Process and system for co-production of catalytic heavy gasoline and catalytic diesel feedstocks and applications

Abstract

The application discloses a method and system for co-production of catalytic heavy gasoline and catalytic diesel oil and application. The method for co-production of catalytic heavy gasoline and catalytic diesel oil comprises the following steps: feeding hydrogenated catalytic diesel oil and diene-saturated catalytic heavy gasoline into a hydrocracking reactor to produce chemical raw materials containing light aromatic hydrocarbons. The method not only solves the comprehensive utilization problem of catalytic heavy gasoline, but also produces liquid phase products with higher light aromatic hydrocarbon yield than using catalytic diesel oil as a production raw material alone, and the total chemical hydrogen consumption obtained by calculation is obviously lower than that required when chemical raw materials are produced from catalytic diesel oil alone, thereby achieving a good technical effect. The technology can be used to produce high-value light aromatic hydrocarbons from catalytic diesel oil and catalytic heavy gasoline, and to produce by-products such as ethylene cracking raw materials and gasoline blending components.

Description

TECHNICAL FIELD

[0001] The present application relates to the technical field of hydrocracking, in particular to an improved method for producing chemical raw materials from catalytic diesel oil, more particularly to a method and system for co-producing chemical raw materials from catalytic heavy gasoline and catalytic diesel oil and applications. BACKGROUND

[0002] Catalytic cracking is one of the most important secondary processing processes in the petroleum refining industry and is also the core process of heavy oil lightening. As an important by-product of the catalytic cracking unit, catalytic diesel oil has the characteristics of high total aromatic content, high sulfur and nitrogen content, and low cetane number, and is not suitable for processing into vehicle diesel. In recent years, with the improvement of the processing capacity of the catalytic cracking unit, the production of catalytic diesel oil has increased year by year, and how to efficiently utilize it has become a bottleneck restricting the processing capacity and efficiency of China's refining enterprises.

[0003] Aromatic hydrocarbons, especially benzene, toluene, and xylene (BTX) and other light aromatic hydrocarbons, are important chemical raw materials, and they are important cornerstones for the huge downstream chemical industry, such as polystyrene, styrene-butadiene rubber, polyester synthetic fiber, polyester, polyurethane, and benzoic acid industries. Currently, the access to BTX is usually through the separation and fractionation unit of the oil refining process device, such as the aromatic recovery unit of catalytic reforming or catalytic cracking. In order to further expand the source of light aromatic hydrocarbon raw materials and maximize the value of catalytic diesel oil, it is an important way for oil refining conversion to utilize the composition characteristics of catalytic diesel oil rich in aromatic hydrocarbons to convert them into light aromatic hydrocarbons and other chemical raw materials with high added value.

[0004] Chinese patent CN103214332A discloses a method for producing light aromatic hydrocarbons and high-quality oil products from catalytic cracking diesel oil. The method extracts catalytic cracking diesel oil with a solvent to obtain extract oil rich in polycyclic aromatic hydrocarbons and raffinate oil rich in alkanes, and performs hydrofining and hydrocracking under hydrogenation reaction conditions to produce light aromatic hydrocarbons and high-octane gasoline fractions. The method has low diesel utilization rate and low value of by-products. Chinese patent CN112574782A solves the problem of low aromatic purity of the hydrocracking product of catalytic diesel oil by performing hydrocracking reaction on the catalytic diesel oil stream after hydrofining, separating impurities, and deep denitrification. However, the process route is long and the process is relatively complicated. Chinese patent CN109777514A discloses a method for producing aromatic hydrocarbons from catalytic diesel oil by coupling hydrocracking and aromatic extraction processes. By setting two different hydrofining reaction zones, the method solves the problem of poor selectivity of monocyclic aromatic hydrocarbons in conventional hydrofining process for high-aromatic diesel oil. However, the method has relatively high chemical hydrogen consumption and requires a large amount of hydrogen resources for enterprises.

[0005] Currently, the industry usually blends the refined catalytic heavy gasoline into the gasoline fuel pool as a blending component, however, the latest environmental regulations further limit the aromatic content in gasoline, making the catalytic heavy gasoline fraction rich in aromatics no longer suitable for blending into motor gasoline. At the same time, due to the easy cracking of heavy gasoline fraction into lighter and lower value liquefied gas and dry gas, separate processing of catalytic heavy gasoline by hydrocracking and other processes is also not what practitioners want to see. Therefore, how to expand the utilization way of catalytic heavy gasoline and improve the supply and demand balance of refining enterprises is also an important problem to be solved. SUMMARY

[0006] To solve the problems in the prior art, the present application provides a method and system for co-producing chemical raw materials from catalytic heavy gasoline and catalytic diesel, and application thereof, aiming to optimize the method for producing light aromatics and other chemical raw materials from catalytic diesel, and to solve the comprehensive utilization of catalytic heavy gasoline. The method for co-producing chemical raw materials from catalytic heavy gasoline and catalytic diesel of the present application comprises the step of feeding the hydrogenated catalytic diesel obtained by hydrogenation refining and the diene-saturated olefin-reduced catalytic heavy gasoline into a hydrocracking reactor in stages to produce chemical raw materials containing light aromatics (including C 6-10 aromatics). The method not only solves the problem of comprehensive utilization of catalytic heavy gasoline, but also produces liquid products with higher light aromatics yield than using catalytic diesel as the only raw material for production, and the total chemical hydrogen consumption calculated is also significantly lower than the chemical hydrogen consumption required when producing chemical raw materials from catalytic diesel alone. The present application is particularly suitable for producing high-value aromatic products such as benzene, toluene, C8 aromatics and a small amount of C9 and C10 aromatics, etc., as well as low-carbon non-aromatic (C 2-5 non-aromatic) products for ethylene cracking and C6 non-aromatic blending components for improving the octane number of gasoline.

[0007] One of the objects of the present application is to provide a method for co-producing chemical raw materials from catalytic heavy gasoline and catalytic diesel, comprising the step of feeding the hydrogenated catalytic diesel obtained by hydrogenation refining and the diene-saturated olefin-reduced catalytic heavy gasoline into a hydrocracking reactor in stages to produce chemical raw materials containing light aromatics.

[0008] In a preferred embodiment of the present application,

[0009] The method comprises:

[0010] 1) catalytic diesel stream is subjected to hydrogenation refining to obtain a hydrogenated catalytic diesel stream;

[0011] 2) the hydrogenated catalytic diesel stream obtained in step 1) is mixed with hydrogen gas and fed into a first bed of a hydrocracking reactor to contact with a first bed catalyst to obtain a first bed effluent;

[0012] 3) catalytic heavy gasoline stream is subjected to diene saturation to obtain an olefin-reduced catalytic heavy gasoline stream;

[0013] 4) the reduced olefins catalytic gas oil stream in step 3) is mixed with the effluent of the first bed of the hydrocracking reactor and contacted with the catalyst of the second bed to obtain a hydrocracking outlet stream;

[0014] 5) the hydrocracking outlet stream in step 4) is separated to obtain a hydrogen-rich gas phase stream and a liquid phase product stream, optionally, the hydrogen-rich gas phase stream is pressurized to obtain a pressurized hydrogen-rich gas phase stream which is mixed with fresh hydrogen to form a mixed hydrogen gas, and then mixed with the hydrocatalytic diesel stream to recycle into the hydrocracking reactor.

[0015] In a preferred embodiment of the present application,

[0016] In step 1),

[0017] The T5 of the catalytic diesel stream is in the range of 150-240°C, preferably in the range of 190-215°C, and the T95 is in the range of 320-360°C, preferably in the range of 320-345°C.

[0018] In a preferred embodiment of the present application,

[0019] The hydrocatalytic diesel obtained in step 1) satisfies the total nitrogen content of less than 10 ppm, the total aromatic hydrocarbon content of greater than 70 wt%, preferably greater than 80 wt%, and more preferably greater than 85 wt%.

[0020] The catalytic diesel stream in step 1) as a raw oil is contacted with a hydrofining catalyst under hydrogen conditions to perform a hydrofining reaction to remove most of the sulfur, nitrogen, oxygen and metal impurities, and the double bond and triple bond unsaturated hydrocarbons in the catalytic diesel are also saturated, and the aromatic hydrocarbons undergo selective saturation reaction with one aromatic ring reserved. After hydrofining, a hydrocatalytic diesel stream is obtained, mainly including hydrocatalytic diesel, gas containing hydrogen sulfide and ammonia. The hydrofining can be performed in any manner and any method known in the art without particular limitation, as long as the hydrocatalytic diesel satisfies the total nitrogen content of less than 10 ppm, the total aromatic hydrocarbon content of greater than 70 wt%, preferably greater than 80 wt%, and more preferably greater than 85 wt%.

[0021] The hydrofining technology in step 1) is the catalytic diesel hydrofining technology known in the prior art. The hydrofining reaction conditions and the hydrofining catalyst can use the reaction conditions realized in the prior art and any type of hydrofining catalyst already available, as long as the purpose of catalytic diesel hydrofining in step 1) can be achieved.

[0022] In the present application, the reaction conditions of the hydrofining in step 1) are preferably:

[0023] a fixed bed reactor, reactor inlet temperature 280-420℃, preferably 300-410℃, more preferably 310-390℃;

[0024] hydrogen to oil volume ratio 500-3000, preferably 800-2000, more preferably 1000-1500;

[0025] hydrogen partial pressure 5-10MPa, preferably 5-8MPa, more preferably 6-7MPa;

[0026] mass space velocity 0.5-2.0h -1 , preferably 0.6-1.5h -1 , more preferably 0.8-1.2h -1 .

[0027] In the present application, the hydrofining catalyst used in step 1) preferably comprises a hydrofining catalyst carrier and a hydrogenation metal active phase; the mass of the hydrofining catalyst carrier accounts for 60-99.9wt%, preferably 65-99.9wt%, more preferably 70-99.9wt%, further preferably 75-99.9wt% of the total mass of the hydrofining catalyst; the mass of the hydrogenation metal active phase accounts for 0.1-40wt%, preferably 0.1-35wt%, more preferably 0.1-30wt%, further preferably 0.1-25wt% of the total mass of the hydrofining catalyst; the hydrofining catalyst carrier comprises: 60-100wt% of alumina and 0-40wt% of silicon oxide based on the total mass of the carrier.

[0028] The hydrogenation metal active phase in the hydrofining catalyst is preferably at least one of nickel, cobalt, molybdenum, tungsten, and iron. The hydrogenation metal active phase is preferably pre-sulfided after being loaded. The pre-sulfidation of the hydrogenation metal can use the conventional catalyst pre-sulfidation method in the prior art.

[0029] The hydrofining catalyst of the present application is preferably a nickel-molybdenum sulfur / alumina bimetallic catalyst or a nickel-molybdenum-tungsten sulfur / alumina trimetallic catalyst, which has good denitrification effect.

[0030] In one preferred embodiment of the present application,

[0031] In step 2),

[0032] The first bed catalyst comprises a molecular sieve acidic carrier A, a metal active component A, and a binder A;

[0033] The molecular sieve acidic carrier A comprises at least one twelve-membered ring molecular sieve, and the mass of the molecular sieve acidic carrier A accounts for 5-80wt%, preferably 20-75wt%, more preferably 30-70wt% of the mass of the first bed catalyst; and / or,

[0034] The metal active component A comprises 0.01-10 wt%, preferably 0.5-8 wt% of Group VIII metal and 3-25 wt%, preferably 5-15 wt% of Group VIB metal oxide and / or sulfide (by mass of oxide) based on the mass of the first bed catalyst; the balance being the binder A. All the component contents above amount to 100 wt%.

[0035] The hydrocatalytically cracked diesel stream in step 2) is contacted with the hydrocracking first bed catalyst under hydrogen atmosphere, mainly C 11+ ring opening, dealkylation, transalkylation of naphthenic monocyclic aromatics, and C 11+ cracking of non-aromatics, etc. Thus, the hydrocracking first bed effluent has a high content of benzene, polyalkylbenzene and low carbon non-aromatics. The hydrocracking reactor is preferably a fixed bed reactor.

[0036] The binder A can be incorporated into the first bed catalyst in any suitable manner, preferably the hydrocracking catalyst carrier A is prepared after kneading, extruding, aging, drying and calcining the molecular sieve acidic carrier A. The metal active component A can be loaded onto the hydrocracking catalyst carrier A in any suitable manner, such as co-precipitation, co-gelation, ion exchange or impregnation, preferably isometric impregnation using water-soluble compounds of the metal active component A. The specific first bed catalyst preparation method in the present application preferably comprises the following steps:

[0037] After mixing the molecular sieve acidic carrier A with the binder A, kneading, extruding, drying at 60-150°C and calcining in air atmosphere at 500-600°C for 3-6h, the desired hydrocracking catalyst carrier A is obtained. The bimetallic water solution is prepared using Group VIII metal compound and Group VIB metal compound, and the hydrocracking catalyst carrier A is impregnated by isometric impregnation method, dried at 60-150°C and calcined in air atmosphere at 450-520°C for 1-4h to obtain the hydrocracking catalyst precursor A. The hydrocracking catalyst precursor A is reduced under hydrogen condition to 300-500°C and maintained for 2-24h (pre-reduction) to obtain the desired first bed catalyst.

[0038] In a preferred embodiment of the present application,

[0039] In step 3),

[0040] The catalytic heavy gasoline stream comprises C 10 The following hydrocarbon materials, wherein the cyclic hydrocarbons (including naphthenes and aromatics) content is greater than 50 wt%, preferably greater than 60 wt%, more preferably greater than 70 wt%; and / or, the T5 of the catalytic heavy gasoline stream is in the range of 70-150°C and the T95 is in the range of 180-240°C.

[0041] In a preferred embodiment of the present application,

[0042] The olefin-reduced catalytic heavy gasoline stream obtained in step 3) has a diene value of less than 0.2 g I2 / 100 g oil.

[0043] In a preferred embodiment of the present application,

[0044] In step 3),

[0045] The flow rate of the catalytic heavy gasoline stream is 0.1-99 wt%, preferably 10-80 wt%, more preferably 15-60 wt%, and further preferably 35-50 wt% of the total feed stream flow rate. The total feed stream flow rate is the sum of the flow rates of the catalytic diesel stream and the catalytic heavy gasoline stream.

[0046] In step 3) of the method of the present application, the catalytic heavy gasoline stream enters the diene saturation reactor to contact the hydrogenation catalyst and undergoes diene saturation. Reducing the olefin content in the catalytic heavy gasoline stream can prevent pipe coking and fouling and increase the pressure drop. The olefin-reduced catalytic heavy gasoline stream mainly comprises olefin-reduced catalytic heavy gasoline, and gas containing hydrogen sulfide and ammonia. The diene saturation can be performed using diene saturation techniques known in the art. The diene saturation reaction conditions and diene saturation catalysts can be any type of diene saturation catalysts and reaction conditions achieved in the prior art, as long as the obtained olefin-reduced catalytic heavy gasoline stream has a diene value of less than 0.2 g I2 / 100 g oil.

[0047] The diene saturation catalyst preferably comprises a metal hydrogenation active phase and a diene saturation catalyst support. The metal hydrogenation active phase preferably comprises at least one of cobalt, nickel, molybdenum, iron, and palladium, and can further comprise one of group IB and IIB metals, such as copper and zinc. The metal hydrogenation active phase accounts for 0.1-40 wt%, preferably 0.1-35 wt%, and more preferably 0.1-25 wt% of the total mass of the diene saturation catalyst in terms of the mass of the oxide. The diene saturation catalyst support is preferably at least one of alumina, silica, silica-alumina composite, and amorphous silica-alumina, and accounts for 60-99.9 wt%, preferably 65-99.9 wt%, and more preferably 75-99.9 wt% of the total mass of the diene saturation catalyst. The metal hydrogenation active phase is preferably pre-sulfided after being supported, and the pre-sulfidation method can be a conventional catalyst pre-sulfidation method in the art.

[0048] The reaction conditions for diene saturation in step 3) of the method of the present application are preferably as follows:

[0049] The diene saturation reactor has a relatively low reaction temperature, and the inlet temperature is preferably 120-230°C, and more preferably 150-200°C.

[0050] The reaction pressure is 2-4.5 MPa, preferably 2.4-4 MPa;

[0051] The hydrogen to oil volume ratio is 100-500, preferably 150-430;

[0052] The mass space velocity is 1-10 h-1, preferably 1.5-5 h-1. -1 -1 .

[0053] In a preferred embodiment of the present application,

[0054] In step 4),

[0055] The second bed catalyst accounts for 10-80 wt% of the total catalyst loading (first bed catalyst + second bed catalyst) of the hydrocracking reactor, preferably 20-70 wt%, more preferably 40-70 wt%.

[0056] In a preferred embodiment of the present application,

[0057] In step 4),

[0058] The second bed catalyst comprises a molecular sieve acidic carrier B, a metal active component B and a binder B;

[0059] The molecular sieve acidic carrier B comprises at least one twelve-membered ring molecular sieve and at least one ten-membered ring molecular sieve; the total mass of the molecular sieve accounts for 10-60 wt%, preferably 35-55 wt%, of the mass of the second bed catalyst; and / or,

[0060] The metal active component B comprises 0.01-10 wt%, preferably 0.5-8 wt%, of Group VIII metal based on the mass of the second bed catalyst and 3-25 wt%, preferably 5-15 wt%, of Group VIB metal oxide and / or sulfide based on the mass of the first bed catalyst; the balance is the binder B. All the component contents above add up to 100 wt%.

[0061] In a preferred embodiment of the present application,

[0062] The twelve-membered ring molecular sieve in the first bed catalyst and the twelve-membered ring molecular sieve in the second bed catalyst are the same or different, and each independently is preferably at least one of Beta molecular sieve, USY molecular sieve, mordenite molecular sieve, more preferably Beta molecular sieve; and / or,

[0063] The ten-membered ring molecular sieve in the second bed catalyst is at least one of ZSM-5, ZSM-11, ZSM-22, SAPO-11, EU-1 molecular sieve, more preferably ZSM-5 molecular sieve; and / or,​

[0064] The Group VIII metal in the first bed catalyst and the Group VIII metal in the second bed catalyst are the same or different, and each is independently preferably at least one of platinum, palladium, cobalt, nickel, iridium, and more preferably at least one of cobalt and nickel; and / or,

[0065] The Group VIB metal oxide and / or sulfide in the first bed catalyst and the Group VIB metal oxide and / or sulfide in the second bed catalyst are the same or different, and each is independently preferably at least one of an oxide and / or sulfide of molybdenum and / or tungsten; and / or,

[0066] The binder A and the binder B are the same or different, and each is independently preferably at least one of alumina, silica, a silica-alumina composite, and amorphous silica-alumina.

[0067] In a preferred embodiment of the present application,

[0068] The mass of the ten-membered ring molecular sieve in the second bed catalyst of the hydrocracking reactor accounts for 10-90wt% of the total mass of the molecular sieve acidic carrier B, preferably 25-75wt%, and more preferably 40-60wt%.

[0069] The optional types and contents of the components in the second bed catalyst of the hydrocracking reactor in step 4) of the method of the present application are basically the same as those in the first bed catalyst of the hydrocracking reactor, with the only difference being that: 1) the molecular sieve acidic carrier in the second bed catalyst contains more than one type of molecular sieve: at least one type of twelve-membered ring molecular sieve and at least one type of ten-membered ring molecular sieve; and 2) the total mass of the molecular sieve accounts for a lower proportion of the mass of the catalyst, i.e. 10-60wt%, preferably 25-45wt%.

[0070] The preparation method of the second bed catalyst of the hydrocracking reactor in step 4) of the method of the present application is the same as that of the first bed catalyst of the hydrocracking reactor.

[0071] The reduced olefin catalytic heavy gasoline stream in step 4) of the method of the present application enters the second bed of the hydrocracking reactor as both a reaction raw material and a quenching medium, thereby reducing the temperature rise of the bed and mixing with the effluent of the first bed of the hydrocracking reactor to undergo moderate hydrocracking.

[0072] In a preferred embodiment of the present application,

[0073] In step 5), the hydrocracking outlet stream is subjected to high-pressure separation to obtain a hydrogen-rich gas phase stream and a liquid phase product stream; and / or,

[0074] The hydrocarbon components of the hydrogen-rich gas phase stream and the liquid phase product stream obtained in step 5) include light aromatic hydrocarbons (C 6-10aromatics), C 10 non-aromatics and C 11+ component (C 11+ component comprises C 11 component);

[0075] the light aromatic yield is greater than 35 wt%, preferably greater than 40 wt%, more preferably greater than 42 wt%; and / or,

[0076] the C 10 the non-aromatics yield is less than 60 wt%, preferably less than 55 wt%; and / or,

[0077] the C 11+ component yield is less than or equal to 20 wt%, preferably less than or equal to 15 wt%, more preferably less than or equal to 8 wt%.

[0078] The hydrogen-rich gas phase stream in step 5) of the process of the present application mainly comprises hydrogen, low carbon gaseous hydrocarbons, hydrogen sulfide and ammonia, wherein the low carbon gaseous hydrocarbons are combined with the liquid phase product stream when making the product material balance. The hydrogen-rich gas phase stream as mixed hydrogen recycle into the hydrocracking reactor can reduce the actual hydrogen consumption compared to using fresh hydrogen once through.

[0079] In a preferred embodiment of the present application,

[0080] The hydrocracking reactor process conditions include:

[0081] a reactor inlet temperature of 280-450°C, preferably 310-430°C, more preferably 350-410°C; and / or,

[0082] a hydrogen partial pressure of 5-10 MPa, preferably 5-9 MPa, more preferably 6-8 MPa; and / or,

[0083] a mass space velocity (calculated as the flow rate of the hydrocracked diesel divided by the total catalyst loading of the reactor) of 0.5-2.0 h -1 , preferably 0.6-1.5 h -1 , more preferably 0.8-1.2 h -1 ; and / or,

[0084] a hydrogen to hydrocracked diesel volume ratio of fresh hydrogen to hydrocracked diesel of 400-1000, preferably 500-800, more preferably 500-600; and / or,

[0085] a hydrogen to hydrocracked diesel volume ratio of mixed hydrogen to hydrocracked diesel of 1000-5000, preferably 2000-4000, more preferably 2000-3000.

[0086] The second object of the present application is to provide a system for producing chemical raw materials by the method of the first object of the present application, comprising:

[0087] a first device configured to receive the catalytic diesel oil stream and discharge the hydrogenated catalytic diesel oil stream;

[0088] a second device configured to receive the catalytic heavy gasoline stream and discharge the olefin-reduced catalytic heavy gasoline stream;

[0089] a third device configured to receive the hydrogenated catalytic diesel oil stream and hydrogen gas in sections and the olefin-reduced catalytic heavy gasoline stream and discharge the hydrocracking outlet stream;

[0090] a fourth device configured to receive the hydrocracking outlet stream and discharge the hydrogen-rich gas phase stream and the liquid phase product stream;

[0091] optionally, a fifth device configured to receive the hydrogen-rich gas phase stream and discharge the pressurized hydrogen-rich gas phase stream.

[0092] Preferably,

[0093] the reactor in the first device is a hydrofining reactor; and / or,

[0094] the reactor in the second device is a diene saturation reactor; and / or,

[0095] the reactor in the third device is a hydrocracking reactor, which comprises a first bed layer filled with the first bed layer catalyst and a second bed layer filled with the second bed layer catalyst; and / or,

[0096] the fourth device comprises a high-pressure separation tank; and / or,

[0097] the fifth device comprises a hydrogen compressor.

[0098] The third object of the present application is to provide an application of the method of the first object of the present application in producing chemical raw materials including light aromatic hydrocarbons from catalytic diesel oil.

[0099] The present application has the following advantages: in the present application, catalytic heavy gasoline is converted into olefin-reduced catalytic heavy gasoline through diene saturation reaction, thus removing most of the olefins and part of the sulfur and nitrogen impurities therein and reducing the possibility of coking and poisoning of the hydrocracking catalyst. Meanwhile, the olefin-reduced catalytic heavy gasoline not only serves as a raw material for co-refining with catalytic diesel oil, but also as a quenching medium for reducing the temperature rise of the hydrocracking reactor, thus fully utilizing resources.

[0100] In addition, generally, due to the consideration that catalytic gasoline is prone to excessive cracking to form low-carbon hydrocarbons such as dry gas and liquefied gas with low added value, the refiners generally avoid subjecting the catalytic gasoline fraction to hydrocracking reaction. However, in the method for co-producing catalytic heavy gasoline and catalytic diesel as chemical raw materials according to the present application, the catalytic heavy gasoline selected has not less than 50% of high-carbon cyclic hydrocarbons such as naphthenes and aromatics, and is mixed with the effluent of the first bed layer having high content of aromatic hydrocarbons such as benzene and alkylbenzene before entering the second bed layer of the hydrocracking reactor. The high-carbon cyclic hydrocarbons such as aromatics are less prone to cracking reaction, and therefore the olefin-reduced catalytic heavy gasoline applied to the second bed layer of the hydrocracking reactor will not be excessively cracked.

[0101] Furthermore, the experimental results show that the method for co-producing catalytic heavy gasoline and catalytic diesel as chemical raw materials according to the present application has higher yield of light aromatics, C 11+ lower yield of components, and lower total chemical hydrogen consumption calculated, compared with the method for producing chemical raw materials by using catalytic diesel alone. BRIEF DESCRIPTION OF DRAWINGS

[0102] Figure 1 Figure 1 is a process flow diagram of the method for co-producing catalytic heavy gasoline and catalytic diesel as chemical raw materials according to the present application, which is intended to illustrate the present application rather than limit it;

[0103] In the figure, 1 is a catalytic diesel stream; 2 is a hydrofining reactor; 3 is a hydrocracked catalytic diesel stream; 4 is hydrogen; 5 is a first bed layer of a hydrocracking reactor; 6 is a catalytic heavy gasoline stream; 7 is a diene saturation reactor; 8 is an olefin-reduced catalytic heavy gasoline stream; 9 is a second bed layer of a hydrocracking reactor; 10 is a hydrocracking outlet stream; 11 is a high-pressure separation tank; 12 is a hydrogen-rich gas phase stream; 13 is a liquid phase product stream; 14 is a pressurized hydrogen-rich gas phase stream; and 15 is fresh hydrogen.

[0104] Figure 2 Figure 2 is a process flow diagram of the method according to the comparative example 1;

[0105] Figure 3 Figure 3 is a process flow diagram of the method according to the comparative example 3; in the figure, 3' is a mixture stream of the olefin-reduced catalytic heavy gasoline and the hydrocracked catalytic diesel. DETAILED DESCRIPTION

[0106] The following specific embodiments and drawings are used to specifically describe the present application, and it is necessary to point out that the following embodiments are only used to further illustrate the present application, and cannot be understood as limiting the protection scope of the present application, and some non-essential improvements and adjustments of the present application made by the person skilled in the art according to the content of the present application still belong to the protection scope of the present application.

[0107] The composition analysis of the catalysts involved in the present application all uses the existing analysis method in the art. For example: the compositions of various catalysts are analyzed by X-ray fluorescence spectrometry (XRF) and inductively coupled plasma emission spectrometry (ICP), and the valence state and composition of the metals in the catalysts are determined by X-ray photoelectron spectroscopy (XPS). The XRF, ICP and XPS characterization respectively selects a Rigaku ZSX100e type X-ray fluorescence spectrometer, a Varian 700-ES series inductively coupled plasma emission spectrometer and a PerkinElmer PHI5000CESCA type X-ray photoelectron spectrometer.

[0108] In the present application, the hydrocracking product hydrocarbon composition is determined by online gas chromatography. The chromatograph model is Agilent 7890A, equipped with an FID detector, the hydrocarbon composition in the hydrogen-rich gas phase stream is analyzed by using an HP-PLOT Q capillary column, and the hydrocarbon composition in the liquid phase product stream is analyzed by using an HP-INNOWax capillary column. The group composition of catalytic diesel, hydrocracking catalytic diesel, catalytic heavy gasoline and olefin-reducing catalytic heavy gasoline is analyzed (multi-dimensional chromatographic analysis) by using a full two-dimensional gas chromatograph / high-throughput time-of-flight mass spectrometer (GCxGC-TOFMS) of LECO Company, USA. The total sulfur and total nitrogen content in the oil product is analyzed by using an Elab6600SN sulfur and nitrogen element analyzer of Elan. The diene value of the olefin-reducing catalytic heavy gasoline stream is determined by using a maleic anhydride method, and the specific operation method is referred to the patent CN103666559A.

[0109] The yield calculation formula of the hydrocarbon component is:

[0110] The yield calculation formula of the hydrocarbon component is:

[0111] = the content of the component in the liquid phase (gas phase) multiplied by the liquid phase yield + the content of the component in the gas phase multiplied by the gas phase yield

[0112]

[0113] (the content of the component in the liquid phase (gas phase) is determined by chromatography; the liquid phase yield is calculated by dividing the mass of the product per hour by the hourly feed amount; the gas phase yield is calculated by dividing the sum of the mass of each tail gas component by the hourly feed amount.)

[0114] The total chemical hydrogen consumption calculation formula of the hydrocracking reaction is: ​

[0115] Total chemical hydrogen consumption of hydrocracking

[0116] =∑ Hydrogen content of each hydrocarbon component in hydrocracking products - ∑ Hydrogen content of hydrocracking feedstock

[0117] Figure 1 This is a schematic diagram of the process flow for the co-production of chemical feedstocks from catalytic heavy gasoline and catalytic diesel of the present invention. Many conventional equipment items, such as pumps, compressors, heat exchangers, extraction devices, and hydrogen pipelines, are omitted from the diagram; however, these items are well-known to those skilled in the art. Figure 1 As shown, the process of the method described in this invention is described in detail below:

[0118] Catalytic diesel stream 1 is hydrorefined in hydrorefining reactor 2 by contacting a hydrorefining catalyst to obtain hydrotreated catalytic diesel stream 3. Hydrotreated catalytic diesel stream 3 is mixed with hydrogen 4 and enters the first bed 5 of hydrocracking reactor, where it is contacted with the first bed catalyst to obtain the first bed effluent. Catalytic heavy gasoline stream 6 is saturated in diene saturation reactor 7 by contacting a diene saturation catalyst to obtain olefin-reducing catalytic heavy gasoline stream 8. olefin-reducing catalytic heavy gasoline stream 8 enters the second bed 9 of hydrocracking reactor, mixes with the first bed effluent, and is contacted with the second bed catalyst to obtain hydrocracking outlet stream 10. Hydrocracking outlet stream 10 is separated in high-pressure separator 11 to obtain hydrogen-rich gaseous stream 12 and liquid product stream 13. Hydrogen-rich gaseous stream 12 is pressurized by hydrogen compressor to obtain pressurized hydrogen-rich gaseous stream 14, which forms mixed hydrogen with fresh hydrogen 15, and then mixes with hydrotreated catalytic diesel stream 3 and enters the hydrocracking reactor for recycling.

[0119] Example 1

[0120] use Figure 1 The process flow and the steps described in the instruction manual are shown. The properties of the catalytic heavy gasoline and catalytic diesel feedstock used are shown in Table 1. The hydrorefining reaction conditions, diene saturation reaction conditions, and hydrocracking reaction conditions used are shown in Table 2. The properties of the resulting de-olefin catalytic heavy gasoline and hydrotreated catalytic diesel are shown in Table 3. After 200 hours of operation, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed, and material balance was performed. The results are shown in Table 11.

[0121] The first bed layer 5 of the hydrocracking reactor is filled with 112.0 g of first bed layer catalyst A1. The catalyst A1 is prepared as follows: 55.6 g of Beta zeolite and 57.6 g of pseudo-boehmite are mixed, extruded into strips, dried at 90°C, and then calcined at 550°C in air for 4 h to obtain a catalyst carrier. A clear solution of 17.34 g of nickel nitrate hexahydrate and 8.4 g of ammonium molybdate tetrahydrate is prepared, and the catalyst carrier is impregnated with the metal solution in an equal volume, dried at 100°C, and then calcined at 520°C for 2 h to obtain a catalyst precursor. The catalyst precursor is reduced in a hydrogen atmosphere at 400°C for 4 h to obtain the first bed layer catalyst A1, which has a composition of 3.5 wt% Ni-5.2 wt% MoO2-1.0 wt% MoO3 / 50.0 wt% Beta zeolite-40.3 wt% Al2O3. The second bed layer 9 is filled with 18.0 g of second bed layer catalyst B1. The catalyst B1 is prepared as follows: 25.0 g of Beta zeolite, 25.0 g of ZSM-5 zeolite, and 64.7 g of pseudo-boehmite are mixed, extruded into strips, dried at 90°C, and then calcined at 550°C in air for 4 h to obtain a catalyst carrier. A clear solution of 17.34 g of nickel nitrate hexahydrate and 8.4 g of ammonium molybdate tetrahydrate is prepared, and the catalyst carrier is impregnated with the metal solution in an equal volume, dried at 100°C, and then calcined at 520°C for 2 h to obtain a catalyst precursor. The catalyst precursor is reduced in a hydrogen atmosphere at 400°C for 4 h to obtain the second bed layer catalyst B1, which has a composition of 3.5 wt% Ni-5.2 wt% MoO2-1.0 wt% MoO3 / 22.5 wt% Beta zeolite-22.5 wt% ZSM-5 zeolite-45.3 wt% Al2O3.

[0122] The hydrofining reactor 2 is filled with 116.7 g of hydrofining catalyst C1. The catalyst C1 is prepared as follows: 107.14 g of pseudo-boehmite is mixed, extruded into strips, dried at 100°C, and then calcined at 550°C for 4 h to obtain a hydrofining catalyst carrier. A clear solution of 9.93 g of nickel nitrate hexahydrate, 11.25 g of ammonium molybdate tetrahydrate, 11.69 g of ammonium metatungstate, and 12 mL of ammonia water is prepared, and the catalyst carrier is impregnated with the metal solution in an equal volume, dried at 100°C, and then calcined at 520°C for 4 h to obtain a catalyst precursor. The catalyst precursor is sulfided in a hydrogen atmosphere by passing in 1.0 wt% carbon disulfide cyclohexane solution, heated to 360°C, and sulfided for 12 h to obtain the hydrofining catalyst C1, which has a composition of 3.1 wt% NiS-10.2 wt% MoS2-11.7 wt% WS2 / 75.0 wt% Al2O3.

[0123] The diene saturated reactor 7 was loaded with 15g of diene saturated catalyst D1. The specific preparation process of catalyst D1 was as follows: 131.43g of boehmite was kneaded, extruded, and dried at 100℃ and calcined at 550℃ for 4h to obtain a diene saturated catalyst support. An impregnation solution was prepared using 4.49g of nickel nitrate hexahydrate, 3.05g of zinc nitrate hexahydrate, and 6.18g of ammonium molybdate tetrahydrate. The catalyst support was then metal-loaded using an equal-volume impregnation method, followed by drying at 100℃ and calcination at 520℃ for 4h to obtain a catalyst precursor. The catalyst precursor was then purged with a 1.0wt% carbon disulfide cyclohexane solution in a hydrogen atmosphere and heated to 360℃ for 9h to obtain diene saturated catalyst D1 with the composition: 1.0wt% ZnS-1.4wt% NiS-5.6wt% MoS2 / 92.0wt% Al2O3.

[0124] Table 1. Properties of feedstocks for catalytic heavy gasoline and catalytic diesel

[0125]

[0126]

[0127] Table 2 Reaction conditions

[0128]

[0129] Note: Calculated based on hydrocatalytic diesel flow rate and total catalyst loading in the reactor.

[0130] Table 3 Properties of Olefin-reducing Catalytic Heavy Gasoline and Hydrogenated Catalytic Diesel

[0131] Hydrogenated catalytic heavy gasoline one Hydrogenated catalytic diesel one Nitrogen (ppm) / 9.98 Aromatics (wt%) 61 76 diene value (g I2 / 100 g) 0.1 /

[0132] Comparative Example 1

[0133] use Figure 2 The process flow shown uses only catalytic diesel as feedstock, and the properties of the catalytic diesel feedstock are the same as in Example 1 (see Table 1). The hydrorefining and hydrocracking reaction conditions used are shown in Table 2. The properties of the obtained hydrotreated catalytic diesel are shown in Table 3. After the unit has been running for 200 hours, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed and material balance was performed. The results are shown in Table 11.

[0134] The first bed 5 of the hydrocracking reactor is filled with 12.0 g of catalyst A1, which is the same as in Example 1; the second bed 9 is filled with 8.0 g of catalyst B1, which is the same as in Example 1.

[0135] The hydrorefining reactor 2 was filled with 16.7g of catalyst C1, which was the same as in Example 1.

[0136] Comparative Example 2

[0137] use Figure 1 The process flow and the method steps described in the instruction manual are shown. The properties of the catalytic diesel feedstock used are as in Example 1 (see Table 1), and the properties of the catalytic heavy gasoline feedstock used are shown in Table 4. The hydrorefining reaction conditions, diene saturation reaction conditions, and hydrocracking reaction conditions used are shown in Table 2. The properties of the obtained hydrotreated catalytic diesel are shown in Table 3, and the properties of the obtained de-olefins catalytic heavy gasoline are shown in Table 5. After the unit has been running for 200 hours, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed and material balance was performed. The results are shown in Table 11.

[0138] The first bed 5 of the hydrocracking reactor is filled with 12.0 g of catalyst A1, which is the same as in Example 1; the second bed 9 is filled with 8.0 g of catalyst B1, which is the same as in Example 1.

[0139] The hydrorefining reactor 2 was filled with 16.7g of catalyst C1, which was the same as in Example 1.

[0140] The diene saturated reactor 7 was filled with 5g of catalyst D1, which was the same as in Example 1.

[0141] Table 4 Properties of feedstock for catalytic heavy gasoline

[0142] Catalytic heavy gasoline two Sulfur (ppm) 610 Nitrogen (ppm) / Aromatics (wt%) 32 Cyclics (wt%) 47 5% distillation point (°C) 103 95% distillation point (°C) 189

[0143] Table 5 Properties of Olefin-Degrading Catalytic Heavy Gasoline

[0144] Hydrogenated catalytic heavy gasoline two Nitrogen (ppm) / Aromatics (wt%) 30 diene value (g I2 / 100 g) 0.05

[0145] Comparative Example 3

[0146] use Figure 3 The process flow shown is the same as... Figure 1 The difference in the process flow shown is that the olefin-reducing catalytic heavy gasoline feedstock is mixed with hydrotreated catalytic diesel and hydrogen before entering the first bed of the hydrocracking reactor, where it sequentially contacts the first bed catalyst and the second bed catalyst. The properties of the catalytic heavy gasoline and catalytic diesel feedstock used are shown in Table 1. The hydrorefining reaction conditions, diene saturation reaction conditions, and hydrocracking reaction conditions used are shown in Table 2. The properties of the resulting olefin-reducing catalytic heavy gasoline and hydrotreated catalytic diesel are shown in Table 3. After 200 hours of unit operation, the hydrocarbon composition in the hydrogen-rich gaseous feedstock 12 and the liquid product feedstock 13 was analyzed, and material balance was performed; the results are shown in Table 11.

[0147] The first bed 5 of the hydrocracking reactor is filled with 12.0 g of catalyst A1, which is the same as in Example 1; the second bed 9 is filled with 8.0 g of catalyst B1, which is the same as in Example 1.

[0148] The hydrorefining reactor 2 was filled with 16.7g of catalyst C1, which was the same as in Example 1.

[0149] The diene saturated reactor 7 was filled with 5g of catalyst D1, which was the same as in Example 1.

[0150] Example 2

[0151] use Figure 1 The process flow and the steps described in the instruction manual are shown. The properties of the catalytic heavy gasoline feedstock used are shown in Table 1, and the properties of the catalytic diesel feedstock are shown in Table 6. The hydrorefining reaction conditions, diene saturation reaction conditions, and hydrocracking reaction conditions used are shown in Table 2. The properties of the resulting de-olefin catalytic heavy gasoline are shown in Table 3, and the properties of the resulting hydrotreated catalytic diesel are shown in Table 7. After 200 hours of unit operation, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed, and material balance was performed. The results are shown in Table 11.

[0152] 12.0 g of catalyst A2 was loaded into the first bed 5 of the hydrocracking reactor. The specific preparation process of catalyst A2 was as follows: 55.6 g of Beta molecular sieve and 58.6 g of pseudoboehmite were mixed, extruded, dried at 90 °C, and then calcined in air at 550 °C for 4 h to obtain the catalyst support. Then, 17.34 g of nickel nitrate hexahydrate and 6.02 g of ammonium metatungstate were prepared into a clear solution. The catalyst support was impregnated with an equal volume of metal solution, dried at 100 °C, and calcined at 520 °C for 2 h to obtain the catalyst precursor. The catalyst precursor was reduced at 450℃ for 4 hours in a hydrogen atmosphere, then cooled to 320℃ and injected with dimethyl disulfide for 4 hours to obtain catalyst A2 with the composition: 3.5wt% Ni, 3.3wt% WO2, 1.7wt% WS2, 0.5wt% WO3 / 50wt% Beta molecular sieve, and 41.0wt% Al2O3. 8.0g of catalyst B2 was loaded into the second bed 9. The specific preparation process of catalyst B2 was as follows: 18.9g of Beta molecular sieve, 25.6g of ZSM-5 molecular sieve, and 72.9g of pseudoboehmite were mixed, extruded, dried at 90℃, and then calcined in air at 550℃ for 4 hours to obtain the catalyst support. Then, 17.34g of nickel nitrate hexahydrate and 5.90g of ammonium metatungstate were prepared into a clear solution. The catalyst support was impregnated with an equal volume of metal solution, dried at 100℃, and calcined at 520℃ for 2 hours to obtain the catalyst precursor. The catalyst precursor was reduced at 450℃ for 4 hours in a hydrogen atmosphere, and then cooled to 320℃ and injected with dimethyl disulfide for 5 hours to obtain catalyst B2 with the following composition: 3.5wt% Ni-2.5wt% WO2-2.5wt% WS2-0.5wt% WO3 / 17.0wt% Beta molecular sieve-23.0wt% ZSM-5 molecular sieve-51wt% Al2O3.

[0153] The hydrorefining reactor 2 was filled with 16.7g of catalyst C1, which was the same as in Example 1.

[0154] The diene saturated reactor 7 was filled with 5g of catalyst D1, which was the same as in Example 1.

[0155] Table 6 Properties of Catalytic Diesel Feedstock 2

[0156] Catalytic diesel two Sulfur (ppm) 1075 Nitrogen (ppm) 641 Aromatics (wt%) 89 5% distillation point (°C) 210 95% distillation point (°C) 352

[0157] Table 7 Properties of Hydrocatalyzed Diesel

[0158] Hydrogenated catalytic diesel two Nitrogen (ppm) 8.63 Aromatics (wt%) 72 diene value (g I2 / 100 g) /

[0159] Example 3

[0160] use Figure 1 The process flow and the steps described in the instruction manual are shown. The properties of the catalytic heavy gasoline feedstock used are shown in Table 1, and the properties of the catalytic diesel feedstock are shown in Table 6. The hydrorefining reaction conditions and diene saturation reaction conditions used are shown in Table 2, and the hydrocracking reaction conditions are shown in Table 8. The properties of the resulting de-olefin catalytic heavy gasoline are shown in Table 3, and the properties of the resulting hydrotreated catalytic diesel are shown in Table 7. After 200 hours of unit operation, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed, and material balance was performed. The results are shown in Table 11.

[0161] The first bed 5 of the hydrocracking reactor was loaded with 12.0 g of catalyst A3. The specific preparation process of catalyst A3 was as follows: 44.4 g of Beta molecular sieve and 76.4 g of pseudoboehmite were mixed, extruded, dried at 90 °C, and then calcined in air at 550 °C for 4 h to obtain the catalyst support. Then, 7.43 g of nickel nitrate hexahydrate and 6.78 g of ammonium molybdate tetrahydrate were prepared into a clear solution. The catalyst support was impregnated with an equal volume of metal solution, dried at 100 °C, and calcined at 520 °C for 2 h to obtain the catalyst precursor. The catalyst precursor was reduced at 380℃ for 4 hours in a hydrogen atmosphere to obtain the first bed catalyst A3, which has the following composition: 1.5wt% Ni, 4.2wt% MoO2, 0.8wt% MoO3 / 40.0wt% Beta molecular sieve, and 53.5wt% Al2O3. The second bed 9 was filled with 8.0g of the second bed catalyst B3. The specific preparation process of catalyst B3 was as follows: 22.2g of Beta molecular sieve, 22.2g of ZSM-5 molecular sieve, and 76.4g of pseudoboehmite were mixed, extruded, dried at 90℃, and then calcined in air at 550℃ for 4 hours to obtain the catalyst support. Then, 7.43g of nickel nitrate hexahydrate and 6.78g of ammonium molybdate tetrahydrate were prepared into a clear solution. The catalyst support was impregnated with an equal volume of the metal solution, dried at 100℃, and calcined at 520℃ for 2 hours to obtain the catalyst precursor. The catalyst precursor was reduced at 380℃ for 4 hours in a hydrogen atmosphere to obtain the second bed catalyst B3 with the following composition: 1.5wt% Ni-4.2wt% MoO2-0.8wt% MoO3 / 20.0wt% Beta molecular sieve-20.0wt% ZSM-5 molecular sieve-53.5wt% Al2O3.

[0162] The hydrorefining reactor 2 was filled with 16.7g of hydrorefining catalyst C1, which was the same as in Example 1.

[0163] Diene saturated reactor 7 is loaded with 5g of diene saturated catalyst D1, which is the same as in Example 1.

[0164] Table 8 Reaction conditions

[0165]

[0166] Example 4

[0167] use Figure 1 The process flow and the steps described in the instruction manual are shown. The properties of the catalytic heavy gasoline and catalytic diesel feedstock used are shown in Table 1. The hydrorefining reaction conditions, diene saturation reaction conditions, and hydrocracking reaction conditions used are shown in Table 2. The properties of the resulting de-olefin catalytic heavy gasoline and hydrotreated catalytic diesel are shown in Table 3. After 200 hours of operation, the hydrocarbon composition in the hydrogen-rich gaseous stream 12 and the liquid product stream 13 was analyzed, and material balance was performed. The results are shown in Table 11.

[0168] The first bed 5 of the hydrocracking reactor is packed with 46.0 g of the first bed catalyst A4. The catalyst A4 is prepared as follows: 55.6 g of Beta zeolite and 57.6 g of pseudo-boehmite are mixed, extruded into strips, dried at 90°C, and then calcined at 550°C in air for 4 h to obtain a catalyst carrier. A clear solution of 17.28 g of cobalt nitrate hexahydrate and 8.40 g of ammonium molybdate tetrahydrate is prepared, and the catalyst carrier is impregnated with the metal solution in an equal volume, dried at 100°C, and then calcined at 520°C for 2 h to obtain a catalyst precursor. The catalyst precursor is reduced at 400°C in a hydrogen atmosphere for 4 h to obtain the first bed catalyst A4, which has the following composition:

[0169] 3.5wt%Co-5.2wt%MoO2-1.0wt%MoO3 / 50.0wt%Beta zeolite-40.3wt%Al2O3; the second bed 9 is packed with 414.0 g of the second bed catalyst B4. The catalyst B4 is prepared as follows: 22.2 g of Beta zeolite, 33.3 g of ZSM-5 zeolite, and 57.6 g of pseudo-boehmite are mixed, extruded into strips, dried at 90°C, and then calcined at 550°C in air for 4 h to obtain a catalyst carrier. A clear solution of 17.28 g of cobalt nitrate hexahydrate and 8.40 g of ammonium molybdate tetrahydrate is prepared, and the catalyst carrier is impregnated with the metal solution in an equal volume, dried at 100°C, and then calcined at 520°C for 2 h to obtain a catalyst precursor. The catalyst precursor is reduced at 400°C in a hydrogen atmosphere for 4 h to obtain the second bed catalyst B4, which has the following composition:

[0170] 3.5wt%Co-5.2wt%MoO2-1.0wt%MoO3 / 20.0wt%Beta zeolite-30.0wt%ZSM-5 zeolite-40.3wt%Al2O3.

[0171] The hydrofining reactor 2 is packed with 116.7 g of the hydrofining catalyst C1, which is the same as in Example 1.

[0172] The diene saturation reactor 7 is packed with 112.0 g of the diene saturation catalyst D1, which is the same as in Example 1.

[0173] Example 5

[0174] The hydrocracking reactor is operated at a temperature of 400°C, a pressure of 20 MPa, a liquid hourly space velocity of 1.0 h-1, and a hydrogen / oil ratio of 400:1. Figure 1The process scheme and the method steps described in the specification. The properties of the catalytic diesel feed used are as in Example 1 (see Table 1), and the properties of the catalytic heavy gasoline feed used are as in Table 9. The hydrofining reaction conditions, diene saturation reaction conditions and hydrocracking reaction conditions used are as in Table 2. The properties of the hydrocracked diesel obtained are as in Table 3, and the properties of the olefin-reduced catalytic heavy gasoline obtained are as in Table 10. The hydrocarbon compositions of the hydrogen-rich gas phase stream 12 and the liquid phase product stream 13 were analyzed and material balances were performed after the unit had been operated for 200 h, and the results are shown in Table 11.

[0175] The first bed 5 of the hydrocracking reactor was packed with 112.0 g of catalyst A1, which is the same as catalyst A1 in Example 1, and the second bed 9 was packed with 18.0 g of catalyst B1, which is the same as catalyst B1 in Example 1.

[0176] The hydrofining reactor 2 was packed with 116.7 g of catalyst C1, which is the same as catalyst C1 in Example 1.

[0177] The diene saturation reactor 7 was packed with 15 g of catalyst D1, which is the same as catalyst D1 in Example 1.

[0178] Table 9 Properties of the three catalytic heavy gasoline feeds

[0179]

[0180]

[0181] Table 10 Properties of the three olefin-reduced catalytic heavy gasoline

[0182] Hydrogenated catalytic heavy gasoline three Nitrogen (ppm) / Aromatics (wt%) 47 diene value (g I2 / 100 g) 0.1

[0183] Table 11 Hydrocarbon yields of the products

[0184]

[0185] C in Example 2 11+ The reason for the higher content of components is mainly affected by the properties of the catalytic diesel feed. The higher the T95, the heavier the oil, and the more heavy components will be produced as a result of the reaction. Although the same catalytic diesel as in Example 2 was used in Example 3, the reaction conditions in Example 3 were more severe, and increasing the reaction temperature and the hydrogen to oil ratio can improve the product quality within a certain range, so the content of components C in Example 3 11+ is lower.

[0186] As can be seen from Example 1, Comparative Example 1 and Table 11, the scheme using only catalytic diesel as the feed obtains a lower light aromatic hydrocarbon yield than the co-refining scheme of catalytic heavy gasoline and catalytic diesel proposed in the present application, but the total calculated chemical hydrogen consumption of the hydrocracking is higher than that of the co-refining scheme of catalytic heavy gasoline and catalytic diesel.

[0187] As can be seen from Example 1, Comparative Example 2 and Table 11, if the selected catalytic heavy gasoline feedstock is not suitable, its content of cyclic hydrocarbons (including naphthenes and aromatics) is less than 50 wt% of the total stream, which will affect the product quality. The light aromatic yield in the product of Comparative Example 2 is much lower than that of Example 1, and the non-aromatic yield, especially the low carbon non-aromatic yield produced by cracking, is significantly increased.

[0188] As can be seen from Example 1, Comparative Example 3 and Table 11, if the olefin-reduced catalytic heavy gasoline is combined with the hydrocatalytic diesel before the first bed of the hydrocracking, and then successively contacts the first bed and the second bed of the hydrocracking catalyst, the light aromatic yield in the final product is lower than that of the method in which the olefin-reduced catalytic heavy gasoline only contacts the second bed of the catalyst in Example 1, and vice versa, the non-aromatic yield, especially the low carbon hydrocarbon yield produced by cracking, and the C 11+ The component yield is higher, and the result is less ideal.

Claims

1. A method for co-conversion of catalytic heavy gas oil and catalytic diesel into chemical feedstocks, comprising the step of feeding a hydrofinished hydrocatalytic diesel and a diene-saturated naphthenic catalytic heavy gas oil into a hydrocracking reactor in stages to produce chemical feedstocks comprising light aromatic hydrocarbons; the method specifically comprising: 1) passing a catalytic diesel stream through a hydrofining unit to produce a hydrofinished hydrocatalytic diesel stream; 2) mixing the hydrofinished hydrocatalytic diesel stream with hydrogen gas in step 1) into a first bed of the hydrocracking reactor to contact a first bed catalyst to produce a first bed effluent; 3) passing a catalytic heavy gas oil stream through a diene saturation unit to produce a naphthenic catalytic heavy gas oil stream; 4) passing the naphthenic catalytic heavy gas oil stream in step 3) into a second bed of the hydrocracking reactor to mix with the first bed effluent and contact a second bed catalyst to produce a hydrocracking outlet stream; 5) passing the hydrocracking outlet stream in step 4) through a separation unit to produce a hydrogen-rich gaseous stream and a liquid product stream.

2. The method of claim 1, wherein in step 1), the catalytic diesel stream has a T5 in the range of 150-240°C and a T95 in the range of 320-360°C.

3. The method of claim 2, wherein in step 1), the catalytic diesel stream has a T5 in the range of 190-215°C and a T95 in the range of 320-345°C.

4. The method of claim 1, wherein the hydrofinished hydrocatalytic diesel stream produced in step 1) has a total nitrogen content of less than 10 ppm and a total aromatic hydrocarbon content of greater than 70 wt%.

5. The method of claim 4, wherein the hydrofinished hydrocatalytic diesel stream produced in step 1) has a total aromatic hydrocarbon content of greater than 80 wt%.

6. The method of claim 5, wherein the hydrofinished hydrocatalytic diesel stream produced in step 1) has a total aromatic hydrocarbon content of greater than 85 wt%.

7. The method of claim 1, wherein in step 2), the first bed catalyst comprises a molecular sieve acidic support A, a metal active component A, and a binder A; the molecular sieve acidic support A comprises at least one twelve-membered ring molecular sieve, and the molecular sieve acidic support A has a mass in the range of 5-80 wt% of the mass of the first bed catalyst; and / or the metal active component A comprises a Group VIII metal in the range of 0.01-10 wt% of the mass of the first bed catalyst and a Group VIB metal oxide and / or sulfide in the range of 3-25 wt% of the mass of the first bed catalyst.

8. The method of claim 7, wherein in step 2), the molecular sieve acidic support A has a mass in the range of 20-75 wt% of the mass of the first bed catalyst; and / or the metal active component A comprises a Group VIII metal in the range of 0.5-8 wt% of the mass of the first bed catalyst and a Group VIB metal oxide and / or sulfide in the range of 5-15 wt% of the mass of the first bed catalyst.

9. The method of claim 8, wherein in step 2), the molecular sieve acidic support A has a mass in the range of 30-70 wt% of the mass of the first bed catalyst. ​ ​ ​ ​ ​ ​ ​ ​ 10. The method of claim 1, wherein: in step 3), The catalytic heavy gasoline stream comprises C 10 The following hydrocarbon species, wherein the cyclic hydrocarbon species content is greater than 50 wt%; and / or, the catalytic heavy gasoline stream has a T5 in the range of 70-150°C and a T95 in the range of 180-240°C.

11. The method of claim 10, wherein: in step 3), The catalytic heavy gasoline stream comprises C 10 The following hydrocarbon material, wherein the cyclic hydrocarbon content is greater than 60 wt%.

12. The method of claim 11, wherein: in step 3), The catalytic heavy gasoline stream comprises C 10 The following hydrocarbon material, wherein the cyclic hydrocarbon content is greater than 70 wt%.

13. The method of claim 1, wherein: the diolefin value of the reduced olefin catalytic heavy gasoline stream obtained in step 3) is less than 0.2 g I2 / 100 g oil.

14. The method of claim 1, wherein: in step 3), the flow rate of the catalytic heavy gasoline stream is 0.1-99 wt% of the total feed stream flow rate.

15. The method of claim 14, wherein: in step 3), the flow rate of the catalytic heavy gasoline stream is 10-80 wt% of the total feed stream flow rate.

16. The method of claim 15, wherein: in step 3), the flow rate of the catalytic heavy gasoline stream is 15-60 wt% of the total feed stream flow rate.

17. The method of claim 1, wherein: in step 4), the second bed catalyst is 10-80 wt% of the total catalyst loading of the hydrocracking reactor.

18. The method of claim 17, wherein: in step 4), the second bed catalyst is 20-70 wt% of the total catalyst loading of the hydrocracking reactor.

19. The method of claim 18, wherein: in step 4), the second bed catalyst is 40-70 wt% of the total catalyst loading of the hydrocracking reactor.

20. The method of claim 7, wherein: in step 4), the second bed catalyst comprises a molecular sieve acidic support B, a metal active component B, and a binder B; the molecular sieve acidic support B comprises at least one twelve-membered ring molecular sieve and at least one ten-membered ring molecular sieve; the total mass of the molecular sieves is 10-60 wt% of the mass of the second bed catalyst; and / or, the metal active component B comprises a Group VIII metal in an amount of 0.01-10 wt% of the mass of the second bed catalyst and a Group VIB metal oxide and / or sulfide in an amount of 3-25 wt% of the mass of the second bed catalyst.

21. The method of claim 20, wherein: in step 4), the molecular sieve acidic support B comprises at least one twelve-membered ring molecular sieve and at least one ten-membered ring molecular sieve; the total mass of the molecular sieves is 35-55 wt% of the mass of the second bed catalyst; and / or, the metal active component B comprises a Group VIII metal in an amount of 0.5-8 wt% of the mass of the second bed catalyst and a Group VIB metal oxide and / or sulfide in an amount of 5-15 wt% of the mass of the second bed catalyst.

22. The method of claim 20, wherein: the twelve-membered ring molecular sieve in the first bed catalyst and the twelve-membered ring molecular sieve in the second bed catalyst are the same or different, each independently at least one of a Beta molecular sieve, a USY molecular sieve, and a mordenite molecular sieve; and / or, the twelve-membered ring molecular sieve in the first bed catalyst and the twelve-membered ring molecular sieve in the second bed catalyst are the same or different, each independently at least one of a Beta molecular sieve, a USY molecular sieve, and a mordenite molecular sieve; and / or, The ten-membered ring molecular sieve in the second bed catalyst is at least one of ZSM-5, ZSM-11, ZSM-22, SAPO-11, EU-1 molecular sieve; and / or, The Group VIII metal in the first bed catalyst and the Group VIII metal in the second bed catalyst are the same or different, and are independently at least one of platinum, palladium, cobalt, nickel, iridium; and / or, The Group VIB metal oxide and / or sulfide in the first bed catalyst and the Group VIB metal oxide and / or sulfide in the second bed catalyst are the same or different, and are independently at least one of oxide and / or sulfide of molybdenum and / or tungsten; and / or, The binder A and the binder B are the same or different, and are independently at least one of alumina, silica, silica-alumina composite, amorphous silica-alumina.

23. The method of claim 20, wherein: The mass of the ten-membered ring molecular sieve in the second bed catalyst accounts for 10-90 wt% of the total mass of the molecular sieve acidic carrier B.

24. The method of claim 23, wherein: The mass of the ten-membered ring molecular sieve in the second bed catalyst accounts for 25-75 wt% of the total mass of the molecular sieve acidic carrier B.

25. The method of claim 24, wherein: The mass of the ten-membered ring molecular sieve in the second bed catalyst accounts for 40-60 wt% of the total mass of the molecular sieve acidic carrier B.

26. The method of claim 1, wherein: In step 5), the hydrocracking outlet stream is subjected to high-pressure separation to obtain a hydrogen-rich gas phase stream and a liquid phase product stream; and / or, The hydrocarbon components of the hydrogen-rich gas phase stream and liquid phase product stream obtained in Step 5) comprise light aromatics, C 10 The following non-aromatic and C 11+ components; The yield of the light aromatic hydrocarbon is greater than 35 wt%; and / or, The C 10 The following non-aromatic hydrocarbon yield is less than 60 wt%; and / or, The C 11+ Component yield is less than or equal to 20 wt%.

27. The method of claim 26, wherein: The yield of the light aromatic hydrocarbon is greater than 40 wt%; and / or, The C 10 The following non-aromatics yield is less than 55 wt%; and / or, The C 11+ The component yield is less than or equal to 15 wt%.

28. The method of claim 27, wherein: The yield of the light aromatic hydrocarbon is greater than 42 wt%; and / or, The C 11+ The component yield is less than or equal to 8 wt%.

29. The method of claim 1, wherein: In step 5), the hydrogen-rich gas phase stream is subjected to pressure increase to obtain a pressurized hydrogen-rich gas phase stream mixed with fresh hydrogen to form mixed hydrogen, which is then mixed with the hydrocracked diesel stream to enter the hydrocracking reactor for recycling.

30. The method of any one of claims 1-29, wherein: The process conditions of the hydrocracking reactor include: The reactor inlet temperature is 280-450°C; and / or, The hydrogen partial pressure is 5-10 MPa; and / or, mass velocity 0.5-2.0 h -1 ; and / or, The hydrogen to oil volume ratio of fresh hydrogen to hydrocracked diesel is 400-1000; and / or, The hydrogen to oil volume ratio of mixed hydrogen to hydrocracked diesel is 1000-5000.

31. The method of claim 30, wherein: The process conditions of the hydrocracking reactor include: The reactor inlet temperature is 310-430°C; and / or, The hydrogen partial pressure is 5-9 MPa; and / or, mass velocity 0.6-1.5 h -1 ; and / or, The hydrogen to oil volume ratio of fresh hydrogen to hydrocracked diesel is 500-800; and / or, The hydrogen to oil volume ratio of mixed hydrogen to hydrocracked diesel is 2000-4000.

32. The method of claim 31, wherein: the process conditions of the hydrocracking reactor include: a reactor inlet temperature of 350-410°C; and / or, a hydrogen partial pressure of 6-8 MPa; and / or, mass velocity 0.8-1.2 h -1 ; and / or, a fresh hydrogen to hydrocatalytic diesel hydrogen to oil volume ratio of 500-600; and / or, a mixed hydrogen to hydrocatalytic diesel hydrogen to oil volume ratio of 2000-3000.

33. A system for producing chemical feedstocks using the method of any one of claims 1-32, comprising: a first device configured to receive the catalytic diesel stream and to discharge the hydrocatalytic diesel stream; a second device configured to receive the catalytic heavy gasoline stream and to discharge the olefin-reduced catalytic heavy gasoline stream; a third device configured to receive the hydrocatalytic diesel stream and hydrogen in stages and the olefin-reduced catalytic heavy gasoline stream and to discharge the hydrocracking outlet stream; a fourth device configured to receive the hydrocracking outlet stream and to discharge the hydrogen-rich vapor stream and the liquid phase product stream.

34. The system of claim 33, further comprising: a fifth device configured to receive the hydrogen-rich vapor stream and to discharge a pressurized hydrogen-rich vapor stream.

35. The system of claim 34, wherein: the reactor in the first device is a hydrofining reactor; and / or, the reactor in the second device is a diene saturation reactor; and / or, the reactor in the third device is a hydrocracking reactor, the hydrocracking reactor including a first bed packed with the first bed catalyst and a second bed packed with the second bed catalyst; and / or, the fourth device includes a high pressure knock-out drum; and / or, the fifth device includes a hydrogen compressor.

36. Use of the method of any one of claims 1-32 to produce chemical feedstocks from catalytic diesel, the chemical feedstocks including light aromatic hydrocarbons.

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