Catalytic cracking light naphtha processing method

Abstract

The invention discloses a catalytic cracking light naphtha processing method. The method comprises the following steps: (I) mixing catalytic cracking light naphtha with hydrogen, and then feeding the mixture into a hydrofining reaction zone for hydrofining reaction to obtain a refined naphtha material flow; (II) the refined naphtha material flow obtained in the step (I) enters a hydro-conversion reaction zone to be subjected to a hydro-conversion reaction, converted catalytic cracking light naphtha is obtained, and the temperature of the hydro-conversion reaction is 100-300 DEG C higher than that of the hydrofining reaction. According to the present invention, the olefin in the catalytic cracking light naphtha product is completely saturated, the normal hydrocarbon content is increased, and the triene yield can be substantially improved when the catalytic cracking light naphtha product is adopted as the ethylene raw material through steam cracking.

Description

Technical Field

[0001] This invention relates to the field of petrochemical technology, specifically to a method for processing catalytic cracking light naphtha (mainly referring to hydrocarbons rich in C5 and C6 hydrocarbons). Background Technology

[0002] Catalytic cracking is the core unit of oil refining. Catalytic cracking products can be classified according to distillation range into dry gas, liquefied petroleum gas, light naphtha, gasoline, diesel, light cycle oil, and catalytic slurry. Among them, light naphtha accounts for 10-20 wt%. Catalytic cracking light naphtha is usually used directly as a gasoline blending component, but due to the influence of saturated vapor pressure, its blending amount is relatively low. At the same time, the demand for ethylene continues to grow, and using catalytic cracking light naphtha as ethylene feedstock meets the "oil conversion" requirement. However, there are two problems when using catalytic cracking light naphtha as ethylene feedstock: (1) Catalytic cracking light naphtha has a high olefin content (greater than 30 wt%), which is prone to coking. After entering the steam cracking ethylene unit, it will shorten the coking cycle and affect the economic benefits of the ethylene unit. (2) Catalytic cracking light naphtha has a high isohydrocarbon content (greater than 50 wt%), which results in a low triene yield in the ethylene unit and a decrease in its economic benefits.

[0003] CN201510379026.6 discloses a method for optimizing the utilization of light naphtha. The method involves first passing the light naphtha through an adsorption separation unit to obtain an adsorbed oil rich in isoalkanes and a desorbed oil rich in n-alkanes. The desorbed oil can be directly used as feedstock for ethylene cracking. The adsorbed oil is then subjected to a distillation unit to obtain gasoline blending material rich in isocyanate C6 and high-purity isopentane. This process, employing adsorption separation, has high energy consumption. Summary of the Invention

[0004] To address the shortcomings of existing technologies, this invention provides a method for processing catalytic cracking light naphtha. The product obtained by processing catalytic cracking light naphtha using this method has fully saturated olefins and increased n-hydrocarbon content, which can significantly improve the yield of trienes (ethylene, propylene, and butadiene) when used as feedstock for ethylene production by steam cracking.

[0005] The first aspect of this invention provides a method for processing catalytic cracking light naphtha, comprising:

[0006] (I) Catalytic cracking light naphtha is mixed with hydrogen and then enters the hydrorefining reaction zone for hydrorefining reaction to obtain refined naphtha stream.

[0007] (II) The refined naphtha stream from step (I) enters the hydroconversion reaction zone for hydroconversion reaction to obtain converted catalytic cracked light naphtha. The temperature of the hydroconversion reaction is 100-300°C higher than that of the hydrorefining reaction.

[0008] Furthermore, the olefin content in the refined naphtha is 0.05 wt% to 1.0 wt%, preferably 0.08 wt% to 1.0 wt%, for example, but not limited to, 0.09 wt%, 0.1 wt%, 0.15 wt%, 0.2 wt%, 0.25 wt%, 0.3 wt%, 0.4 wt%, 0.45 wt%, 0.5 wt%, 0.55 wt%, 0.6 wt%, 0.65 wt%, 0.7 wt%, 0.75 wt%, 0.8 wt%, 0.85 wt%, 0.9 wt%, 0.95 wt%, 1.0 wt%, and any range between any two values.

[0009] Further, the catalytic cracking light naphtha has the following properties: an initial boiling point of 15–30°C and a final boiling point of 60–90°C. The catalytic cracking light naphtha contains 30 wt%–60 wt% olefins and 30 wt%–66 wt% alkanes, based on a total weight of 100 wt% of the catalytic cracking light naphtha. The catalytic cracking light naphtha contains 50 wt%–90 wt% C5-C6 isomers, preferably 60 wt%–80 wt%, based on a total weight of 100 wt% of the catalytic cracking light naphtha; wherein, the C5-C6 isomer content refers to the sum of the proportions of C5-C6 isomers in the alkanes and the proportions of C5-C6 isomers in the olefins of the light naphtha.

[0010] Further, in step (I), the hydrorefining reaction zone is filled with a hydrorefining catalyst, which can be a conventional hydrorefining agent in the art, generally comprising a hydrorefining active metal component and a support. The hydrorefining active metal is a Group VIB metal or a Group VIII metal; the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel. The support is an alumina support. Generally, based on the weight of the catalyst, the content of the Group VIB metal (calculated as oxide) is 20%–30%, and the content of the Group VIII metal (calculated as oxide) is 2%–10%. The hydrorefining catalyst can be an FF series catalyst developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd., such as FF-66.

[0011] Further, in step (I), the hydrorefining conditions are: reaction temperature 150–250°C, reaction pressure 2–8 MPa, and liquid hourly space velocity 3.0–15.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1.

[0012] Further, in step (II), the hydrogenation conversion conditions are: reaction temperature 300–450°C, reaction pressure 2–8 MPa, and liquid hourly space velocity (LHSV) 0.5–5.0 h⁻¹. -1 For example, but not limited to: 1.0h-1 2.0h -1 3.0h -1 4.0h -1 4.5h -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1.

[0013] Further, in step (II), the hydroconversion zone is filled with a hydroconversion catalyst. The hydroconversion catalyst comprises ZSM-5 molecular sieve and macroporous alumina supported on Group VIB and Group VIII metals. Further, the amount of Brønsted acid in the hydroconversion catalyst is I... A Total Group VIII metals in the catalyst phase I mt molar ratio I A / I mt The value is 5.0 to 0.5, preferably 1.5 to 3.0.

[0014] Furthermore, the amount of Brønsted acid (B acid) in the hydrogenation conversion catalyst is I. A The molar ratio of the total amount of Group VIB and Group VIII metals in the catalyst bulk phase I A / I mt In this context, the unit for the amount of Brønsted acid is mol / g, and the total amount of Group VIB and Group VIII metals in the catalyst bulk phase refers to the total molar amount of Group VIB and Group VIII metals per gram of catalyst, expressed in atomic terms. Furthermore, the Group VIB and Group VIII metals exist in the catalyst in the form of oxides.

[0015] Further, in the hydroconversion catalyst, based on the weight of the catalyst, the content of ZSM-5 molecular sieve is 30wt% to 80wt%, the content of macroporous alumina is 15wt% to 65wt%, and the total content of Group VIB and Group VIII metals (calculated as oxides) is 0.5wt% to 8.0wt%; preferably, based on the weight of the catalyst, the content of ZSM-5 molecular sieve is 50wt% to 75wt%, the content of macroporous alumina is 20wt% to 30wt%, and the total content of Group VIB and Group VIII metals (calculated as oxides) is 1.0wt% to 5.0wt%. Further, in the macroporous alumina loaded with Group VIB and Group VIII metals, the mass ratio of Group VIB metals (calculated as oxides) to Group VIII metals (calculated as oxides) is 2.0 to 6.0.

[0016] Furthermore, in the macroporous alumina loaded with Group VIB and Group VIII metals, the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel.

[0017] Furthermore, the total amount of Group VIB and Group VIII metals on the surface of the hydroconversion catalyst is I ms The total amount of Group VIB and Group VIII metals in the bulk catalyst phase I mt The molar ratio is 0.60–0.90. The total amount of Group VIB and Group VIII metals on the catalyst surface refers to the total molar amount of Group VIB and Group VIII metals per gram of catalyst surface, expressed in atomic quantities. Further, in the hydroconversion catalyst, the pore volume of the macroporous alumina is 0.8–1.5 cm³. 3 / g, specific surface area of ​​200-400m² 2 / g. The pore volume of the ZSM-5 molecular sieve is 0.18–0.26 cm³. 3 / g, specific surface area of ​​300-450m² 2 / g, the silicon-aluminum molar ratio SiO2:Al2O3 is 20-100.

[0018] Furthermore, the preparation method of the hydrogenation conversion catalyst includes the following steps:

[0019] (1) Macroporous alumina was impregnated with an impregnation solution containing Group VIB and Group VIII metals and then calcined to obtain macroporous alumina loaded with Group VIB and Group VIII metals.

[0020] (2) The macroporous alumina loaded with Group VIB and Group VIII metals obtained in step (1) is mixed with ZSM-5 molecular sieve, shaped, and calcined to obtain a light naphtha conversion catalyst.

[0021] Furthermore, the impregnation in step (1) is performed using a pressure impregnation method.

[0022] Furthermore, the pressure impregnation includes: mixing macroporous alumina with an impregnation solution containing Group VIB and Group VIII metals, and impregnating in a closed atmosphere under nitrogen.

[0023] Furthermore, the impregnation conditions include: a pressure of 0.5–3.0 MPa, a temperature of 60–80 °C, and an impregnation time of 1–4 h.

[0024] Further, the impregnation is preferably carried out under stirring at a stirring rate of 100–300 rpm. Further, after impregnation, stirring and heating are stopped, and the mixture is allowed to stand for 2–4 hours until cooled to room temperature. The resulting material is first subjected to conventional filtration and drying, and then calcined to obtain macroporous alumina loaded with Group VIB and Group VIII metals. The drying temperature is 100°C–150°C, and the drying time is 2–4 hours; the calcination temperature is 400°C–600°C, and the calcination time is 3–5 hours.

[0025] Furthermore, in step (1), the pore volume of the macroporous alumina is 0.8–1.5 cm³. 3 / g, specific surface area of ​​200-400m² 2 / g.

[0026] Further, in step (1), the impregnation solution containing Group VIB metal and Group VIII metal is provided with one or more of tungsten salt or molybdenum salt as the source of Group VIB metal and one or more of nickel salt or cobalt salt as the source of Group VIII metal.

[0027] Furthermore, in the impregnation solution containing Group VIB metals and Group VIII metals, the mass content of Group VIB metals, calculated as oxides, is 3.0 to 30.0 g / 100 mL, and the mass content of Group VIII metals, calculated as oxides, is 1.0 to 10.0 g / 100 mL.

[0028] Further, in step (1), the pH value of the impregnation solution containing Group VIB metals and Group VIII metals is 8.0 to 10.0. Further, there are no particular limitations on the pH adjuster used; conventional pH adjusters in the art can be used.

[0029] Furthermore, in step (2), the mass ratio of macroporous alumina loaded with Group VIB and Group VIII metals to ZSM-5 molecular sieve is 0.2 to 1.0.

[0030] Furthermore, the pore volume of the ZSM-5 molecular sieve is 0.18–0.26 cm³. 3 / g, specific surface area of ​​300-450m² 2 / g, the silicon-aluminum molar ratio SiO2:Al2O3 is 20-100.

[0031] Further, in step (2), the molding can be carried out using conventional methods in the art, such as extrusion molding. During the molding process, conventional molding aids can be added, such as at least one of extrusion aids, pectinic acids, and binders; the extrusion aid can be guar gum powder, the pectinic acid can be at least one of citric acid and nitric acid, preferably at least one of citric acid and nitric acid; the binder can be aluminum sol or silica sol prepared from porous alumina, preferably aluminum sol. Further, the amount of molding aid added accounts for 1.0% to 5.0% of the total mass of the obtained catalyst.

[0032] Further, in step (2), after molding, the catalyst is dried and calcined to obtain the catalyst. The drying temperature is 100℃~150℃, and the drying time is 2~4h; the calcination temperature is 400℃~600℃, and the calcination time is 3~5h.

[0033] Compared with the prior art, the present invention has the following beneficial effects:

[0034] In the catalytic cracking light naphtha processing method of the present invention, the saturation and conversion of catalytic cracking light naphtha are completed by combining hydrorefining and hydroconversion processes, resulting in a hydrocarbon stream product that is free of olefins and has a high content of n-alkanes. When used as ethylene feedstock, the yield of ethylene and trienes is higher.

[0035] In the catalytic cracking light naphtha processing method of this invention, the catalytic cracking light naphtha first undergoes hydrorefining. The olefin content in the refined naphtha stream is controlled so that trace amounts of olefins act as initiators, rapidly reacting with the acidic centers of the catalyst to generate carbocations. These carbocations continue to propagate after the reaction is complete, allowing the reaction to proceed continuously, thereby improving the conversion capacity of the light naphtha conversion stage. Then, the naphtha passes through a hydroconversion reaction zone. The hydroconversion catalyst possesses Group VIII and Group VIB bimetallic centers with superior hydrorefining activity, and the Brønsted acid content in the catalyst is I... A The total amount of Group VIB and Group VIII metals in the bulk catalyst phase I m molar ratio I A / I mt It has a higher efficiency than traditional hydrocracking catalysts, which can improve the conversion activity of light naphtha in catalytic cracking.

[0036] In this invention, the hydrogenation metals (Group VIII and Group VIB bimetals) in the hydroconversion catalyst are mainly supported on macroporous alumina. This results in low hydrogenation activity within the molecular sieve channels, reducing carbocation saturation and promoting the conversion of light naphtha in catalytic cracking. After diffusing out of the molecular sieve channels, the n-olefins or small molecule olefins obtained from the light hydrocarbon conversion rapidly contact the hydrogenation centers on the alumina, undergoing a saturation reaction, preventing carbon deposition, and improving catalyst stability.

[0037] When catalytic cracking light naphtha is processed using the method of this invention, the conversion rate of C5-C6 isomeric hydrocarbons can reach over 45.0%, and the proportion of normal hydrocarbons in the obtained product can reach over 40%. Detailed Implementation

[0038] The following examples and comparative examples further illustrate the role and effect of the technical solution of the present invention, but the following examples do not constitute a limitation on the scope of protection of the present invention.

[0039] In this invention, the total amount of Group VIB and Group VIII metals on the catalyst surface is I msThe elemental composition and state of the catalyst surface (3 nm–10 nm) were determined by X-ray photoelectron spectroscopy (XPS) using a Thermofisher Multilab 2000 electron spectrometer. The excitation source was Mg Kα, and the cathode voltage and current were 13 kV and 20 mA, respectively. The electron binding energy was calibrated using Cls (284.6 eV). The data obtained directly from XPS are mass percentages. ms The values ​​are converted to mol / g.

[0040] In this invention, the total amount of Group VIB and Group VIII metals in the bulk phase of the catalyst is I mt The results were obtained by X-ray fluorescence spectroscopy (XRF) analysis using a ZSX100e X-ray fluorescence spectrometer. The spectral line was Kα, the crystal was LiF1, the target material was Rh, the detector was SC scintillation, the timing was 20 s, and the optical atmosphere was vacuum. The data obtained directly from XRF are mass percentages. mt The values ​​are converted to mol / g. The total amount of Group VIB and Group VIII metals in the bulk catalyst refers to the total amount of Group VIB and Group VIII metals in the entire catalyst.

[0041] In this invention, the amount of Brønsted acid in the catalyst is I. A Measurements were taken using a pyridine infrared spectroscopy device. The powdered catalyst was compressed into tablets, evacuated, and degassed at 450°C for 2 hours. After the temperature dropped to room temperature, the kinetic diameter was measured... Pyridine molecules are used as probe molecules to measure the infrared spectrum of chemical desorption and calculate the amount of Brønsted acid (B acid) adsorbed. Because the diameter of pyridine molecules is smaller than that of molecular sieve channels, this method can obtain the amount of B acid in mol / g.

[0042] In this invention, the amount of carbon deposited on the catalyst is measured by an infrared sulfur and carbon analyzer. A high-frequency infrared carbon and sulfur analyzer of EMIA-820V from HORIBA Corporation of Japan is used. The high-frequency induction heating furnace has an output power of 2300W and a frequency of 20MHz. The sample heating temperature is 1800℃. The carbon content of the catalyst is calculated based on the peak area of ​​the vibrational spectrum of CO2 generated by the combustion of carbon deposits.

[0043] In this invention, unless otherwise specified, all percentages (%) in the embodiments and comparative examples refer to mass fractions.

[0044] In this invention, the hydroconversion catalysts in each embodiment are represented by CAT-B followed by a number, such as CAT-B1, CAT-B2, CAT-B3, and CAT-B4, and their properties are shown in Table 1. The hydroconversion catalysts in each comparative example are represented by DCAT-B followed by a number, such as DCAT-B1, DCAT-B2, DCAT-B3, and DCAT-B4, and their properties are shown in Table 1. The ZSM-5 molecular sieve used in the hydroconversion catalyst of this invention is a commercially available product with a pyridine infrared acidity of 0.52 mmol / g and a pore volume of 0.18 cm³. 3 / g, specific surface area is 191cm³ 2 / g, the silicon-to-aluminum molar ratio SiO2:Al2O3 is 55; the macroporous alumina used is a commercially available product with a pore volume of 1.1 cm³. 3 / g, specific surface area is 350cm² 2 / g.

[0045] The hydrorefining catalyst involved in the embodiments and comparative examples of this invention is FF-66, a commercial catalyst developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd.

[0046] In this invention, the calculation methods for the conversion rate of catalytic cracking light naphtha and the proportion of n-alkane in the products are as follows:

[0047] Conversion rate = (mass of isomeric C5 and C6 hydrocarbons in feedstock - mass of isomeric C5 and C6 hydrocarbons in product) / (mass of isomeric C5 and C6 hydrocarbons in feedstock) × 100%;

[0048] The proportion of n-alkane = mass of n-alkane in the product / total mass of the product × 100%, where n-alkane includes ethane, propane, n-butane, n-pentane and n-hexane.

[0049] The catalytic cracking light naphtha conversion evaluation test of this invention was carried out in a fixed-bed device, wherein the first reactor was filled with hydrorefining catalyst FF-66 and served as the hydrorefining reaction zone for the hydrorefining reaction (first reaction); the second reactor was filled with hydroconversion catalyst and served as the hydroconversion reaction zone for the hydroconversion reaction (second reaction).

[0050] In this invention, in each example, the nickel source used is nickel nitrate, and the tungsten source is ammonium metatungstate.

[0051] Example 1

[0052] 100g of macroporous alumina was impregnated for 2h in 200mL of an aqueous solution with a nickel content of 2.0g / 100mL (nickel as oxide, nickel source: nickel nitrate) and a tungsten content of 10.0g / 100mL (tungsten as oxide, ammonium metatungstate as tungsten source) and a pH of 9.0. The impregnation pressure was 1.0MPa and the impregnation temperature was 60℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-1. 100g of ZSM-5 molecular sieve and 30g of A-NW-1 were mixed, and 85g of pectinic acid (concentration: 1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3h, and calcined at 500℃ for 4h to obtain a hydroconversion catalyst, denoted as CAT-B1, the properties of which are shown in Table 1.

[0053] Comparative Example 1

[0054] 100g of ZSM-5 molecular sieve and 45g of macroporous alumina were mixed, and 85g of pectinic acid (1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then crushed, extruded into strips, dried at 120℃ for 3 hours, and calcined at 500℃ for 4 hours to obtain DA-Z-1. 100g of DA-Z-1 was impregnated in 200mL of an aqueous solution containing 1.5g / 100mL of nickel and 8.0g / 100mL of tungsten (based on oxides) for 2 hours. After filtration, drying at 120℃ for 3 hours, and calcining at 500℃ for 4 hours, a comparative hydroconversion catalyst, denoted as DCAT-B1, was obtained. Its properties are shown in Table 1.

[0055] Comparative Example 2

[0056] 100g of macroporous alumina was impregnated for 2h in 200mL of an aqueous solution containing 8.0g / 100mL nickel and 40.0g / 100mL tungsten (based on oxides), with a pH of 9.0. The impregnation pressure was 1.0MPa and the impregnation temperature was 60℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as DA-NW-2. 100g of ZSM-5 molecular sieve and 30g of DA-NW-2 were mixed, and 85g of pectinic acid (1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3h, and calcined at 500℃ for 4h to obtain a comparative hydroconversion catalyst, denoted as DCAT-B2. Its properties are shown in Table 1.

[0057] Comparative Example 3

[0058] 100g of macroporous alumina was impregnated in 200mL of an aqueous solution (pH 5.4) containing 2.0g / 100mL of nickel and 10.0g / 100mL of tungsten (based on oxides) for 2 hours. After filtration, drying at 120℃ for 3 hours, and calcination at 450℃ for 3 hours, macroporous alumina loaded with nickel oxide and tungsten oxide was obtained, denoted as DA-NW-3. 100g of ZSM-5 molecular sieve and 30g of DA-NW-3 were mixed, and 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3 hours, and calcined at 500℃ for 4 hours to obtain a catalyst, denoted as DCAT-B3, the properties of which are shown in Table 1.

[0059] Comparative Example 4

[0060] Take 100g of small-pore alumina (pore volume 0.48cm³). 3 / g, specific surface area is 251cm³ 2 The nickel oxide and tungsten oxide were impregnated for 2 hours in 200 mL of an aqueous solution containing 2.0 g / 100 mL of nickel and 10.0 g / 100 mL of tungsten (calculated as oxides) with a pH of 9.0. The impregnation pressure was 1.0 MPa and the impregnation temperature was 60 °C. After impregnation, the mixture was cooled to room temperature, filtered, dried at 120 °C for 3 hours, and calcined at 450 °C for 3 hours to obtain small-pore alumina loaded with nickel oxide and tungsten oxide, denoted as DA-NW-4. 100 g of ZSM-5 molecular sieve and 30 g of DA-NW-4 were mixed, and 85 g of pectinic acid (concentration of 1 wt%) and 1 g of extrusion aid guar gum powder were added. The mixture was then crushed, extruded, dried at 120 °C for 3 hours, and calcined at 500 °C for 4 hours to obtain a catalyst, denoted as DCAT-B4, the properties of which are shown in Table 1.

[0061] Example 2

[0062] 100g of macroporous alumina was impregnated for 2h in 200mL of an aqueous solution containing 1.0g / 100mL nickel and 5.0g / 100mL tungsten (based on oxides), with a pH of 9.0. The impregnation pressure was 0.5MPa and the impregnation temperature was 70℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-2. 100g of ZSM-5 molecular sieve and 30g of A-NW-2 were mixed, and 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3h, and calcined at 500℃ for 4h to obtain a hydroconversion catalyst, denoted as CAT-B2, whose properties are shown in Table 1.

[0063] Example 3

[0064] 100g of macroporous alumina was impregnated for 2h in 200mL of an aqueous solution containing 2.0g / 100mL nickel and 10.0g / 100mL tungsten (based on oxides), with a pH of 10.0. The impregnation pressure was 2.0MPa and the impregnation temperature was 80℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-3. 80g of ZSM-5 molecular sieve and 50g of A-NW-3 were mixed, and 90g of pectinic acid (1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3h, and calcined at 500℃ for 4h to obtain a hydroconversion catalyst, denoted as CAT-B3, whose properties are shown in Table 1.

[0065] Example 4

[0066] 100g of macroporous alumina was impregnated for 2h in 200mL of an aqueous solution containing 1.0g / 100mL nickel and 5.0g / 100mL tungsten (based on oxides), with a pH of 8.0. The impregnation pressure was 2.0MPa and the impregnation temperature was 80℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-4. 50g of ZSM-5 molecular sieve and 50g of A-NW-4 were mixed, and 75g of pectinic acid (1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded, dried at 120℃ for 3h, and calcined at 500℃ for 4h to obtain a hydroconversion catalyst, denoted as CAT-B4, whose properties are shown in Table 1.

[0067] Table 1. Composition and properties of catalysts in the examples and comparative examples.

[0068]

[0069] Note: The alumina content of DCAT-B4 is calculated based on microporous alumina, while the content of the other catalysts is calculated based on macroporous alumina.

[0070] Example 6

[0071] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst CAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0072] Example 7

[0073] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B2 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst CAT-B2 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0074] Example 8

[0075] 0.7 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst CAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 14.3 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0076] Example 9

[0077] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B3 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst CAT-B3 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0078] Example 10

[0079] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B4 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst CAT-B4 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0080] Example 11

[0081] 3.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst CAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 3.3 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 3.

[0082] Comparative Example 5

[0083] The first reactor of the fixed-bed reactor was not loaded with catalyst. 10.0 mL of catalyst CAT-B1 was placed in the second reactor of the fixed-bed reactor. The reaction was carried out at a pressure of 4.0 MPa, a hydrogen-to-oil volume ratio of 400:1, and a liquid hourly space velocity of 1.0 h⁻¹. -1 The light naphtha conversion reaction was carried out at a reaction temperature of 400℃. The properties of the feedstock are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0084] Comparative Example 6

[0085] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst DCAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst DCAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0086] Comparative Example 7

[0087] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst DCAT-B2 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst DCAT-B2 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0088] Comparative Example 8

[0089] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst DCAT-B3 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst DCAT-B3 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0090] Comparative Example 9

[0091] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst DCAT-B4 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting effluent was then introduced into the second reactor and contacted with catalyst DCAT-B4 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 10.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0092] Comparative Example 10

[0093] 0.4 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst CAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 25.0 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0094] Comparative Example 11

[0095] 6.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 10.0 mL of catalyst CAT-B1 was placed in the second reactor. Catalytic cracking light naphtha was introduced into the first reactor and contacted with catalyst FF-66 to undergo the first reaction. The resulting reactor effluent was then introduced into the second reactor and contacted with catalyst CAT-B1 to undergo the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity (LHSV) 1.7 h⁻¹. -1 The reaction temperature was 200℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and liquid hourly space velocity 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the olefin content, conversion rate of each component, yield of each component, and catalyst carbon deposition in the first reaction effluent are shown in Table 4.

[0096] Table 2 Composition and Oil Properties of Light Naphtha from Catalytic Cracking Feedstock

[0097] composition Content, wt% Alkanes 36.7 Olefins 58.9 Cycloalkanes 2.5 Aromatics 1.9 <![CDATA[C5-C6 Isoparaffin Content]]> 66.7 Initial boiling point, ℃ 22.8 Final boiling point, ℃ 65.7

[0098] Table 3 Evaluation results of catalysts in each example

[0099]

[0100] Table 4 Evaluation results of catalysts for each comparative example

[0101]

[0102] The specific embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combining the various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for processing catalytic cracking light naphtha, comprising: (I) Catalytic cracking light naphtha is mixed with hydrogen and then enters the hydrorefining reaction zone for hydrorefining reaction to obtain refined naphtha stream. (II) The refined naphtha stream from step (I) enters the hydroconversion reaction zone for hydroconversion reaction to obtain converted catalytic cracked light naphtha. The temperature of the hydroconversion reaction is 100-300°C higher than that of the hydrorefining reaction.

2. The processing method according to claim 1, characterized in that: The initial boiling point of the catalytic cracked light naphtha is 15–30°C, and the final boiling point is 60–90°C. Preferably, the olefin content in the catalytic cracking light naphtha is 30wt% to 60wt%, and the alkane content is 30wt% to 66wt%, based on a total weight of 100wt% of the catalytic cracking light naphtha. Preferably, the C5-C6 isohydrocarbon content in the catalytic cracking light naphtha is 50wt% to 90wt%, more preferably 60wt% to 80wt%, based on a total weight of 100wt% of the catalytic cracking light naphtha.

3. The processing method according to claim 1, characterized in that: The hydrogenation purification conditions are as follows: reaction temperature 150–250℃, reaction pressure 2–8 MPa, and liquid hourly space velocity (LISH) 3.0–15.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1; And / or, the hydrogenation conversion conditions are: reaction temperature 300–450°C, reaction pressure 2–8 MPa, and liquid hourly space velocity 0.5–5.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1; And / or, the olefin content in the refined naphtha is 0.05 wt% to 1.0 wt%, preferably 0.08 wt% to 1.0 wt%.

4. The processing method according to claim 1, characterized in that: The hydrorefining reaction zone is filled with a hydrorefining catalyst, and the hydroconversion zone is filled with a hydroconversion catalyst; Preferably, the hydroconversion catalyst comprises ZSM-5 molecular sieve and macroporous alumina supported on Group VIB and Group VIII metals; the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel.

5. The processing method according to claim 4, characterized in that: The amount of Brønsted acid in the hydroconversion catalyst I A Total amount of Group VIII metals in the catalyst phase I mt molar ratio I A / I mt The value is 5.0 to 0.5, preferably 1.5 to 3.

0.

6. The processing method according to claim 4 or 5, characterized in that: In the hydrogenation conversion catalyst, based on the weight of the catalyst, the content of ZSM-5 molecular sieve is 30wt% to 80wt%, the content of macroporous alumina is 15wt% to 65wt%, and the total content of Group VIB and Group VIII metals as oxides is 0.5wt% to 8.0wt%. Preferably, based on the weight of the catalyst, the content of ZSM-5 molecular sieve is 50wt% to 75wt%, the content of macroporous alumina is 20wt% to 30wt%, and the total content of Group VIB and Group VIII metals as oxides is 1.0wt% to 5.0wt%.

7. The processing method according to claim 4 or 5, characterized in that: In macroporous alumina loaded with Group VIB and Group VIII metals, the mass ratio of Group VIB metals as oxides to Group VIII metals as oxides is 2.0 to 6.

0.

8. The processing method according to claim 4, characterized in that: The total amount of Group VIB and Group VIII metals on the surface of the hydroconversion catalyst is I ms The total amount of Group VIB and Group VIII metals in the bulk catalyst phase I mt The molar ratio is 0.60 to 0.

90.

9. The processing method according to claim 4, characterized in that: In the hydroconversion catalyst, the macroporous alumina has a pore volume of 0.8–1.5 cm³. 3 / g, specific surface area of ​​200-400m² 2 / g.

10. The processing method according to claim 4, characterized in that: The ZSM-5 molecular sieve has a pore volume of 0.18–0.26 cm³. 3 / g, specific surface area of ​​300-450m² 2 / g, the silicon-aluminum molar ratio SiO2:Al2O3 is 20-100.

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