separated to obtain liquefied petroleum gas product.

[0021] In the process of this invention, the fractionation in step (3) is generally carried out in a fractionation device, such as a fractionation tower. If there is liquid product in the gas separator, it will directly enter the fractionation device and be fractionated together with the separated liquid product into light fraction and heavy fraction.

[0022] In the process of this invention, the liquid phase product and / or part of the light distillate obtained in step (3) are returned to the first dissolved gas equipment as circulating oil; the volume ratio of circulating oil to fresh catalytic gasoline feedstock is 1:10~2:1, preferably 1:5~1:1.

[0023] In the process of this invention, the second dissolved gas device in step (4) is any mixing device capable of premixing the catalytic gasoline feedstock and hydrogen. The mixing device may be one or more of the following: static mixer, dissolved gas pump, mechanical stirring device, colloid mill, microporous plate nano / micro hydrogen dispersion component, microbubble generator, ceramic membrane nano / micro hydrogen dispersion component, jet mixer, and microchannel mixer.

[0024] In the process of this invention, the selective hydrogenation reaction conditions in step (4) are as follows: reaction pressure of 1.0~4.0 MPa, preferably 1.6~3.2 MPa, reaction temperature of 230~320℃, preferably 250~300℃, and volume hourly space velocity of the heavy catalytic gasoline of 1.0~10.0 h⁻¹. -1 Preferably, it is 2.0~6.0h. -1 The mass ratio of hydrogen to heavy fraction is 200:1 to 10:1, preferably 100:1 to 50:1.

[0025] In the process of this invention, the selective hydrogenation catalyst mentioned in step (4) can be a catalyst with selective hydrogenation desulfurization function well known in the art. It can be a commercially available product or prepared according to existing technology, such as Axens RT-225, ExxonMobil HR-806 catalyst, and Sinopec (Dalian) Petrochemical Co., Ltd.'s FGH-21 and FGH-31 catalysts. The selective hydrogenation catalyst generally includes a hydrogenation active component and a catalyst support. The hydrogenation active component is at least one Group VIB metal molybdenum and at least one Group VIII metal cobalt. The weight content of molybdenum and cobalt based on oxides is 4% to 40%. The catalyst support is generally alumina, amorphous aluminum silicate, silicon oxide, titanium oxide, etc., and may also contain other additives such as P, Si, B, Ti, Zr, Ca, etc.

[0026] In the process of this invention, part of the heavy catalytic gasoline product obtained in step (4) is returned to the second dissolved gas equipment as circulating oil, with a circulating mass ratio of 1:10 to 2:1, preferably 1:5 to 1:1.

[0027] In the process of this invention, the unrecycled light fraction in step (3) is mixed with the heavy catalytic gasoline product obtained in step (4) to obtain the finished catalytic gasoline.

[0028] In existing technologies, although hydrodesulfurization technology can remove mercaptans from liquefied petroleum gas (LPG), it easily leads to olefin saturation, thereby reducing the utilization value of LPG products. Catalytic cracking gasoline has high sulfur and olefin content. To achieve gasoline product quality upgrades, traditional desulfurization technologies saturate high-octane olefin components, generating low-octane alkanes, resulting in a significant decrease in octane number.

[0029] This invention provides a combined LPG-catalytic gasoline process that can selectively hydrogenate both LPG and catalytic gasoline simultaneously, enabling the simultaneous production of high-value LPG and gasoline with minimal octane number loss. By comparing the commonalities of LPG and catalytic gasoline hydrogenation, and considering the simultaneous need for desulfurization and reduced olefin saturation in both LPG and catalytic gasoline, the inventors employ a dissolved hydrogen method to maximize the hydrogen supply for hydrodesulfurization without excessive hydrogen supply that could lead to oversaturation of olefins. Two different functional catalysts are used in a thiol transfer reactor to hydrogenate and saturate dienes and isomerize double bonds. Small-molecule thiols and sulfides are converted into large-molecule sulfides. A separation system yields low-sulfur LPG with minimal olefin loss. The separated liquid phase is fractionated into light catalytic gasoline and heavy catalytic gasoline. The heavy catalytic gasoline enters a gasoline selective hydrogenation reactor, where it undergoes selective deep hydrogenation desulfurization using a heavy catalytic gasoline selective hydrogenation catalyst. By minimizing olefin saturation, minimal octane number loss is achieved. The inventors also employed a unique sulfidation start-up method. After sulfidation, a large amount of acidic CO2 molecules are adsorbed onto the active sites of the catalyst's hydrodesulfurization reaction. Catalytic gasoline feedstock is then introduced at a relatively high reaction temperature to accelerate carbon deposition on the catalyst. Utilizing the different active sites on the catalyst for the hydrodesulfurization reaction and the olefin hydrogenation saturation reaction, more carbon deposits cover the active sites of the olefin hydrogenation saturation reaction, thus reducing its activity. When the CO2 content in the hydrogen returns to ≤20 μg / g, the hydrodesulfurization activity of the catalyst is restored due to the recoverable adsorption of CO2 molecules on the catalyst. However, the olefin hydrogenation saturation reaction activity is now suppressed, significantly improving the selectivity of LPG and catalytic gasoline hydrodesulfurization. This integrated processing of the two feedstocks reduces equipment investment costs, improves production efficiency, and effectively solves the problem of high temperature rise during LPG hydrogenation. The dissolved hydrogen method also reduces hydrogen consumption. Attached Figure Description

[0030] Figure 1This is a flowchart of the combined process of liquefied petroleum gas-catalytic gasoline hydrogenation of the present invention.

[0031] Wherein: 1-Catalytic gasoline feedstock, 2-Feedstock pump, 3-Liquefied petroleum gas feedstock, 4-Hydrogen, 5-First gas-dissolving equipment, 6-First fixed-bed reactor, 7-Pressure reducing valve, 8-First reactant stream gas-liquid separator, 9-Partial liquid product of the first reactant stream, 10-One-way valve, 11-Gas product, 12-Gas separator, 13-Separated hydrogen, 14-Liquefied petroleum gas product, 15-Liquid product, 16-Fracturing tower, 17-Light catalytic gasoline product obtained from fractionation, 18-Light catalytic gasoline product obtained from partial distillation, 19-One-way valve, 20-Heavy catalytic gasoline obtained from fractionation, 21-Heavy catalytic gasoline feedstock pump, 22-Second gas-dissolving equipment, 23-Second fixed-bed reactor, 24-Pressure reducing valve, 25-Second reactant stream gas-liquid separator, 26-Heavy catalytic gasoline product, 27-Partial liquid product obtained from the separation of the second reactant stream, 28-One-way valve, 29-Gas product obtained from the separation of the second reactant stream. Detailed Implementation

[0032] The following examples illustrate the process and effects of the liquefied petroleum gas-catalytic gasoline hydrogenation combined process of the present invention, but do not constitute a limitation on the method of the present invention.

[0033] The mercaptan transfer catalyst is a catalyst with alumina as a support and molybdenum and nickel as hydrogenation active components, with a total metal content of 20 wt% based on oxides. The heavy catalytic gasoline selective hydrogenation catalyst is a catalyst with alumina as a support and molybdenum and cobalt as hydrogenation active components, with a total metal content of 16 wt% based on oxides. Table 1 shows the properties of liquefied petroleum gas feedstock, and Table 2 shows the properties of catalytic gasoline feedstock.

[0034] Table 1 Properties of Liquefied Petroleum Gas Feedstock

[0035] composition content Sulfur content, μg / g 38 Ethane, % (V / V) 0.6 Ethylene, % (V / V) 0.1 Propane, % (V / V) 9.76 Propylene, % (v / v) 39.47 Butane, % (V / V) 20.91 Butene, % (V / V) 28.91 Butadiene, % (V / V) 0.07 <![CDATA[C5 + , %(V / V)]]> 0.27

[0036] Table 2 Properties of Catalytic Gasoline Feedstock

[0037] project Catalytic gasoline Octane number, RON 94.1 Sulfur content, μg / g 417.1 Diene, % 1.14 Multidimensional chromatographic composition Olefins, % 36.9 Example 1

[0038] During the sulfidation process, the CO2 content in the hydrogen was controlled at 5 μg / g. After establishing hydrogen circulation, the inlet temperature of the first and second fixed-bed reactors was raised to 230℃ and maintained at that temperature for 8 hours. Then, the reactor inlet temperature was raised to 280℃ and maintained at that temperature for 8 hours. The circulating hydrogen H2S concentration was 11000 mg / m³. 3 After sulfidation, the concentration of hydrogen sulfide in the replaced hydrogen gas is 90 mg / m³. 3The CO2 content is 6900 μg / g, and it is fed into a catalytic gasoline feedstock with a volume hourly space velocity of 6 h⁻¹. -1 The inlet temperature of the first fixed-bed reactor is controlled at 190℃, and the inlet temperature of the second fixed-bed reactor is controlled at 300℃. After 240 hours of introducing catalytic gasoline feedstock, the CO2 content in the hydrogen is restored to a controlled level of 5μg / g, and the inlet temperatures of the first and second fixed-bed reactors drop to the actual reaction temperatures for normal production. Catalytic gasoline feedstock 1 is mixed with liquefied petroleum gas feedstock 3 and hydrogen 4 in the first dissolved gas equipment 5 and then enters the first fixed-bed reactor 6 for reaction. The reaction pressure is 2.3MPa, the reaction temperature is 150℃, and the volume hourly space velocity (VHSV) of the catalytic gasoline feedstock and liquefied petroleum gas feedstock is 5h⁻¹. -1 The volume ratio of hydrogen to catalytic gasoline feedstock and liquefied petroleum gas feedstock is 150:1, and the volume ratio of catalytic gasoline feedstock to liquefied petroleum gas is 4:1. The reaction effluent passes through pressure reducing valve 7 and enters the gas-liquid separator 9 of the first reactor. A portion of the liquid product 15 from the first reactor effluent is returned to the feedstock pump 2 via check valve 10, with a volume ratio of circulating oil to fresh catalytic gasoline feedstock of 1:3. The gaseous product 11 enters the gas separator 12 for separation to obtain separated hydrogen 13 and liquefied petroleum gas product 14. The liquid product 15 is fractionated in the fractionation tower 16 to obtain light catalytic gasoline product 17. A portion of the light catalytic gasoline product 18 obtained from fractionation is returned to the feedstock pump 2 via check valve 19. The heavy catalytic gasoline 20 obtained from fractionation is mixed with hydrogen 4 in the second dissolved gas unit 22 via the heavy catalytic gasoline feedstock pump 21 and then enters the second fixed-bed reactor 23. The reaction pressure is 1.6 MPa, the reaction temperature is 260℃, and the volume hourly space velocity (VHSV) of the heavy catalytic gasoline is 3.0 h⁻¹. -1 The mass ratio of hydrogen to heavy distillate is 70:1. After passing through pressure reducing valve 24, the mixture enters the second reaction stream gas-liquid separator 25 to obtain heavy catalytic gasoline product 26. Part of the liquid product 27 obtained from the second reaction stream separation is returned to the heavy catalytic gasoline feedstock pump 21 via check valve 28. The gaseous product 29 obtained from the second reactor effluent separation enters the gas separator 12. The light catalytic gasoline product 17 and the heavy catalytic gasoline product 26 are mixed to obtain the finished catalytic gasoline. Example 2

[0039] All implementation processes are the same as in Example 1, except that the CO2 content in the hydrogen is controlled at 10 μg / g during sulfidation and 5000 μg / g after sulfidation. The inlet temperature of the first fixed-bed reactor is controlled at 200℃, and the inlet temperature of the second fixed-bed reactor is controlled at 290℃. After 360 hours of introducing catalytic gasoline feedstock, the CO2 content in the hydrogen is restored to the controlled 10 μg / g. The volume hourly space velocity (VHSV) of the catalytic gasoline feedstock and liquefied petroleum gas feedstock in the first fixed-bed reactor is 0.5 h⁻¹. -1The volume ratio of hydrogen to catalytic gasoline feedstock and liquefied petroleum gas feedstock is 12:1, and the volume ratio of catalytic gasoline feedstock to liquefied petroleum gas is 3:1. The mass ratio of recycled oil to fresh catalytic gasoline feedstock is 1:15. The volume hourly space velocity (VHSV) of the mixture in the second fixed-bed reactor is 2.5 h⁻¹. -1 The mass ratio of hydrogen to heavy distillate is 90:1.

[0040] Comparative Example 1

[0041] All implementation processes are the same as in Example 1, except that the inlet temperature of the first fixed-bed reactor is controlled at 160°C and the inlet temperature of the second fixed-bed reactor is controlled at 260°C during the stabilization process; after 36 hours of introducing catalytic gasoline feedstock, the volume hourly space velocity (VHSV) of the catalytic gasoline feedstock and liquefied petroleum gas feedstock in the first fixed-bed reactor is 6 h⁻¹. -1 The volume ratio of hydrogen to catalytic gasoline feedstock and liquefied petroleum gas feedstock is 12:1, and the volume ratio of catalytic gasoline feedstock to liquefied petroleum gas is 25:1. The mass ratio of recycled oil to fresh catalytic gasoline feedstock is 1:4. The volume hourly space velocity (VHSV) of the mixture in the second fixed-bed reactor is 12 h⁻¹. -1 The mass ratio of hydrogen to heavy distillate is 500:1.

[0042] Comparative Example 2

[0043] All implementation processes are the same as in Example 1, except that the CO2 content is 5 μg / g after vulcanization.

[0044] Table 3 Comparison of LPG and Catalytic Gasoline

[0045] project Example 1 Example 2 Comparative Example 1 Comparative Example 2 Properties of liquefied petroleum gas products Sulfur content, μg / g 1.6 1.8 2.8 1.5 Ethane, % (V / V) 0.7 0.6 0.9 0.8 Ethylene, % (V / V) 0 0 0 0 Propane, % (V / V) 10.22 10.18 14.23 12.26 Propylene, % (v / v) 39.07 39.19 35.06 37.07 Butane, % (V / V) 21.29 21.21 19.20 23.39 Butene, % (V / V) 28.59 28.63 26.57 26.50 Butadiene, % (V / V) 0.01 0.01 0.01 0.01 <![CDATA[C5 + , % (V / V)]]> 0.21 0.22 0.28 0.11 Catalytic gasoline product properties Octane number, RON 93.0 92.9 92.0 92.4 Sulfur content, μg / g 7.1 7.2 12.1 6.8 Diene, % 0.17 0.19 0.15 0.13 Multidimensional chromatographic composition Olefins, % (V / V) 34.8 34.7 32.6 33.6

[0046] As can be seen from the comparison results of the examples and comparative examples in Table 3, the hydrogenation combination process of the present invention can be used to simultaneously produce high-value liquefied petroleum gas products and gasoline products with minimal octane number loss for catalytic gasoline.

Claims

1. A combined process for liquefied petroleum gas-catalytic gasoline hydrogenation, comprising the following: (1) a first fixed-bed reactor and a second fixed-bed reactor connected in series, wherein the first fixed-bed reactor is filled with a mercaptan transfer catalyst and the second fixed-bed reactor is filled with a selective hydrodesulfurization catalyst, and the two reactors are simultaneously subjected to sulfidation treatment; after sulfidation, CO2 is introduced into the two reactors to maintain the CO2 concentration in the reaction system at ≥3000μg / g, and catalytic gasoline feedstock is introduced to react for a period of time, the inlet temperature of the first fixed-bed reactor is 180~220℃, the inlet temperature of the second fixed-bed reactor is 280~320℃, and the atmosphere is switched to hydrogen. (2) The catalytic gasoline feedstock, liquefied petroleum gas and hydrogen are mixed in the first dissolved gas device. The mixture enters the first fixed bed reactor and reacts with the mercaptan transfer catalyst to obtain the first reaction stream. (3) The first reaction stream is separated into gas phase product and liquid phase product. The gas phase product is further desulfurized to obtain liquefied petroleum gas product. The liquid phase product is fractionated to obtain light fraction and heavy fraction. (4) The heavy fraction and hydrogen are mixed in the second dissolved gas device. The mixture enters the second fixed bed reactor and reacts with the selective hydrodesulfurization catalyst to obtain the second reaction stream. The second reaction stream is separated into gas and liquid to obtain the heavy catalytic gasoline product.

2. The process according to claim 1, characterized in that: Before the sulfidation treatment described in step (1), the first and second fixed-bed reactors have completed nitrogen gas sealing, catalyst drying, hydrogen replacement, hydrogen gas sealing and established hydrogen circulation.

3. The process according to claim 1, characterized in that: The sulfidation conditions described in step (1) are as follows: the amount of sulfiding agent introduced is 90wt%~150wt% of the theoretical sulfur required by the catalyst; the sulfidation process adopts programmed temperature rise, and the temperature is raised to 200~230℃ and kept at a constant temperature for 4~16h, and then raised to 260~290℃ and kept at a constant temperature for 4~16h; the sulfiding agent is one or more of carbon disulfide, dimethyl disulfide, methyl sulfide, and n-butyl sulfide.

4. The process according to claim 1, characterized in that: The CO2 injection rate in step (1) is ≥6000 μg / g, calculated based on its concentration in hydrogen.

5. The process according to claim 1, characterized in that: The feed rate of the catalytic gasoline feedstock in step (1) is a volume hourly space velocity (VHSV) of 2 to 10 h⁻¹. -1 The reaction time is 48~600 hours, the inlet temperature of the first fixed-bed reactor is 190~210℃, and the inlet temperature of the second fixed-bed reactor is 290~310℃.

6. The process according to claim 1, characterized in that: After step (2), the hydrogen sulfide concentration in the hydrogen gas is replaced by hydrogen circulation, with a hydrogen concentration ≤ 300 mg / m³. 3 .

7. The process according to claim 1, characterized in that: In the hydrogen atmosphere described in step (1), the CO2 content is ≤20μg / g.

8. The process according to claim 1, characterized in that: The catalytic gasoline feedstock mentioned in step (2) includes fresh catalytic gasoline feedstock and recycled oil, wherein the properties of the fresh catalytic gasoline feedstock are: sulfur content 100~2000μg / g, olefin content 20v%~60v%.

9. The process according to claim 1, characterized in that: The liquefied petroleum gas mentioned in step (2) has a sulfur content of 10~300μg / g and an olefin content of 20v%~80v.

10. The process according to claim 1, characterized in that: The first dissolved gas device mentioned in step (2) is one or more of the following: static mixer, dissolved gas pump, mechanical stirring device, microporous plate nano / micro hydrogen dispersion component, microbubble generator, ceramic membrane nano / micro hydrogen dispersion component, jet mixer and microchannel mixer.

11. The process according to claim 1, characterized in that: The thiol transfer conditions in step (2) are: reaction pressure of 1.0~4.0 MPa, reaction temperature of 90~200℃, and volume hourly space velocity of the mixture of 1.0~10.0 h⁻¹. -1 The volume ratio of hydrogen to catalytic gasoline feedstock and liquefied petroleum gas feedstock is 100:1 to 1:1, and the volume ratio of catalytic gasoline feedstock to liquefied petroleum gas is 0.5:1 to 20:

1.

12. The process according to claim 1, characterized in that: The thiol transfer catalyst described in step (2) includes a hydrogenation active component and a catalyst support. The hydrogenation active component includes molybdenum and nickel, with a weight content of 3% to 40% based on oxides. The catalyst support is one or more of alumina, amorphous aluminosilicate, silicon oxide, and titanium oxide.

13. The process according to claim 12, characterized in that: The thiol transfer catalyst described in step (2) contains one or more of the following promoters: P, Si, B, Ti, Zr, and Ca.

14. The process according to claim 1, characterized in that: After hydrogen sulfide is removed from the gaseous product obtained in step (3) by a gas separator, hydrogen is separated to obtain liquefied petroleum gas product.

15. The process according to claim 14, characterized in that: The fractionation in step (3) is carried out in a fractionation device. If there is liquid product in the gas separator, it is directly fed into the fractionation device and fractionated together with the liquid product obtained from the gas-liquid separation into light and heavy fractions.

16. The process according to claim 8, characterized in that: The liquid phase product and / or part of the light fraction obtained in step (3) are returned to the first dissolved gas unit as circulating oil; the volume ratio of circulating oil to fresh catalytic gasoline feedstock is 1:10~2:

1.

17. The process according to claim 1, characterized in that: The second dissolved gas device in step (4) adopts one or more of the following: static mixer, dissolved gas pump, mechanical stirring equipment, microporous plate nano / micro hydrogen dispersion component, microbubble generator, ceramic membrane nano / micro hydrogen dispersion component, jet mixer and microchannel mixer.

18. The process according to claim 1, characterized in that: The selective hydrogenation reaction conditions in step (4) are: reaction pressure of 1.0~4.0MPa, reaction temperature of 230~320℃, and mass ratio of hydrogen to heavy distillate of 200:1~10:

1.

19. The process according to claim 1, characterized in that: The selective hydrodesulfurization catalyst described in step (4) includes a hydrogenation active component and a catalyst support, wherein the hydrogenation active component is molybdenum and cobalt, and the weight content of molybdenum and cobalt based on oxides is 4% to 40%, and the catalyst support is one or more of alumina, amorphous aluminosilicate, silicon oxide, and titanium oxide.

20. The process according to claim 19, characterized in that: The selective hydrodesulfurization catalyst described in step (4) contains one or more of the following additives: P, Si, B, Ti, Zr, and Ca.

21. The process according to claim 1, characterized in that: The heavy catalytic gasoline product obtained in step (4) is partially returned to the second dissolved gas equipment as circulating oil, with a circulating mass ratio of 1:10 to 2:

1.

22. The process according to claim 16, characterized in that: The unrecycled light fraction in step (3) is mixed with the heavy catalytic gasoline product obtained in step (4) to obtain the finished catalytic gasoline.

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