COMPLETE CONVERSION METHOD AND DEVICE FOR PRODUCING LIGHT AROMATIC HYDROCARBONS FROM LIGHT CYCLE OIL

Abstract

The present invention relates to a total conversion method and a device for producing light aromatic hydrocarbons from catalytic diesel; the technical solution of the present invention comprises: subjecting a stream of catalytic diesel material to hydrogenation refining and impurity separation thereof, then subjecting the same to a selective conversion reaction and subjecting the resulting mixed aromatic hydrocarbons to separation so that light aromatic hydrocarbons such as benzene-toluene and xylene, and a C9A aromatic hydrocarbon, a C10A aromatic hydrocarbon and a heavy tailing oil at the bottom of a tower are separated sequentially;and the heavy tailings oil at the bottom of the tower enters a selective post-saturation reactor, subjecting it to a highly selective hydrogenation saturation under low temperature and low pressure conditions to obtain a product having a benzene ring, and then returning the same to a selective conversion reactor; the present invention achieves the complete conversion of fractions to produce light aromatic hydrocarbons from a catalytic diesel, and has the technical effect of a relatively high yield of monocyclic aromatic hydrocarbons, such as benzene-toluene, xylene, C9A aromatic hydrocarbon, and a C10A aromatic hydrocarbon;

Description

COMPLETE CONVERSION METHOD AND DEVICE FOR PRODUCING LIGHT AROMATIC HYDROCARBONS FROM LIGHT CYCLE OIL TECHNICAL FIELD The invention relates to the production of light aromatic hydrocarbons in the field of catalytic petroleum cracking, in particular to a process and a device for producing light aromatic hydrocarbons from light cycle oil. BACKGROUND OF THE INVENTION Light aromatic hydrocarbons, such as benzene, toluene, xylene and the like, are important basic organic chemical raw materials, widely used in the fields of synthetic materials and the like, and closely related to economic development and people's daily lives. At present, the aromatic hydrocarbon source mainly comprises two processing routes: first, naphtha undergoes catalytic reforming and aromatic extraction to provide an aromatic hydrocarbon feedstock; and secondly, the by-product of the ethylene process, mainly pyrolysis gasoline is hydrogenated and extracted to provide an aromatic hydrocarbon feedstock. Light cycle oil (LCO) mainly comprises Cn+alkylbenzene and polycyclic aromatic hydrocarbon. Due to the presence of a large amount of polycyclic aromatic hydrocarbon, the diesel processed by LCO is poor in economy and can only be used as gasoline by some companies. Along with the slowing increase in diesel demand, the development of efficient conversion technology is urgently needed, by which LCO is converted into light aromatic hydrocarbon through hydrocracking reaction, and achieving reducing costs and improving the efficiency of the aromatic hydrocarbon industry through refining and conversion. Currently, widely adopted LCO upgrading media include hydrofining, hydroenhancing, and light oil-type hydrocracking. Diesel hydrofining comprises saturation of olefin hydrogenation, desulfurization, denitrification and partial saturation of aromatic hydrocarbon under medium or low pressure conditions to improve its color and stability. However, with respect to the LCO obtained from a catalytic device for processing inferior raw materials, hydrofining is insufficient to meet the product requirements in cetane number. Hydro-upgrading processes, such as the UOP unicracking process (US 5026472), target diesel with high cetane number. The process has good performance of aromatic hydrocarbon hydrogenation saturation and aperture selectivity. MA / Ί I OI ring, high conversion of aromatic hydrocarbons and the greatest improvement of cetane number and higher diesel performance. Light oil type hydrocracking serves to provide reformed naphtha fraction or gasoline fraction by refining light diesel components and then carrying out violent saturation hydrogenation, which process also has the problem of low feedstock conversion efficiency into aromatic hydrocarbons. If the naphtha fraction is used to reform aromatics, the cycloalkanes and paraffins produced after super-saturation need to be converted to aromatics in the reformer, which is not an economical route. The light oil hydrocracking process as described in CN 101684415 does not directly produce aromatics and the maximum potential aromatics content of heavy naphtha is only 57%. CN1955262A describes a two-stage hydrocracking process, the hydrocracking catalyst of which contains Pt and / Pd noble metals and non-noble metals, as well as Y zeolite and alumina, and the feedstock is LCO. However, the maximum potential aromatic content of naphtha products is only 76.8%, and the purity of aromatic hydrocarbon is not high, which cannot meet the requirements of an integrated aromatic hydrocarbon unit. CN103897731A describes a process for producing light aromatics by mixing catalytic cracking diesel and Cio+ distillate oil, where the product is cut through hydrofining and hydrocracking, the fraction greater than 195 °C is used as a mixing component For clean diesel, the fraction less than 195 °C is fed into an aromatics plant to produce light aromatics and clean gasoline blending components, and the yield of aromatics products is relatively low. In the existing LCO conversion technology, the heavy tail oil is discharged as a diesel component or particularly recycled to the hydrofining reactor and cannot be effectively and completely utilized to increase the yield of light aromatic hydrocarbons. Furthermore, the hydrofining reaction on the metal sulfide type hydrofining catalyst must be carried out under severe operating conditions involving high temperature and high pressure, where the reaction is limited by the thermodynamic equilibrium, the selectivity of the saturation reaction Partiality of polycyclic aromatic hydrocarbon is poor, and the retention rate of aromatic hydrocarbon after hydrofining of LCO is less than 90%. The content of polycyclic aromatic hydrocarbons in heavy chemical wood pulping byproducts from the production of light aromatic hydrocarbons by diesel is greater than 90%, the sulfur and nitrogen content is low, and the recycling back to a reactor Hydrofining can cause problems of supersaturation and loss of aromatic hydrocarbons. MA / t / ZUZZ / UÓΊ l OI BRIEF DESCRIPTION OF THE INVENTION Aiming at the problems of the prior art, the inventors carry out a series of investigations and find that an LCO stream can be subjected to hydrofining and impurity separation, followed by a selective conversion reaction including hydrocracking, the generated mixed aromatic hydrocarbons. can be separated sequentially into benzene-toluene, xylene, a stream containing C9 aromatic hydrocarbons and Cío aromatic hydrocarbons, and lower heavy tail oil rich in C10+ heavy aromatics; wherein the bottom heavy bottom oil is fed to a post-saturation selective reactor and subjected to high selectivity saturation hydrogenation under low temperature and low pressure conditions to provide a product having a benzene ring, which is then is fed to a selective conversion reaction, to achieve full conversion of light aromatic hydrocarbon production from LCO, resulting in high yields of light aromatic hydrocarbons. A light aromatic hydrocarbon according to the present invention refers to an aromatic hydrocarbon having a carbon number less than or equal to 10, comprising C6 aromatic hydrocarbons, such as benzene; C7 aromatic hydrocarbons, such as toluene; C8 aromatic hydrocarbons, such as ethylbenzene, xylenes; C9 aromatic hydrocarbons, such as methylethylbenzene, propylbenzene, trimethylbenzene; CIO aromatic hydrocarbons, such as tetramethylbenzene, dimethylethylbenzene, diethylbenzene and the like. Correspondingly, C10+ heavy aromatics refer to an aromatic hydrocarbon having a carbon number greater than 10. One of the objectives of the present invention is to provide a total conversion process for producing light aromatics from LCO. The total conversion process to produce light aromatic hydrocarbons from LCO of the invention comprises the steps of: 1) feeding LCO into a first reaction zone for hydrofining, to provide a first stream; 2) feeding the first stream to a second reaction zone for selective conversion to provide a second stream, wherein the first stream is optionally subjected to separation of impurities in a second separation zone before being fed to the second reaction zone; 3) subjecting the second stream to a first separation in a first separation zone, to provide a third stream comprising C10+ heavy aromatics at the bottom of the first separation zone; 4) feed the third stream to a selective reaction zone after the MA / t / ZUZZ / UÓΊ ÍOÍ saturation for saturation by hydrogenation, to provide a fourth stream; 5) recycle the fourth stream to the second reaction zone. In an exemplary embodiment, the present invention provides a total conversion process for producing LCO fuel light aromatics comprising the steps of: 1) feeding the LCO into a first reaction zone and contacting a hydrofining catalyst in the presence of hydrogen to provide a first stream; the first reaction zone is subjected to a hydrofining reaction; 2) feeding the first stream into a second reaction zone after removal of impurities, and contacting a selective conversion catalyst in the presence of hydrogen to provide a second stream; wherein the selective conversion comprises a hydrocracking reaction; 3) separating the second stream to provide fractions that include a Ce-Cs aromatic hydrocarbon stream, a stream containing Cg aromatic hydrocarbons and Cío aromatic hydrocarbons, and a third stream containing Cw+ heavy aromatics; 4) feeding the third stream into a post-saturation selective reaction zone and contacting a post-saturation selective catalyst in the presence of hydrogen, to provide a fourth stream; wherein post-saturation selection is carried out by using a hydrogenation saturation reaction; and 5) recycle the fourth stream to the second reaction zone. According to one aspect of the invention: in step 1) of the process of the invention, the LCO as the crude oil is subjected to hydrofining in a first reaction zone in the presence of hydrogen, where the LCO stream and the hydrogen They are contacted with a hydrofining catalyst to perform desulfurization and denitrification, and a selective saturation reaction of polycyclic aromatic hydrocarbons resulting in a retained aromatic ring. Hydrofining can be carried out in any manner by any process conventionally known in the art, and is not particularly limited, as long as the LCO fuel is subjected to desulfurization and denitrification, and the polycyclic aromatic hydrocarbon therein is subjected to saturation. by hydrogenation to retain an aromatic ring. The first stream obtained after hydrofining the LCO mainly comprises refined LCO with most of the sulfur and nitrogen impurities removed, and a gas phase containing hydrogen sulfide and ammonia. In step 1) of the process of the invention, LCO and hydrogen as crude oil are contacted with a hydrofining catalyst in a first reaction zone to carry out the hydrofining reaction. The hydrofining reaction is an LCO hydrofining technique well known in the art. The hydrofining reaction can be carried out under reaction conditions for LCO hydrofining known in the art; wherein the hydrofining catalyst can use any type of hydrofining catalyst existing in the field as long as the purpose of hydrofining the LCO in step 1) can be met. In step 1) of the process of the present invention, the conditions for the hydrofining reaction of the first reaction zone preferably comprises: a volume ratio of hydrogen to oil of 500-3000 Nm3 / m3, preferably 800-2000 Nm3 / m3, and more preferably 1000-1500 Nm3 / m3; a reactor inlet temperature of 280-420°C, preferably 300-410°C, and more preferably 310-390°C; a hydrogen partial pressure of 5-10 MPa, preferably 5-8 MPa, and more preferably 6-7 MPa; and / or a space speed of 0.5-2.0 h1, preferably 0.6-1.5 h1, more preferably 0.8-1.2 h'1. In the process of the present invention, the hydrofining catalyst of step 1) can preferably: comprising the following components in parts by weight: a) 60-99.9 parts, preferably 65-99.9 parts, preferably 70-99.9 parts, more preferably 75-99.9 parts of a support; and bl) a hydrogenating metal oxide, in an amount of 0.1 to 40 parts, preferably 0.1 to 35 parts, preferably 0.1 to 30 parts, and more preferably 0.1 to 25 parts; based on the weight of the support and the hydrogenating metal oxide. In an exemplary embodiment, the support comprises, in parts by weight: 60-100 parts of alumina; and 0-40 parts of silica; based on the total weight of alumina and silica. In an exemplary embodiment, the hydrogenating metal is at least one selected from the group consisting of nickel, cobalt, molybdenum, tungsten, and iron. The hydrogenating metal is sulfidized before being supported. The hydrofining catalyst of the present invention can be prepared by any process known in the art, for example, the support can be prepared by extrusion, lamination or oil column formation known in the art. In one embodiment, the catalyst can be prepared by forming the support and then impregnating the metal. Preferably, the first stream obtained by hydrofining in step 1) is subjected to separation of impurities and then the impurities such as hydrogen sulfide and ammonia contained in the first stream are separated, the first stream separated from the impurities is separated and is fed to the second reaction zone. The separation of impurities preferably comprises gas-liquid separation and extraction of hydrogen sulfide to ΜΛ / t / ZUZZ / UÓΊ I OI provide the first separated impurity stream from which impurities such as hydrogen sulfide and ammonia are separated. More specifically, separation techniques conventional in the art may be employed, such as gas-liquid separation by injecting gas phase water to remove ammonia, liquid phase extraction to remove hydrogen sulfide, and the like. According to one aspect of the invention: in step 2) of the process of the present invention, the first stream after removal of impurities is subjected to selective conversion in a second reaction zone in the presence of hydrogen by reactions including hydrocracking . For example, selective conversion includes a hydrocracking reaction that selectively converts a first stream obtained after hydrofining to a second stream. The second stream obtained in step 2) mainly comprises dry gases (including methane and ethane), light C3-C5 hydrocarbons, a benzene-toluene fraction, a xylene fraction, a C9-C10 fraction and a heavy tail oil. . One purpose of the selective conversion in step 2) is to perform hydrocracking with the condition of retaining an aromatic ring of the polycyclic aromatic hydrocarbons in the heavy aromatics of the first stream, effectively controlling the degree of saturation and the opening position of the ring, and simultaneously allow the isomerization and cracking of the macromolecular non-aromatic hydrocarbons of the first stream; to maximize the production of light aromatics with economical hydrogen consumption. The selective conversion reaction of this step can be carried out according to any process known in the art for hydrogenation reactions, provided that the first stream is selectively converted to the second stream. In the process of the present invention, the reaction conditions for the second reaction zone of step 2) can use the conventional reaction conditions for field hydrocracking reaction. In the present invention, the reaction conditions for the second reaction zone preferably comprise: a volume ratio of hydrogen to oil of 800-5000 Nm3 / m3, preferably 1000-4000 Nm3 / m3, and more preferably 1500-3000 Nm3 / m3; a reactor inlet temperature of 280-450°C, preferably 300-430°C, and more preferably 310-400°C; a hydrogen partial pressure of 5-10 MPa, preferably 5-9 MPa, and more preferably 6-8 MPa; and / or a space rate of 0.5-2.0 h'1, preferably 0.6-1.5 h'1, more preferably 0.8-1.2 h'1. The selective conversion catalyst in step 2) can be any type of MA / t / ZUZZ / UÓΊ fOÍ hydrocracking catalyst known in the art, provided that the above purpose of step 2) is achieved. In order to achieve the conversion of the first stream to the second stream effectively, the selective conversion catalyst of the invention is preferably a catalyst provided in Chinese patent application ZL 201810153543.5. The contents of Chinese patent application ZL 201810153543.5 are incorporated by reference herein in their entirety. A preferred selective conversion catalyst comprises: in parts by weight: a2) 5-80 parts of solid acid zeolite; b2) 0.05 to 8 parts of a Group VIII metal; c2) 3-25 parts of a Group VIB metal oxide; d2) 0.1-2 parts of a Group VIB metal sulfide; e2) 20-95 parts of a first binder; The parts by weight of the components are based on the total part by weight of the catalyst. The selective conversion catalyst of the present invention may comprise, in addition to the above main components, other auxiliary agents commonly used for catalysts in the art, such as diatomaceous earth, activated clay and the like. The useful amount may be a conventional amount. Preferably, the solid acid zeolite is at least one of mordenite, β-zeolite, ZSM zeolite, EU-1 zeolite, SAPO zeolite and Y zeolite. Preferably, the solid acidic zeolite has a crystallite diameter of less than 500 nm, preferably less than 400 nm, more preferably less than 300 nm, more preferably less than 200 nm. Preferably, the solid acid zeolite has a silicon-aluminum molecular ratio of 10-500, preferably 10-200, more preferably 11-80, and more preferably 20-60. Preferably, the Group VIII metal is at least one of platinum, palladium, cobalt, nickel and iridium. Preferably, the Group VIB metal oxide is at least one of molybdenum oxide and tungsten oxide. Preferably, the Group VIB metal sulfide is at least one of molybdenum sulfide and tungsten sulfide. Preferably, the first binder is at least one of alumina, a silica-alumina compound, a titania-alumina compound and a magnesia-alumina compound. The selective conversion catalyst of the present invention can be prepared by any method known in the art, for example, the support can be prepared by extrusion, lamination or oil column formation and the like. In one embodiment, the catalyst can be prepared by forming the support and then impregnating the metal. In one embodiment, the catalyst MA / Ί I OI selective conversion can be prepared by a process comprising the steps of: mixing the solid acid zeolite with a first adhesive, then kneading, extruding, baking at 60-150 °C, and calcining in an atmosphere of air at 500-600 °C for 3-6 hours, to provide the necessary catalyst support; prepare a composite metal aqueous solution of a Group VIII metal compound and a Group VIB metal compound, impregnate the catalyst support by the solution through a isometric impregnation method, bake at 60-150 °C, and calcining in an air atmosphere at 450-520°C for 1-4 hours, to provide a catalyst precursor; and reduce the catalyst precursor to 400-500 °C in the presence of hydrogen, and maintain the temperature for 2-24 hours (pre-reduction), followed by cooling to 300-380 °C, inject a vulcanizing agent for vulcanization for 4-24 hours, to provide the required hydrocracking catalyst. According to one aspect of the invention, in step 3) of the process of the invention, the second stream is subjected to a first separation in a first separation zone, and the Ce-Cs aromatic stream obtained comprises at least fractions of benzene, toluene, xylene, etc. For the process according to the invention, in step 3), the first separation of the second stream preferably comprises gas-liquid separation and rectification of the second stream. More preferably, the benzene-toluene fraction obtained after rectification is subjected to extraction separation. Specifically, the second stream is subjected to gas-liquid separation to separate a dry gas and a liquid phase, where the dry gas is discharged to the outside, and the liquid phase is fed to a depentanizer for depentanization. Depentanization separates a light fraction of C3-C5 hydrocarbons for discharge to the outside and a depentanizer bottom stream, which depentanizer bottom stream is fed to a deheptanizer. The deheptanizer separates a stream rich in a benzene-toluene fraction and a lower deheptanizer stream, which lower deheptanizer stream is fed to a xylene column. The xylene column separates a mixed xylene product and a xylene-stripped bottom stream, which xylene-stripped bottom stream is subjected to heavy aromatic removal. Removal of heavy aromatics separates a C9-C10 stream for discharge to the outside and a third stream separated from the bottom. The third stream is a heavy tail oil containing the Cio+ heavy aromatics. The heavy tail oil is fed to a selective post-saturation reactor. The deheptanizer separates a stream rich in benzene-toluene fraction, which stream is preferably subjected to extraction to separate the mixed aromatic hydrocarbon from pure benzene-toluene, while the non-aromatic hydrocarbons IVIA / t / ZUZZ / UÓ I ÍOÍ separated by extraction are discharged outside. The above-described gas-liquid separation and rectification can be carried out by extraction and rectification methods commonly used in the art. The aromatic content in the third stream obtained by separating the second stream obtained by the selective conversion is preferably higher than that of the non-aromatic content; wherein the third stream of the present invention more preferably has an aromatic content of 80% by weight or more, more preferably 90% by weight or more. According to one aspect of the invention, in step 4) of the process of the invention, the third stream containing Cio+ heavy aromatics obtained in step 3) is subjected to a high selectivity hydrogenation saturation reaction in a zone of selective post-saturation reaction in the presence of hydrogen at low temperature and low pressure, to provide a product having a benzene ring, forming a fourth stream containing the product, mainly, a fraction with a distillation point of more than 210°C. Saturation hydrogenation can be carried out according to any conventionally known method known in the art, as long as the effect of the selective post-saturation reaction described above can be achieved. The saturation hydrogenation in the post-saturation selective reaction zone in process step 4) is preferably a liquid hydrogenation reaction, to simplify the flow sheet, reduce equipment and reduce energy consumption. The reaction conditions may be those for conventional hydrogenation saturation reactions in the art, and preferably comprise: a volume ratio of hydrogen to oil of 200-3000 Nm3 / m3, preferably 300-1500 Nm3 / m3, and more preferably 300-1000 Nm3 / m3; a reactor inlet temperature of 100-300°C, preferably 120-280°C, and more preferably 150-250°C; a hydrogen partial pressure of 1.0-4.0 MPa, preferably 1.2-3.0 MPa; and / or a space rate of 0.1-5.0 h'1, preferably 0.5-4.0 h'1, more preferably 0.6-2.0 h'1. In step 4), the third stream is contacted with the post-saturation selective catalyst in the post-saturation selective reaction zone to perform a hydrogenation saturation reaction, wherein the post-saturation selective catalyst can be an existing hydrogenation saturation catalyst in the art, provided that the purpose of hydrogenation saturation in step 4 is achieved). For example, the catalyst for saturation hydrogenation of aromatic hydrocarbons described in CN 103041832A can be used. ΜΛ / t / ZUZZ / UÓΊ I OI The selective post-saturation catalyst of step 4) according to the invention may comprise: in parts by weight: a3) 10-90 parts of amorphous silica-alumina, where the silica content is 3-20% by weight; b3) 0.1 to 5.0 parts of a Group VIII metal; c3) 5-80 parts of a second binder; based on the total weight of the amorphous silica-alumina, the Group VIII metal and the second binder. In one embodiment, the Group VIII metal is at least one selected from the group consisting of platinum, palladium, cobalt, nickel and iridium. In one embodiment, the second binder is selected from alumina. The selective post-saturation catalyst of the present invention can be prepared by any method known in the art. For example, the support may be prepared by extrusion, lamination or oil column formation and the like in the art. In one embodiment, the catalyst can be prepared by forming the support and then impregnating the metal. For the total conversion process to produce light aromatic hydrocarbons from LCO, the LCO serving as crude oil can come from a catalytic cracking device in the field, which has an initial boiling point of 160-210 °C. The composition of LCO is not particularly limited and LCO can be derived from crude oils from different places of origin and have different compositions. However, by way of example, LCO fuel predominantly contains paraffinics, cycloalkanes, olefins, sulfur-containing hydrocarbons, nitrogen-containing hydrocarbons, Cn+ alkylbenzenes, and polycyclic aromatic hydrocarbons. Among others, the content of Cn+ alkylbenzenes is 10-40% by weight, the content of polycyclic aromatic hydrocarbon is 15-50% by weight, the content of sulfur is 200-15000 ppm by weight, the content of nitrogen is 100- 1500 ppm by weight, and the rest are high boiling alkanes, cycloalkanes and olefins. Another purpose of the invention is to provide a device for the total conversion process to produce light aromatic hydrocarbons from LCO. The invention provides a device for producing light aromatic hydrocarbons from LCO, comprising: a first reaction zone for hydrofining; configured to receive the LCO and to discharge a first stream; a second reaction zone for selective conversion (including hydrocracking); configured to receive the first stream and to discharge a second stream; a first separation zone; configured to receive a second stream; and to discharge the third stream at the bottom; a selective post-saturation reaction zone for saturation by MA / t / ZUZZ / UÓΊ I OI hydrogenation; configured to receive the third stream and to discharge a fourth stream; and a first pipe; configured to recycle the fourth stream to the second reaction zone. In one embodiment, the device for producing light aromatics from LCO fuel comprising: a first reaction zone; configured to receive the LCO and to discharge a first stream; a second reaction zone; configured to receive the first stream and to discharge a second stream; a first separation zone; configured to receive the second stream; and to discharge fractions comprising the Ce-Cs aromatic hydrocarbon stream, a stream containing C9 aromatic hydrocarbons and Cío aromatic hydrocarbons and a third stream containing C10+ heavy aromatics; a selective post-saturation reaction zone; configured to receive the third stream and to discharge a fourth stream; a first pipe; configured to recycle the fourth stream to the second reaction zone. Specifically, in one embodiment: For one embodiment of the device of the present invention, the first reaction zone is equipped with a hydrofining unit where the hydrofining reactor used is a fixed bed reaction system. Specifically, a fixed bed reaction system known in the art can be used, and a fixed bed reaction system equipped with a hydrogen recycling system is more preferable. The hydrofining reactor may have an inlet temperature of 250-450 °C. For one embodiment of the device of the present invention, the second reaction zone is equipped with a hydrocracking reactor for selective conversion, where the hydrocracking reactor used is a fixed bed reaction system. Specifically, a fixed bed reaction system known in the art can be used, and a fixed bed reaction system equipped with a hydrogen recycling system is more preferable. The selective conversion (hydrocracking) reactor can have an inlet temperature of 280-450 °C. For one embodiment of the device of the present invention, the post-saturation selective reaction zone is equipped with a hydrosaturation unit, wherein the post-saturation selective reactor used is a fixed bed reaction system; more preferably a liquid hydrogenation fixed bed reaction system without a hydrogen recycling system. Specifically, a fixed bed reaction system can be used ΜΛ / Ί I OI known in the art. Post-saturation (hydrosaturation) can have an inlet temperature between 100°C and 300°C, and a hydrogen partial pressure between 1.0 and 4.0 MPa. For one embodiment of the device of the present invention, the first separation zone comprises a gas-liquid separator and a rectification column, optionally connected in sequence, where the rectification column preferably comprises a depentanizer (first rectification column), a deheptanizer (second rectification column), a xylene removal column (third rectification column) and a heavy aromatics removal column (fourth rectification column), optionally connected in sequence, to sequentially separate fractions comprising a stream rich in benzene-toluene fraction, a xylene stream, the stream contains Cg aromatic hydrocarbons and Cío aromatic hydrocarbons, and the third stream contains Cio+ heavy aromatics. Further preferably, the second stream is fed to a gas-liquid separator to separate a dry gas stream and a liquid phase stream, wherein the liquid phase stream is fed to a depentanizer to separate a C3 light hydrocarbon stream. -C5 at the top of the depentanizer and a lower depentanizer stream, which lower depentanizer stream is fed to a deheptanizer. The deheptanizer separates a stream rich in benzene-toluene fraction at the top and a bottom stream from the deheptanizer. Preferably, the stream rich in benzene-toluene fraction is fed to an extraction device, to separate mixed aromatic hydrocarbons from pure benzene-toluene, while non-aromatic hydrocarbons separated by extraction are poured out. The deheptanizer bottom stream is fed into a xylene column to directly separate a mixed xylene product and a xylene column bottom stream. The bottom stream of the xylene column is fed into a heavy aromatics column, from which a C9-C10 is separated at the top for discharge to the outside and a third stream is separated at the bottom. The third stream is fed to a selective post-saturation reactor. Extraction and rectification can be carried out with the extraction and rectification method commonly used in the field. The gas-liquid separator, rectification column and extraction device can be the conventional equipment in the field. In one embodiment of the device according to the invention, a second separation zone is placed between the first reaction zone and the second reaction zone to separate impurities comprising sulfides and / or nitrides in the first stream. The second separation zone is configured to receive the first stream and to discharge a gas phase, a hydrogen sulfide and an ammonia stream, and a separate first stream of impurities. The separation device of the second separation zone may be a separation device conventionally used in the field, such as a gas-liquid separator (equipped with washing ΜΛ / Ί fOÍ of ammonia by injection of water in gas phase), an extraction device (such as an extraction column of a hydrogen sulfide liquid phase extraction device) and the like. The process of the invention removes sulfur and nitrogen impurities in the LCO stream of a catalytic cracking device by hydrofining the LCO stream through a first reaction zone, and subjects the fused ring aromatic hydrocarbon and the polycyclic aromatic hydrocarbon to a selective hydrogenation saturation reaction, whereby hydrogenation is carried out until only products having an aromatic ring, such as tetrahydronaphthalene, indene and polyalkylbenzene, are retained. The process stream is then fed to a second reaction zone after impurities are optionally removed, for selective conversion hydrocracking reaction, to generate a stream rich in light aromatic hydrocarbons, such as benzene, toluene, xylene, aromatic hydrocarbons. from Cg, aromatic hydrocarbons from Cío, and the like. Then, after removing the light components before benzene, benzene-toluene, xylene, Cg aromatic hydrocarbons, Cío aromatic hydrocarbons and lower heavy tail oil (containing mainly heavy aromatics) are sequentially separated from the stream of the product. The bottom heavy bottom oil is fed to a selective post-saturation reactor, for high selectivity saturation hydrogenation under low temperature and low pressure conditions, to provide a product having a retained aromatic ring. The product was fed to a second reaction zone for a selective conversion hydrocracking reaction, to achieve a full conversion process to produce light aromatic hydrocarbons from LCO. Through the inventive process, the yield of light aromatic hydrocarbon is improved, the loss of aromatic hydrocarbon is reduced, and hydrogen consumption is reduced. The process solves the problems of the prior art well and achieves good technical effect to increase the yield of aromatic hydrocarbon products. In the technical solution of the invention, through hydrofining in the first reaction zone, the LCO stream has a polycyclic aromatic hydrocarbon saturation rate of more than 50%, a sulfur content reduced to less than 1OOppm, a of nitrogen reduced to less than 15ppm, and a final distillation point reduced by more than 10°C. The LCO stream is converted to monocyclic aromatic hydrocarbons with ten or less carbon atoms after sequentially subjecting the hydrofining device in the first reaction zone and the selective conversion device in the second reaction zone, with a conversion rate of More than 50%. The stream after the post-saturation selective reactor shows high hydrogenation saturation selectivity, with an aromatic hydrocarbon retention rate greater than 98%. ΜΛ / Ί l OI Compared with the prior art, the technical solution of the invention adopts a two-stage hydrofining-selective conversion process, with two catalysts used in two stages in series (hydrofining catalyst and selective conversion catalyst), comprising hydrofining, conversion selective and a process after the saturation of the heavy tail oil. By means of the invention, the technical problems of incomplete conversion of low-yield full-range LCO of light aromatic hydrocarbon in the conversion process in the prior art are solved. The heavy tail oil from the light aromatic hydrocarbon production by LCO is fed into an aftertreatment reactor, for a selective saturation reaction under mild pressure and temperature conditions. The selectivity of hydrogenation saturation is greatly improved to more than 98% or even more, so that the problem of excessive hydrogenation saturation is solved; which also helps reduce the hydrogen consumption of the cracking reaction that takes place when the non-aromatic hydrocarbon is generated by over-hydrogenation in the selective conversion reactor. The process improves the technical and economic indexes of the comprehensive process to produce light aromatic hydrocarbons by LCO, and realizes the complete fraction conversion of LCO. Compared with a conventional two-stage selective hydrofining-conversion process to produce light aromatic hydrocarbon by LCO, the yield of benzene, toluene, xylene and monocyclic light aromatics such as C9 and Cío can be improved by at least 2% or more, preferably 5% or more. BRIEF DESCRIPTION OF THE DRAWINGS Figure 1 is a schematic process flow diagram of the present invention for a total conversion process to produce LCO light aromatics. Description of reference numbers: indicates feedstock oil - LCO indicates a first reaction zone indicates an exit stream from the first reaction zone - first stream 4 indicates a gas-liquid separator indicates a gas phase stream containing hydrogen sulfide and ammonia 6 indicates a liquid phase stream after gas-liquid separation indicates an extraction column for hydrogen sulfide indicates an extracted stream containing hydrogen sulfide indicates the hydrofined LCO after removing hydrogen sulfide - the first stream with separated impurities indicates a selective conversion reactor MA / t / ZUZZ / UÓ 1 fOÍ indicates a product of the selective conversion reaction - the second stream indicates a first separation zone, comprising, for example, a gas-liquid separator, a depentanizer, a deheptanizer, a column of xylene, a column of heavy aromatics and similar rectification columns, and a device for extracting benzene-toluene fractions indicates a stream of dry gas and light C3-C5 hydrocarbons separated from the first separation zone indicates a stream of benzene -toluene separated from the first separation zone indicates a xylene stream separated from the first separation zone indicates a stream containing C9 aromatic hydrocarbons and Cío aromatic hydrocarbons separated from the first separation zone indicates a heavy tail oil stream , separated from the first separation zone - the third stream indicates a post-saturation selective reactor indicates an output stream of the post-saturation selective reactor - the fourth stream indicates a second separation zone (within the line of points). Modalities The present invention will be described in detail with reference to the Drawings and Examples, although it should be understood that the following Examples are merely illustrative of the present invention and should not be taken as limiting the scope of the present invention. Instead, those skilled in the art may realize that modifications and variations thereof that would occur to those skilled in the art when reading the present description are still covered by the scope of protection of the invention. All publications, patent applications, patents, and other references mentioned in this specification are incorporated herein by reference in their entirety. Unless specifically defined, all technical and scientific terms used herein have the same meaning as commonly understood by one skilled in the art to which this invention pertains. In case of conflict, this specification, including definitions, will control. When this specification mentions a material, substance, method, step, device or component, etc. with derivative words known to those skilled in the art, prior art or the like, the term derivative is intended to cover those used MA / t / ZUZZ / UÓΊ ÍOÍ conventionally in the field of the present application, but also cover those that are not currently known, while they will be known in the art to be useful for similar purposes. The end points of the scales and any values ​​described in the text of this application are not limited to the precise scale or value, but should be understood to encompass values ​​close to these scales or values. For numerical scales, each scale between its endpoints and individual point values, and each individual point value may be combined with each other to provide one or more new numerical scales, and such numerical scales should be construed as specifically described herein. The various technical solutions can then, in principle, be combined with each other to provide new technical solutions, which should also be considered specifically described herein. Preferred embodiments of the present invention have been described in detail; However, the present invention was not limited to the specific details of the above embodiments, and several simple modifications can be made to the technical solution of the present invention within the technical idea of ​​the present invention, and these simple modifications were all within the scope of protection of the present invention. It should be noted that the various features to be described in the embodiments below may be combined in any suitable manner and, to avoid unnecessary repetition, the present invention does not separately describe various possible combinations. Furthermore, any combination of the various embodiments of the present invention can be made, and the same should be considered as the content of the present invention as long as the idea of ​​the present invention is not violated. Unless otherwise specifically stated, reference to pressure in this specification indicates a gauge pressure. Unless otherwise specifically specified, reference to space velocity in this specification indicates a liquid hourly space velocity LHSV. Unless otherwise specifically stated, all percentages, shares, ratios, etc. involved in this specification are indicated by weight, unless the weight basis does not conform to conventional compression by those skilled in the art. Figure 1 is a schematic process flow diagram of an exemplary embodiment of the LCO fuel light aromatic hydrocarbon production process according to the present invention, wherein various conventional devices such as pumps, compressors, heat exchangers, devices extraction, hydrogen pipes, etc. are omitted, while such devices are well known to those skilled in the art. As shown in Figure 1, the flow of an exemplary embodiment of the process of the present invention is described in detail as follows: ΜΛ / Ί I OI The LCO 1 serving as feedstock oil is fed to a hydrofining device of the first reaction zone 2, to provide hydrofined LCO containing hydrogen sulfide and ammonia, mainly, the output stream of the first reaction zone 3 (the first stream); the first stream is fed to a gas-liquid separator 4 and a hydrogen sulfide extraction column 7 from the second separation zone 20, to separate hydrogen sulfide and ammonia (via a gas phase stream 5 containing hydrogen sulfide and ammonia, and an extraction stream 8 containing hydrogen sulfide) obtained by denitrification and desulfurization in the hydrofining process, to provide a first stream separated from the impurities 9. This stream is fed into the extraction device. selective conversion of the second reaction zone 10. An output stream of the second reaction zone 11 (the second stream), rich in light aromatic hydrocarbons, such as benzene, toluene and xylene, C9A C10A fractions, and bottom oil heavy is fed into a first separation zone 12, and separated to provide dry gas and light C3-C5 hydrocarbon stream 13, a benzene-toluene stream 14, a xylene stream 15, a hydrocarbon-containing stream 16 Cg aromatics in Cío aromatic hydrocarbons and a third stream 17 of a heavy tail oil containing Cio+ heavy aromatics. The third stream 17 is fed into the post-saturation selective reactor 18 of the post-saturation selective reaction zone, and the output stream of the post-saturation selective reactor 19 (the fourth stream), without separation, is recycled. to the selective conversion device of the second reaction zone 10. Specifically, the first separation zone 12 comprises a gas-liquid separator, a depentanizer, a deheptanizer, a xylene column, a heavy aromatics column and other rectification columns, and a benzene-toluene fraction extraction device ( not shown in the figure) connected in series. The analysis of the composition of the catalysts involved in the present invention is carried out by means of analytical methods known in the art. For example, the catalyst composition can be analyzed by ICP (inductively coupled plasma) and XRF (X-ray fluorescence) methods for selective conversion catalyst. The composition ratio of metal oxide and metal sulfide, Group VIB, is determined by XPS (X-ray photoelectron spectroscopy). ICP is measured using the Varian 700-ES series XPS instrument. XRF is measured using a Rigaku model ZSX lOOe XRF instrument. XPS test conditions comprise: a Perkin Elmer PHI 5000C ESCA model 4.0x10-8 Pa vacuum. In the invention, the familiar composition of LCO and hydrofined LCO, as well as the MA / t / ZUZZ / UÓΊ ÍOÍ family composition of heavy tail oil and selectively saturated heavy tail oil, is analyzed (multidimensional chromatographic analysis) using a two-dimensional gas chromatography / high-flow time-of-flight mass spectrometer (GCxGC -TOFMS) from the LECO company, America. In the present invention, the composition of the reactant stream (e.g., selectively converted product, etc.) is determined by gas chromatography. The chromatography model is Agilent 7890A, equipped with an FID detector, a FFAP capillary chromatographic column is used for separation. The chromatographic column is operated in a temperature-programmed model, with an initial temperature of 90°C, held for 15 minutes, then heated to 220°C at a rate of 15°C / minute and held for 45 minutes. . In the process of the invention, the retention rate of the aromatic hydrocarbon in the hydrofining and selective saturation (post-saturation) processes is calculated as follows: T T, _ , . , , .. aromatic content in the hydrofining or selective saturation product Vel. Aromatic retention = ---------------------------------------------- ---------------* aromatic content in the raw material for hydrofining or selective saturation 100% (i) The yield of monocyclic light aromatics such as benzene-toluene, xylene, C9A aromatics, C10A aromatics and the like is calculated as follows: BTX output + C9A + C10A Yield of BTX + C9A + C10A = ----——-------------* 100% quantity of raw materials The catalyst raw materials for the inventive and comparative examples were each commercially available. COMPARATIVE EXAMPLE 1 An LCO was processed by a two-step process of hydrofining-selective conversion, mainly, the LCO serving as feedstock oil was subjected to hydrofining and impurity separation followed by hydrocracking and then the hydrocracking product was subjected to separation gas-liquid and rectification systems, to provide products of benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, heavy tail oil and the like through separation. The process flow of Comparative Example 1 did not involve selective hydrosaturation of a heavy bottom oil at >210°C in a selective post-saturation reaction zone. The analytical data of the LCO feedstock and the hydrofined product are MA / 1 / Oí showed in Table 1, where the LCO had an aromatic content of 87.15% by weight. The hydrofining catalyst, selective conversion (hydrocracking) catalyst and reaction conditions used are listed in Table 2. TABLE 1 Element LCO raw materials Hydrofining product Density (4 °C) 0.953 0.932 Sulfur (ppm by weight) 1070 87 Nitrogen (ppm by weight) 632 8.6 Non-aromatic hydrocarbons (weight) 10.85 20.62 Monocyclic aromatic hydrocarbons (% by weight) 37.40 53.7 1 Polycyclic aromatic hydrocarbons (% by weight) 51.75 25.67 Distillation test (D-86) ° C ° C Initial boiling point 193 188 5% 212 210 10% 235 232 30% 246 237 50% 288 275 70% 315 313 ​​90% 345 337 Distillation end point 372 363 MA / t / ZUZZ / UÓΊ I OI TABLE 2 Hydrofining catalyst Al 3.0% by weightN¡O-10.5% by weightMoO3-12.7% by weightWO3 / 73.8% by weightAbOa Al' 3.1% by weightNiS-10.2% by weightMoS2-13.2% by weightWS2 / 73.5% by weightAbOs Partial pressure of hydrogen for hydrofining 6.5 MPa Hydrofining reaction temperature 315 °C at inlet Hydrofining LHSV space velocity 1.2 h-1 Hydrogen to hydrofining oil ratio 1500 (v / v) Selective conversion catalyst B1 0.1 part Pt / 60 parts of USY zeolite-39.9 parts of AI2O3 Partial pressure of hydrogen for selective conversion 7.0 MPa Reaction temperature of selective conversion 340 °C at inlet Space velocity of LHSV of selective conversion 1.0 h-1 Ratio of hydrogen to oil of selective conversion 1800 ( v / v) Preparation of the hydrofining catalyst Al used: 2 g of sesbania powder, 9 ml of nitric acid and 60 ml of water were added to 100 g of pseudoboehmite, kneaded into a cluster and extruded into strips, kept at room temperature for 24 hours, dried at 100 °C for hours and calcined at 550 °C in an air atmosphere for 3 hours, to provide a hydrofining catalyst support. 7.90 g of nickel nitrate hexahydrate, 8.71 g of ammonium molybdate, 9.18 g of ammonium metatungstate and 10 ml of aqueous ammonia were dissolved in water to provide 50 ml of a clear solution. 50g of the hydrofined catalyst support was added into said 50ml of soaking solution for 3 hours in an isovolumetric soaking mode, dried at a temperature of 110 °C for 12 hours and calcined at a temperature of 500 °C in an atmosphere of air for 4 hours, to provide an Al hydrofining catalyst. The Al catalyst comprised 3.0% by weight NiO-lO.5% by weight MoO3-12.7% by weight WO3 / 73.8% by weight AI2O3, mainly, comprising three metals of nickel, molybdenum and tungsten. A cyclohexane solution containing 0.5% carbon disulfide was injected into a fixed bed reactor with the hydrofining catalyst Al, heated at room temperature to a vulcanization end point temperature of 360 °C according to with the program of 10 °C / h, and was maintained for 12h to finish the prevulcanization of the hydrofining catalyst, to provide a vulcanized hydrofining catalyst ΑΓ, which comprised: 3.1% by weight NiS-10.2% by weight MoS2-13.2% by weight WS2 / 73.5% by weight AI2O3 where the Group VIB and Group VIII metals were present in the sulfide state. The LCO and hydrogen were mixed and then fed into a hydrofining reactor to remove most of the sulfur and nitrogen impurities therein, where the polycyclic aromatic hydrocarbons were saturated to become hydrocarbons containing an aromatic ring. Table 1 also lists the sulfur and nitrogen content, density, aromatic hydrocarbon content, and distillate distribution of the hydrofining product. The first stream obtained after hydrofining the LCO was subjected to impurity separation, which comprises the steps of carrying out gas-liquid separation in the first stream and extraction using nitrogen under normal pressure for 3 hours, to completely remove the hydrogen sulfide dissolved in the first stream. The sulfur content and nitrogen contents of the hydrofined product (the first separated impurity stream in liquid phase) were 87 ppm and 8.6 ppm, respectively. The retention rate of polycyclic aromatic hydrocarbons during hydrofining was 89.04 wt%, which was calculated from the aromatic composition data. Table 2 also lists the composition of the selective conversion catalyst B1 used for hydrocracking and the reaction conditions used. The USY zeolite and alumina were kneaded, extruded and formed, to provide the support for the selective conversion catalyst. Then, an appropriate amount of chloroplatinic acid was formulated in a calara solution, to immerse the support in a ¡sovolumetric soaking mode, dried and calcined in air at 500 °C for 2 hours, to provide a precursor of the conversion catalyst. selective. The selective conversion catalyst precursor was reduced to 450 °C in the presence of hydrogen to provide the desired selective conversion catalyst Bl, comprising: 0.1 part Pt-60 parts USY zeolite-39.9 parts AI2O3. The catalyst bed was cooled to 340 °C, and the hydrofined product after extraction (the first separated stream of impurities) was mixed with hydrogen, and fed into a selective conversion reactor, and the reaction product was fed to a gas-liquid rectification and separation system. After gas-liquid separation and a rectification system, benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons were obtained through separation and yield of monocyclic light aromatic hydrocarbons such as benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, was 21.48% by weight through calculation. The bottom oil weighed at >210 °C had a yield of 38.27 wt%, a specific gravity of 0.935, and a sulfur and nitrogen content of respectively 19.5 ppm and 1.5 ppm. A multidimensional chromatographic analysis was carried out to provide a family composition of the third stream of the following ΜΛ / Ί I OI way: 41.98% by weight of non-aromatic hydrocarbons, 26.38% by weight of monocyclic aromatic hydrocarbons and 31.64% by weight of polycyclic aromatic hydrocarbons. EXAMPLE 1 In the Example, a total conversion process to produce LCO light aromatics was carried out according to the flow chart shown in Figure 1. The LCO was subjected to hydrofining and separation of impurities, followed by selective conversion ( hydrocracking) and then the bottom oil weighed at > 210 °C was fed into a post-saturation selective reaction zone for selective hydrosaturation, specifically: The raw materials, the hydrofining catalyst, and the hydrofining reaction conditions were the same as those of Comparative Example 1, and the selective conversion catalyst B2 (hydrocracking catalyst) and the selective conversion reaction conditions are shown in the Table 3. MA / 1 / O / TABLE 3 Selective conversion catalyst B2 3.50 parts Ni-5.00 parts WO3 -0.27 parts WS2 / 50 parts p-zeolite-41.23 parts AI2O3 Hydrogen partial pressure for selective conversion 7.0 MPa Temperature for selective conversion 340 °C in input LHSV space velocity for selective conversion 1.0 h1 Hydrogen to oil ratio for selective conversion 1600 (v / v) The composition of the selective conversion catalyst B2 and the reaction conditions used are listed in Table 3. Selective conversion catalyst B2 was prepared as follows: 70 wt% β-zeolite (with silicon-aluminum molecular ratio SAR = 25) and 30 wt% alumina were kneaded, extruded and molded to provide support for the selective conversion catalyst. Then, an appropriate amount of nickel nitrate and ammonium tungstate were formulated in calara solution, to immerse the support in a volumetric soaking mode, dried at 100 °C, and calcined in air at 500 °C for 2 hours, to provide a precursor to the selective conversion catalyst. The selective conversion catalyst precursor was reduced to 450 °C for 4 hours in the presence of hydrogen, cooled to 330 °C, and then dimethyl disulfide was injected for vulcanization for 4 hours, to provide the desired selective conversion catalyst B2. Based on 100 parts by weight of the total weight of the catalyst, catalyst B2 comprised: 3.5 parts Ni-5.0 parts WO3 -0.27 part WS2 -50 parts p-zeolite-41.23 parts AI2O3. The first stream obtained from the LCO after hydrofining was subjected to impurity separation, where the first stream was subjected to gas-liquid separation and extracted with nitrogen for 3 hours under normal pressure, to completely remove the dissolved hydrogen sulfide. in the first stream. The hydrofined product after extraction (the first separated stream of impurities) was mixed with hydrogen, fed into a selective conversion reactor, and the reaction product was fed to a gas-liquid separation and rectification system. After gas-liquid separation and a rectification system, benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons were obtained through separation and yield of monocyclic light aromatic hydrocarbons such as benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, was 32.27% by weight through calculation. The bottom oil weighed at >210 °C (the third stream) had a yield of 24.75% by weight, a specific gravity of 0.957, and a sulfur and nitrogen content of respectively 25.4ppm and 1.6ppm. Multidimensional chromatographic analysis was carried out to provide a familiar composition of the third stream as shown in Table 4: 8.54 wt% non-aromatic hydrocarbons, 37.56 wt% monocyclic aromatic hydrocarbons, and 53.90 wt% aromatic hydrocarbons. polycyclical. MA / t / ZUZZ / UÓΊ fOÍ TABLE 4 Heavy fraction at > 210°C Density (4 °C) 0.957 Sulfur (ppm by weight) 25.4 Nitrogen (ppm by weight) 1.6 Non-aromatic hydrocarbons (% by weight) 8.54 Monocyclic aromatic hydrocarbons (% by weight) 37.56 Polycyclic aromatic hydrocarbons ( % by weight) 53.90 The selective post-saturation catalyst C2 for treating heavy tail oil at >210°C comprised: 0.05 wt%Pt-0.15 wt%Pd -4.5 wt%SiOz -95.3 wt%Al2O3. The post-saturation selective catalyst C2 was prepared as follows: a commercial amorphous silica-alumina material with 20 wt% SIO2 was mixed with pseudo-boehmite, then a nitric acid peptizing agent was added, a extrusion auxiliary of sesbania powder and a suitable amount of water, kneaded, extruded and formed, dried in air at 100 °C for 24 hours and then calcined in air at 550 °C for 4 hours, to provide the catalyst support. An appropriate amount of chloroplatinic acid and palladium chloride was dissolved in water to provide a metal impregnation solution, to immerse the catalyst support in an isovolumetric soaking mode, air dried at 80 °C for 48 hours and then calcined in air at 480 °C for 2 hours, to provide the selective catalyst after saturation C2. The post-saturation selective catalyst C2 was reduced in the presence of hydrogen with a final reduction temperature of 450 °C, and held for two hours. A supersaturated amount of hydrogen was dissolved in the heavy tail oil at >210 °C through a hydrogen mixer and fed into a selective saturation reactor, with reaction conditions of: a volume ratio of hydrogen to tail oil. 450Nm3 / m3, a reactor inlet temperature of 180°C, a hydrogen partial pressure of 1.5MPa, and a feed volume space velocity of 1.0 h'1. After equilibrating the entire reaction system, the analytical results for the selective saturation products were as shown in Table 5, where the sulfur and nitrogen contents were 16.8ppm and 1.2ppm, respectively. The retention rate of polycyclic aromatic hydrocarbons for the selective saturation process was 99.43% by weight, which was calculated from the aromatic composition data. ΜΛ / Ί I OI TABLE 5 Heavy fraction at > 210°C Density (4 °C) 0.939 Sulfur (ppm by weight) 16.8 Nitrogen (ppm by weight) 1.2 Non-aromatic hydrocarbons (% by weight) 8.96 Monocyclic aromatic hydrocarbons (% by weight) 59.82 Polycyclic aromatic hydrocarbons ( % by weight) 31.22 The bottom oil weighed at >210 °C after selective saturation was returned to the selective conversion reactor and steady stream equilibrium was established. After gas-liquid separation and a rectification system, benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons were obtained through separation and yield of monocyclic light aromatic hydrocarbons such as benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, was 46.35% by weight through calculation. EXAMPLE 2 In the Example, a total conversion process to produce LCO light aromatics was carried out according to the flow chart shown in Figure 1. The LCO was subjected to hydrofining and separation of impurities, followed by selective conversion ( hydrocracking) and then the bottom oil weighed at > 210 °C was fed into a post-saturation selective reaction zone for selective hydrosaturation, specifically: The raw materials, the hydrofining catalyst, and the hydrofining reaction conditions were the same as those of Comparative Example 1, and the selective conversion catalyst B3 (hydrocracking catalyst) and the selective conversion reaction conditions are shown in the Table 6. MA / t / ZUZZ / UÓΊ I OI TABLE 6 Selective conversion catalyst B3 0.2 part of Pd-6.5 parts of Ni-4.2 parts of MOO2 -7.9 parts of MoOs -1.1 parts of M0S2 / 35 parts of mordenite-10 parts of β-zeolite-l 1 parts of ZSM-5- 24.1 parts AI2O3 Hydrogen partial pressure for selective conversion 8.0 MPa Temperature for selective conversion 360 °C at inlet LHSV space velocity for selective conversion 1.2 h1 Hydrogen to oil ratio for selective conversion 2000 (v / v) The composition of the selective conversion catalyst B3 and the reaction conditions used are listed in Table 6. Selective conversion catalyst B3 was prepared as follows: hydrogen mordenite (SAR = 45), hydrogen β zeolite (SAR = 25), hydrogen ZSM-5 (SAR = 27) and pseudo-boehmite were thoroughly mixed, kneaded, extruded, dried at 120 °C, and calcined in an air atmosphere at 550 °C for 4 hours, to provide the required selective conversion catalyst support. Palladium chloride, nickel nitrate and ammonium molybdate were prepared in a trimetallic solution, to immerse the catalyst support in an isovolumetric soaking mode, dried at 120 °C and then calcined in an air atmosphere at 500 °C for 2 hours, to provide a selective conversion catalyst precursor. The selective conversion catalyst precursor was reduced to 450 °C and held for 8 hours in the presence of hydrogen, cooled to 330 °C, and then dimethyl disulfide was injected for vulcanization and held for 8 hours, to provide the desired selective conversion catalyst B3. Based on 100 parts by weight of the total weight of the catalyst, catalyst B3 comprised: 0.2 part Pd-6.5 parts Ni-4.2 parts MoCh -7.9 parts MOO3 -1.1 parts M0S2 -35 parts mordenite-10 parts of β-zeolite-l 1 parts of ZSM-5-24.1 parts of AI2O3. The first stream obtained from the LCO after hydrofining was subjected to impurity separation: where the first stream was subjected to gas-liquid separation and extracted with nitrogen for 3 hours under normal pressure, to completely eliminate the dissolved hydrogen sulfide in the first stream. The hydrofined product after extraction (the first separated stream of impurities) was mixed with hydrogen, fed into a selective conversion reactor, and the reaction product was fed to a gas-liquid separation and rectification system. After gas-liquid separation and a rectification system, benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons were obtained through separation and yield of monocyclic light aromatic hydrocarbons such as benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, was 30.08% by weight through calculation. The bottom oil weighed at >210 °C (the third stream) had a yield of 33.15% by weight, a specific gravity of 0.961, and a sulfur and nitrogen content of respectively 16.4ppm and 0.8ppm. Multidimensional chromatographic analysis was carried out to provide a familiar composition of the third stream as shown in Table 7: 7.58 wt% non-aromatic hydrocarbons, 38.12 wt% monocyclic aromatic hydrocarbons, and 54.30 wt% aromatic hydrocarbons. polycyclical. MA / t / ZUZZ / UÓΊ I OI TABLE 7 Heavy fraction at > 210°C Density (4 °C) 0.951 Sulfur (ppm by weight) 16.4 Nitrogen (ppm by weight) 0.8 Non-aromatic hydrocarbons (% by weight) 7.58 Monocyclic aromatic hydrocarbons (% by weight) 38.12 Polycyclic aromatic hydrocarbons ( % by weight) 54.30 The selective post-saturation catalyst C3 for treating heavy tail oil at >210°C comprised: 0.10 wt% Pt-0.30% Pd-4.0 wt%¡-6.0 wt% S¡02-89.6 wt% AI2O3. The selective post-saturation catalyst C3 was prepared as follows: a commercial amorphous silica-alumina material with 9 wt% SiOz was mixed with pseudo-boehmite, then a peptizing agent of nitric acid, an auxiliary of extrusion of sesbania powder and a suitable amount of water, kneaded, extruded and formed, dried in air at 100 °C for 24 hours and then calcined in air at 550 °C for 4 hours, to provide the support of the catalyst. Appropriate amounts of chloroplatinic acid, palladium chloride and nickel acetate were dissolved in water to provide a metal impregnation solution, to immerse the catalyst support in a ¡sovolumetric soaking mode, air dried at 100 °C for 18 hours, then calcined in air at 500 °C for 2 hours, to provide the selective post-saturation catalyst C3. The post-saturation selective catalyst C3 was reduced in the presence of hydrogen with a final reduction temperature of 450 °C, and held for two hours. A supersaturated amount of hydrogen was dissolved in the heavy tail oil at >210 °C through a hydrogen mixer and fed into a selective saturation reactor, with reaction conditions of: a volume ratio of hydrogen to tail oil. 600Nm3 / m3, a reactor inlet temperature of 150°C, a hydrogen partial pressure of 2.0MPa, and a feed volume space velocity of 1.5 hours'1. After equilibrating the entire reaction system, the analytical results for the selective saturation products were as shown in Table 8, where the sulfur and nitrogen contents were 11.3ppm and 0.6ppm, respectively. The retention rate of polycyclic aromatic hydrocarbons for the selective saturation process was 99.63% by weight, which was calculated from the aromatic composition data. MA / Ί I OI TABLE 8 Heavy fraction at : selective saturation > 210°C for Density (4 °C) 0.942 Sulfur (ppm by weight) 11.3 Nitrogen (ppm by weight) 0.6 Non-aromatic hydrocarbons (% by weight) 7.92 Monocyclic aromatic hydrocarbons (% by weight) 66.35 Polycyclic aromatic hydrocarbons (% by weight) 25.73 The bottom oil weighed at >210 °C after selective saturation was returned to the selective conversion reactor and steady stream equilibrium was established. After gas-liquid separation and a rectification system, benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons were obtained through separation and yield of monocyclic light aromatic hydrocarbons such as benzene-toluene, xylene, C9A aromatic hydrocarbons and C10A aromatic hydrocarbons, was 47.98% by weight through calculation.

Claims

1. A process for producing light aromatics from light cycle oil, i.e., LCO, characterized in that it comprises the steps of: 1) feeding LCO into a first reaction zone for hydrofining, to provide a first stream; 2) feeding the first stream to a second reaction zone for selective conversion, to provide a second stream, wherein the first stream is optionally subjected to impurity separation in a second separation zone before being fed to the second reaction zone; 3) subjecting the second stream to a first separation in a first separation zone, to provide a third stream comprising heavy aromatics of Cio+ at the bottom of the first separation zone; 4) feeding the third stream to a post-saturation selective reaction zone for hydrogenation saturation, to provide a fourth stream; 5) recycling the fourth stream to the second reaction zone.

2. The process according to claim 1, further characterized in that: in addition to the third stream, step 3) also provides fractions including a Ce-Cs aromatic hydrocarbon stream, and a stream containing C9 aromatic hydrocarbons and C10 aromatic hydrocarbons, wherein the Ce-Cs aromatic hydrocarbon stream comprises at least one of benzene, toluene, and xylene.

3. The process according to claim 1 or 2, further characterized in that: in step 2), the separation of impurities is carried out comprising subjecting the first stream to gas-liquid separation and the extraction of hydrogen sulfide.

4. The process according to any of the preceding claims, further characterized in that: in step 3), the first separation of the second stream comprises gas-liquid separation and rectification; and the rectification preferably comprises depentanization, deheptanization, xylene removal and heavy aromatic removal; wherein a stream rich in a benzene-toluene fraction obtained from desheptanization is preferably subjected to extraction separation. 5 - The process in accordance with any of the preceding claims, further characterized in that: the reaction conditions for the first reaction zone comprise: a hydrogen-to-oil volume ratio of 500-3000 Nm3 / m3, preferably 800-2000 Nm3 / m3, and more preferably 1000-1500 Nm3 / m3; and / or a reactor inlet temperature of 280-420 °C, preferably 300-410 °C, and more preferably 310-390 °C; and / or a hydrogen partial pressure of 5-10 MPa, preferably 5-8 MPa, and more preferably 6-7 MPa; and / or a space velocity of 0.5-2.0 h'1, preferably 0.6-1.5 h'1, more preferably 0.8-1.2 h'1.

6. The process in accordance with any of the preceding claims, further characterized in that: in step 2), the selective conversion is carried out in the presence of a selective conversion catalyst comprising, in parts by weight: a2) 5-80 parts of solid acid zeolite; b2) 0.05 to 8 parts of a Group VIII metal; c2) 3-25 parts of a Group VIB metal oxide; d2) 0.1-2 parts of a Group VIB metal sulfide; and e2) 20-95 parts of a first binder.

7. The process according to claim 6, further characterized in that: the solid acid zeolite is at least one of mordenite, β-zeolite, ZSM zeolite, EU-1 zeolite, SAPO zeolite and Y zeolite; the Group VIII metal is at least one of platinum, palladium, cobalt, nickel and iridium; the Group VIB metal oxide is at least one of molybdenum oxide and tungsten oxide; the Group VIB metal sulfide is at least one of molybdenum sulfide and tungsten sulfide; and the first binder is at least one of alumina, a silica-alumina compound, a titania-alumina compound and a magnesia-alumina compound.

8. The process in accordance with any of the preceding claims, further characterized in that: the reaction conditions for the second reaction zone comprise: a hydrogen-to-oil volume ratio of 800-5000 Nm3 / m3, preferably 1000-4000 Nm3 / m3, and more preferably 1500-3000 Nm3 / m3; and / or a reactor inlet temperature of 280-450 °C, preferably 300-430 °C, and more preferably 310-400 °C; and / or a hydrogen partial pressure of 5-10 MPa, preferably 5-9 MPa, and more preferably 6-8 MPa; and / or a space velocity of 0.5-2.0 h1, preferably 0.6-1.5 h1, more preferably 0.8-1.2 h1.

9. The process according to any of the preceding claims, further characterized in that: in step 4), the hydrogenation saturation is carried out in the presence of a post-saturation selective catalyst, comprising, in parts by weight: a3) 10-90 parts of amorphous silica-alumina, wherein the silica content of the amorphous silica-alumina is from 3 to 20% by weight; b3) 0.1 to 5.0 parts of a Group VIII metal; and c3) 5-80 parts of a second binder; the Group VIII metal is preferably at least one selected from platinum, palladium, cobalt, nickel, and iridium; and the second binder is selected from alumina. 10 - The process according to any of the preceding claims, further characterized in that: the reaction conditions for the post-saturation selective reaction zone comprise: a hydrogen-to-oil volume ratio of 200-3000 Nm3 / m3, preferably 300-1500 Nm3 / m3, and more preferably 300-1000 Nm3 / m3; and / or a reactor inlet temperature of 100-300 °C, preferably 120-280 °C, and more preferably 150-250 °C; and / or a hydrogen partial pressure of 1.0-4.0 MPa, preferably 1.2-3.0 MPa; and / or a space velocity of 0.1-5.0 h'1, preferably 0.5-4.0 h'1, more preferably 0.6-2.0 h'1.

11. A device for carrying out the process of any of claims 1 to 10, for producing light aromatics from LCO, characterized in that it comprises: a first reaction zone for hydrofining; configured to receive the LCO and to discharge a first stream; a second reaction zone for selective conversion; configured to receive the first stream and to discharge a second stream; a first separation zone; configured to receive a second stream; and to discharge the third stream at the bottom; a post-saturation selective reaction zone for hydrogenation saturation; configured to receive the third stream and to discharge a fourth stream; and a first pipe; configured to recycle the fourth stream to the second reaction zone.

12. The device according to claim 11, further characterized in that: the reactor of the first reaction zone is a fixed bed reaction system; and / or the reactor of the second reaction zone is a fixed bed reaction system; and / or the reactor of the post-saturation selective reaction zone is a fixed bed reaction system.

13. The device according to claim 12, further characterized in that: the fixed bed reaction system of the first reaction zone is equipped with a hydrogen recycling system; and / or the fixed bed reaction system of the second reaction zone is equipped with a hydrogen recycling system; and / or; the fixed bed reaction system of the post-saturation selective reaction zone is a liquid-phase hydrogenation reaction system without a hydrogen recycling system.

14. The device according to any of claims 11 to 13, further characterized in that: the first separation zone comprises a gas-liquid separator and a rectification column, which are optionally connected in series, to sequentially separate fractions comprising a benzene-toluene stream, a xylene stream, a stream containing C9 aromatic hydrocarbons and CIO aromatic hydrocarbons, and the third stream containing C10+ heavy aromatics; the rectification column preferably comprises, optionally connected in series, a deheptanizer, a xylene column, and a heavy aromatic compounds column. 15.- The device according to claim 14, further characterized in that: the first separation zone comprises a deheptanizer and, downstream of the same, a device for extracting the benzene-toluene fraction, to separate the stream rich in benzene-toluene fraction separated from the deheptanizer.

16. The device according to any of claims 11 to 15, further characterized in that: a second separation zone is placed between the first reaction zone and the second reaction zone, to separate impurities comprising hydrogen sulfide and ammonia from the first stream; and the second separation zone is configured to receive the first stream and to discharge a gas phase, a hydrogen sulfide and ammonia stream, and a first stream separated from impurities.

17. The device according to claim 16, further characterized in that: the second separation zone comprises a gas-liquid separator and an extraction device.

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