Process for producing benzene-rich naphtha

Abstract

The present invention relates to the field of catalytic cracking, and discloses a method for producing benzene-rich naphtha. The method includes: hydrogen and heavy diesel are contacted with a hydrorefining catalyst to obtain a hydrotreating effluent having an organic nitrogen content of 5-50 μg / g; the hydrotreating effluent is contacted with a hydrocracking catalyst to obtain a hydrocracking effluent, the hydrocracking catalyst including a carrier and a VIB group metal component loaded on the carrier, the carrier containing an MFI molecular sieve; the weight ratio of the VIB group metal component to the MFI molecular sieve in terms of oxide is 0.15-1: 1; the hydrocracking effluent is fractionated to obtain a naphtha fraction. The method provided by the present invention has a simple process flow, a low reaction temperature, and relatively low energy consumption; in the naphtha fraction obtained, the aromatics yield is improved, and in the particularly light naphtha fraction (C5-C7), the benzene yield reaches 17.2%, which is significantly improved.

Description

Technical Field

[0001] The present invention relates to the field of catalytic cracking, and in particular to a method for producing benzene-rich naphtha. Background Art

[0002] With the development of the economy, the growth rate of demand for aromatics has accelerated, which makes the conversion of heavy aromatics into low-carbon aromatics through hydrocracking have real economic value. The demand for low-carbon aromatics is large, and the products with industrial value mainly include benzene, toluene, xylene, ethylbenzene, etc. The most important use of benzene in industry is the production of chemical raw materials, such as styrene, phenol, cyclohexane, maleic anhydride, nitrobenzene, alkylbenzene, caprolactam, etc. The existing industrial methods mainly produce it by separation from catalytic reforming production oil, coal tar, and cracking oil. It can also be produced by disproportionation and toluene dealkylation reactions. These processes have disadvantages such as high reaction temperature (500-600℃) and complex raw materials.

[0003] CN104560164A discloses a hydroreforming method for producing a high-octane gasoline component or BTX feedstock. This method involves mixing low-quality diesel with a hydrogen-rich gas, which is then introduced into a hydrorefining reaction zone. The mixture then contacts and reacts with a hydrorefining catalyst. The reaction effluent from the hydrorefining reaction zone enters a hydroreforming reaction zone without passing through any intermediate separation facilities, where it sequentially contacts and reacts with a first hydroreforming catalyst and a second hydroreforming catalyst. The effluent from the hydroreforming reaction zone is cooled and separated to produce a hydrogen-rich gas and a liquid product. This method uses low-quality diesel fractions as a feedstock and can produce high-octane gasoline. The BTX content in the high-octane gasoline fraction can reach over 40%.

[0004] CN111100705A discloses a method for maximizing the production of aromatics by catalytic diesel hydroconversion, comprising the following steps: (1) mixing high-aromatic catalytic diesel with recycled hydrogen and entering a hydrorefining reaction zone for reaction, wherein the hydrorefining reaction zone is provided with at least two catalyst beds, and the sulfidation degree of the catalyst tends to decrease along the direction of the flow; (2) the product oil obtained in step (1) enters a hydrocracking reaction zone and contacts at least two sulfided hydrocracking catalysts to carry out a ring-opening conversion reaction of polycyclic aromatic hydrocarbons; (3) the product oil obtained in step (2) is passed through a separation system to obtain gas, gasoline, and diesel fractions, and the diesel fraction is entirely recycled to the hydrorefining reaction zone; (4) the naphtha obtained in step (3) enters an aromatics extraction unit and is extracted with a solvent to obtain a BTX product. This method balances the initial activity of the catalysts by using graded hydrogenation catalysts with different sulfidation degrees, thereby resolving the problem of a long initial adjustment period and a long-term low octane number of the gasoline fraction, the main target product, during the application of the technology.

[0005] These methods utilize a graded hydrocracking zone to produce a specific amount of naphtha fraction with a high BTX content, but the overall operational process is highly complex. Furthermore, because the process utilizes a bimetallic hydrogenation component, the catalyst exhibits a high hydrogenation function and a strong ability to saturate aromatics due to the promoter effect of the Group VIII metal on the Group VIB metal in its sulfided state. Furthermore, the mutual support effect of the bimetallic sulfided metal system allows for the system to accommodate more active hydrogen, preventing the low-carbon aromatics produced during the ring-opening and side-chain scission processes from being fully unsaturated. This results in a limited total aromatics yield in the final product, particularly benzene yield. Summary of the Invention

[0006] The purpose of the present invention is to overcome the problem of low benzene yield in naphtha fractions in the prior art and provide a method for producing benzene-rich naphtha. Among the fractions generated by this method, the benzene yield in the light naphtha fraction (C5-C7) reaches 17.2%, which is a significant improvement.

[0007] In order to achieve the above object, the present invention provides a method for producing benzene-rich naphtha, which comprises the following steps:

[0008] (1) subjecting hydrogen and heavy diesel to a first contact reaction with a hydrotreating catalyst to obtain a hydrotreating effluent, wherein the content of organic nitrogen in the hydrotreating effluent is 5-50 μg / g;

[0009] (2) subjecting the hydrotreated effluent to a second contact reaction with a hydrocracking catalyst to obtain a hydrocracking effluent, wherein the hydrocracking catalyst comprises a carrier and a Group VIB metal component supported on the carrier, wherein the carrier comprises an MFI molecular sieve; the weight ratio of the Group VIB metal component to the MFI molecular sieve, calculated as oxide, is 0.15-1:1; the conditions of the second contact reaction include: a reaction pressure of 5.0-7.5 MPa; a reaction temperature of 380-420° C.; and a hydrogen-to-oil volume ratio of 500-1200:1;

[0010] (3) fractionating the hydrocracking effluent to obtain a naphtha fraction;

[0011] Wherein, the density of the heavy diesel is 0.9-1.05g / cm 3 , the distillation range is 180-380° C., and the content of paraffins in the heavy diesel is less than 20% by weight.

[0012] Existing industrial hydrocracking has the characteristics of low reaction temperature, simple feed and process. Its raw materials include heavy and low-quality fractions such as VGO, which have a wide distillation range and contain a large amount of aromatics and cycloalkanes. These aromatics and cycloalkanes usually exist in the form of side chains, thereby maintaining a certain distillation range in the mixture. The reactions occurring at the acid center in the conventional hydrocracking process are usually based on the positive carbon ion mechanism, and the feed macromolecules are converted into small molecules through ring opening and side chain scission reactions. However, the carbon-carbon bond breakage in this process is usually based on Beta cleavage, that is, the breakage of the adjacent chemical bonds of the positive carbon ion coordination chemical bond, which makes the reaction products of this process usually contain certain branches, making it difficult to obtain low-branched or even unbranched low-carbon aromatic compounds (BTEX).

[0013] The present invention adopts a heavy diesel fraction with a lower distillation range and rich in aromatics as a raw material, wherein the density of the heavy diesel is 0.9-1.05 g / cm 3 , a distillation range of 180-380°C, and the content of paraffins in the heavy diesel is less than 20% by weight, and the content of organic nitrogen in the hydrotreatment effluent is controlled to be 5-50 μg / g. At the same time, combined with a specific hydrocracking catalyst, the hydrocracking catalyst includes a carrier and a VIB Group metal component supported on the carrier, wherein the carrier contains an MFI molecular sieve; the weight ratio of the VIB Group metal component to the MFI molecular sieve, calculated as oxide, is 0.15-1:1, and specific second contact reaction conditions can effectively avoid the problem that it is difficult to obtain low-branched or even unbranched low-carbon aromatic compounds (BTEX) using heavy and inferior fractions such as VGO as raw materials. In addition, the aromatics yield in the naphtha fraction obtained by the method provided by the present invention is improved, especially the benzene yield in the light naphtha fraction (C5-C7) reaches 17.2%, which is a significant improvement.

[0014] Through the above technical solution, the present invention has the following advantages:

[0015] (1) Compared with existing industrial methods such as separation from catalytic reforming oil, coal tar, and cracking oil or preparation through disproportionation and toluene dealkylation, the method provided by the present invention has a simple process flow, a low reaction temperature, and relatively low energy consumption;

[0016] (2) Compared with the existing method of using a bimetallic hydrogenation component as a hydrocracking catalyst, the method provided by the present invention uses a monometallic hydrogenation component as the hydrocracking catalyst in combination with a carrier containing an MFI molecular sieve, which is beneficial for making the low-carbon aromatics generated by the reaction as unsaturated as possible during the ring opening and side chain scission process, thereby improving the yield of aromatics in the final product, especially the yield of benzene. DETAILED DESCRIPTION

[0017] The endpoints of the ranges and any values ​​disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0018] In the present invention, the dry basis of a substance refers to a solid product obtained by calcining the substance at 600°C for 3 hours.

[0019] In the present invention, "optional" means not essential and can be understood as including or not including.

[0020] In the present invention, the density of the heavy diesel is the density at 20°C.

[0021] The present invention provides a method for producing benzene-rich naphtha, the method comprising the following steps:

[0022] (1) subjecting hydrogen and heavy diesel to a first contact reaction with a hydrotreating catalyst to obtain a hydrotreating effluent, wherein the content of organic nitrogen in the hydrotreating effluent is 5-50 μg / g;

[0023] (2) subjecting the hydrotreated effluent to a second contact reaction with a hydrocracking catalyst to obtain a hydrocracking effluent, wherein the hydrocracking catalyst comprises a carrier and a Group VIB metal component supported on the carrier, wherein the carrier comprises an MFI molecular sieve; the weight ratio of the Group VIB metal component to the MFI molecular sieve, calculated as oxide, is 0.15-1:1; the conditions of the second contact reaction include: a reaction pressure of 5.0-7.5 MPa; a reaction temperature of 380-420° C.; and a hydrogen-to-oil volume ratio of 500-1200:1;

[0024] (3) fractionating the hydrocracking effluent to obtain a naphtha fraction;

[0025] Wherein, the density of the heavy diesel is 0.9-1.05g / cm 3 , the distillation range is 180-380° C., and the content of paraffins in the heavy diesel is less than 20% by weight.

[0026] According to some embodiments of the present invention, the raw material used in the method is heavy diesel with a lower distillation range and rich in aromatics, and the density of the heavy diesel is 0.9-1.05 g / cm 3, a distillation range of 180-380°C, and a paraffin content of less than 20% by weight in the heavy diesel. Conventional feedstocks used in the prior art have the disadvantage of having a wide distillation range and a low content of active components that can be converted into low-carbon aromatics. The method provided by the present invention utilizes heavy diesel that meets these requirements, thereby improving the yield of the target product. Any heavy diesel that meets these requirements is within the applicable scope of the present invention.

[0027] According to some embodiments of the present invention, the heavy diesel may further contain other fractions, wherein the other fractions are selected from at least one of catalytic cracking cycle oil, coal tar, coal liquefaction oil, naphthenic diesel, heavy oil ebullated bed diesel fraction, slurry bed diesel fraction, and suspended bed diesel fraction. The catalytic cracking cycle oil, coal tar, coal liquefaction oil, naphthenic diesel, heavy oil ebullated bed diesel fraction, slurry bed diesel fraction, and suspended bed diesel fraction have conventional definitions in the art and are not further described herein.

[0028] According to some embodiments of the present invention, the density of the heavy diesel is 0.9-1.05 g / cm 3 , preferably 0.92-0.97 g / cm 3 The use of heavy diesel with the preferred density is more conducive to further improving the yield of benzene in the product.

[0029] According to some embodiments of the present invention, the heavy diesel further contains sulfur and / or nitrogen, and there is no particular limitation on the content of sulfur and / or nitrogen in the heavy diesel. To facilitate subsequent conversion, preferably, the heavy diesel does not contain nitrogen.

[0030] According to some embodiments of the present invention, preferably, the method further comprises: removing mechanical impurities from the heavy diesel before performing step (1). The present invention does not particularly limit the method for removing the mechanical impurities, and conventional removal methods in the art, such as filtration, can be used.

[0031] According to some embodiments of the present invention, in step (1), the organic nitrogen content in the hydrotreatment effluent is 5-50 μg / g. The organic nitrogen content in the hydrotreatment effluent can be adjusted by controlling the conditions of the first contact reaction, that is, by adopting optional conventional diesel hydrofining conditions. The present invention has a wide range of selection conditions for the first contact reaction, based on whether the hydrofining reaction can occur and whether the organic nitrogen content in the resulting hydrotreatment effluent meets the above requirements.

[0032] According to some embodiments of the present invention, preferably, the conditions of the first contact reaction include: reaction pressure of 2-8 MPa; reaction temperature of 350-440°C; hydrogen-to-oil volume ratio of 100-1400; liquid hourly space velocity of 0.5-3h-1 .

[0033] According to some embodiments of the present invention, in step (2), the conditions of the second contact reaction include: a reaction pressure of 5.0-7.5 MPa; a reaction temperature of 380-420°C; a hydrogen-to-oil volume ratio of 500-1200:1; preferably, the liquid hourly space velocity of the second contact reaction is 0.5-3.0 h -1 .

[0034] According to some embodiments of the present invention, the hydrocracking reaction process is a fixed-bed continuous flow one-time process flow, which satisfies the requirement that during the hydrocracking process, the inlet logistics and the outlet logistics flow continuously upstream and downstream, and as the reaction proceeds, the reaction pressure, reaction temperature, hydrogen-to-oil volume ratio, and liquid hourly space velocity can be maintained at fixed values, and changes in these parameters only need to meet specific process conditions set externally.

[0035] According to some embodiments of the present invention, the second contact reaction can be carried out in any conventional reaction device that is sufficient to cause the hydrotreatment effluent to undergo a hydrocracking reaction with the hydrocracking catalyst, for example, a fixed bed reactor, a moving bed reactor, an ebullient bed reactor or a slurry bed reactor, and the present invention has no particular limitation on this.

[0036] In the method provided by the present invention, the reaction pressure and reaction temperature are relatively low, and the above preferred embodiment is conducive to further improving the yield of benzene in the product.

[0037] According to some embodiments of the present invention, the hydrocracking catalyst comprises a support and a Group VIB metal component supported on the support, wherein the support comprises an MFI molecular sieve; and the weight ratio of the Group VIB metal component to the MFI molecular sieve, calculated as oxide, is 0.15-1:1. To further increase the yield of aromatics in the naphtha fraction, particularly the yield of benzene in the light naphtha fraction (C5-C7), the hydrocracking catalyst preferably does not contain a Group VIII metal component.

[0038] According to some embodiments of the present invention, preferably, the content of the Group VIB metal component is 5-30 wt%, preferably 10-25 wt%, based on the total dry weight of the hydrocracking catalyst and calculated as oxide.

[0039] According to some embodiments of the present invention, preferably, the content of the carrier is 70-95 wt%, preferably 75-90 wt%, based on the total dry weight of the hydrocracking catalyst and calculated on a dry basis.

[0040] According to some embodiments of the present invention, preferably, in the hydrocracking catalyst, the weight ratio of the Group VIB metal component, calculated as oxide, to the MFI molecular sieve is 0.2-0.5:1. A weight ratio of the Group VIB metal component to the MFI molecular sieve within the preferred range is beneficial for further improving the yield of benzene in the product.

[0041] According to some embodiments of the present invention, preferably, the Group VIB metal component is Mo and / or W, more preferably W.

[0042] According to some embodiments of the present invention, preferably, the MFI molecular sieve is selected from ZSM-5 molecular sieve and / or ZRP molecular sieve, more preferably ZSM-5 molecular sieve.

[0043] According to some embodiments of the present invention, preferably, the silicon-aluminum ratio of the MFI molecular sieve is 15-300, preferably 20-80.

[0044] According to some embodiments of the present invention, preferably, the specific surface area of ​​the MFI molecular sieve is 180-650m 2 / g, preferably 300-450m 2 / g.

[0045] According to some embodiments of the present invention, preferably, the pore volume of the MFI molecular sieve is 0.1-0.6 mL / g, preferably 0.2-0.5 mL / g.

[0046] According to some embodiments of the present invention, the specific surface area and pore volume of the MFI molecular sieve can be measured by a static low-temperature adsorption capacity method.

[0047] According to some embodiments of the present invention, preferably, the carrier also contains a heat-resistant inorganic oxide. The heat-resistant inorganic oxide refers to a porous material with a maximum operating temperature of not less than 600°C. The heat-resistant inorganic oxide can increase the strength of the hydrocracking catalyst, and improve and adjust the physicochemical properties of the hydrocracking catalyst, such as improving the pore structure of the catalyst. The heat-resistant inorganic oxide can be an inorganic oxide commonly used for hydrogenation catalyst carriers, for example, it can be selected from at least one of silicon oxide, aluminum oxide and titanium oxide. Preferably, the heat-resistant inorganic oxide is selected from silicon oxide and / or aluminum oxide, more preferably aluminum oxide, and further preferably, the specific surface area of ​​the aluminum oxide is 150-400m 2 / g, preferably 200-350m 2 / g; the pore volume is 0.45-1.3 mL / g, preferably 0.9-1.2 mL / g. The alumina may include gibbsite, such as gibbsite, bayerite, and nordstrandite, and the alumina may also include diaspore, such as boehmite and pseudoboehmite.

[0048] According to some embodiments of the present invention, the content of the MFI molecular sieve and the heat-resistant inorganic oxide in the carrier is selected in a wide range. Preferably, based on the total amount of the carrier, the content of the MFI molecular sieve is 40-90% by weight, preferably 50-80% by weight; the content of the heat-resistant inorganic oxide is 10-60% by weight, preferably 20-50% by weight.

[0049] According to some embodiments of the present invention, the method for preparing the hydrocracking catalyst has a wide range of options. Preferably, the method for preparing the hydrocracking catalyst comprises the following steps:

[0050] (a) mixing an MFI molecular sieve with an optional heat-resistant inorganic oxide to prepare a carrier;

[0051] (b) introducing the Group VIB metal component onto the support by impregnation.

[0052] According to some embodiments of the present invention, in step (a), the method for preparing the carrier is well known to those skilled in the art and is not particularly limited in the present invention. For example, the method may include: mixing the MFI molecular sieve with an optional heat-resistant inorganic oxide, and then shaping and drying to obtain the carrier. The shaping may be performed using any conventional method in the art, such as tableting, ball rolling, or extrusion. There is also no particular limitation on the shape of the carrier, which may be spherical, bar-shaped (including solid or hollow bars), block-shaped, etc. The bar-shaped may be a multi-leaf clover-shaped, four-leaf clover-shaped, or a variant thereof.

[0053] According to some embodiments of the present invention, in step (b), the step of introducing the Group VIB metal component onto the support by impregnation includes: contacting the support with an impregnation solution, the impregnation solution comprising a metal precursor and a solvent, the metal precursor comprising a Group VIB element, preferably tungsten, more preferably, the metal precursor being selected from at least one of tungstic acid, metatungstic acid, ethylmetatungstic acid, tungstate, metatungstate, and ethylmetatungstate, further preferably metatungstate, and even more preferably ammonium metatungstate; and the solvent is preferably water. There is also no particular limitation on the amount of the solvent used, and it can be used in accordance with existing techniques, as long as the loading amount of the Group VIB metal component is guaranteed.

[0054] According to some embodiments of the present invention, there is no particular limitation on the contact temperature, and it can be any temperature that the impregnation liquid can reach. There is no particular limitation on the contact time, as long as the required amount of metal precursor can be loaded on the carrier. In general, the higher the contact temperature and the greater the concentration of the impregnation liquid, the shorter the time required to reach the same impregnation amount (i.e., the weight difference between the carrier after impregnation and before impregnation); and vice versa. Once the required impregnation amount and conditions are determined, it is easy to select a suitable contact time. The specific operation of the impregnation method is well known to those skilled in the art, and the impregnation method can be a saturated impregnation method or a supersaturated impregnation method. There is no particular limitation on the environment of the impregnation method. According to conventional methods in the art, it can be carried out under sealed conditions or in an open environment. The lost solvent can be replenished during the contact process, or it may not be replenished. Various gases, such as air, nitrogen, water vapor, etc., can be introduced during the contact process, or no new components may be introduced.

[0055] According to some embodiments of the present invention, after contacting the support with the impregnation solution, preferably, the steps of drying and calcining the resulting material may also be included. Drying and calcining are conventional steps in catalyst preparation and are not particularly limited. For example, the drying conditions may be: a temperature of 80-350°C, preferably 100-300°C, and a time of 0.5-24 hours, preferably 1-12 hours. The calcining conditions may be: a temperature of 350-600°C, preferably 400-550°C, and a time of 0.2-12 hours, preferably 1-10 hours.

[0056] According to some embodiments of the present invention, preferably, before using the hydrocracking catalyst, the process further includes a step of pre-sulfurizing the hydrocracking catalyst with a sulfur-containing feedstock (e.g., sulfur, hydrogen sulfide) at a temperature of 140-370°C in the presence of hydrogen. The pre-sulfurization can be carried out outside the reactor or in situ in the reactor to convert the Group VIB metal component from an oxidized state to a sulfided state.

[0057] According to some embodiments of the present invention, the hydrotreating catalyst may be a catalyst having catalytic activity for aromatic saturation, hydrodesulfurization, and hydrodenitrogenation, and may be a precious metal catalyst or a non-precious metal catalyst. Preferably, the hydrotreating catalyst is a non-precious metal catalyst.

[0058] According to some embodiments of the present invention, the hydrorefining catalyst preferably comprises a support and a Group VIB metal component and a Group VIII metal component supported on the support. Based on the total amount of the hydrorefining catalyst and calculated as oxide, the Group VIII metal component may be present in an amount of 1-10% by weight, preferably 1.5-7% by weight; the Group VIB metal component may be present in an amount of 5-50% by weight, preferably 7-35% by weight. Preferably, the hydrorefining catalyst may further comprise at least one additive, which may be at least one of phosphorus, fluorine, and boron. Based on the total amount of the hydrorefining catalyst and calculated as an element, the additive may be present in an amount of 1-10% by weight.

[0059] According to some embodiments of the present invention, preferably, in the hydrotreating catalyst, the Group VIB metal is Mo; and the Group VIII metal component is Ni.

[0060] According to some embodiments of the present invention, preferably, in the hydrorefining catalyst, the carrier is selected from alumina and / or silica.

[0061] According to some embodiments of the present invention, in the hydrotreating catalyst, the Group VIB metal component and the Group VIII metal component may both exist in the form of oxides.

[0062] According to some embodiments of the present invention, the hydrotreating catalyst may be a commercially available industrial catalyst or may be prepared by itself, and the present invention has no particular limitation thereto.

[0063] According to some embodiments of the present invention, there is no particular limitation on the method for preparing the hydrorefining catalyst, and the method can be prepared with reference to the prior art. For example, the method can be prepared with reference to the method disclosed in the prior art "Lv Weichao et al. Effect of modification of citric acid and phosphorus on the hydrodenitrogenation performance of Ni-Mo / Al2O3 coker wax oil [J]. Acta Petrolei Sinica (Petroleum Processing), 2014." It can also be prepared with reference to the method disclosed in the prior art "Hu Anpeng et al. Effect of citric acid on the effect of additive Ni in NiMoγ-Al2O3 catalyst [J]. Petroleum Refining and Chemical Industry, 2018, 49(10): 6." As long as a hydrorefining catalyst that meets the above requirements can be obtained, the hydrorefining catalyst can be prepared by, for example, an impregnation method. The impregnation method can be a co-impregnation method or a step-by-step impregnation method. The specific operation is well known to those skilled in the art, and the present invention will not be described in detail here.

[0064] According to some embodiments of the present invention, the volume ratio of the hydrorefining catalyst to the hydrocracking catalyst can be selected based on the specific reaction conditions, and the selectable range is relatively wide, all of which can achieve the objectives of the present invention to a certain extent. Preferably, the volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 0.2-3, preferably 0.4-1.5. Adoption of the above preferred embodiment is conducive to further improving the yield of benzene in the product.

[0065] According to some embodiments of the present invention, in step (3), the hydrocracking effluent is fractionated to obtain a naphtha fraction. The present invention does not particularly limit the specific operation steps of the fractionation, and various methods conventionally used in the art can be used for the fractionation. For example, the hydrocracking effluent can be introduced into a fractionation tower for fractionation to obtain the target product.

[0066] According to some embodiments of the present invention, in step (3), the products obtained by fractionating the hydrocracking effluent further include a gas component and / or a diesel fraction. The method can select the distillation range of each product according to the specific use of the fraction.

[0067] According to some embodiments of the present invention, preferably, the distillation range of the naphtha fraction is 45-85°C. In the naphtha fraction within this preferred distillation range, the yield of aromatics is high, and in particular, the yield of benzene in the light naphtha fraction (C5-C7) is significantly improved. It should be noted that the distillation range of 45-85°C does not mean that the initial distillation point is 45°C and the final distillation point is 85°C, but rather that the initial distillation point is not lower than 45°C and the final distillation point is not higher than 85°C.

[0068] According to some embodiments of the present invention, the hydrorefining catalyst and the hydrocracking catalyst may be loaded in different reaction zones of the same reactor, or may be located in different reactors. The selection may be made according to the specific device, and the present invention has no particular limitation on this.

[0069] According to some embodiments of the present invention, preferably, the yield of benzene in the naphtha fraction is greater than 14 wt %.

[0070] The present invention will be described in detail below through examples.

[0071] In the following examples and comparative examples, the hydrorefining catalyst used in the hydrorefining section is HDN-1, which is prepared by referring to the method disclosed in the prior art "Lv Weichao et al. Effect of Modification of Citric Acid and Phosphorus on the Hydrodenitrogenation Performance of Ni-Mo / Al2O3 Coker Wax Oil [J]. Acta Petrolei Sinica (Petroleum Processing), 2014." In this hydrorefining catalyst, the NiO content is 4.5wt%, the MoO3 content is 25wt%, and the rest is an alumina support.

[0072] Pseudoboehmite was purchased from Sinopec Catalyst Changling Branch under the trade name PB100, with a specific surface area of ​​263 m 2 / g, pore volume 1.03mL / g, dry basis 70wt%.

[0073] ZSM-5 molecular sieve was purchased from Sinopec Catalyst Jianchang Branch Company, with a specific surface area of ​​325m 2 / g, the framework silicon-aluminum ratio is 70, the pore volume is 0.31 mL / g, and the dry basis is 80 weight%.

[0074] In the following preparation examples and comparative preparation examples, the pore volume and specific surface area of ​​the molecular sieves were measured using an ASAP 2400 automatic adsorption apparatus from Micromertics Instruments, Inc., USA, by the static low-temperature adsorption capacity method (in accordance with the national standard GB / T5816-1995). The specific method is as follows: the molecular sieve to be tested is vacuum degassed at 250°C and 1.33 Pa for 4 hours, and then contacted with nitrogen as the adsorbate at -196°C to achieve adsorption equilibrium by static adsorption; the amount of nitrogen adsorbed by the adsorbent is calculated from the difference between the amount of nitrogen inlet and the amount remaining in the gas phase after adsorption, and the pore size distribution is then calculated using the BJH formula, and the specific surface area and pore volume are calculated using the BET formula.

[0075] Preparation Examples 1-2 are used to illustrate hydrocracking catalysts and their preparation methods.

[0076] Preparation Example 1

[0077] 35.7 g of pseudo-boehmite was mixed with ZSM-5 molecular sieve and extruded into trilobal strips with a circumscribed circle diameter of 1.6 mm. The strips were dried at 120° C. for 3 h and calcined at 600° C. for 3 h to obtain a carrier CS-1. After cooling to room temperature, ammonium metatungstate (tungsten oxide content: 91 wt %) was prepared into 60 mL of an impregnation solution (ammonium metatungstate aqueous solution). Then, 100 g of the carrier CS-1 was impregnated with the prepared ammonium metatungstate aqueous solution, dried at 120° C. for 3 h, and calcined at 480° C. for 4 h to obtain a hydrocracking catalyst, designated as C-1.

[0078] The hydrocracking catalyst comprises 21% by weight of tungsten oxide and 79% by weight of the carrier. The weight ratio of tungsten oxide to ZSM-5 molecular sieve is 0.35:1. Based on the total amount of the carrier, the ZSM-5 molecular sieve content in the carrier is 75% by weight and the aluminum oxide content is 25% by weight.

[0079] Preparation Example 2

[0080] A hydrocracking catalyst was prepared according to the method of Preparation Example 1, except that the amounts of the raw materials were different. The rest was the same as Preparation Example 1, to obtain a hydrocracking catalyst, which was designated as C-2.

[0081] The hydrocracking catalyst comprises 10% by weight of tungsten oxide and 90% by weight of the carrier. The weight ratio of tungsten oxide to ZSM-5 molecular sieve is 0.25:1. Based on the total amount of the carrier, the ZSM-5 molecular sieve content in the carrier is 45% by weight and the aluminum oxide content is 55% by weight.

[0082] Comparative Preparation Example 1

[0083] A hydrocracking catalyst was prepared according to the method of Preparation Example 1, except that the amount of ammonium metatungstate was such that the weight ratio of the Group VIB metal component (tungsten oxide) to the MFI molecular sieve (ZSM-5 molecular sieve) in the prepared hydrocracking catalyst, calculated as oxide, was 0.1:1. The types, amounts, steps and conditions of the remaining raw materials were the same as those in Preparation Example 1, to obtain a hydrocracking catalyst, which was designated as R-1.

[0084] Comparative Preparation Example 2

[0085] A hydrocracking catalyst was prepared according to the method of Preparation Example 1, except that the impregnation solution also contained nickel nitrate. The rest was the same as Preparation Example 1 to obtain a hydrocracking catalyst, which was designated as R-2.

[0086] In the hydrocracking catalyst, the weight ratio of nickel oxide to tungsten oxide is 0.25.

[0087] Examples 1-2 are used to illustrate the method for producing benzene-rich naphtha

[0088] Example 1

[0089] The hydrogenation reaction zone was sequentially loaded with hydrofining catalyst HDN-1 and hydrocracking catalyst C-1. The volume ratio of hydrofining catalyst to hydrocracking catalyst was 0.428. Catalytic diesel fuel 1 (feedstock properties are shown in Table 1) was mixed with hydrogen and introduced into the hydrogenation reaction zone. Specific reaction conditions and product properties are shown in Table 2.

[0090] Example 2

[0091] The method of Example 1 was followed, except that hydrocracking catalyst C-1 was replaced with hydrocracking catalyst C-2, and the feed oil was replaced with catalytic diesel 2 from catalytic diesel 1. Specific reaction conditions and product properties are shown in Table 2.

[0092] Comparative Example 1

[0093] The method of Example 1 was followed, except that hydrocracking catalyst C-1 was replaced with hydrocracking catalyst R-1. Specific reaction conditions and product properties are shown in Table 2.

[0094] Comparative Example 2

[0095] The method of Example 1 was followed, except that the hydrocracking catalyst C-1 was replaced with the hydrocracking catalyst R-2. Specific reaction conditions and product properties are shown in Table 2.

[0096] Comparative Example 3

[0097] The method of Example 1 was followed, except that catalytic diesel 1 was replaced with a comparative feedstock oil (properties of which are shown in Table 1). Specific reaction conditions and product properties are shown in Table 2.

[0098] Comparative Example 4

[0099] The method of Example 1 was followed, except for the reaction conditions in the hydrocracking stage. The specific conditions and product properties are shown in Table 2.

[0100] Table 1

[0101] Raw oil name Catalytic diesel 1 Catalytic diesel 2 Comparison of crude oil <![CDATA[Density (20 °C), g / cm 3 > 0.966 0.944 0.890 S / ％ 0.748 1.67 1.73 N / (μg / g) 556 582 432 Distillation range (D-1160), ℃ 217-344 185-370 132-365 Mass fraction of paraffins / % 12 11 32

[0102] Table 2

[0103]

[0104] Table 2 (continued)

[0105]

[0106] The above results show that the yield of aromatics in the naphtha fraction obtained by the method provided by the present invention is improved, especially the benzene yield in the light naphtha fraction (C5-C7) reaches 17.2%, which is more than 50% higher than the selectivity of 10.7% in the comparative example, showing a significant improvement.

[0107] By comparing the results of Example 1 and Comparative Examples 1-2, it can be seen that the weight ratio of the VIB Group metal component, calculated as oxide, to the MFI molecular sieve in the hydrocracking catalyst used in the method of Comparative Example 1 is outside the range defined by the present invention, or the hydrocracking catalyst used in the method of Comparative Example 2 is outside the range defined by the present invention, both of which cannot increase the yield of aromatics in the obtained naphtha fraction, especially the yield of benzene in the light naphtha fraction (C5-C7).

[0108] Comparing the results of Example 1 and Comparative Examples 3-4, it can be seen that using feedstock oil outside the range specified by the present invention, or using the conditions of the hydrocracking reaction (i.e., the second contact reaction) outside the range specified by the present invention, also cannot increase the yield of aromatics in the obtained naphtha fraction, especially the yield of benzene in the light naphtha fraction (C5-C7).

[0109] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. A method for producing benzene-rich naphtha, characterized in that: The method comprises the following steps: (1) subjecting hydrogen and heavy diesel to a first contact reaction with a hydrotreating catalyst to obtain a hydrotreating effluent, wherein the content of organic nitrogen in the hydrotreating effluent is 5-50 μg / g; (2) subjecting the hydrotreated effluent to a second contact reaction with a hydrocracking catalyst to obtain a hydrocracking effluent, wherein the hydrocracking catalyst is a carrier and a Group VIB metal component supported on the carrier, wherein the carrier contains an MFI molecular sieve; the weight ratio of the Group VIB metal component to the MFI molecular sieve, calculated as oxide, is 0.15-1:1; the conditions of the second contact reaction include: a reaction pressure of 5.0-7.5 MPa; a reaction temperature of 380-420°C; and a hydrogen-to-oil volume ratio of 500-1200:1; (3) fractionating the hydrocracking effluent to obtain a naphtha fraction; the distillation range of the naphtha fraction is 45-85° C.; and the yield of benzene in the naphtha fraction is greater than 14% by weight; The density of the heavy diesel at 20°C is 0.9-1.05 g / cm 3 , the distillation range is 180-380° C., and the content of paraffins in the heavy diesel is less than 20% by weight.

2. The method according to claim 1, wherein The content of the Group VIB metal component is 5-30% by weight based on the total dry weight of the hydrocracking catalyst and calculated as oxide; and / or The content of the carrier is 70-95% by weight based on the total dry weight of the hydrocracking catalyst.

3. The method according to claim 2, wherein: The content of the Group VIB metal component is 10-25% by weight based on the total dry weight of the hydrocracking catalyst and calculated as oxide; and / or The content of the carrier is 75-90 wt % based on the total dry weight of the hydrocracking catalyst and calculated on a dry basis.

4. The method according to claim 1, wherein In the hydrocracking catalyst, the weight ratio of the VIB Group metal component to the MFI molecular sieve, calculated as oxide, is 0.2-0.5:

1.

5. The method according to claim 1, wherein The VIB Group metal component is Mo and / or W; and / or The MFI molecular sieve is selected from ZSM-5 molecular sieve and / or ZRP molecular sieve; and / or The MFI molecular sieve has a silicon-aluminum ratio of 15-300 and a specific surface area of ​​180-650 m 2 / g; pore volume is 0.1-0.6 mL / g.

6. The method according to claim 5, wherein: The Group VIB metal component is W; and / or The MFI molecular sieve is a ZSM-5 molecular sieve; and / or The MFI molecular sieve has a silicon-aluminum ratio of 20-80 and a specific surface area of ​​300-450 m 2 / g; pore volume is 0.2-0.5 mL / g.

7. The method according to claim 1, wherein The carrier further contains a heat-resistant inorganic oxide; based on the total amount of the carrier, the content of the MFI molecular sieve is 40-90% by weight; and the content of the heat-resistant inorganic oxide is 10-60% by weight.

8. The method according to claim 7, wherein: Based on the total amount of the carrier, the content of the MFI molecular sieve is 50-80% by weight; the content of the heat-resistant inorganic oxide is 20-50% by weight.

9. The method according to claim 7, wherein: The heat-resistant inorganic oxide is selected from at least one of silicon oxide, aluminum oxide and titanium oxide.

10. The method according to claim 9, wherein: The heat-resistant inorganic oxide is silicon oxide and / or aluminum oxide.

11. The method according to claim 10, wherein: The heat-resistant inorganic oxide is aluminum oxide.

12. The method according to claim 11, wherein The specific surface area of ​​the alumina is 150-400 m 2 / g; pore volume is 0.45-1.3 mL / g.

13. The method according to claim 12, wherein: The specific surface area of ​​the alumina is 200-350 m 2 / g; pore volume is 0.9-1.2 mL / g.

14. The method according to any one of claims 1 to 13, wherein: The method for preparing the hydrocracking catalyst comprises the following steps: (a) mixing an MFI molecular sieve with an optional heat-resistant inorganic oxide to prepare a support; (b) introducing the Group VIB metal component onto the support by impregnation.

15. The method according to any one of claims 1 to 13, wherein: In step (1), the density of the heavy diesel at 20°C is 0.92-0.97 g / cm 3 .

16. The method according to any one of claims 1 to 13, wherein: In step (2), the liquid hourly space velocity of the second contact reaction is 0.5-3.0 h -1 .

17. The method according to any one of claims 1 to 13, wherein: The hydrotreating catalyst contains a carrier and a Group VIB metal component and a Group VIII metal component supported on the carrier.

18. The method according to claim 17, wherein Based on the total amount of the hydrorefining catalyst and calculated as oxide, the content of the Group VIII metal component is 1-10 wt %, and the content of the Group VIB metal component is 5-50 wt %.

19. The method according to claim 17, wherein In the hydrotreating catalyst, the Group VIB metal is Mo; and the Group VIII metal component is Ni.

20. The method according to claim 17, wherein In the hydrorefining catalyst, the carrier is selected from alumina and / or silicon oxide.

21. The method according to any one of claims 1 to 13, wherein: The volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 0.2-3; and / or In step (3), the products obtained by fractionating the hydrocracking effluent also include a gas component and / or a diesel fraction.

22. The method according to claim 21, wherein The volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 0.4-1.5.

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