Integration process for the refining of pyrolysis oil with maximum BTX yield

Abstract

A process for producing aromatic compounds from pyrolysis oil, comprising: refining the pyrolysis oil to pyrolysis gasoline in a multi-stage reactor comprising a slurry phase reactor and a fixed-bed reactor, wherein the slurry phase reactor comprises a mixed metal oxide catalyst and the fixed-bed reactor comprises a mesoporous zeolite-supported metal catalyst; aromatizing the pyrolysis gasoline in an aromatization unit; hydrodealkylation and transalkylation of a product from the aromatization unit in a hydrodealkylation-transalkylation unit, thereby generating an aromatic stream; and processing the aromatic stream in an aromatic recovery complex to produce the aromatic compounds comprising benzene, toluene, and xylenes (BTX).

Description

Cross-reference to related application

[0001] This application claims priority to US Application No. 17 / 703,381, filed March 24, 2022, the entire disclosure of which is hereby incorporated by reference. Technical area

[0002] Embodiments of the present disclosure generally relate to the refining and refining of Hydrocarbon oil and particularly relate to a process for producing aromatic compounds from Pyrolysis gasoline. Technical background

[0003] To maximize the benefits of refining and petrochemical integration and to meet the stringent specifications for fuels, the recycling of petrochemical intermediates / refinery streams into petrochemical building blocks have gained attention in recent years. Meanwhile, market demand for aromatics, which serve as basic building blocks and intermediates in the petrochemical industry, have steadily increased, and the The market price for these petrochemical compounds has become noticeably dependent on the naphtha and crude oil prices. This trend has prompted researchers to look for alternative and cheaper raw materials for the production of aromatics, creating an opportunity and a potential market for Technologies that produce aromatic building blocks from cheap industrial intermediate streams.

[0004] In a typical steam pyrolysis process, a heavy by-product stream emerges as a bottom layer from the steam cracker, which is called pyrolysis oil. This heavy stream has a low sulfur content, but very rich in aromatics; however, it has limited applications and is normally used as heating oil Due to its high aromatic content, this inexpensive pyrolysis oil can be an ideal starting material for Production of valuable petrochemical intermediates such as benzene, toluene and xylenes (BTX). Summary

[0005] Therefore, there is a constant need for systems and methods for recovering a larger amount of Aromatics. The present embodiments are directed to the upgrading of pyrolysis oil in a continuous system to maximize the BTX yield. The system enables the increase of pyrolysis gasoline production at Utilization of pyrolysis oil to maximize BTX production.

[0006] According to one embodiment, a process for producing aromatic compounds from pyrolysis oil comprises The following: Refining the pyrolysis oil to pyrolysis gasoline in a multi-stage reactor, which slurry phase reactor and a fixed bed reactor, wherein the slurry phase reactor comprises a Mixed metal oxide catalyst and the fixed bed reactor comprises a mesoporous zeolite-supported metal catalyst comprises; aromatization of the pyrolysis gasoline in an aromatization unit; hydrodealkylation and transalkylation a product from the aromatization unit in a hydrodealkylation-transalkylation unit, whereby a aromatic stream is generated; and processing the aromatic stream in an aromatics recovery complex, for the production of aromatic compounds including benzene, toluene and xylenes (BTX).

[0007] Further features and advantages of the embodiments described herein are described in the following detailed description and are partly obvious to the person skilled in the art from this description or are obtained by practical application of the described embodiments, including the detailed description and claims provided below. Brief description of the drawings

[0008] The following detailed description of certain embodiments of the present disclosure is best to be understood when read in conjunction with the following drawings, in which Figure 1 is a schematic representation of the systems and methods of the present disclosure; Fig. 2 is another schematic representation of the systems and methods of the present disclosure; Fig. 3 is a diagram of the chemical processes used in the upgrading of pyrolysis oil according to one or several embodiments of the present disclosure; and Fig. 4 is a diagram of the chemical processes used in the upgrading of pyrolysis gasoline according to one or several embodiments of the present disclosure. Detailed description

[0009] The term ´´hydrocarbon oil´´ or ´´hydrocarbon feedstock´´ as used herein refers to a oily liquid consisting primarily of a mixture of hydrocarbon compounds. Hydrocarbon oil may include refined oil derived from crude oil, synthetic crude oil, bitumen, oil sands, shale oil or coal oil The term ´´refined oil´´ includes vacuum gas oil (VGO), deasphalted oil (DAO) from a Solvent deasphalting process, demetallized oil (DMO), light and / or heavy coker gas oil from a coker process, cycle oil from a fluid catalytic cracking (FCC) process and gas oil from a visbreaking Procedure.

[0010] The term ´´hydrocarbon´´ used here refers to a chemical compound which is exclusively consists of carbon and hydrogen atoms. A term like ´´C -C hydrocarbon´´ refers to a xy Hydrocarbon with x to y carbon atoms. A C -C hydrocarbon includes, for example, methane, ethane, 1 6 Propane, butanes and pentanes.

[0011] The term ´´liquid flow rate´´ or ´´LHSV´´ used here refers to the relationship between the liquid flow rate of the hydrocarbon feed and the catalyst volume or mass.

[0012] The term ´´mass flow rate´´ or ´´WHSV´´ used here refers to the relationship between the Weight of the reagent flow and the weight of the catalyst.

[0013] As used herein, the term ´´conduit´´ includes casings, linings, pipes, hoses, coiled tubes and mechanical structures with internal cavities.

[0014] The term ´´hydrogen / oil ratio´´ or ´´hydrogen-to-oil ratio´´ or ´´hydrogen / hydrocarbon ratio'' refers to a standard measure of the volume rate of the hydrocarbon circulating through the reactor. Hydrogen in relation to the volume of feedstock. The hydrogen / oil ratio can be determined by comparing the Flow volume of the hydrogen gas stream and the flow volume of the hydrocarbon feed become.

[0015] The term ´´reduced content´´ of a substance used here means that the concentration of the substance before after passing through a stage of the procedure under review than after passing through the stage. The term ´´increased content´´ of a substance means that the concentration of the substance after passing through a stage of the procedure under investigation is higher than before passing through the stage.

[0016] According to the formulation used here, any stream which is described as ´´rich´´ in a chemical substance, 50% or more of that chemical substance by volume.

[0017] As used throughout this disclosure, the term ''zeolites'' may refer to microporous inorganic Materials with regular intracrystalline cavities and channels of molecular size. Zeolites include generally a crystalline structure, as opposed to an amorphous structure as found in some porous materials such as amorphous silica. Zeolites generally contain a microporous Framework that can be identified by a framework type. The microporous structure of zeolites (e.g., 0.3 nm to 2 nm pore size) can lead to large areas and desirable size / shape selectivity, which is important for the catalysis of The zeolites described can be, for example, aluminosilicates, titanium silicates or pure silicates In embodiments, the described zeolites may contain micropores (present in the microstructure structure of a zeolite) and additionally contain mesopores. As used in this disclosure, Micropores refer to pores in a structure having a diameter of more than or equal to 0.1 nm and less than or equal to 2 nm, and mesopores refer to pores in a structure having a diameter of more than 2 nm and less than or equal to 50 nm. Unless otherwise stated here, the ´´pore size´´ of a material refers to the average pore size, but materials can additionally contain micropores and / or mesopores with a certain size that is not identical to the average pore size.

[0018] According to one or more embodiments, a process for producing aromatic compounds comprises from pyrolysis oil, the refinement of the pyrolysis oil to pyrolysis gasoline in a multi-stage reactor, which has a slurry phase reactor and a fixed bed reactor, wherein the slurry phase reactor comprises a Mixed metal oxide catalyst and the fixed bed reactor comprises a mesoporous zeolite-supported metal catalyst The process further comprises aromatizing the pyrolysis gasoline in an aromatization unit; Hydrodealkylation and transalkylation of a product from the aromatization unit in a hydrodealkylation transalkylation unit, thereby producing an aromatic stream; and processing the aromatic stream in an aromatics recovery complex to produce the aromatic compounds containing benzene, toluene and xylenes (BTX).

[0019] As shown in Fig. 1, a hydrocarbon feedstock 1 is supplied to a steam cracker and a fractionation unit 100. The steam cracker and fractionation unit 100 produce several product streams, which contain hydrogen 2, fuel gas 3, ethylene 4, propylene 5, butane 6, pyrolysis gasoline 7 and pyrolysis oil 8. The pyrolysis oil 8 is drained from the bottom of the steam cracker 100.

[0020] The pyrolysis oil 8 is fed to a multi-stage reactor 200, which produces pyrolysis gas 10. As shown in Fig. 2 As shown, the multi-stage reactor 200 comprises a slurry phase reactor 110 and a fixed bed reactor 120, which is connected downstream of the slurry phase reactor 110. The slurry phase reactor 110 may have a mixed metal oxide catalyst, and the fixed bed reactor comprises a mesoporous zeolite-supported metal catalyst, both of which are described in more detail below.

[0021] The pyrolysis oil 8, optionally with a diluent, is fed into the slurry phase reactor 110 The pyrolysis oil 8, which comprises polyaromatics, is brought into contact with the mixed metal oxide catalysts, while hydrogen gas 9 is supplied to the slurry phase reactor 110. In embodiments, the Mixed metal oxide catalysts are added simultaneously with the pyrolysis oil 8, so that a liquid and a solid Phase to produce the slurry.

[0022] In embodiments, the pyrolysis oil, a diluent, hydrogen and the mixed metal oxide Catalyst is added to a slurry phase reactor. In some embodiments, the components can simultaneously or, in other embodiments, at different times. The reactor can In some embodiments, the stirring may include stirring. In other In some embodiments, the agitation may include shaking. In some embodiments, the reactor may be Stir and shake to agitate.

[0023] During operation, the pyrolysis oil 8 is subjected to continuous selective hydrogenation, selective opening of saturated Rings, hydrodealkylation, transalkylation and disproportionation, whereby in Slurry phase reactor 110, 27 di-aromatic intermediates are formed in the effluent. An example of a A cascade of such reactions is shown in Fig. 3, which starts with pyrene and leads to a methylated naphthalene. Formation of methylated naphthalene from pyrene is of course only one example of the reaction cascade, and there are also other starting materials and products are conceivable.

[0024] Referring again to Fig. 2, after the reaction, the sludge effluent 27 can be discharged into a separator 130 to remove the spent catalyst and heavy residues, thereby creating a Fixed bed reactor feed 23, which comprises the di-aromatic intermediates, and a heavy stream 24, which contains the spent catalyst and the heavy residues with the remaining polyaromatic compounds, The heavier stream 24 can be fed to a catalyst recovery unit 140 in which the used catalyst 28 is separated from the heavy residues 29. Optionally, the heavy residue 29 can be used as Feedstock for the slurry phase reactor 110.

[0025] The feed to the fixed-bed reactor 23 may contain di-aromatic and mono-aromatic compounds. Examples for di-aromatic compounds include, but are not limited to, naphthalene and tetralin, both of which may be unsubstituted or substituted with straight-chain or branched hydrocarbon substituents. Examples of monoaromatic compounds include, but are not limited to, benzene which is unsubstituted or substituted with straight-chain or branched hydrocarbon substituents

[0026] In embodiments, the pyrolysis oil may contain one or more polyaromatic compounds. Polyaromatic compounds may contain one or more aromatic carbon atoms (C +), for example C 16 C polyaromatic compounds. For example, the polyaromatic compounds CC , CC , CC 16- 110 16- 100 16- 90 16- , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C , C -C 60 16 70 16 60 16 50 16 40 16 30 16 20 20 110 30 110 40 110 60 110 60 110 70 110 60 , C -C , or even C -C In embodiments, the polyaromatic compounds may have several 110 90 110 100 110 contain condensed aromatic rings, for example 3, 4, 5, 6, 7, 8, 9 or 10 condensed benzene rings.

[0027] In embodiments, the diluent may be an organic solvent, for example toluene, benzene or a combination of toluene and benzene. Without wishing to be bound to any particular theory, It is assumed that the diluent increases the fluidity of the pyrolysis oil feed so that the raw pyrolysis oil has a has better contact with the mixed metal oxide catalysts than would be possible without the diluent.

[0028] The conditions of the slurry phase reactor 110, such as the flow rate into the reactor, the The temperature in the reactor and the pressure in the reactor can be varied to control the reactions. In In some embodiments, the flow rate of the hydrogen gas can be between 15 ml / min and 35 ml / min, for example 20 ml / min up to 35 ml / min, at 25 ml / min up to 35 ml / min, at 30 ml / min up to 35 ml / min, at 15 ml / min up to 30 ml / min, at 15 ml / min up to 25 ml / min or even 15 ml / min to 20 ml / min. It is conceivable that the flow rate of hydrogen can be between one of the specified lower limits and one of the upper limits specified here. Without being tied to a specific To be bound by theory, it is assumed that at a hydrogen flow rate of less than 15 ml / min no sufficient amounts of hydrogen can enter the reactor. A hydrogen flow rate of more than 35 ml / min on the other hand, can lead to too much hydrogen circulating in the system, so that an unacceptable amount of hydrogen is consumed.

[0029] In embodiments, the pyrolysis oil, the diluent, the hydrogen and the mixed metal oxide Catalyst may remain in the slurry phase reactor 110 for a period of 1.5 h to 7.5 h. For example, the Components in the slurry phase reactor 1.5 h to 7 h, 1.5 h to 6.5 h, 1.5 h to 6 h, 1.5 h to 5.5 h, 1.5 h to 5 h, 1.5 h to 4.5 h, 1.5 h to 4 h, 1.5 h to 3.5 h, 1.5 h to 3 h, 1.5 h to 2.5 h, 1.5 h to 2 h, 2 h to 7.5 h, 2.5 h to 7.5 h, 3 h to 7.5 h, 3.5 h to 7.5 h, 4 h to 7.5 h, 4.5 h to 7.5 h, 5 h to 7.5 h, 5.5 h to 7.5 h, 6 h to 7.5 h, 6.5 h to 7.5 h, or even 7 h to 7.5 h. The flow rate of the pyrolysis oil and the diluent is intended to be between one of the lower limits specified here and one of the upper limits specified here. Without to be bound to a particular theory, it is assumed that if the components are less than 1.5 hours remain in the slurry phase reactor, the continuous selective hydrogenation, the selective opening of saturated Rings, hydrodealkylation, transalkylation and / or disproportionation do not have enough time to However, undesirable by-products may be produced if this time exceeds 7.5 hours. extends beyond.

[0030] In embodiments, the slurry phase reactor 110 may be operated at a temperature of 350°C to 450°C, such as for example from 360°C to 450°C, from 370°C to 450°C, from 380°C to 450°C, from 390°C to 450°C, from 400°C to 450°C, from 410°C to 450°C, from 420°C to 450°C, from 430°C to 450°C, from 440°C to 450°C, from 350°C to 440°C, from 350°C to 430°C, from 350°C to 420°C, from 350°C to 410°C, from 350°C to 400°C, from 350°C to 390°C, from 350°C to 380°C, from 350°C to 370°C or even from 350°C to 360°C. It can be provided that the temperature is between one of the lower limits specified here and one of the upper limits specified here. Without To be bound to a specific theory, it is assumed that a reactor temperature below 350°C will lead to can be that the continuous selective hydrogenation, the selective opening of saturated rings, the hydrodealkylation, the Transalkylation and / or disproportionation are too slow to be economically viable, while Reactor temperature above 450°C can cause one or more of these reactions to proceed too quickly, resulting in a runaway reaction or premature deactivation of the catalyst.

[0031] In embodiments, the slurry phase reactor 110 may be operated at a pressure of 3 MPa to 18 MPa, from 3.5 MPa to 18 MPa, from 4 MPa to 18 MPa, from 4.5 MPa to 18 MPa, from 5 MPa to 18 MPa, from 5.5 MPa to 18 MPa, from 6 MPa to 18 MPa, from 6.5 MPa to 18 MPa, from 7 MPa to 18 MPa, from 7.5 MPa to 18 MPa, from 8 MPa to 18 MPa, from 8.5 MPa to 18 MPa, from 9 MPa to 18 MPa, from 9.5 MPa to 18 MPa, from 10 MPa to 18 MPa, from 10.5 MPa to 18 MPa, from 11 MPa to 18 MPa, from 11.5 MPa to 18 MPa, from 12 MPa to 18 MPa, from 12.5 MPa to 18 MPa, from 13 MPa to 18 MPa, from 13.5 MPa to 18 MPa, from 14 MPa to 18 MPa, from 14.5 MPa to 18 MPa, from 15 MPa to 18 MPa, from 15.5 MPa to 18 MPa, from 16 MPa to 18 MPa, from 3 MPa to 17.5 MPa, from 3 MPa to 17 MPa, from 3 MPa to 16.5 MPa, from 3 MPa to 16 MPa, from 3 MPa to 15.5 MPa, from 3 MPa to 15 MPa, from 3 MPa to 14.5 MPa, from 3 MPa to 14 MPa, from 3 MPa to 13.5 MPa, from 3 MPa to 13 MPa, from 3 MPa to 12.5 MPa, from 3 MPa to 12 MPa, from 3 MPa to 11.5 MPa, from 3 MPa to 11 MPa, from 3 MPa to 10.5 MPa, from 3 MPa to 10 MPa, from 3 MPa to 9.5 MPa, from 3 MPa to 9 MPa, from 3 MPa to 8.5 MPa, from 3 MPa to 8 MPa, from 3 MPa to 7.5 MPa, from 3 MPa to 7 MPa, from 3 MPa to 6.5 MPa, from 3 MPa to 6 MPa, from 3 MPa to 5.5 MPa, or even from 3 MPa up to 5 MPa. It is conceivable that the pressure may be between one of the lower limits disclosed here and one of the upper limits disclosed here. Without being bound to any particular theory, assumed that a pressure below 3 MPa is not sufficient for the continuous selective hydrogenation, the selective Opening of saturated rings, hydrodealkylation, transalkylation and / or disproportionation take place However, at pressures above 18 MPa, special high-pressure equipment may be required, which would increase the cost of carrying out the reaction.

[0032] As already mentioned, the slurry phase reactor 110 includes a mixed metal oxide catalyst. In embodiments, the mixed metal oxide catalyst comprises two or more of FeO, ZrO, CeO, AlO, TiO, MoO, Co 2 3 2 2 2 3 2 3 2 O and NiO. In embodiments, the mixed metal oxide catalyst may contain 70 wt% to 90 wt% Fe O , 5 wt% to 60 3 2 3 Wt.% ZrO , 1 wt.% to 4 wt.% CeO and 5 wt.% to 10 wt.% Al O , wherein the wt.% is based on 2 2 2 3 the total amount of oxides. In embodiments, the mixed metal oxide catalyst may contain 10 to 50 wt. -% TiO ; 10 to 15 wt.% MoO ; 1 to 10 wt.% Co O ; and 1 to 5 wt.% NiO, wherein the wt.% in relation to 2 3 2 3 on the total amount of oxides.

[0033] In embodiments, 90% to 100% of the pyrolysis oil can be converted into intermediates. For example, 90% to 99%, 90% to 98%, 90% to 97%, 90% to 96%, 90% to 95%, 90% to 94%, 90% to 93%, 90% to 92%, 90% to 91%, 91% to 100%, 92% to 100%, 93% to 100%, 94% to 100%, 95% to 100%, 96% to 100%, 97% to 100%, 98% to 100% or even 99% to 100% of the pyrolysis oil can be converted into the intermediate products. It is intended that the yield of intermediate products is between one of the lower limits mentioned here and one of the upper limits mentioned here.

[0034] As shown in Fig. 2, the fixed bed reactor 120 may contain a mesoporous zeolite-supported metal catalyst which is described in more detail below. The feed of the fixed bed reactor 23 can be carried out with the mesoporous Zeolite-supported metal catalysts are brought into contact while hydrogen gas 25 is supplied. Optionally the fixed bed reactor feedstock 23 can be treated before contact with the mesoporous zeolite-supported metal catalysts through a heat exchanger (not shown). The optional heat exchanger can help reduce the Temperature of the liquid intermediates to a temperature sufficient to prevent condensation of the To avoid intermediates between the slurry phase reactor 110 and the fixed bed reactor 120.

[0035] After the reaction with the mesoporous zeolite-supported metal catalyst in the fixed bed reactor 120, a Separator 150 separates the pyrolysis gasoline 10 from other components in the effluent of the fixed bed reactor 120, so that the Pyrolysis gas can be further refined in the aromatization unit 300.

[0036] During operation, the contact of the intermediate products in the fixed bed reactor 120 with the mesoporous zeolite supported metal catalysts, which are described in more detail below, BTX by selective hydrogenation, selective Opening of the saturated ring followed by hydrodealkylation, transalkylation and disproportionation. This reaction cascade, which starts with naphthalene and leads to BTX, is shown in Fig. 4. In the case of overcracking saturated hydrocarbons can also be formed. The formation of BTX from naphthalene is of course only one example of the reaction cascade, and other starting materials are also conceivable.

[0037] The conditions of the reactor, such as the flow rate into the reactor, the temperature in the reactor and the pressure in the Reactor, can be varied to control the reaction in the fixed bed reactor 120. In embodiments, the Flow rate of hydrogen gas is 15 ml / min to 35 ml / min, for example 20 ml / min to 35 ml / min, 25 ml / min up to 35 ml / min, at 30 ml / min up to 35 ml / min, at 15 ml / min up to 30 ml / min, at 15 ml / min up to 25 ml / min or even at 15 ml / min to 20 ml / min. It is conceivable that the hydrogen flow rate can be between one of the lower limits and one of the upper limits specified here. Without being bound to any particular theory, It is assumed that a hydrogen flow rate into the reactor of less than 15 ml / min does not sufficient amounts of hydrogen into the fixed bed reactor 120. A hydrogen flow rate of However, more than 35 ml / min can lead to too much hydrogen circulating in the system, resulting in an unacceptable amount of hydrogen is consumed.

[0038] In embodiments, the flow rate into the fixed bed reactor 120 may be 0.2 g / min to 1 g / min, such as Example at 0.3 g / min to 1 g / min, at 0.4 g / min to 1 g / min, at 0.5 g / min to 1 g / min, at 0.6 g / min to 1 g / min, at 0.7 g / min to 1 g / min, at 0.8 g / min to 1 g / min, at 0.9 g / min to 1 g / min, at 0.2 g / min to 0.9 g / min, at 0.2 g / min to 0.8 g / min, at 0.2 g / min to 0.7 g / min, at 0.2 g / min to 0.6 g / min, at 0.2 g / min to 0.5 g / min, at 0.2 g / min to 0.4 g / min or even at 0.2 g / min to 0.3 g / min. It is conceivable that the flow rate may be between one of the lower limits and one of the upper limits specified here. Any suitable device for controlling the Flow rate, for example a liquid pump, can be used.

[0039] In embodiments, the fixed bed reactor 120 may be operated at a temperature of 350°C to 450°C, such as from 360°C to 450°C, from 370°C to 450°C, from 380°C to 450°C, from 390°C to 450°C, from 400°C to 450°C, from 410°C to 450°C, from 420°C to 450°C, from 430°C to 450°C, from 440°C to 450°C, from 350°C to 440°C, from 350°C to 430°C, from 350°C to 420°C, from 350°C to 410°C, from 350°C to 400°C, from 350°C to 390°C, from 350°C to 380°C, from 350°C to 370° C or even from 350°C to 360°C. It is conceivable that the temperature can be between one of the lower limits disclosed and one of the upper limits disclosed here. The temperature within of the fixed bed reactor can, but does not have to, be the same as the temperature inside the Slurry phase reactor. Without being bound to any particular theory, it is assumed that a Reactor temperature below 350°C can lead to the continuous selective hydrogenation, the selective opening saturated rings, hydrodealkylation, transalkylation and / or disproportionation are too slow to to be economically viable, but that a reactor temperature above 450°C can lead to one or more these reactions occur too quickly, resulting in a runaway reaction or premature deactivation of the catalyst can lead to.

[0040] In embodiments, the fixed bed reactor 120 may be operated at a pressure of 3 MPa to 18 MPa, from 3.5 MPa to 18 MPa, from 4 MPa to 18 MPa, from 4.5 MPa to 18 MPa, from 5 MPa to 18 MPa, from 5.5 MPa to 18 MPa, from 6 MPa to 18 MPa, from 6.5 MPa to 18 MPa, from 7 MPa to 18 MPa, from 7.5 MPa to 18 MPa, from 8 MPa to 18 MPa, from 8.5 MPa to 18 MPa, from 9 MPa to 18 MPa, from 9.5 MPa to 18 MPa, from 10 MPa to 18 MPa, from 10.5 MPa to 18 MPa, from 11 MPa to 18 MPa, from 11.5 MPa to 18 MPa, from 12 MPa to 18 MPa, from 12.5 MPa to 18 MPa, from 13 MPa to 18 MPa, from 13.5 MPa to 18 MPa, from 14 MPa to 18 MPa, from 14.5 MPa to 18 MPa, from 15 MPa to 18 MPa, from 15.5 MPa to 18 MPa, from 16 MPa to 18 MPa, from 3 MPa to 17.5 MPa, from 3 MPa to 17 MPa, from 3 MPa to 16.5 MPa, from 3 MPa to 16 MPa, from 3 MPa to 15.5 MPa, from 3 MPa to 15 MPa, from 3 MPa to 14.5 MPa, from 3 MPa to 14 MPa, from 3 MPa to 13.5 MPa, from 3 MPa to 13 MPa, from 3 MPa to 12.5 MPa, from 3 MPa to 12 MPa, from 3 MPa to 11.5 MPa, from 3 MPa to 11 MPa, from 3 MPa to 10.5 MPa, from 3 MPa to 10 MPa, from 3 MPa to 9.5 MPa, from 3 MPa to 9 MPa, from 3 MPa to 8.5 MPa, from 3 MPa to 8 MPa, from 3 MPa to 7.5 MPa, from 3 MPa to 7 MPa, from 3 MPa to 6.5 MPa, from 3 MPa to 6 MPa, from 3 MPa to 5.5 MPa, or even from 3 MPa up to 5 MPa. It is conceivable that the pressure may be between one of the lower limits disclosed here and one of the upper limits disclosed here. Without being bound to any particular theory, assumed that a pressure below 3 MPa is not sufficient to ensure continuous selective hydrogenation, selective Opening of saturated rings, hydrodealkylation, transalkylation and / or disproportionation. However, at pressures above 18 MPa, special high-pressure equipment may be required, which increases the cost for carrying out the reaction.

[0041] The fixed bed reactor 120 includes a mesoporous zeolite-supported metal catalyst. In general Zeolites can be characterized by a framework type that defines their microporous structure. The zeolites described are characterized in one or more embodiments by the framework type not particularly restricted. The framework types are described, for example, in the "Atlas of Zeolite Framework Types" by Ch. Baerlocher et al., fifth revised edition, 2001, which is referred to here. In embodiments The zeolites may contain microstructures (containing micropores) which are formed by zeolites of the *BEA framework type (for example, but not limited to, zeolite Beta), zeolites of the FAU framework type (for example, but not limited to limited, zeolite Y), zeolites of the MOR framework type or zeolites of the MFI framework type (for example, but not limited to limited, ZSM-5). It is understood that *BEA, MFI, MOR and FAU refer to zeolite framework types as identified by their respective three-letter codes established by the International Zeolite Association (IZA) Other types of frameworks are contemplated in the presently disclosed embodiments. In Embodiments, the mesoporous zeolite of the catalyst can be zeolite beta, ZSM-5, mordenite, zeolite Y and Combinations of two or more of these. The metal of the catalyst can be derived from a heteropolyacid. Exemplary heteropolyacids include Keggin-type heteropolyacids, for example phosphorus-tungsten Heteropolyacid (H PW O ), phosphorus-molybdenum heteropolyacid (H PMo O ), silicon-tungsten heteropolyacid (H 3 12 40 3 12 40 4 SiW O ), silicon-molybdenum heteropolyacid (H SiMo O ) or combinations thereof. 12 40 4 12 40

[0042] As shown in Fig. 2, the pyrolysis gasoline 10 is then aromatized in an aromatization unit 300. In a further embodiment, the pyrolysis gas 7 from a steam cracker 100 can also be Aromatization unit 300 to increase the aromatics yield.

[0043] An exemplary pygas formulation may contain 15 wt% to 20 wt% paraffins, 1.5 wt% to 3 wt% naphthenes, 50 wt% to 70 wt% aromatic hydrocarbons, 1 wt% to 2 wt% di-aromatic hydrocarbons, 5 % to 10 wt. % olefins and 7 wt. % to 9 wt. % diolefins. In embodiments, the sum the concentration of each of the above-mentioned components of the pygas is about 100 wt.%, for example 99.5 wt.% to 100.5 wt%, which takes into account the presence of impurities and experimental variations.

[0044] In the aromatization unit 300, the paraffins are at least partially converted into BTX by Cyclization, dealkylation and / or hydrodealkylation reactions. In embodiments, the aromatization involves contacting the hydrotreated pyrolysis gasoline stream with a catalyst, comprising a zeolite. In embodiments, the zeolite may be a Y-type zeolite, a ZSM-5 type zeolite or a combination of Y-type zeolite and ZSM-5 type zeolite. In embodiments, a A single reactor can be used for the entire cyclization, dealkylation, and hydrodealkylation. In other Embodiments may use more than one reactor, for example two reactors, for cyclization, dealkylation and Hydrodealkylation can be used. For example, one reactor can be used for cyclization and another reactor for dealkylation and / or hydrodealkylation. In embodiments where more than one reactor used, the catalyst in each reactor can be the same or different. In addition, the Temperature, pressure and WHSV may be the same or different in each reactor.

[0045] In embodiments, the flavoring unit 300 may be operated at a temperature of 200°C to 700°C, from 200°C up to 650°C, from 200°C to 600°C, from 200°C to 550°C, from 200°C to 500°C, from 200°C to 450°C, from 200°C to 400°C, from 200°C to 350°C, from 200°C to 300°C or from 200°C to 250°C, from 250°C to 700°C, from 300°C to 950°C, from 350° C to 700°C, from 400°C to 700°C, from 450°C to 700°C, from 450°C to 650°C, from 500°C to 700°C, from 500°C to 600°C, from 525°C to 575°C, from 550°C to 700°C, from 600°C to 700°C or even at a temperature of 650°C to 700°C It is understood that the temperature may be between any lower limit for the temperature and any upper temperature limit disclosed herein. Without being tied to a specific To be bound by theory, it is assumed that a reactor temperature below 200°C can lead to the Cyclization, dealkylation and / or hydrodealkylation reactions are too slow to be economically viable However, a reactor temperature above 700°C may cause the polymerization of species in the reactor which causes the catalyst to carbonize.

[0046] In embodiments, the aromatization unit 300 may be operated at a pressure of 0.1 MPa to 3 MPa, from 0.1 MPa to 2.9 MPa, from 0.1 MPa to 2.8 MPa, from 0.1 MPa to 2.7 MPa, from 0.1 MPa to 2.6 MPa, from 0.1 MPa to 2.5 MPa, from 0.1 MPa up to 2.4 MPa, from 0.1 MPa to 2.3 MPa, from 0.1 MPa to 2.2 MPa, from 0.1 MPa to 2.1 MPa, from 0.1 MPa to 2 MPa, from 0.1 MPa to 1.9 MPa, from 0.1 MPa to 1.8 MPa, from 0.1 MPa to 1.7 MPa, from 0.1 MPa to 1.6 MPa, from 0.1 MPa to 1.5 MPa, from 0.1 MPa to 1.4 MPa, from 0.1 MPa to 1.3 MPa, from 0.1 MPa to 1.2 MPa, from 0.1 MPa to 1.1 MPa, from 0.1 MPa to 1 MPa, from 0.1 MPa to 0.9 MPa, from 0.1 MPa to 0.8 MPa, from 0.1 MPa to 0.7 MPa, from 0.1 MPa to 0.6 MPa, from 0.1 MPa to 0.5 MPa, from 0.1 MPa to 0.4 MPa, from 0.1 MPa to 0.3 MPa, from 0.1 MPa to 0.2 MPa, from 0.2 MPa to 3 MPa, from 0.3 MPa to 3 MPa, from 0.4 MPa to 3 MPa, from 0.5 MPa to 3 MPa, from 0.6 MPa to 3 MPa, from 0.7 MPa to 3 MPa, from 0.8 MPa to 3 MPa, from 0.9 MPa to 3 MPa, from 1 MPa to 3 MPa, from 1.1 MPa to 3 MPa, from 1.2 MPa to 3 MPa, from 1.3 MPa to 3 MPa, from 1.4 MPa to 3 MPa, from 1.5 MPa to 3 MPa, from 1.6 MPa to 3 MPa, from 1.7 MPa to 3 MPa, from 1.8 MPa to 3 MPa, from 1.9 MPa to 3 MPa, from 2 MPa to 3 MPa, from 2.1 MPa to 3 MPa, from 2.2 MPa to 3 MPa, from 2.3 MPa to 3 MPa, from 2.4 MPa to 3 MPa, from 2.5 MPa up to 3 MPa, from 2.6 MPa to 3 MPa, from 2.7 MPa to 3 MPa, from 2.8 MPa to 3 MPa or even at a pressure of 2.9 MPa to 3 MPa. It is understood that the operating pressure may be between any lower pressure limit and any upper pressure limit disclosed herein. Without being To be bound by a particular theory, it is assumed that a pressure below 0.1 MPa is not sufficient for the Cyclization, dealkylation and / or hydrodealkylation can occur. At a pressure above 3 MPa, However, special high-pressure equipment may be required, which increases the cost of carrying out the reaction would. -1 -1 - 1

[0047] In embodiments, the aromatization unit 300 may be provided with a WHSV of 0.1 h to 20 h, from 0.1 h to 19 -1 -1 -1 -1 -1 -1 -1 -1 -1 - 1 -1 -1 h , from 0.1 h to 18 h , from 0.1 h to 17 h , from 0.1 h to 16 h , from 0.1 h to 15 h , from 0.1 h to 14 h , from 0.1 -1 -1 -1 -1 -1 -1 - 1 -1 -1 -1 -1 -1 h to 13 h , from 0.1 h to 12 h , from 0.1 h to 11 h , from 0.1 h to 10 h , from 0.1 h to 9 h , from 0.1 h to 8 h , -1 -1 -1 - 1 -1 -1 -1 -1 -1 -1 -1 from 0.1 h to 7 h , from 0.1 h to 6 h , from 0.1 h to 5 h , from 0.1 h to 4 h , from 0.1 h to 3 h , from 0.1 h to 2 h -1 - 1 -1 -1 -1 -1 -1 -1 -1 -1 -1 - 1 -1 , from 0.1 h to 1 h , from 0.5 h to 5 h , from 0.75 h to 1.25 h , from 1 h to 20 h , from 2 h to 20 h , from 3 h -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 until 20 h , from 4 h to 20 h , from 5 h to 20 h , from 6 h to 20 h , from 7 h to 20 h , from 8 h to 20 h , from 9 h -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 to 20 h , from 10 h to 20 h , from 11 h to 20 h , from 12 h to 20 h , from 13 h to 20 h , from 14 h to 20 h , from -1 -1 -1 -1 -1 -1 -1 -1 -1 -1 15 h to 20 h , from 16 h to 20 h , from 17 h to 20 h , from 18 h to 20 h , or even from 19 h to 20 h It is understood that the WHSV may vary from any lower WHSV limit disclosed herein to to any upper limit for the WHSV disclosed here. Without being bound to any particular theory -1 to be, it is assumed that a WHSV below 0.1 h requires a catalyst volume and / or a reactor volume that is too large to be commercially viable, or that a longer residence time of the reactants may be required - 1 which leads to a shorter time to catalyst deactivation. However, a WHSV of more than 20 h can lead to lead to the residence time of the reactants in the reactor being too short to carry out the hydrogenation.

[0048] Referring to Figures 1 and 2, the product 13 after aromatization is obtained from the Aromatization unit 300 in a hydrodealkylation-transalkylation unit 400 hydrodealkylated and transalkylated, to produce an aromatic stream 16.

[0049] In embodiments, the hydrodealkylation-transalkylation unit 400 may be a The fixed bed reactor can be charged with a catalyst composition that has a active metal on a mesoporous zeolite. The active metal can be, for example, nickel, Molybdenum, tungsten, platinum, palladium or a mixture of two or more of these metals. The carrier can for example, a mesoporous zeolite such as beta-mordenite or ZSM-5.

[0050] In embodiments, the hydrodealkylation-transalkylation unit 400 may be operated at a temperature of 300°C up to 500°C, such as at a temperature of 300°C to 490°C, from 300°C to 480°C, from 300°C up to 470°C, from 300°C to 460°C, from 300°C to 450°C, from 300°C to 440°C, from 300°C to 430°C, from 300°C to 420°C, from 300°C to 410°C, from 300°C to 400°C, from 300°C to 390°C, from 300°C to 380°C, from 300°C to 370°C, from 300°C to 360°C, from 300°C to 350°C, from 300°C to 340°C, from 300°C to 330°C, from 300°C to 320°C, from 300°C to 310°C, from 310°C to 500°C, from 320°C to 500°C, from 330°C to 500°C, from 340°C to 500°C, from 350°C to 500°C, from 360°C to 500° C, from 370°C to 500°C, from 380°C to 500°C, from 380°C to 480°C, from 390°C to 500°C, from 400°C to 500°C, from 410°C up to 500°C, from 420°C to 500°C, from 430°C to 500°C, from 440°C to 500°C, from 450°C to 500°C, from 460°C to 500°C, from 470°C to 500°C, from 480°C to 500°C or even from 490°C to 500°C. It is understood that the temperature between any lower limit for the temperature disclosed herein and any upper limit for the temperature. Without being bound to a particular theory, it is assumed that a Reactor temperature below 300°C can lead to the hydrodealkylation and transalkylation reactions slowly to be economically viable, but that a reactor temperature above 500°C leads to an increase unwanted by-products, thereby reducing the effectiveness of this step.

[0051] In embodiments, the hydrodealkylation-transalkylation reactor 400 may be operated at a pressure of 1.5 MPa to 6 MPa such as at a pressure of 1.5 MPa to 5.9 MPa, from 1.5 MPa to 5.8 MPa, from 1.5 MPa to 5.7 MPa, from 1.5 MPa to 5.6 MPa, from 1.5 MPa to 5.5 MPa, from 1.5 MPa to 5.4 MPa, from 1.5 MPa to 5.3 MPa, from 1.5 MPa to 5.2 MPa, from 1.5 MPa to 5.1 MPa, from 1.5 MPa to 5 MPa, from 1.5 MPa to 4.9 MPa, from 1.5 MPa to 4.8 MPa, from 1.5 MPa up to 4.7 MPa, from 1.5 MPa to 4.6 MPa, from 1.5 MPa to 4.5 MPa, from 1.5 MPa to 4.4 MPa, from 1.5 MPa to 4.3 MPa, from 1.5 MPa to 4.2 MPa, from 1.5 MPa to 4.1 MPa, from 1.5 MPa to 4 MPa, from 1.5 MPa to 3.9 MPa, from 1.5 MPa to 3.8 MPa, from 1.5 MPa to 3.7 MPa, from 1.5 MPa to 3.6 MPa, from 1.5 MPa to 3.5 MPa, from 1.5 MPa to 3.4 MPa, from 1.5 MPa to 3.3 MPa, from 1.5 MPa to 3.2 MPa, from 1.5 MPa to 3.1 MPa, from 1.5 MPa to 3 MPa, from 1.5 MPa to 2.9 MPa, from 1.5 MPa to 2.8 MPa, from 1.5 MPa to 2.7 MPa, from 1.5 MPa to 2.6 MPa, from 1.5 MPa to 2.5 MPa, from 1.5 MPa to 2.4 MPa, from 1.5 MPa to 2.3 MPa, from 1.5 MPa to 2.2 MPa, from 1.5 MPa to 2.1 MPa, from 1.5 MPa to 2 MPa, from 1.5 MPa to 1.9 MPa, from 1.5 MPa up to 1.8 MPa, from 1.5 MPa to 1.7 MPa, from 1.5 MPa to 1.6 MPa, from 1.6 MPa to 6 MPa, from 1.7 MPa to 6 MPa, from 1.8 MPa up to 6 MPa, from 1.9 MPa to 6 MPa, from 2 MPa to 6 MPa, from 2.1 MPa to 6 MPa, from 2.2 MPa to 6 MPa, from 2.3 MPa to 6 MPa, from 2.4 MPa to 6 MPa, from 2.5 MPa to 6 MPa, from 2.6 MPa to 6 MPa, from 2.7 MPa to 6 MPa, from 2.8 MPa to 6 MPa, from 2.9 MPa to 6 MPa, from 3 MPa to 6 MPa, from 3.1 MPa to 6 MPa, from 3.2 MPa to 6 MPa, from 3.3 MPa to 6 MPa, from 3.4 MPa to 6 MPa, from 3.5 MPa to 6 MPa, from 3.6 MPa to 6 MPa, from 3.7 MPa to 6 MPa, from 3.8 MPa to 6 MPa, from 3.9 MPa up to 6 MPa, from 4 MPa to 6 MPa, from 4.1 MPa to 6 MPa, from 4.2 MPa to 6 MPa, from 4.3 MPa to 6 MPa, from 4.4 MPa to 6 MPa, from 4.5 MPa to 6 MPa, from 4.6 MPa to 6 MPa, from 4.7 MPa to 6 MPa, from 4.8 MPa to 6 MPa, from 4.9 MPa to 6 MPa, from 5 MPa to 6 MPa, from 5.1 MPa to 6 MPa, from 5.2 MPa to 6 MPa, from 5.3 MPa to 6 MPa, from 5.4 MPa to 6 MPa, from 5.5 MPa to 6 MPa, from 5.6 MPa to 6 MPa, from 5.7 MPa to 6 MPa, from 5.8 MPa to 6 MPa, or even from 5.9 MPa to 6 MPa. It is understood that the pressure of the hydrodealkylation-transalkylation unit 400 can be adjusted between any desired disclosed lower limit for pressure and any upper limit for pressure disclosed herein. Without being bound to any particular theory, it is assumed that a pressure below 1.5 MPa is not sufficient for hydrodealkylation and / or transalkylation can take place. However, at a pressure of more than 6 MPa, special high-pressure equipment would be required, which would increase the cost of carrying out the reaction.

[0052] Referring again to Figures 1 and 2, the aromatic product produced in the aromatization unit 400 Stream 16 is then processed in an Aromatics Recovery Complex (ARC) 500 to recover aromatic compounds which include benzene 18, toluene 19 and xylenes 20 (BTX). Raffinate 17 can be used at the upper end of the ARC 500 Heavy aromatics (C9 / C9+) 21 can be removed from the bottom of the ARC 500. In further Embodiments as shown in Fig. 1, at least a portion of the heavy aromatic stream containing C aromatics 9 or higher, from the ARC 500 to the hydrodealkylation-transalkylation unit 400 (see stream 22). In addition, a separator 410 can be installed upstream of the ARC 500 to separate some lighter components. before feeding them to the ARC 500.

[0053] Within the ARC 500, several process steps can be carried out to produce high-quality products, e.g. xylenes and benzene, and converting lower-value products, such as toluene, into higher-value products. For example, the aromatics present can be separated into different fractions according to the number of carbon atoms, e.g., benzene, Toluene, xylenes and ethylbenzene etc. Fraction C can then be subjected to a process for the production of para-xylene (´´p-xylene´´) 8 which is a high-quality product. P-xylene can be obtained in high purity from fraction C, 8 by selective adsorption or crystallization of the p-xylene from ortho-xylene (´´o-xylene´´), meta-xylene (´´m-xylene´´) and Ethylbenzene is separated. The o-xylene and m-xylene remaining after the separation of p-xylene can be isomerized to to obtain an equilibrium mixture of xylenes. The ethylbenzene can be isomerized to xylenes or to benzene and ethane The p-xylene of the equilibrium mixture can also be dealkylated by adsorption or crystallization from the o-xylene and separated from m-xylene, and the p-xylene-depleted stream can be sent to the isomerization unit and then to the p-xylene recovery unit until all o-xylene and m-xylene are converted to p-xylene and recovered, or at least until the attempt to convert further p-xylene is no longer economically viable makes sense. Examples

[0054] Using the embodiments described above, an exemplary finishing system and - The procedure is simulated with the HYSIS v 10.0 simulator as follows. The following examples show the are for illustrative purposes only and should not be construed as limiting the scope of the present disclosure be understood. settlement proceedings

[0055] Fig. 5 shows a process flow diagram of a conventional process. In this embodiment, the pyrolysis gasoline is hydrotreated to remove di- and mono-olefins, which are responsible for the xylene separation used adsorbents are harmful, and then the deolefinated stream is fed into the Aromatics recovery complex. In addition, the pyrolysis oil is not converted into pyrolysis gasoline, but -1 simply disposed of. 380°C, 30 bar, WHSV 1.2 h. The material balance for this process is shown in Table 1 as follows. Table 1 - Material balance for the state of the art FIG. 5 Current# 7 8 9 10 1011 1012 1013 1014 1015 Flow rate. Kg / h 500 100 2.5 502.5 190.4 171.3 69.1 19.0 52.6 Composition, wt.% H 100% 0.3% 0.9% 2 (C1-C4) C5+ non-aromatic 37.5% 99.1% Benzene 34.1% 100% Toluene 13.8% 100% EB 1.3% 34.2% Xylene 2.5% 65.8% C9+ aromatics 10.5% 100% Pygas 100% Pyoil 100% Fuel gas heating oil Total 100% 100% 100% 100% 100% 100% 100% 100% 100%

[0062] Example of the invention

[0056] For the process of Fig. 1, Table 2 contains the process conditions and Table 3 the material balance. Table 2 - Process conditions for the process of Fig. 1. reactor Reactor type Reaction conditions Catalyst number Temperature: 400°C Mixed metal oxide 140 bar pressure Catalyst ZrO -CeO-Al O 110 Slurry Phase Reactor 2 2 Length of stay: 4 -FeO 3 times Hours Temperature: 400°C Zeolite-supported H pressure during 2 120 fixed bed reactor metal catalyst CoMoP Reaction: 30 - bar / Beta -1 WHSV: 1.2 hours Temperature: 450 -650°C H pressure during Fixed bed metal modified 2 300 Aromatization unit reaction: 1 - 10 bar zeolite catalyst -1 WHSV: 0.5-5 hours Temp: 300 - 450°C Zeolite catalyst Fixed bed hydrodealkylation 400 H pressure during - mordenite and / Transalkylation unit 2 medium-pore ZSM-5 Reaction: 20 - 50 bar -1 LHSV: 0.5-5 h Table 3 - Material balance of FIG. 1 Current# 7 8 9 10 11 12 13 14 Flow rate, kg / hr 500 100 1.2 94.2 1.4 4.4 594.2 1.9 Composition, wt.% H2 100% 0% 100% (C1-C4) 9% C5+ non-aromatic 5% Benzene 9.4% 35% Toluene 38.2% 28% EB 1% Xylene 17.3% 5% C9+ aromatics 35.0% 16% Pygas 100% Pyoil 100% Fuel Gas 100% Heating oil 100% Total 100% 100% 100% 100% 100% 100% 100% 100% Table 3 (continued) - Material balance of FIG. 1 Current# 15 16 17 18 19 20 21 22 Flow rate, kg / hr 62.9 533.2 32.5 279.1 0 159.9 58.1 0 Composition, wt.% H2 3.4% (C1-C4) 96.6% C5+ non-aromatic 6.1% 100% Benzene 41.1% 100% Toluene 22.1% EB 2.5% Xylene 17.3% 100% C9+ aromatics 10.9% 100% Pygas Fuel gas heating oil Total 100% 100% 100% 100% 0% 100% 100% 0%

[0057] The numbers of the streams are shown in Fig. 1 and Fig. 5. From Table 1 and Fig. 5 it can be seen that With a pygas feed stream of 500 kg / h and 100 kg / h of pyro oil, approximately 259 kg / h of BTX can be produced. Table 3 and Fig. 1 show, however, that the addition of pyrolysis oil to refine the pyrolysis gas, to aromatize the pygas and for hydrodealkylation / transalkylation, the BTX capacity increased to 439 kg / h.

[0058] According to one aspect, either alone or in combination with another aspect, a method for Production of aromatic compounds from pyrolysis oil: Refining of the pyrolysis oil to pyrolysis gasoline in a multi-stage reactor comprising a slurry phase reactor and a fixed bed reactor, wherein the Slurry phase reactor comprises a mixed metal oxide catalyst and the fixed bed reactor comprises a mesoporous Zeolite-supported metal catalyst; aromatization of the pyrolysis gasoline in an aromatization unit; Hydrodealkylation and transalkylation of a product from the aromatization unit in a hydrodealkylation Transalkylation unit, producing an aromatic stream; and processing the aromatic stream in an aromatics recovery complex to produce the aromatic compounds containing benzene, toluene and xylenes (BTX).

[0059] According to a second aspect, either alone or in combination with one of the other aspects, the The method further comprises supplying pyrolysis gas from a steam cracker to the aromatization unit.

[0060] According to a third aspect, either alone or in combination with one of the other aspects, the Pyrolysis oil from a steam cracker bottom product.

[0061] According to a fourth aspect, either alone or in combination with one of the other aspects, the The process further comprises recycling a heavy aromatics stream comprising C aromatics or higher from the 9 Aromatic recovery complex into the hydrodealkylation-transalkylation unit.

[0062] According to a fifth aspect, either alone or in combination with one of the other aspects, the Aromatization step one or more of the reactions cyclization, dealkylation and hydrodealkylation.

[0063] According to a sixth aspect, either alone or in combination with one of the other aspects, the -1 Aromatization in a reactor with a temperature of 200°C to 700°C, a mass flow rate of 0.1 h to 20 -1 h and a pressure of 0.1 MPa to 3 MPa.

[0064] According to a seventh aspect, either alone or in combination with one of the other aspects, the Aromatization involves bringing the pyrolysis gasoline into contact with a catalyst comprising a Y-type zeolite, a ZSM-5 type zeolite or a combination of Y type zeolite and ZSM-5 type zeolite.

[0065] According to an eighth aspect, either alone or in combination with one of the other aspects, the Slurry phase reactor at a temperature of 350°C to 450°C and a pressure of 3 MPa to 18 MPa operated.

[0066] According to a ninth aspect, either alone or in combination with one of the other aspects, the Mixed metal oxide catalyst two or more of FeO, ZrO, CeO, AlO, TiO, MoO, CoO, and NiO. 2 3 2 2 2 3 2 3 2 3

[0067] According to a tenth aspect, either alone or in combination with one of the other aspects, the Mixed metal oxide catalyst: from 70 wt% to 90 wt% Fe O ; from 5 wt% to 60 wt% 2 3 % ZrO ; from 1 wt.% to 4 wt.% CeO ; and from 5 wt.% to 10 wt.% Al O , the wt.% being based on the 2 2 2 3 Total amount of oxides can be calculated.

[0068] According to an eleventh aspect, either alone or in combination with one of the other aspects, the Mixed metal oxide catalyst: from 10 wt% to 50 wt% TiO ; from 10 wt% to 15 wt% MoO ; from 1 wt% to 10 2 3 wt% Co O ; and from 1 wt% to 5 wt% NiO, the wt% being calculated on the basis of the total amount of oxides 2 3 become.

[0069] According to a twelfth aspect, either alone or in combination with one of the other aspects, the Fixed bed reactor operated at a temperature of 350°C to 450°C and a pressure of 3 MPa to 18 MPa.

[0070] According to a thirteenth aspect, either alone or in combination with one of the other aspects, the Metal of the mesoporous zeolite-supported metal catalyst is a heteropolyacid.

[0071] According to a fourteenth aspect, either alone or in combination with one of the other aspects, the Heteropolyacid at least one Keggin-type heteropolyacid selected from the group consisting of Phosphorus-tungsten heteropolyacid (H PWO ), phosphorus-molybdenum heteropolyacid (H PMo O ), silicon-tungsten 3 12 40 3 12 40 Heteropolyacid (H SiW O ), silicon-molybdenum heteropolyacid (H SiMo O ) and combinations thereof. 4 12 40 4 12 40

[0072] According to a fifteenth aspect, either alone or in combination with one of the other aspects, the Zeolite carrier a zeolite selected from the group consisting of zeolite beta, ZSM-5, mordenite, zeolite Y and Combinations of two or more of them.

[0073] It is pointed out that the indications contained in the present disclosure that a Component of the present disclosure is ´´functional´´ or ´´sufficient´´ in a particular manner to to embody a certain property or to function in a certain way, structural clues are, in Contrary to references to the intended use. More specifically, the references in the present disclosure on the way in which a component is ´´functional´´ or ´´sufficient´´, refer to an existing physical state of the component and are as such considered to be unambiguous representations of the structural features of the component to understand.

[0074] Having described the subject matter of the present disclosure in detail and with reference to certain Embodiments, it is noted that the various embodiments described in the present disclosure details disclosed should not be understood as relating to elements that are material Components of the various embodiments described in the present disclosure. In addition, It will be appreciated that modifications and variations are possible without departing from the scope of the present disclosure which includes the embodiments defined in the appended claims, but not is limited.

[0075] The singulars ´´ein´´, ´´die´´ and ´´der´´ also include the plural, unless the context clearly indicates something else emerges.

[0076] Ranges are indicated throughout this disclosure. It is contemplated that each discrete value that encompassed by the ranges is also included. In addition, the ranges defined by each discrete value which is encompassed by the expressly disclosed areas.

[0077] As used in this disclosure and in the appended claims, the words ''comprising'' are intended to ´´have´´ and ´´include´´ and all grammatical variations thereof each have an open, non-limiting meaning that does not exclude additional elements or steps.

[0078] In this disclosure, terms such as ´´first´´ and ´´second´´ are used arbitrarily and are used only to Distinction between two or more circumstances or components. The terms 'first' and 'second' serve no other purpose and are neither part of the name or description of the component, nor do they define necessarily a relative location, position or order of the component. Furthermore, it is to be understood that the mere Use of the terms ´´first´´ and ´´second´´ does not presuppose that there is a ´´third´´ component, although this Possibility is also taken into consideration in the context of this disclosure. QUOTES CONTAINED IN THE DESCRIPTION

[0000] This list of documents submitted by the applicant was generated automatically and is intended solely for This list is not part of the German patent or trademark law. Utility model application. The DPMA assumes no liability for any errors or omissions. Cited patent literature

[0000] US 17 / 703381

[0001]

Claims

1. A process for producing aromatic compounds from pyrolysis oil, the process comprising: Upgrading of the pyrolysis oil to pyrolysis gasoline in a multi-stage reactor comprising a slurry phase reactor and a fixed bed reactor, wherein the slurry phase reactor comprises a mixed metal oxide catalyst, and the fixed bed reactor comprises a mesoporous zeolite-supported metal catalyst; Aromatization of the pyrolysis gasoline in an aromatization unit; Hydrodealkylation and transalkylation of a product from the aromatization unit in a hydrodealkylation Transalkylation unit, producing an aromatic stream; and Processing the aromatic stream in an aromatics recovery complex to produce the aromatic Compounds including benzene, toluene and xylenes (BTX).

2. The process of claim 1, further comprising supplying pyrolysis gas from a steam cracker to the Flavoring unit.

3. A process according to any one of claims 1 or 2, wherein the pyrolysis oil is obtained from a bottom product of a steam cracker comes from.

4. A process according to any one of the preceding claims, further comprising recycling a heavy aromatic stream comprising C or higher aromatics from the aromatics recovery complex into the hydrodealkylation 9 Transalkylation unit.

5. A process according to any one of the preceding claims, wherein the flavoring step comprises one or more of the Reactions include cyclization, dealkylation and hydrodealkylation.

6. A process according to any one of the preceding claims, wherein the aromatization is carried out in a reactor having a temperature -1 - 1 from 200°C to 700°C, a mass flow rate of 0.1 h to 20 h and a pressure of 0.1 MPa to 3 MPa is carried out.

7. A process according to any one of the preceding claims, wherein the aromatization comprises contacting the Pyrolysis gasoline with a catalyst comprising a Y-type zeolite, a ZSM-5 type zeolite or a Combination of Y-type zeolite and ZSM-5-type zeolite.

8. A process according to any one of the preceding claims, wherein the slurry phase reactor is operated at a temperature from 350°C to 450°C and a pressure of 3 MPa to 18 MPa.

9. A process according to any one of the preceding claims, wherein the mixed metal oxide catalyst comprises two or more of Fe O , 2 3 ZrO , CeO , Al O , TiO , MoO , Co O and NiO. 2 2 2 3 2 3 2 3 10. A process according to any one of the preceding claims, wherein the mixed metal oxide catalyst comprises: from 70 wt% to 90 wt% Fe O ; z3 from 5 wt% to 60 wt% ZrO ; 2 from 1 wt% to 4 wt% CeO ; and 2 from 5 wt% to 10 wt% Al O , where the wt% is calculated based on the total amount of oxides. 2 3 11. A process according to any one of the preceding claims, wherein the mixed metal oxide catalyst comprises: from 10 wt% to 50 wt% TiO ; 2 from 10 wt% to 15 wt% MoO ; 3 from 1 wt% to 10 wt% Co O ; and 2 3 from 1 wt% to 5 wt% NiO, where the wt% is calculated based on the total amount of oxides.

12. A process according to any one of the preceding claims, wherein the fixed bed reactor is operated at a temperature of 350°C to 450°C and a pressure of 3 MPa to 18 MPa.

13. A process according to any one of the preceding claims, wherein the metal of the mesoporous zeolite-supported metal catalyst comprises a heteropolyacid.

14. The process according to claim 13, wherein the heteropolyacid comprises at least one Keggin-type heteropolyacid which is selected from the group consisting of phosphorus-tungsten heteropolyacid (H PW O ), phosphorus-molybdenum- 3 12 40 Heteropolyacid (H PMo O ), silicon-tungsten heteropolyacid (H SiW O ), silicon-molybdenum heteropolyacid (H 3 12 40 4 12 40 4 SiMo O ) and combinations thereof. 12 40 15. A process according to any one of claims 1 to 14, wherein the zeolite support comprises a zeolite selected from the group consisting of Group consisting of zeolite beta, ZSM-5, mordenite, zeolite Y and combinations of two or more thereof.