JP2015519377A - Process for converting a hydrocarbon-based material into a fluid hydrocarbon product comprising p-xylene - Google Patents

Abstract

The present invention relates to a process for producing a flowable hydrocarbon product. More specifically, the present invention relates to a process for producing a fluid hydrocarbon product through catalytic pyrolysis. The reactant comprises a hydrocarbon-based material (eg, biomass). The catalyst comprises a zeolite catalyst treated with a silicone compound. The product comprises p-xylene.

Description

  This application claims priority to US Provisional Application No. 61/655605, dated June 5, 2012, under 35 USC 119 (e). The disclosure of the provisional application is incorporated herein by reference.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH OR DEVELOPMENT The U.S. Government has issued a patent under restricted circumstances, as provided for in the license paid for the present invention and award CBET-0747996 awarded by the National Science Foundation I have the right to require the right holder to license to others on reasonable terms.

  The present invention relates to a process for converting (or converting or converting) a hydrocarbonaceous material into a fluid hydrocarbon product comprising p-xylene. More specifically, the present invention relates to a process for converting a hydrocarbonaceous material to a fluid hydrocarbon product comprising p-xylene through pyrolysis of a catalyst.

  p-Xylene is used as a starting material for plasticizers and polyester fibers. The oxide of p-xylene is used for the commercial synthesis of terephthalic acid. Furthermore, dimethyl terephthalate is formed by esterification of terephthalic acid and methanol. Both monomers are used in the production of polyethylene terephthalate (PET) bottles, plastic bottles, and polyester garments.

  Catalyst pyrolysis comprising rapid pyrolysis of the catalyst (CFP) is used to convert hydrocarbonaceous materials (eg biomass) to fluid hydrocarbon products at a rapid heating rate in the presence of the catalyst. A process (or method or process) that can be performed. By using the method of the present invention, the flowable hydrocarbon product may comprise p-xylene and may further comprise additional aromatics, olefins, and the like.

In the present invention, a hydrocarbon-based material is supplied to a reactor (for example, a fluidized bed reactor). The hydrocarbonaceous material may be first pyrolyzed to form one or more pyrolysis products. The pyrolysis product may comprise one or more pyrolysis vapors. These pyrolysis products are modified (or modified or changed or modified) one or more olefinic compounds in the presence of a zeolite catalyst, CO, may react to form CO ₂ and an aromatic compound. The modified zeolite catalyst has pores with small size pore openings. The pyrolysis product that reacts enters the pores in the modified zeolite catalyst. At this time, the thermal decomposition products formed in the pores of the catalyst can diffuse out of the pores. The aromatic compound may comprise p-xylene or xylene having a relatively high selectivity relative to p-xylene. The advantages of this method are: 1) all desired chemistries may occur in a single step process, 2) the method may use a relatively inexpensive zeolite catalyst, and 3) p- Xylene can be produced at relatively high production levels.

  p-Xylene can be the most valuable xylene (ie, o-, m- and p-xylene). However, during catalytic pyrolysis of various hydrocarbon-based materials, the xylene is m-xylene and / or o-xylene with a selectivity equal to or higher than the selectivity of p-xylene. It is formed. Also, the produced p-xylene can be isomerized with respect to m-xylene and / or o-xylene. As a result, xylene with an unnecessarily high selectivity relative to m-xylene and / or o-xylene can be formed. As a result, the problem provides a method that can produce p-xylene or xylene with relatively high selectivity to p-xylene. The present invention provides a solution to this problem. The present invention provides for the use of a zeolite catalyst that has been modified to improve selectivity to p-xylene. Zeolite catalysts can consist of pores with catalyst sites on the outer surface of the catalyst (or catalytic sites) and pore openings. The pores may comprise catalyst sites on the inside, some of the catalyst sites may be located near the pore openings, and some of the catalyst sites may be located away from the pore openings . An effective amount of silicone compound to reduce the size of the pore openings and make at least some of the catalyst sites inaccessible (or inaccessible) to the reactants on the outer surface of the catalyst By treating the zeolite catalyst with the catalyst, the catalyst can be modified. The treatment process can also be used to render at least some of the catalyst sites inaccessible to the reactants in the pores near the pore openings.

  Zeolite catalysts can be treated by applying a silicone compound to the surface of the catalyst. The application of the silicone compound reduces the size of the pore openings in the catalyst, and the catalyst sites provided on the outer surface of the catalyst and the catalyst sites provided inside the pores in the vicinity of the pore openings. Can be covered or invisible. By coating the catalyst site with a treatment layer (or treatment layer), the catalytic activity of the catalyst site can be suppressed and / or extinguished. It is believed that the catalyst sites on the outer surface of the catalyst and the catalyst sites in the pores near the pore openings are not selective for xylene formation. On the other hand, catalytic sites in the pores that are separated from the pore openings increase the selectivity of the para. Suppression or extinction of the activity of the catalyst sites in the catalyst outer surface and in the catalyst pores near the pore openings increases the rate of catalytic reaction occurring in the pores of the zeolite catalyst away from the pore openings, thereby -Increased selectivity to xylene.

  Pore with reduced pore size can increase xylene para selectivity. This is because, by reducing the size of the pore openings, p-xylene can diffuse out of the pores, limiting the diffusion of m-xylene and o-xylene. This is illustrated in FIG. FIG. 2 shows a zeolitic catalyst having pore openings reduced in size so that p-xylene rather than m-xylene or o-xylene can diffuse out of the pores of the zeolite catalyst.

  The present invention relates to a process for producing a flowable hydrocarbon product comprising p-xylene from a hydrocarbon-based material. The method includes supplying a hydrocarbon-based material to the reactor, pyrolyzing at least a portion of the hydrocarbon-based material in the reactor under reaction conditions sufficient to produce a pyrolysis product, and Including the step of catalytically reacting at least a portion of the pyrolysis product under reaction conditions in which a zeolite catalyst was present to produce a fluid hydrocarbon product. The zeolite catalyst comprises pores having pore openings, catalyst sites on the outer surface of the catalyst, and a treatment layer, wherein the treatment layer reduces the size of the pore openings and the catalyst sites on the outer surface of the catalyst. An effective amount of the treatment layer derived from the silicone compound to render at least some of the thermal decomposition products inaccessible. The catalyst sites may be located in the pores near the pore openings and the treatment layer may render at least some of these catalyst sites inaccessible to pyrolysis products.

  In the present invention, when the active catalyst site is only in the catalyst pores and the xylene diffusing from the catalyst pores is only p-xylene, the selectivity of p-xylene in xylene can be up to 100%. In some embodiments, p-xylene may be produced in preference to o-xylene and / or m-xylene, however some o-xylene and / or m-xylene may be produced.

  In the present invention, the flowable hydrocarbon product produced using the method described above comprises xylene, and the selectivity of p-xylene to xylene is at least about 40%, at least about 45%, at least about 50%, at least about 55%, at least about 60%, at least about 65%, at least about 70%, at least about 75%, at least about 80%, at least about 85%, or at least about 90% obtain.

  Zeolite catalysts can be treated with silicone compounds to reduce the size of the pore openings and cover or hide the catalyst sites on the outer surface of the catalyst. This treatment process can also be used to cover or obscure catalyst sites in the pores near the pore openings. Catalytic sites can also be referred to as acid sites. A catalyst site that is covered or disappeared may be referred to as a deactivated catalyst site. The silicone compound may have a molecular size that cannot enter the pores of the catalyst. During the catalyst treatment step, the silicone compound can be applied to the catalyst and subsequently calcined. This process can be repeated until the desired level of processing is provided. Some of the catalyst sites on the outer surface of the catalyst that can be deactivated by treatment with the silicone compound are at least about 15%, at least about 25%, at least about 35%, at least about 45% of the available catalyst sites. At least about 55%, at least about 65%, at least about 75%, at least about 85%, at least about 90%, at least about 95%, at least about 98%, or at least about 99%. Some of the catalyst sites in the catalyst pores near the pore openings that can be deactivated by treatment with the silicone compound are at least about 20%, at least about 25%, at least about 30% of the available catalyst sites, At least about 33%, at least about 35%, at least about 40%, at least about 45%, at least about 49%, at least about 50%, at least about 55%, at least about 60%, at least about 65%, at least about 70%, It may be at least about 75%, at least about 80%, at least about 85%, at least about 90%, at least about 95%, or at least about 99%.

  The zeolite catalyst can comprise silica and alumina. The molar ratio of silica to alumina may be about 10: 1 to about 50: 1, about 10: 1 to about 40: 1, about 10: 1 to about 20: 1, or about 15: 1. The zeolite catalyst further comprises nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper, chromium, gallium, one or more oxides thereof, or a mixture of two or more thereof. It's okay.

The silicone compound may comprise at least one group represented by the following formula:

（Ｒ_１およびＲ_２は水素、ハロゲン、ヒドロキシル、アルキル、アルコキシル、ハロゲン化アルキル、アリール、ハロゲン化アリール、アラルキル、ハロゲン化アラルキル、アルカリル又はハロゲン化アルカリルを独立して含んで成り、ｎは少なくとも２である数である。Ｒ_１および/またはＲ_２はメチル、エチル、又はフェニルを含んで成り得る。ｎは約３〜約１０００の数であってよい。 The silicone compound may be represented by the following formula:

(R ₁ and R ₂ independently comprise hydrogen, halogen, hydroxyl, alkyl, alkoxyl, halogenated alkyl, aryl, aryl halide, aralkyl, halogenated aralkyl, alkaryl or alkaryl, and n is at least 2 R ₁ and / or R ₂ can comprise methyl, ethyl, or phenyl, n can be a number from about 3 to about 1000.

  The silicone compound may have a number average molecular weight of about 80 to about 20,000 or about 150 to 10,000.

  Silicone compounds are dimethyl silicone, diethyl silicone, phenyl methyl silicone, methyl hydrogen silicone, ethyl hydrogen silicone, phenyl hydrogen silicone, methyl ethyl silicone, phenyl ethyl silicone, diphenyl silicone, methyl trifluoropropyl silicone, ethyl trifluoropropyl silicone, poly Dimethyl silicone, tetrachlorophenyl methyl silicone, tetrachlorophenyl ethyl silicone, tetrachlorophenyl hydrogen silicone, tetrachlorophenyl phenyl silicone, methyl vinyl silicone, hexamethylcyclotrisiloxane, octamethylcyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenylcyclotetra Rokisan or it may comprise mixtures of two or more of these.

  The silicone compound may comprise tetraorthosilicate. The silicone compound may comprise tetramethylorthosilicate, tetraethylorthosilicate, or a mixture thereof.

  The reactor may comprise a continuous stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor, or a fluidized bed reactor. A fluidized bed reactor may be particularly advantageous.

  The hydrocarbon material may comprise a solid hydrocarbon material, a semi-solid hydrocarbon material, a liquid hydrocarbon material, or a mixture of two or more thereof. The hydrocarbonaceous material may comprise biomass. Hydrocarbon materials include plastic waste, recycled plastic, agricultural solid waste, municipal solid waste, food waste, animal waste, carbohydrates, lignocellulosic material, xylitol, glucose, cellobiose, hemicellulose, lignin, sugarcane bagasse , Glucose, wood, corn stover, or a mixture of two or more thereof. The hydrocarbonaceous material may comprise furan and / or 2-methylfuran. The hydrocarbon-based material may comprise pine. The hydrocarbon-based material may be a pyrolysis oil derived from biomass, a carbohydrate derived from biomass, an alcohol derived from biomass, a biomass extract, a pretreated biomass, a digested biomass product, or two or more of these It may comprise a mixture. A mixture of two or more of any of the foregoing may be used as the hydrocarbonaceous feedstock.

  The reactor may be operated at a temperature of about 400 ° C to about 600 ° C, about 425 ° C to about 500 ° C, or about 440 ° C to about 460 ° C.

The hydrocarbon-based material may be up to about 3 hours- ¹ (or less), up to about 2 hours- ^1, up to about 1.5 hours- ^1, up to about 0.9 hours- ^1, up to about 0.01 hours- ¹ . About 3 hours ⁻¹ , about 0.01 hours ⁻¹ to about 2 hours ⁻¹ , about 0.01 hours ⁻¹ to about 1.5 hours ⁻¹ , about 0.01 hours ⁻¹ to about 0.9 hours ^{− 1} , about 0.01 hour ⁻¹ to about 0.5 hour ⁻¹ , about 0.1 hour ⁻¹ to about 0.9 hour ⁻¹ , or about 0.1 hour ⁻¹ to about 0.5 hour ⁻¹ May be fed to the reactor at a normalized mass space velocity of

  The reactor may be operated at a pressure of at least about 100 kPa, at least about 200 kPa, at least about 300 kPa, or at least about 400 kPa. The reactor may be operated at a pressure of less than about 600 kPa, less than about 400 kPa, or less than about 200 kPa. The reactor may be operated at a pressure of about 100 to about 600 kPa, about 100 to about 400 kPa, or about 100 to about 200 kPa.

  The process of the present invention may be carried out under reaction conditions that minimize coke formation. A pyrolysis product may be formed wherein less than about 30%, less than about 25%, less than about 20%, less than about 15%, or less than about 10% by weight of the pyrolysis product is coke. .

  The method of the present invention may further comprise the step of recovering the fluid hydrocarbon product. The flowable hydrocarbon product may further comprise other aromatic and / or olefinic compounds in addition to p-xylene. Fluid hydrocarbon products are benzene, toluene, ethylbenzene, methylethylbenzene, trimethylbenzene, o-xylene, m-xylene, indane, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrinden, dimethylhydrindene. Or a mixture of two or more of these.

  The aromatic carbon yield in the flowable hydrocarbon product may be at least about 10%, at least about 15%, or at least about 22%. The carbon yield of the olefin in the flowable hydrocarbon product may be at least about 3%, at least about 6%, at least about 9%, or at least about 12%. The mass yield of p-xylene may be at least about 1.5 wt%, at least about 2 wt%, at least about 2.5 wt%, or at least about 3 wt%.

  The catalytic reaction step may comprise dehydration, decarbonylation, decarboxylation, isomerization, oligomerization, and / or dehydrogenation reaction.

  The pyrolysis step and the catalytic reaction step may be performed in a single vessel (or tank or vessel). Alternatively, the pyrolysis step and the catalytic reaction step may be performed in separate vessels.

  Other advantages and novel features of the invention will become apparent from the following detailed description of various non-limiting aspects of the invention when considered in conjunction with the accompanying drawings. In cases where the present specification and a document incorporated by reference include conflicting and / or inconsistent disclosure, the present specification shall control.

  Non-limiting aspects of the present invention will now be described by way of example with reference to the accompanying drawings, which are schematic and are not intended to be drawn to scale. In the drawings, each identical or nearly identical component that is illustrated is typically illustrated by a single numeral. For clarity, not all components are shown in all drawings, and all components of each aspect of the invention are not shown where illustration is not necessary since those skilled in the art can understand the invention.

FIG. A shows the selectivity of p-xylene obtained from the conversion of 2-methylfuran (2MF) + propylene over ZSM, ZSM ^* , GaZSM and GaZSM ^* . FIG. B shows the selectivity of p-xylene obtained from the conversion of furan with ZSM, ZSM ^* , GaZSM and GaZSM ^* . FIG. 1A is a schematic diagram illustrating a CFP process for converting a solid hydrocarbonaceous material into a flowable hydrocarbon product. FIG. 1B is a schematic diagram illustrating a CFP process for converting solid hydrocarbonaceous material into a flowable hydrocarbon product. FIG. 2 is a schematic diagram showing pores with reduced pore openings and diffusion of p-xylene outside the pores.

  All range and ratio limitations disclosed in the specification and claims may be combined in any manner. Unless specifically stated, a reference to “a”, “an” and / or “the” may include one or more, and a reference to a single item includes multiple items. It will be understood that it is good.

  The phrase “and / or” should be understood to mean “either or both” of the connected elements, that is, elements that are connected in some cases and elements that are not connected in other cases. It is. Unless specifically stated to be contrary, other elements may be more specific in the “and / or” section, regardless of whether they are related to the element specifically identified by the “and / or” section. It may exist arbitrarily other than the elements specified in. Thus, as a non-limiting example, a reference to “A and / or B” when used in conjunction with an open-ended word such as “comprising”, in certain embodiments, is A without B (optionally other than B). A) comprising the other elements, in another embodiment B without A (optionally B comprising elements other than A), and in yet another aspect (optionally including other elements) It is possible to refer to both A and B).

  The term “or” should be understood to have the same meaning as “and / or” as defined above. For example, when separating items in a list, it is “or” or “and / or” inclusive, that is, not only at least one element but also one or more elements, multiple elements, or list elements. Including, and optionally, additional, unlisted items. Only terms that are clearly contradicted such as “only one” or “exactly one” refer to comprising exactly one of a number of elements or elements of a list. It's okay. Generally, the term “or” as used herein is exclusive, such as “any”, “one of them”, “only one”, or “exactly one”. When preceded by a term, it is only to be interpreted as indicating an exclusive alternative expression (ie, “one or the other, not both”).

  The phrase “at least one” with respect to a list of one or more elements does not necessarily include each element specifically listed in the list of elements and at least one of all elements, and It should be understood that any combination of elements in the list of elements is not excluded, but means at least one element selected from any one or more elements in the list of elements. This provision also means that the phrase “at least one” may optionally exist other than that element, whether or not it relates to an element specifically identified in the list of referenced elements. Good things can be done. Thus, as a non-limiting example, “at least one of A and B” (equivalently “at least one of A or B” or equivalently “at least one of A and / or B”) Where B is absent (and optionally comprises elements other than B), and refers to at least one of optionally comprising more than one A; in another embodiment, A is absent ( And optionally other elements than A) at least one optionally comprising more than one B, and in yet another embodiment, optionally comprising more than one A. At least one and at least one of those optionally including more than one B (and optionally including other elements) may be referred to.

  “Comprising”, “comprising”, “carrying”, “having”, “containing”, “comprising” It will be understood that transitional terms or phrases such as “involving”, “holding”, etc. are open-ended, ie, include but are not limited to.

  The terms "pyrolysis" and "pyrolysis" refer to heat without oxygen or other oxidants, or with large amounts of oxygen or other oxidants, and with or without the use of a catalyst. Refers to changing one material (eg, solid hydrocarbonaceous material) to one or more materials (eg, volatile organic compounds, gas, coke, etc.).

  The term “catalytic pyrolysis” refers to pyrolysis performed in the presence of a catalyst.

  The term “aromatic” or “aromatic compound” refers to a single aromatic ring system (eg, benzyl, phenyl, etc.) and / or (eg, naphthyl, 1,2,3,4-tetrahydronaphthyl, etc.). ) Refers to a hydrocarbon compound comprising one or more aromatic groups such as a molten polycyclic aromatic ring system. Examples of aromatic compounds include, but are not limited to, toluene, indane, indene, 2-ethyltoluene, 3-ethyltoluene, 4-ethyltoluene, trimethylbenzene (for example, 1,3,5-trimethyl). Benzene, 1,2,4-trimethylbenzene, 1,2,3-trimethylbenzene, etc.), ethylbenzene, styrene, cumene, methylbenzene, propylbenzene, xylene (for example, p-xylene, m-xylene, o-xylene, etc.) ), Naphthalene, methyl-naphthalene (eg 1-methylnaphthalene, anthracene, 9,10-dimethylanthracene, pyrene, phenanthrene, dimethyl-naphthalene (eg 1,5-dimethyl-naphthalene, 1,6-dimethylnaphthalene, 2 , 5-dimethyl-naphthalene, etc.), ethyl-naphth Ren, hydrindene, methyl -. Hydrindene, and dimethyl hydrate Linden and the like In certain embodiments, monocyclic and / or polycyclic aromatic may be generated.

  The term “petrochemicals” is used herein to refer to chemicals, chemical precursors, chemical intermediates, etc. that are traditionally obtained from petroleum sources. Petrochemicals comprise paraffins, olefins, aromatic compounds and the like. For purposes of this application, when these materials are obtained from biomass and other non-petroleum sources (eg, replastics, municipal solid waste, sugarcane bagasse, wood, etc.), the term “petrochemicals” May be used despite the fact that chemical precursors, chemical intermediates, etc. are not obtained directly from petroleum.

  The term “biomass” refers to biological material that is present and has recently failed. According to the method of the present invention, biomass is converted into, for example, liquid fuel (eg, biofuel or biodiesel) or other fluid hydrocarbon product. Biomass consists of trees (eg, wood) and other plants; agricultural products and agricultural waste (eg, corn stover, bagasse, fruit, litter, silage, etc.); energy crops (eg, switchgrass, pampas grass); algae and others Marine plants; metabolic waste (eg, fertilizer, filth); and cellulose municipal waste. Biomass may be thought of as comprising material that has recently participated in the carbon cycle so that the release of carbon in the combustion process does not result in a net increase on average in some short time.

  For this reason, peat, lignite, coal, shale oil or petroleum consists of carbon that does not participate in the carbon cycle for long periods of time, so that combustion may result in a net increase in carbon dioxide in the atmosphere. It is not considered to be. The term biomass may refer to plant material grown for use as a biofuel, and also includes plant or animal material used for the production of fiber, chemicals, heat, etc. It may be made. Biomass may also comprise biodegradable waste or by-products that can be burned or converted into chemicals as fuel. These include municipal waste, green waste (biodegradable waste consisting of garden or park waste such as turf or flower cutting, hedge trimming, etc.), livestock by-products containing animal fertilizer, food processing waste Sewage sludge, black liquor from wood pulp or algae. Biomass may be obtained from a plant comprising Japanese pampas grass, euphorbia, sunflower, switchgrass, hemp, corn (or corn), poplar, willow, sugarcane, oil palm and the like. Biomass may be obtained from roots, stems, leaves, seed shells, fruits and the like. Although the processing of biomass may vary depending on reactor requirements and biomass shape, the particular plant or other biomass source used is important for the fluid hydrocarbon product produced according to the method of the present invention. It is not necessary.

  The hydrocarbon-based feed for the method of the present invention may comprise a solid hydrocarbon-based material, a semi-solid hydrocarbon-based material, a liquid hydrocarbon-based material, or a mixture of two or more thereof. The solids content of the hydrocarbon-based feedstock is up to about 100%, about 30% to about 100%, about 50% to about 100%, about 70% to about 100%, about 90%. % To about 100%, about 95% to about 100%, about 98% to about 100%, about 30% to about 95%, about 50% to about 95%, about 70% % To about 95%, about 80% to about 95%, about 85% to about 95%, or about 90% to about 95% by weight. The hydrocarbonaceous material may comprise biomass. Hydrocarbon materials include plastic waste, recycled plastic, agricultural solid waste, municipal solid waste, food waste, animal waste, carbohydrates, lignocellulosic material, xylitol, glucose, cellobiose, hemicellulose, lignin, sugarcane bagasse , Glucose, wood, corn stover, or a mixture of two or more thereof. The hydrocarbonaceous material may comprise furan and / or 2-methylfuran. The hydrocarbon-based material may comprise pine. The hydrocarbon-based material may be pyrolysis oil obtained from biomass, carbohydrate obtained from biomass, alcohol obtained from biomass, biomass extract, pretreated biomass, digested biomass product, or a mixture of two or more thereof. May be included. A mixture of two or more of any of the foregoing may be used.

  The method of the present invention may include the step of feeding a hydrocarbon-based material to the reactor. At least a portion of the hydrocarbonaceous material may be pyrolyzed in the reactor under reaction conditions sufficient to produce one or more pyrolysis products. At least a portion of the pyrolysis product can be catalytically reacted under conditions sufficient to produce a fluid hydrocarbon product. The reactor may include a continuous stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor, or a fluidized bed reactor. Suitably the reactor may comprise a fluidized bed reactor. The catalytic reaction step may be accomplished by feeding the catalyst with a hydrocarbon-based material. The catalyst may be supplied separately. Part of the catalyst may be supplied with the hydrocarbon feed, and part of the catalyst may be supplied separately.

  The methods of the invention can be used for the production of fluid (eg, liquid, supercritical fluid, and / or gas) hydrocarbon products through a catalytic pyrolysis process (eg, a CFP process). The flowable hydrocarbon product or part of the flowable hydrocarbon product may comprise a liquid at standard atmospheric temperature and atmospheric pressure (SATP—ie 25 ° C. and 100 kpa absolute pressure). The hydrocarbon feed may be pyrolyzed at an intermediate temperature (eg, about 400 ° C. to about 600 ° C.) compared to the temperatures typically used in the past. The pyrolysis step may be performed for a time effective to produce separate identifiable fluid hydrocarbon products. The method of the present invention may include heating the hydrocarbonaceous material (and optionally the catalyst) to the reaction temperature at a relatively high heating rate (eg, greater than about 50 ° C./second).

  The method of the present invention may include the use of special catalysts. These catalysts are zeolite catalysts comprising silica and alumina. The catalyst is characterized by pores with reduced pore openings and coated or invisible catalyst sites on the outer surface of the catalyst. The catalyst can have coated or hidden catalyst sites in the pores near the pore openings. The catalyst is treated with a silicone compound to reduce the pore opening size and to coat or obscure the catalyst sites within the catalyst outer surface and the catalyst pores near the pore opening. obtain. The catalyst may comprise relatively small particles that may be agglomerated. The composition fed to the reactor is a relatively high mass ratio of catalyst to hydrocarbonaceous material (e.g., about 0.33: 1 to about 20: 1, about 0.5: 1 to about 20: 1). Or about 2: 1 to about 10: 1).

  The expression that the catalyst site is located “in the pore near the pore opening” refers to the catalyst site in the pore that becomes inaccessible to pyrolysis products as a result of treatment with the silicone compound. In some embodiments, these catalytic sites may be located within the pores at a depth of only about 10 angstroms, about 7 angstroms, about 5 angstroms, or about 2 angstroms from the pore openings. These catalytic sites are suppressed or fail as a result of the silicone compound treatment. That is, these catalysts can be deactivated as a result of treatment with a silicone compound.

  The method of the present invention may include a single stage (or single stage or single stage) process for the pyrolysis of hydrocarbonaceous materials. The method may include providing or using a single stage pyrolysis apparatus. A single stage pyrolysis apparatus may be an apparatus in which pyrolysis and subsequent catalytic reactions are performed in a single vessel (or tank). The single stage pyrolysis apparatus may include a continuous stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor, or a fluidized bed reactor. A multi-stage apparatus may also be used for the production of fluid hydrocarbon products according to the present invention.

  The hydrocarbonaceous material can comprise any suitable size solid (or solid). In some cases it may be advantageous to use hydrocarbon solids having a relatively small particle size. In one example, small particle solids may react faster than large particle solids due to the relatively high surface area to volume ratio compared to large particle solids. Furthermore, the small particle size can make heat transfer within each particle and / or within the reactor volume more efficient. Thereby, the formation of undesirable reaction products can be suppressed or reduced. Moreover, the small particle size can increase solid-gas and solid-solid contact, which can improve heat and mass transfer. The average particle size of the solid hydrocarbonaceous material is less than about 5 mm, less than about 2 mm, less than about 1 mm, less than about 500 microns, less than about 60 mesh (250 microns), less than about 100 mesh (149 microns), about 140 mesh (105 microns), less than about 170 mesh (88 microns), less than about 200 mesh (74 microns), less than about 270 mesh (53 microns), or less than about 400 mesh (37 microns), or smaller average particles It can be size.

  In order to reduce the pressure required to pass the solid hydrocarbon feed through the reactor, it is desirable to use a feed with an average particle size above the minimum particle size. For example, at least about 400 mesh (37 microns), at least about 270 mesh (53 microns), at least about 200 mesh (74 microns), at least about 170 mesh (88 microns), at least about 140 mesh (105 microns), at least about 100 Solid hydrocarbon feed with an average particle size of mesh (149 microns), at least about 60 mesh (250 microns), at least about 500 microns, at least about 1 mm, at least about 2 mm, at least about 5 mm, or larger It is desirable to use materials.

  The hydrocarbonaceous material may comprise biomass. The hydrocarbon-based material may be plastic waste, recycled plastic, agricultural and / or municipal solid waste, food waste, animal waste, carbohydrates, lignocellulosic material (eg, wood chips or waste), or two of these. It may comprise more than one mixture. The hydrocarbonaceous material may comprise xylitol, glucose, cellobiose, cellulose, hemicellulose, lignin, or a mixture of two or more thereof. The hydrocarbonaceous material may comprise sugarcane bagasse, glucose, wood, corn stover, or a mixture of two or more thereof. The hydrocarbon-based material may comprise wood.

  A biomass pyrolysis liquid or bio-oil may be formed during the pyrolysis step of the method of the invention. The biomass pyrolysis liquid may be dark brown and may approximate biomass with an elemental composition. The biomass pyrolysis liquid may consist of a fairly complex mixture of oxygenated (or oxidized) hydrocarbons and a significant proportion of water from both the original moisture and reaction products. Compositionally, biomass pyrolysis oils may vary depending on the type of biomass, but oxygenated low molecular weight alcohols (eg furfuryl alcohol), aldehydes (aromatic aldehydes), ketones (furanones), phenols (methoxy) It is known to comprise phenol) and water. There may also be a solid char suspended (or suspended) in the oil. The liquid can be formed by rapidly cooling the intermediate product of flash pyrolysis of hemi-cellulose, cellulose and lignin in the biomass. Chemically, the oil may comprise hundreds of different chemicals in a wide variety of proportions ranging from formaldehyde and acetic acid to complex high molecular weight phenols, anhydro sugars (or anhydro sugars) and other oligosaccharides. . The oil may have a unique odor from low molecular weight aldehydes and acids and may be acidic with a pH of about 1.5 to about 3.8 and may be irritating.

  The residence time of the catalyst in the reactor can be defined as the volume of the reactor filled with the catalyst divided by the volume flow rate of the catalyst through the reactor. For example, if a 3 liter reactor contains 2 liters of catalyst and the catalyst is fed through the reactor at a flow rate (or flow rate) of 0.4 liters / minute, that is, fed and removed, the catalyst residence time is 2 / 0.4 minutes or 5 minutes.

  The residence time of the catalyst in the reactor is at least about 1 minute, at least about 2 minutes, at least about 5 minutes, at least about 7 minutes, at least about 10 minutes, at least about 15 minutes, at least about 20 minutes, at least about 25 minutes, It may be at least about 30 minutes, at least about 60 minutes, or at least about 120 minutes. In some cases, the residence time of the catalyst in the reactor is less than about 120 minutes, about 1 minute to about 120 minutes, about 2 minutes to about 120 minutes, about 5 minutes to about 120 minutes, about 7 minutes to about 120 minutes. About 10 minutes to about 120 minutes, about 12 minutes to about 120 minutes, about 15 minutes to about 120 minutes, about 20 minutes to about 120 minutes, about 30 minutes to about 120 minutes, or about 60 minutes to about 120 minutes It may be. In some cases, the relatively long residence time allows additional chemical reactions to form the desired product. Long catalyst residence times can be achieved, for example, by increasing the volume of the reactor and / or decreasing the volumetric flow rate of the catalyst. The residence time of the catalyst may be relatively short, for example less than about 120 minutes, or less than about 60 minutes.

  Contact time may be defined as the residence time of the material in the reactor or other device when measured or calculated under standard temperature and pressure conditions (ie, 0 ° C. and 100 kPa absolute pressure). For example, a 2 liter reactor fed with a standard 3 liter / min gas has a contact time of 2/3 min, ie 40 seconds, for that gas. Due to the chemical reaction, the contact time or residence time is based on the volume of the reactor in which the substantial reaction occurs, and excludes the volume in which no substantial reaction occurs, such as the inlet conduit or the discharge conduit. Due to the catalytic reaction, the volume of the reaction chamber is the volume where the catalyst is present.

  The term “reactant conversion” refers to the difference in moles or mass of reactants between the material flowing into the reactor and the material flowing out of the reactor, the moles of reactants in the material flowing into the reactor or Refers to the value divided by mass. For example, when 100 g of ethylene is supplied to the reactor and 30 g of ethylene is discharged from the reactor, the ethylene conversion is [(100-30) / 100] = 70%.

  The term “fluid” refers to a gas, a liquid, a mixture of gas and liquid, or a gas or liquid comprising dispersed solids, droplets and / or gaseous bubbles. The terms “gas” and “steam” have the same meaning and are sometimes used interchangeably. In certain embodiments, it may be advantageous to control the residence time of the fluidizing fluid within the reactor. The fluidization residence time of the fluidizing fluid is defined as the reactor volume divided by the volumetric flow rate of the fluidizing fluid under temperature and pressure processing conditions.

  The term “fluidized bed reactor” can be used to refer to a reactor comprising a vessel comprising a particulate solid material (eg, silica particles, etc.). The fluid (eg, gas or liquid) passes through the particulate solid material at a sufficiently high rate to suspend the solid material and behave as if the solid material is a fluid. The term “circulating fluidized bed reactor” is used to refer to a fluidized bed reactor in which particulate solid material exits the reactor, is circulated through a line in fluid communication with the reactor, and is recycled back into the reactor. Can be done.

  Bubbling fluidized bed reactors and turbulent fluidized bed reactors may be used. In a bubbling fluidized bed reactor, the fluid flow used to fluidize the particulate solid material is operated at a sufficiently low flow rate so that bubbles and voids (or voids) are contained within the bulk (or volume or volume) of the fluidized bed during operation. Can be observed. In a turbulent fluidized bed reactor, the flow rate of the fluidized stream is greater than that used in the bubbling fluidized bed reactor. Specific examples of fluidized bed reactors, circulating fluidized bed reactors, bubbling fluidized bed reactors, and turbulent fluidized bed reactors can be found in Kirkusmer Chemical Technology (Online) Encyclopedia, Vol. 11, Hoboken, NJ: Williinter Science, 2001, 791-825. These pages are incorporated herein for reference.

  The terms “olefin” or “olefin compound” (ie, “alkene”) may be used to indicate an unsaturated hydrocarbon comprising one or more pairs of carbon atoms joined by a double bond. Olefin may comprise both cyclic and acyclic (aliphatic), double bonds between carbon atoms forming part of a cyclic (ring-closed) group or between carbon atoms forming part of an open chain group. Each is provided. Further, the olefin may comprise any suitable number of double bonds (eg, monoolefin, diolefin, triolefin, etc.). Examples of olefin compounds may include ethene, propene, allene (propadiene), 1-butene, 2-butene, isobutene (2-methylpropene), butadiene and isoprene, among others. Examples of cyclic olefins may include cyclopentene, cyclohexane, cycloheptene, among others. Aromatic compounds such as toluene are not considered olefins. However, olefins comprising aromatic moieties such as benzyl acrylate or styrene are considered olefins.

  The pore size is related to the size of molecules or atoms that can penetrate into the pores of the material. As used herein, the term “pore size” for zeolite refers to the Norman radius adjusted pore size. The Norman radius-adjusted pore size is defined, for example, in 409-414, 1 (1999) of the 12th International Zeolite Conference Meeting held July 5-10, 1998 in Baltimore, Cook, M .; Conner, W. C., “How big are the pores in the zeolite?” And is incorporated herein by reference. As a specific typical calculation, the atomic radius of ZSM-5 pores is about 5.5-5.6 angstroms as measured by X-ray diffraction. In order to adjust the repulsion effect between oxygen atoms in the catalyst, Cook and Conner show that the Norman tuning radius is 0.7 Angstroms, which is larger than the atomic radius (about 6.2-6.3 Angstroms). Yes.

  Those skilled in the art understand how to determine the pore size (eg, minimum pore size, average minimum pore size) within the catalyst. For example, X-ray diffraction (XRD) can be used to determine atomic coordination (or coordinates). XRD techniques for pore size determination are described in, for example, “Fundamentals of powder diffraction and structural analysis of materials” by Springer Science + Business Media, Inc., Pecharsky.VK et al. All of which are incorporated herein by reference. Other techniques that can help to determine the pore size (eg, zeolite pore size) include, for example, helium pycnometry or low pressure argon adsorption techniques. These or other techniques are described on pages 185-195 of “Fluid catalytic cracking: science and technology” by Magee, J.S et al., Elsevier Publishing, July 1, 1993, for reference. All of which are incorporated herein. The pore size of the mesoporous catalyst is described in, for example, “Adsorption, surface area and porosity” by Gregg, SJ et al., 1982 Academic Press, 2nd edition, New York, which is incorporated herein by reference. , 1998, Roquerol, F et al., Academic Press, New York, et al., “Adsorption, Principles, Procedures and Applications with Powders and Porous Materials”.

  The screening method can be used to select a catalyst with an appropriate pore size for conversion of a particular pyrolysis product molecule. The screening method can include determining the size of the desired pyrolysis product molecule to catalyze (eg, the molecular dynamics diameter (or dynamic molecular size) of the pyrolysis product molecule). One skilled in the art can calculate, for example, the kinetic diameter of a given molecule. In doing so, a catalyst type can be selected that results in catalyst pores that are sufficiently large (eg, a Norman-adjusted minimum radius) so that the pyrolysis product extends into the catalyst and / or can react with the catalyst. In some embodiments, a catalyst may be selected that results in a pore size that is sufficiently small to contain and / or not react with undesirable pyrolysis products.

  The catalyst may comprise any catalyst suitable for carrying out the catalytic reaction step of the process of the present invention. The catalyst is a product or intermediate distribution to reduce the activation energy of the reaction carried out in the catalytic reaction step (increase the reaction rate) and / or during the reaction (e.g. of a shape selective catalyst). Can be used to improve. Examples of reactions that can act as a catalyst include dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization, decarbonylation, decarboxylation, aldol condensation, and combinations thereof. The catalyst component may be acidic, neutral or basic.

  The method of the present invention may include a CFP process. For the CFP process, particularly advantageous catalysts are pore sizes (eg, mesoporous and pore sizes typically associated with zeolites), eg, less than about 100 angstroms, less than about 50 angstroms, less than about 20 angstroms, about 10 angstroms. The catalyst may comprise an internal porosity that is selected according to a pore size of less than, less than about 5 angstroms, or smaller. In some embodiments, a catalyst having an average pore size of about 5 angstroms to about 100 angstroms may be used. In some embodiments, a catalyst having an average pore size of about 5.5 angstroms to about 6.5 angstroms or about 5.9 angstroms to about 6.3 angstroms may be used. In some cases, a catalyst having an average pore size of about 7 angstroms to about 8 angstroms, or about 7.2 angstroms to about 7.8 angstroms may be used.

  The catalyst may be selected from naturally occurring zeolites, synthetic zeolites and combinations thereof. The catalyst may comprise a ZSM-5 zeolite catalyst. The catalyst may comprise an acid site. These acid sites may also be referred to as catalytically active sites. Other zeolite catalysts that may be used include ferrierite, zeolite Y, zeolite beta, mordenite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S) AIPO -31, SSZ-23, and the like. The catalyst comprises silica and alumina, and may further comprise one or more additional metals and / or metal oxides. Suitable metals and / or metal oxides include, for example, nickel, palladium, platinum, titanium, vanadium, chromium, manganese, iron, cobalt, zinc, copper, gallium, and / or any of these oxides, among others. Can be mentioned. In some cases, rare earth elements, i.e. elements 57-71, cerium, zirconium, or oxides thereof, or combinations thereof may be included to alter the activity, structure, and / or stability of the catalyst. Further, in some cases, the properties of the catalyst (eg, pore structure, type, and / or number of catalyst sites) may be selected to selectively produce the desired product.

  The catalyst is reduced once with a silicone compound to reduce the size of the pore openings in the catalyst and to cover or hide the catalyst sites on the outer surface of the catalyst and the catalyst sites in the catalyst pores near the pore openings. The above treatment or impregnation may be performed. By covering the catalyst sites with the treatment layer, the catalytic activity can be suppressed and / or extinguished. In order to facilitate a more controlled application of the silicone compound, the silicone compound may be dispersed in a carrier such as, for example, an aqueous or organic liquid carrier.

  At each stage of the catalyst treatment process, the silicone compound may be deposited on the outer surface of the catalyst by any suitable method. For example, the silicone compound may be dissolved in an organic support, mixed with the catalyst, and then dried by evaporation or vacuum distillation. The catalyst may be contacted with the silicone compound in a weight ratio of about 1000: 1 to about 1:10 catalyst to silicone compound.

  The silicone compound may be provided in the form of a solution or an emulsion under conditions of contact with the catalyst. The deposited silicone compound may cover and spread substantially exclusively on the outer surface of the catalyst, blocking the outer site and partially blocking the site in or near the narrow mouth and pore openings. Examples of methods for depositing silicone compounds on the surface of zeolites can be found in US Pat. Nos. 4,090,981; 5,243,117; 5,403,800 and 5,590,098, which are incorporated herein by reference.

  The catalyst is treated with one or more coatings with a silicone compound, and each coating is then fired. The coated (or coated or coated) catalyst may include silica and alumina. The molar ratio of silica to alumina may be about 10: 1 to about 50: 1, about 10: 1 to about 40: 1, about 10: 1 to about 20: 1, or about 15: 1. The coated catalyst may further comprise nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper, chromium, gallium, one or more oxides thereof, or a mixture of two or more thereof.

で表される少なくとも１つの基を含んで成る化合物を含んで成ってよい。
シリコーン化合物は式

で表されてよい。Ｒ_１およびＲ_２は独立して、水素、ハロゲン、ヒドロキシル、アルキル、アルコキシル、ハロゲン化アルキル、アリール、ハロゲン化アリール、アラルキル、ハロゲン化アラルキル、アルカリル又はハロゲン化アルカリルを含んで成る。ｎは少なくとも２である数である。Ｒ_１および/またはＲ_２はメチル、エチルまたはフェニルを含んで成ってよい。ｎは約３〜約１０００の数であってよい。 The term “silicone compound” is used herein to refer to any compound comprising one or more Si—O groups. The silicone compound may be a silicate (or silicate) comprising one or more SiO ₄ ⁴⁻ , Si ₂ O ₇ ⁶⁻ or Si ₆ O ₁₈ ¹²⁻ groups. These may include one or more tetraorthosilicates. The silicone compound comprises one or more siloxanes comprising one or more silicon-oxygen skeletons (—Si—O—Si—O—) having organic (eg, hydrocarbon) side groups bonded to silicon atoms. It's okay. These may comprise one or more siloxane polymers (eg, polydimethylsiloxane). The silicone compound may be a linear, branched or cyclic compound. The silicone compound may be a monomer, oligomer or polymer. Silicone compound has the formula

It may comprise a compound comprising at least one group represented by:
Silicone compounds are of formula

It may be expressed as R ₁ and R ₂ independently comprise hydrogen, halogen, hydroxyl, alkyl, alkoxyl, alkyl halide, aryl, aryl halide, aralkyl, aralkyl halide, alkaryl or alkaryl halide. n is a number that is at least 2. R ₁ and / or R ₂ may comprise methyl, ethyl or phenyl. n may be a number from about 3 to about 1000.

  The silicone compound may have a number average molecular weight of about 80 to about 20,000 or about 150 to 10,000.

  Silicone compounds are dimethyl silicone, diethyl silicone, phenyl methyl silicone, methyl hydrogen silicone, ethyl hydrogen silicone, phenyl hydrogen silicone, methyl ethyl silicone, phenyl ethyl silicone, diphenyl silicone, methyl trifluoropropyl silicone, ethyl trifluoropropyl silicone, poly Dimethyl silicone, tetrachlorophenyl methyl silicone, tetrachlorophenyl ethyl silicone, tetrachlorophenyl hydrogen silicone, tetrachlorophenyl phenyl silicone, methyl vinyl silicone, hexamethylcyclotrisiloxane, octamethylcyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenylcyclotetra Rokisan or may comprise of two or more of these mixtures.

  The silicone compound may include tetraorthosilicate. The silicone compound may comprise tetramethylorthosilicate, tetraethylorthosilicate, or a mixture thereof.

  The kinetic diameter of the silicone compound may be larger than the pore diameter of the catalyst to avoid simultaneous entry of the silicone compound into the pores and a simultaneous decrease in the internal activity of the catalyst.

  Organic supports for silicone compounds may comprise hydrocarbons such as straight chain, branched chain, and cyclic alkanes having 5 or more carbons. The carrier may comprise a linear, branched or cyclic alkane having a boiling point greater than about 70 ° C. and comprising about 6 or more carbons. Optionally, a mixture of low volatility organic compounds such as hydrocracked circulating oil may be used as the carrier. The low volatility hydrocarbon carrier (or carrier) for the silicone compound may include decane, dodecane, mixtures thereof, and the like.

  After depositing each of the silicone compounds, the catalyst may be calcined to decompose the molecular or polymer species into solid state species. The catalyst may be calcined at a rate of about 0.2 ° C./min to about 5 ° C./min above about 200 ° C. but below a temperature at which the crystallinity of the zeolite may be adversely affected. Generally, such temperatures are below about 600 ° C. The firing temperature may be from about 350 ° C to about 550 ° C. The catalyst may be maintained at the calcination temperature for about 1 to 24 hours, or about 2 to about 6 hours.

  The catalyst may be treated with tetraorthosilicate using a chemical liquid deposition (CLD) process. The tetraorthosilicate may comprise tetramethylorthosilicate, tetraethylorthosilicate (TEOS), or mixtures thereof. The CLD process involves dispersing the catalyst in a liquid medium (eg, hexane, alkane, aromatic or other non-polar organic solvent) at a concentration of about 0.1 to about 20% by weight or about 1 to about 10% by weight. Providing tetraorthosilicate to provide a mixture comprising from about 0.01 to about 5% by weight or from about 0.1 to about 1% by weight of tetraorthosilicate; with stirring (eg, from about 50 to Refluxing the resulting mixture at an elevated temperature (about 150 ° C. or about 90 ° C.); recovering the catalyst from the liquid medium (eg, through centrifugation); drying the catalyst; and about 1 to about 24 Calcination of the catalyst in air at a temperature of about 100 to about 550 ° C. or about 150 ° C. to about 325 ° C. for a period of time or about 2 to about 12 hours. This procedure may be repeated any desired number of times (eg, 1, 2, 3 additions, etc.) to provide the desired treatment layer obtained from tetraorthosilicate. In using TEOS as a tetraorthosilicate, this process can be referred to as TEOS CLD silylation.

  While not wishing to be bound by theory, the benefits of treatment with silicone compounds are that the active catalyst sites on the outer surface of the catalyst are substantially inaccessible to the reactants and the size of the pore opening is reduced. It is believed to be obtained in part by increasing the twist. Active catalytic sites present on the outer surface of the catalyst are believed to isomerize the para-isomer until it reaches an equilibrium level with the other two isomers. Thus, by reducing the usefulness (or availability) of these active catalyst sites, a relatively high proportion of para-xylene can be maintained. The silicone compounds of the present invention can block these external catalyst sites by chemically modifying, coating or hiding the active catalyst sites, or make these external catalyst sites unavailable for para-isomers. It is thought that it becomes.

  Reactants (eg, solid hydrocarbonaceous materials and / or non-solid reactants) in the reactor and / or under defined reaction conditions (under conditions where the reactants are pyrolyzed or catalyzed in a given reactor system) It is also beneficial to control the residence time of the catalyst.

  The term “overall residence time” refers to any reactor or device exiting a reactor or device comprising fluidized gas, product, and impurities, measured or calculated at the average temperature of the reactor or device and the outlet pressure of the reactor or device. The gas outlet flow rate refers to the reactor or device or the volume of a specific part of the reactor or device divided.

  The term “reactant residence time” of a reactant in the reactor is defined as the amount of time that the reactant spends in the reactor. The residence time is based on the supply rate of the reactants and may be independent of the reaction rate. The reactant residence time of the reactants in the reactor may be calculated using different methods depending on the type of reactor used. For gaseous reactants with known flow rates in the reactor, this is typically a simple calculation. In the case of solid reactants, where the reactor comprises a packed bed reactor in which only the reactants are fed continuously (ie, no carrier or fluid flow is utilized), the residence time of the reactants in the reactor Can be calculated by dividing the volume of the reactor by the volumetric flow rate of the hydrocarbonaceous material and the fluid hydrocarbon product exiting the reactor.

  When a reaction occurs in a reactor (eg, a batch reactor) that is closed to a large flow during operation, the batch (or batch) residence time of the reactants in such a reactor is determined within the reactor comprising the reactants. The time at which the temperature reaches a level sufficient to initiate a pyrolysis reaction (e.g., typically about 300 ° C. to about 1000 ° C. for many typical hydrocarbonaceous materials in the case of CFP) and The reactor is quenched (e.g., cooled to a temperature low enough to support further pyrolysis, e.g., typically from about 300 <0> C to about 1000 <0> C for many hydrocarbonaceous materials) It may be defined as the amount of elapsed time between hours.

  In some cases, for example, for certain fluidized bed reactors, the reactor feed stream may comprise a feed stream comprising auxiliary materials (ie, substances other than solid hydrocarbonaceous materials and / or non-solid reactants). For example, in some cases where a fluidized bed is used as a reactor, the feed stream may contain a fluidizing fluid. When a circulating fluidized bed is used, both the catalyst and fluidizing fluid may be fed to the reactor, recycled, or fed and recycled. In such cases, the reactant residence time of the reactants in the reactor is the reactor volume divided by the volume flow rate of reactants and reaction product gas exiting the reactor, similar to the packed bed situation described above. Can be determined. However, since it may not be convenient to directly determine the volume flow of reactant and reaction product gas exiting the reactor, the volume flow of reactant and reaction product gas exiting the reactor is dependent on the total volume of the gas stream exiting the reactor. It may be estimated by subtracting the supply volume flow rate of auxiliary material (eg, fluidizing fluid, catalyst, contaminants, etc.) in the reaction product gas reactor from the flow rate.

The term “selectivity” refers to the amount of production of a particular product compared to product selection. Selectivity for a product can be calculated by dividing a particular product amount by the amount of product produced. For example, if 75 grams of aromatics are produced in the reaction and 20 grams of benzene is found in these aromatics, the selectivity for benzene in the aromatic product on a mass basis is 20/75 = 26.7% It is. Selectivity is calculated on a mass basis as in the example above, or if selectivity is calculated by dividing the amount of carbon found in a particular product by the amount of carbon found in product selection, Selectivity can be calculated on a carbon basis. Unless otherwise specified, selectivity is mass based for reactions comprising biomass as a reactant. For reactions that involve the conversion of certain molecular reactants (eg, ethene), selectivity is calculated by dividing the selected product by all products produced (unless otherwise specified, on a mass basis). ) Percentage. The selectivity of the various materials can be determined using the following formula:
(1) Total selectivity = (carbon mole in product / carbon mole in all product) × 100%
(2) Aromatic selectivity = (carbon mole in aromatic product / carbon mole in total aromatic product) × 100%
(3) Olefin selectivity = (carbon mole in olefin product / carbon mole in all olefin product) × 100%
(4) Selectivity of p-xylene in xylene = (mole of p-xylene isomer / mole of all xylene isomers) × 100%

The term “yield” is used herein to refer to the amount of reactant flowing into the reactor divided by the amount of product leaving the reactor, and is usually expressed as a percentage or fraction. Yields are often calculated based on mass basis, carbon basis, or specific feed component basis. Mass yield is the mass of a particular product divided by the weight of the raw material used to prepare the product. For example, if 500 g of biomass is fed to the reactor and 45 g of p-xylene is produced, the mass yield of p-xylene is 45/500 = 9%. Carbon yield is the mass of carbon found in a particular product divided by the mass of carbon in the feed to the reactor. For example, if 500 g of biomass containing 40% carbon reacts to produce 45 g of p-xylene containing 90.6% carbon, the carbon yield is [(45 × 0.906. ) / (500 × 0.40)] = 20.4%. Carbon yield from biomass is the mass of carbon found in a particular product divided by the mass of carbon fed to the reactor in a particular feed component. For example, 500 grams of 40% by reacting C0 ₂ biomass 100 grams comprising carbon, and (comprising 90.6% carbon) when generating a p- xylene 40 g, the carbon of the biomass The yield is [(40 × 0.906) / (500 × 0.40)] = 18.1%. Note that the mass of CO ₂ is not taken into account.

  The embodiments described herein may also include a chemical process design used to perform catalytic pyrolysis. The process may include the use of one or more fluidized bed reactors (eg, circulating fluidized bed reactor, turbulent fluidized bed reactor, bubbling fluidized bed reactor, etc.). The process design described herein may optionally include special handling of materials fed to one or more reactors. For example, the feed material may be dried, cooled, and / or ground prior to feeding the material to the reactor. Other aspects of the invention may relate to product compositions produced using the process design described herein.

Without being bound to a particular action mode or sequence of steps in the overall thermal / catalytic conversion process, catalytic pyrolysis can produce one or more pyrolysis products (eg, volatile organics, gas, solid coke, etc.). One using a catalyst under reaction conditions sufficient to produce at least partial pyrolysis of a hydrocarbonaceous material to produce (eg, solid biomass such as cellulose) and a fluid hydrocarbon product It is believed to comprise at least a portion of the catalytic reaction of the above pyrolysis product. Catalytic reactions occur in catalysts where volatile organics are converted to p-xylene and other hydrocarbons, such as aromatics, and olefins (eg, zeolite catalysts) in addition to carbon monoxide, carbon dioxide, water and coke, for example. May contain volatile organics. During or upon contact with the catalyst, the pyrolysis product undergoes a series of dehydration, decarbonylation, decarboxylation, isomerization, oligomerization, and dehydrogenation reactions leading to olefins, CO, CO ₂ and water. You can receive it.

  The catalyst provided herein is at least about 40%, at least about 45%, at least about 50%, or at least about 55%, or at least about 60%, at least about 65, relative to p-xylene in xylene. %, At least about 70%, at least about 75%, about at least about 80%, at least about 85%, or at least 90% may be particularly suitable for producing xylene having a relatively high selectivity. .

  FIG. 1A comprises a schematic diagram of a typical chemical process design used to perform catalytic pyrolysis according to the method of the present invention. The process can include a CFP process. Referring to FIG. 1A, the feed stream 10 includes solid hydrocarbonaceous material that can be fed to the reactor 20. The solid hydrocarbonaceous material may generally comprise at least carbon and hydrogen. In certain solid hydrocarbonaceous materials (eg wood), carbon may be the most abundant component in mass, and in other cases (eg glucose etc.) there may be more oxygen than carbon. Further, the specific solid hydrocarbon-based material may include a material having a relatively small ratio of other elements such as nitrogen and sulfur.

  The feed stream to the reactor may be free of olefins, or (eg, olefins are less than about 1%, less than about 0.1%, or about 0.01% by weight of the total weight of reactants fed to the reactor. It may comprise a small amount of olefins so that it comprises less than%. However, in other embodiments, olefins may be present in one or more reactant feed streams.

  The solid hydrocarbonaceous material feed composition (in feed stream 10 of FIG. 1A) may comprise a mixture of solid hydrocarbonaceous material and catalyst. The mixture may comprise, for example, a solid catalyst and a solid hydrocarbonaceous material. In other embodiments, the catalyst may be provided separately from the solid hydrocarbonaceous material (eg, by supplying the catalyst via a separate catalyst inlet). A variety of catalysts may be used. For example, in some instances, zeolites with varying silica to alumina molar ratios and / or varying pore size and / or pore opening size and / or varying catalytically active metals and / or metal oxides A catalyst may be used.

  Moisture 12 may be removed from the solid hydrocarbonaceous feed composition prior to being fed to the reactor, for example by an optional dryer 14. The removal of moisture from the solid hydrocarbonaceous material feed stream can be advantageous for several reasons. For example, moisture in the feed stream may require additional energy input to heat the solid hydrocarbonaceous material to a temperature high enough to achieve pyrolysis. Changes in the moisture content of the solid hydrocarbon feed can lead to difficulties in controlling the reactor temperature. Furthermore, the removal of moisture from the solid hydrocarbon feed can reduce or eliminate the need to treat the water during subsequent processing steps.

  The solid hydrocarbon feedstock composition until the solid hydrocarbon feedstock composition comprises less than about 10 wt%, less than about 5 wt%, less than about 2 wt%, or less than about 1 wt% water. May be dried. Suitable equipment that can remove water from the feed composition is known to those skilled in the art. By way of example, the dryer may be at a specific temperature (eg, at least about 80 ° C., at least about 100 ° C., at least about at least) where the solid hydrocarbonaceous feed composition passes continuously, semi-continuously, or periodically. It may comprise an oven heated to a temperature of 150 ° C. or higher. The dryer may comprise a vacuum chamber in which solid hydrocarbon feedstock composition may be processed as a batch. Drying may combine high temperature with vacuum operation. The dryer may be integrally connected to the reactor or may be provided as a separate unit from the reactor.

  The particle size of the solid hydrocarbon feedstock composition may be reduced with an optional grinding system 16 prior to passing the solid hydrocarbon feedstock through the reactor. The average diameter of the pulverized solid hydrocarbon feedstock composition exiting the pulverization system is about 50% or less, about 25% or less, about 10% or less of the average diameter of the raw material composition fed to the pulverization system, It may contain about 5% or less, about 2% or less. Large particle solid hydrocarbon feeds may be easier to transport and handle than small particle feeds. On the other hand, in some cases it may be advantageous to feed the reactor with small particulate solid hydrocarbonaceous material. The use of a grinding system allows the transport of large particles of solid hydrocarbonaceous feedstock between the source and the process and allows the supply of small particles to the reactor.

  Suitable equipment capable of grinding the solid hydrocarbon feedstock composition is known to those skilled in the art. For example, the grinding system may include an industrial mill (eg, hammer mill, ball mill, etc.), a bladed unit (eg, chipper, shredder, etc.), or any other suitable type of grinding system. The grinding system may include a cooling system (eg, an active cooling system such as a pumped fluid heat exchanger, an active cooling system comprising fins, etc.) and introduces a solid hydrocarbonaceous feed composition into the reactor. It may be used to maintain a solid hydrocarbonaceous feed composition at a relatively low temperature (eg, ambient temperature) previously. The grinding system may be integrally connected to the reactor or may be provided as a unit separate from the reactor. Although the grinding process is shown after the drying process of FIG. 1A, the order of these operations may be reversed in certain embodiments. In yet another aspect, the drying and grinding steps may be done using an integral unit.

  Grinding and cooling of the solid hydrocarbonaceous material may be accomplished using separate units. Cooling of the solid hydrocarbonaceous material is desirable, for example, to reduce or prevent undesirable decomposition of the solid hydrocarbonaceous feedstock prior to passing the solid hydrocarbonaceous material through the reactor. The solid hydrocarbonaceous material may be passed through a grinding system to produce a ground solid hydrocarbonaceous material. The ground solid hydrocarbonaceous material may be cooled from the grinding system through a cooling system. The solid hydrocarbonaceous material is less than about 300 ° C, about 200 ° C, about 100 ° C, about 75 ° C, about 50 ° C, about 35 ° C or about 20 ° C before introducing the solid hydrocarbonaceous material into the reactor. It may be cooled to temperature. The cooling system may comprise an active cooling unit (eg, a heat exchanger) that can reduce the temperature of the solid hydrocarbonaceous material. Two or more drying, grinding systems, and cooling systems may be combined in a single unit. The cooling system may be directly integrated with one or more reactors.

  The hydrocarbonaceous material may be transferred to the reactor 20. In one example, the reactor is for performing catalytic pyrolysis of at least a portion of a first reactant comprising a hydrocarbonaceous material under reaction conditions sufficient to produce one or more pyrolysis products. May be used. In the exemplary embodiment of FIG. 1A, the reactor comprises any suitable reactor known to those skilled in the art. For example, in certain examples, the reactor may include, among other things, a continuous stirred tank reactor (CSTR), a batch reactor, a semi-batch reactor, or a fixed bed reactor. In one example, the reactor includes a fluidized bed reactor, such as a circulating fluidized bed reactor. In one example, the fluidized bed reactor may provide improved mixing of the catalyst and solid hydrocarbonaceous material during pyrolysis and / or subsequent reaction, which may lead to improved control of the reaction products formed. Also, the use of a fluidized bed reactor can result in improved heat transfer within the reactor. Further, in certain instances, improved mixing in a fluidized bed reactor may lead to a reduction in the amount of coke deposited on the catalyst, thereby reducing catalyst deactivation.

  The reactor may have any suitable size for performing the methods described herein. For example, the reactor is about 0.1-1 L, 1-50 L, 50-100 L, 100-250 L, 250-500 L, 500-1000 L, 1000-5000 L, 5000-10,000 L, or 10,000-50,000 L. May have a volume of. In some examples, the reactor may have a volume that is greater than about 1L, and in other examples, greater than about 10L, 50L, 100L, 250L, 500L, 1000L, or 10,000L. A reactor volume greater than about 50,000 L may also be possible. The reactor may be cylindrical, spherical, or any other suitable shape.

  High yields of desired product formation, low yields of coke formation, and / or more controlled production when certain combinations of reaction conditions and system components are performed in the methods and systems described herein. Product formation (high production of p-xylene over other products) can be achieved. For example, normalized mass space velocity (for solid hydrocarbonaceous materials and / or fluidizing fluid), reactor and / or solid separator (or separator) temperature, reactor pressure, feed stream heating rate, catalyst and solid hydrocarbon The mass ratio of the system material, the residence time of the hydrocarbon material in the reactor, the residence time of the reaction product in the solid separator, and / or the type of catalyst (as well as the molar ratio of silica and alumina and the pore size (Aperture size) may be controlled to achieve beneficial results, as described below.

  The reactor may be operated at any suitable temperature. In certain instances, it may be desirable to operate the reactor at an intermediate temperature compared to the temperatures typically used in many conventional catalytic pyrolysis systems. For example, the reactor may be operated at a temperature of about 400 ° C to about 600 ° C, about 425 ° C to about 500 ° C, or about 440 ° C to about 460 ° C. By operating the reactor at these intermediate temperatures, it is possible to maximize the amount of desired product. However, the present invention is not limited to the use of such intermediate temperatures, and in other embodiments, lower and / or higher temperatures can be used.

  The reactor may also be operated at any suitable pressure. The reactor may be operated at a pressure of at least about 100 kPa, or at least about 200 kPa, at least about 300 kPa, or at least about 400 kPa. The reactor may be operated at a pressure of less than about 600 kPa, less than about 400 kPa, or less than about 200 kPa. The reactor may be operated at a pressure of about 100 to about 600 kPa, about 100 to about 400 kPa, or about 100 to about 200 kPa. However, the invention is not limited to the use of such pressures, and in other embodiments lower and / or higher pressures may be used.

  As the feed stream enters the reactor, it may be advantageous to heat the feed stream at a relatively fast rate. A fast (or high) heating rate may be advantageous for a number of reasons. For example, a fast heating rate can increase the mass transfer rate of reactants from the bulk solid hydrocarbonaceous material to the catalytic reaction site. This may be due to, for example, the pyrolysis of the solid hydrocarbonaceous material in the catalyst prior to complete pyrolysis of the solid hydrocarbonaceous material and / or the second reactant to a generally undesirable product (eg, coke). The introduction of volatile organic compounds formed therebetween can be facilitated. Furthermore, the fast heating rate can reduce the amount of time that the reactants are exposed to a low temperature (ie, a temperature between the feed temperature and the desired reaction temperature). Prolonged exposure of the reactants to low temperatures can lead to undesirable product formation through undesirable degradation and / or reaction pathways. Examples of suitable heating rates for heating the feed stream upon entering the reactor include greater than about 50 ° C / second, greater than about 100 ° C / second, greater than about 200 ° C / second, about 300 ° C / second. Greater than about 400 ° C / second, greater than about 500 ° C / second, greater than about 600 ° C / second, greater than about 700 ° C / second, greater than about 800 ° C / second, greater than about 900 ° C / second Heating rates greater than or greater than about 1000 ° C./second. In some cases, the reactants may be heated at a heating rate between about 500 ° C./second and about 1000 ° C./second. In certain embodiments, the heating rate for heating the feed stream upon entering the reactor may be from about 50 ° C./second to about 1000 ° C./second, or from about 50 ° C./second to about 400 ° C./second. However, the present invention is not limited to using such heating rates, and in other embodiments, lower and / or higher heating rates can be used.

  The normalized mass space velocity of the hydrocarbonaceous material may be selected to selectively produce the desired array (or array) of flowable hydrocarbon products. As used herein, the term “normalized mass space velocity” refers to the mass flow rate of a component (eg, measured in g / hour) into the reactor (eg, measured in grams). It is defined as the catalyst mass in the reactor divided by and has units of reciprocal time. For example, the normalized mass space velocity of the solid hydrocarbonaceous material fed to the reactor can be calculated as the mass flow rate of the solid hydrocarbonaceous material into the reactor divided by the mass of the catalyst in the reactor. The normalized mass space velocity of the components in the reactor (eg, hydrocarbonaceous material) can be calculated using different methods depending on the type of reactor used. For example, in a system using a batch or semi-batch reactor, solid hydrocarbonaceous material is not continuously fed to the reactor and the solid hydrocarbonaceous material does not have a normalized mass space velocity. For systems where the catalyst is fed into and / or removed from the reactor (eg, a circulating fluidized bed reactor) during the reaction, the normalized mass space velocity is the value of the reactor during a period of operation (eg, steady state operation). It can be determined by calculating the average amount of catalyst in the volume.

The normalized mass space velocity of the hydrocarbonaceous material fed to the reactor is about 3 hours- ¹ or less, about 2 hours- ¹ or less, about 1.5 hours- ¹ or less, about 0.9 hours- ¹ or less, about 0.01 hours- ¹ to about 3 hours- ¹ , about 0.01 hours- ¹ to about 2 hours- ¹ , about 0.01 hours- ¹ to about 1.5 hours- ¹ , about 0.01 hours- ¹ About 0.9 hours ⁻¹ , about 0.01 hours ⁻¹ to about 0.5 hours ⁻¹ , about 0.1 hours ⁻¹ to about 0.9 hours ⁻¹ , or about 0.1 hours ⁻¹ The reactor is fed with a normalized mass space velocity of about 0.5 hours- ¹ . However, the invention is not limited to using such normalized mass space velocities, and in other embodiments, lower and / or higher normalized mass space velocities can be used.

The residence time (ie, reactant residence time) of the reactant (eg, hydrocarbonaceous material) within the reactor is at least about 1 second, at least about 2 seconds, at least about 5 seconds, at least about 7 seconds, at least about 10 seconds. At least about 15 seconds, at least about 20 seconds, at least about 25 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds. In some cases, the residence time of the reactant (eg, hydrocarbonaceous material) in the reactor is less than about 5 minutes, about 1 second to about 4 minutes, about 2 seconds to about 4 minutes, about 5 seconds to about 4 minutes. From about 7 seconds to about 4 minutes, from about 10 seconds to about 4 minutes, from about 12 seconds to about 4 minutes, from about 15 seconds to about 4 minutes, from about 20 seconds to about 4 minutes, from about 30 seconds to about 4 minutes, or It may be from about 60 seconds to about 4 minutes.
The aforementioned “rapid pyrolysis” studies often used systems with very short reactant residence times (eg, less than 2 seconds). However, in some cases, the use of relatively long residence times may allow additional chemical reactions to form the desired product. Long residence times can be achieved, for example, by increasing the volume of the reactor and / or decreasing the volumetric flow rate of the hydrocarbonaceous material. However, in certain aspects described herein, the residence time of the reactant (eg, hydrocarbonaceous material) may be relatively short, for example, less than about 2 seconds, or less than about 1 second.

  The contact time between the pyrolysis product (eg, pyrolysis steam) and the catalyst in the reactor is at least about 1 second, at least about 2 seconds, at least about 5 seconds, at least about 7 seconds, at least about 10 seconds, at least about 15 Second, at least about 20 seconds, at least about 25 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds. The contact time is less than about 5 minutes, about 1 second to about 4 minutes, about 2 seconds to about 4 minutes, about 5 seconds to about 4 minutes, about 7 seconds to about 4 minutes, about 10 seconds to about 4 minutes, It may be from about 12 seconds to about 4 minutes, from about 15 seconds to about 4 minutes, from about 20 seconds to about 4 minutes, from about 30 seconds to about 4 minutes, or from about 60 seconds to about 4 minutes.

  In one example where a fluidized bed reactor is used, the feedstock (eg, solid hydrocarbonaceous material) in the reactor may be fluidized by flowing a fluid stream through the reactor. In the exemplary embodiment of FIG. 1A, fluid stream 44 is used to fluidize the feedstock within reactor 20. Fluid may be supplied to the fluid stream from the fluid source 24 and / or from the reactor product stream through the compressor 26. As used herein, the term “fluid” means a raw material that is generally in a liquid, supercritical or gaseous state. However, the fluid may comprise solids such as, for example, suspended particles or colloidal particles. In certain embodiments, it may be advantageous to control the residence time of the fluidizing fluid in the reactor. The residence time of the fluidizing fluid can be defined as the volume of the reactor divided by the volumetric flow rate of the fluidizing fluid. The fluidizing fluid residence time is at least about 0.1 seconds, at least about 0.2 seconds, at least about 0.5 seconds, at least about 1 second, at least about 2 seconds, at least about 3 seconds, at least about 4 seconds, at least It may be about 5 seconds, at least about 6 seconds, at least about 8 seconds, at least about 10 seconds, at least about 12 seconds, at least about 24 seconds, or at least about 48 seconds. The fluidizing fluid residence time is about 0.1 second to about 48 seconds, about 0.2 seconds to about 48 seconds, about 0.5 seconds to about 480 seconds, about 1 second to about 48 seconds, about 3 seconds to About 48 seconds, about 5 seconds to about 48 seconds, about 6 seconds to about 48 seconds, about 8 seconds to about 48 seconds, about 10 seconds to about 48 seconds, about 12 seconds to about 48 seconds, or about 24 seconds to about It may be 48 seconds.

  Suitable fluidizing fluids that can be used in the present invention include, for example, inert gases (eg, helium, argon, neon, etc.), hydrogen, nitrogen, carbon monoxide, and carbon dioxide.

  As shown in the exemplary embodiment of FIG. 1A, products (eg, fluidized hydrocarbon products) formed during the reaction of reactants (eg, solid hydrocarbonaceous material) exit the reactor through product stream 30. In addition to the reaction product, the product stream may in some cases comprise unreacted reactants, fluidizing fluid, and / or catalyst. In certain embodiments, desired reaction products (eg, liquid aromatic hydrocarbons, olefinic hydrocarbons, gaseous products, etc.) may be recovered from the reactor effluent stream.

  As shown in the exemplary embodiment of FIG. 1A, the product stream 30 may be fed to an optional solids separation device 32. A solid separator may be used in some cases to separate reaction products from a catalyst (eg, an at least partially deactivated catalyst) present in the product stream. Furthermore, the solids separator may be used in some instances to remove coke and / or ash from the catalyst. In certain embodiments, the solid separator may include an optional purge stream 33 that may be used to purge coke, ash, and / or catalyst from the solid separator.

  The equipment required to accomplish the solid separation and / or decarbonization process can be easily designed by those skilled in the art. For example, the solid separation device may include a container including a mesh material that defines a container holding portion and a permeation portion. The mesh can serve to retain the catalyst in the retainer and allow reaction products to pass through the permeate. The catalyst exits the solids separator through the mesh holding port while the reaction product may exit the mesh permeation port. Other examples of solid state separators and / or decarbonizers are Kirkusmer's Encyclopedia of Chemical Technology (online), Vol. 11, Hoboken, NJ. Willie-Interscience, 2001, pages 700-734, and C. D. Cooper and F. C. Array Air Pollution Prevention, Design Approach, 2nd Edition Prospect Heights, Illinois: Weibrand Press, Inc., l994, pages 127-149, which are described in more detail Which is incorporated herein by reference.

  The solid separation device may be operated at any suitable temperature. In some embodiments, the solid separation device may be operated at an elevated temperature. For certain reactions, the solids separator can be elevated to allow additional modification and / or reaction of the compound from the reactor. This can increase the desired product formation. While not being bound by any theory, it is believed that sufficient energy can be provided to cause the endothermic reforming reaction by raising the temperature in the solid separator. The solid separation device may be operated at a temperature of, for example, about 25 ° C to about 200 ° C, about 200 ° C to about 500 ° C, about 500 ° C to about 600 ° C, or about 600 ° C to about 800 ° C. In some cases, the solid separation device may be operated at a temperature of at least about 500 ° C, at least about 600 ° C, at least 700 ° C, at least 800 ° C, or higher.

It may be beneficial to control the residence time of the catalyst in the solids separator. The residence time of the catalyst in the solid separator can be defined as the volume of the solid separator divided by the volume flow of the catalyst through the solid separator. In some cases, a relatively long residence time of the catalyst in the solids separator may be desired to facilitate removal of a sufficient amount of ash, coke, and / or other undesirable products from the catalyst. . Further, by employing a relatively long residence time of the catalyst in the solids separator, the pyrolysis product may be further reacted to produce the desired product.
The residence time and temperature in the solids separator may be selected integrally so that the desired product stream is produced. The residence time of the catalyst in the solid separator is at least about 1 second, at least about 5 seconds, at least about 7 seconds, at least about 10 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds. , At least about 300 seconds, at least about 600 seconds, or at least 1200 seconds. Methods for controlling the residence time of the catalyst in the solid separator are known to those skilled in the art. For example, in some cases, the inner wall of the solid separator has a baffle that serves to limit the flow of catalyst through the solid separator and / or increase the path length of the fluid flow within the solid separator. You can do it. Additionally or alternatively, the residence time of the catalyst in the solid separation device controls the flow rate of the catalyst through the solid separation device (eg, by controlling the flow rate of the fluidizing fluid through the reactor). It may be controlled by.

  The solid separation device may have any suitable size. For example, the solid separator has a volume of about 0.1 to 1L, 1 to 50L, 50 to 100L, 100 to 250L, 250 to 500L, 500 to 1000L, 1000 to 5000L, 5000 to 10,000L. Good. In some examples, the solid separation device may have a volume greater than about 1 L, and in other examples, the solids separation device has a volume greater than about 10 L, 50 L, 100 L, 250 L, 500 L, 1,000 L, or 10,000 L. You can do it. Also, a solid separator volume greater than 50,000 L is possible. The solid separation device may be cylindrical, spherical, or any other shape, and may be cyclic or non-circular. In some embodiments, the solid separation device may include vessels or other unit operations similar to those used for one or more reactors used in the process. The catalyst flow path in the solid state separation device may have any suitable shape. For example, the flow path may be substantially straight. In some cases, the solid separation device may comprise a flow path having a meander, bend, spiral, or any other suitable shape. The ratio of the flow path length of the solid separator (or, in certain embodiments, the path length of the catalyst through the solid separator) to the average diameter of the solid separator channel includes any suitable ratio. Good. The ratio may be at least about 2: 1, at least 5: 1, at least 10: 1, at least 50: 1, at least 100: 1, or more.

  A solid separation device may not be required in all embodiments. For example, in situations where a fixed catalyst bed reactor is used, the catalyst may be retained in the reactor and the reaction product may exit the reactor that is substantially free of catalyst, thereby eliminating the need for a separation step.

  The separated catalyst may exit the solids separator via stream 34. A portion of the separated catalyst may be returned to the reactor via a return pipe (or return pipe), not shown in FIG. 1A. The catalyst exiting the solids separator may be at least partially deactivated. The separated catalyst may be fed to a regenerator 36 that can reactivate any catalyst that has been at least partially deactivated. The regenerator may include an optional purge stream 37 that may be used to purge coke, ash, and / or catalyst from the regenerator. Methods for activating catalysts are described, for example, in the Kirk Osmer Chemical Technology Encyclopedia (online), Vol. 5, Hoboken, NJ: Willi Interscience, 2001-255, which is incorporated herein by reference. It is well known to those skilled in the art as described on pages 322.

  A portion of the catalyst may be removed from the reactor through the catalyst outlet (not shown in FIG. 1A). The catalyst removed from the reactor may be partially deactivated and passed through a conduit to a regenerator 36 or another regenerator (not shown in FIG. 1A). The regenerated removed catalyst may be returned to the reactor via stream 47 or may be returned to the reactor separately from the fluidizing gas via another stream (not shown in FIG. 1A). Good.

  For example, as shown in FIG. 1A, oxidant may be supplied to the regenerator via stream 38. The oxidant may be derived from any source comprising, for example, oxygen, air, steam tanks, among others. The reactivated catalyst may comprise residual carbon and / or coke that may be removed via reaction with an oxidant in the regenerator. The regenerator in FIG. 1A comprises a vent stream 40 comprising a regeneration reaction product, residual oxidant, and the like.

  The regenerator may be of any suitable size as described above in connection with the reactor or solid separation device. Further, the regenerator may be operated at elevated temperatures (eg, temperatures of at least about 300 ° C, 400 ° C, 500 ° C, 600 ° C, 700 ° C, 800 ° C, or higher) in some cases. The residence time of the catalyst in the regenerator may be controlled using methods well known to those skilled in the art, including those outlined above. The mass flow rate of the catalyst through the regenerator may be related to the flow rate in the reactor and / or solids separator to maintain the mass balance in the system.

  Regenerated catalyst may exit the regenerator through stream 42. The regenerated catalyst may be recycled back to the reactor through the regeneration stream 47. In some cases, the catalyst may be lost from the system or may be intentionally removed during operation. Additional “replenished (or created)” catalyst may be added to the system via a replenishment stream 46. The regeneration and make-up catalyst may be fed to the reactor along with the fluidizing fluid via regeneration stream 47. Alternatively, the catalyst and fluidizing fluid may be fed to the reactor via separate streams.

  Referring to solid separator 32 in FIG. 1A, the reaction product (eg, fluidized hydrocarbon product) may exit the solid separator via stream 48. In some cases, a portion of stream 48 may be purged via purge stream 60. The contents of the purge stream may be supplied to, for example, a combustor or a water-gas shift reactor to recover energy lost from the system. In some cases, the reaction product in stream 48 may be fed to optional condenser 50. The condenser may include a heat exchanger that condenses at least a part of the reaction product from a gaseous state to a liquid state. A condenser may be used to separate the reaction product into gaseous, liquid, and solid fractions (or fractions). The operation of condensers is well known to those skilled in the art. Examples of condensers that may be used are described in more detail in the Perry Chemical Engineer Handbook, Section 11: “Heat Transfer Device”, 8th Edition, New York: McGraw Hill, 2008, which is incorporated herein by reference. Have been described.

The condenser may utilize pressure changes to condense a portion of the product stream. In FIG. 1A, stream 54 may comprise a liquid fraction of reaction products (eg, water, aromatics, olefin compounds, etc.) and stream 74 may be a gaseous fraction of reaction products (eg, CO 2). , _{CO 2, H} 2, etc.). In some embodiments, the gaseous fraction may be supplied to the steam recovery system 70. The steam recovery system may be used, for example, to recover any desired steam in stream 74 and transport the steam through stream 72. Further, stream 76 may be used to transport CO, CO ₂ , and / or other non-recoverable gases from the steam recovery system. Optional steam recovery systems may be located elsewhere. For example, in certain aspects, the steam recovery system may be located downstream of the purge stream 54. One skilled in the art can select an appropriate arrangement for the steam recovery system.

Other products (eg, excess gas) may be transported via stream 56 to any compressor 26, which other products may be compressed in the reactor and used as fluidizing gas, and / or Alternatively, the other product may assist in transporting the hydrocarbonaceous material to the reactor (stream 58) or may be used to transport the catalyst to the reactor (not shown) for additional It may be used to transport non-solid raw materials.
In certain instances, the liquid fraction may be further processed, for example, separating the aqueous phase from the organic phase or separating individual compounds.

  The set of embodiments described by FIG. 1A includes reactors, solid state separators, regenerators, condensers, etc., but not all embodiments comprise the use of these elements. For example, in certain embodiments, a feed stream may be fed to and reacted with a catalyst fixed bed reactor and the reaction product may be recovered directly from the reactor and cooled without the use of a dedicated condenser. In certain instances, dryers, grinding systems, solid separators, regenerators, condensers, and / or compressors may be used as part of the process, one or more of these elements being fluidly and / or integral. May have a separate unit not connected to the reactor. In other embodiments, one or more of a dryer, a grinding system, a solids separator, a regenerator, a condenser, and / or a compressor may not be present. In certain embodiments, the desired reaction product may be recovered at any point during the production process (eg, after passage through the reactor, after separation, after condensation, etc.).

  The process may involve the use of more than one reactor. For example, as shown in the exemplary embodiment of FIG. 1B, multiple reactors may be connected in fluid communication with each other for operation in series and / or in parallel. The process includes providing a solid hydrocarbonaceous material in the first reactor, and at least a solid hydrocarbonaceous material under reaction conditions sufficient to produce one or more pyrolysis products in the first reactor. Comprising the step of thermally decomposing a portion. The catalyst may be provided to the first reactor and at least a portion of the one or more pyrolysis products in the first reactor is under reaction conditions sufficient to produce one or more fluid hydrocarbon products. Can be reacted catalytically with a catalyst. The process comprises reacting at least a portion of the one or more pyrolysis products in the second reactor with the catalyst under reaction conditions sufficient to produce one or more fluid hydrocarbon products. May further be included. After catalytically reacting at least a portion of the one or more pyrolysis products in the second reactor, the process includes one or more fluid hydrocarbon products from the first reactor in the second reactor. Further reacting to produce at least a portion of one or more other hydrocarbon products.

  In FIG. 1B, the reaction product from the reactor 20 may be transported to the second reactor 20 ′. Those skilled in the art are familiar with the use of multiple reactor systems for the pyrolysis of organic materials to produce organic products, such systems are known in the art. Although FIG. 1B shows a set of aspects in which the reactors are in fluid communication with each other, in one example, the two reactors may not be in fluid communication. For example, the first reactor may be used to produce a first reaction product that may be transported to a separate facility (or facility) for reaction in the second reactor. In one example, a composition comprising a solid hydrocarbonaceous material (with or without a catalyst) may be heated in a first reactor to pyrolyze at least a portion of the solid hydrocarbonaceous material. , Pyrolysis products (and optionally at least partially deactivated catalysts (or deactivated catalysts) may be produced. The first pyrolysis products may be in the form of liquids and / or gases. The composition comprising the first pyrolysis product may then be heated in a second reactor in fluid communication with the first reactor or not in fluid communication with the first reactor. After the process, a second pyrolysis product from the second reactor may be recovered, which may be in the form of a liquid and / or gas, in some cases fed to the first reactor. Hydrocarbon materials The first pyrolysis product produced from the first reactor has a chemical composition greater than the second pyrolysis product, for example, comprising a mixture of solid hydrocarbonaceous material and solid catalyst. The amount, state (eg, gas to fluid) may be different, for example, the first pyrolysis product may comprise substantially liquid while the second pyrolysis product substantially comprises gas. In another example, the first pyrolysis product may include a flowable product (eg, bio-oil, sugar) and the second pyrolysis product is the first pyrolysis product. In certain instances, the first pyrolysis product may comprise a flowable product (eg, comprising an aromatic compound). The second pyrolysis product has a relatively higher amount of olefin than the first pyrolysis product. In yet another example, the first pyrolysis product may comprise a flowable product (eg, bio-oil, sugar) and the second pyrolysis product is a first pyrolysis product. It may comprise relatively more oxygenated aromatics than pyrolysis products.

  One or more reactors in a multiple reactor configuration may be a fluidized bed reactor (eg, a circulating fluidized bed reactor, a turbulent fluidized bed reactor, etc.), or in other examples, any other type of reactor (eg, as described above). Optional reactor). For example, the first reactor may comprise a circulating fluidized bed reactor or a turbulent fluidized bed reactor, and the second reactor comprises a circulating fluidized bed reactor or a turbulent fluidized bed reactor in fluid communication with the first reactor. Further, the configuration of the plurality of reactors may include additional processing steps and / or any of the devices described herein (eg, solid separation devices, regenerators, condensers, etc.). The reactor and / or additional processing equipment may be operated using any of the processing parameters described herein (eg, temperature, residence time, etc.).

  The catalyst components useful in the construction of the present invention are selected from any catalyst known in the art or understood by one of ordinary skill in the art that is aware of the present invention. Functionally, the catalyst is related to or related to dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization, decarbonylation, decarboxylation, aldol condensation, and / or pyrolysis of hydrocarbonaceous materials. May be limited only by the ability of any such material to facilitate and / or achieve any other reaction or process. The catalyst component can be considered acidic, neutral or basic, as will be appreciated by those skilled in the art.

  In some cases, the catalyst particles described herein may comprise polycrystalline solids (eg, polycrystalline particles). Further, the catalyst particles may include a single crystal in some embodiments. In certain cases, the particles may be separate and distinct physical objects. In other cases, the particles may comprise an aggregate of a plurality of individual particles in intimate contact with each other at least at certain times during preparation and / or use.

  The catalysts described herein, for example, used in the feed stream, in the reactor, etc. may be of any suitable size. In some cases, as described above, certain embodiments use a catalyst comprising relatively small catalyst particles that may be in the form of a larger catalyst body composed of a plurality of agglomerated catalyst particles. It can be advantageous. In some embodiments, the use of small catalyst particles, for example, increases the surface area of the outer catalyst and decreases the diffusion distance through the catalyst, thereby reducing the extent to which the hydrocarbonaceous material contacts the surface portion of the catalyst. Can be increased. In some cases, the catalyst size and / or catalyst particle size may be selected based at least in part on, for example, the desired fluid flow type and catalyst life.

  In certain embodiments, the average diameter of the catalyst body (measured by conventional sieving analysis) may include about a single catalyst particle each in a particular example, or may include an aggregate of a plurality of particles in other examples. <5 mm, <2 mm, <1 mm, <500 microns, <60 mesh (250 microns), <100 mesh (149 microns), <140 mesh (105 microns), <170 mesh (88 microns) Less than, less than about 200 mesh (74 microns), less than about 270 mesh (53 microns), less than about 400 mesh (37 microns), or smaller.

  The catalyst may be less than about 5 microns, less than about 1 micron, less than about 500 nm, less than about 100 nm, about 100 nm to about 5 microns, about 500 nm to about 5 microns, about 100 nm to about 1 micron, or about 500 nm to about 1 micron. It may comprise particles having a maximum cross-sectional dimension. Catalyst particles having dimensions within the ranges stated above may be agglomerated to form individual catalyst objects having dimensions within the ranges stated above. As used herein, the “maximum cross-sectional dimension” of a particle refers to the maximum dimension between the two boundaries of the particle. One skilled in the art can determine the maximum cross-sectional dimension of the particles, for example, by analyzing a scanning electron micrograph (SEM) of catalyst preparation. In embodiments comprising agglomerated particles, the particles should be considered separately when determining the maximum cross-sectional dimension. In such cases, the measurement establishes a virtual boundary between each of the agglomerated particles and measures the maximum cross-sectional dimension of the virtual individualized particle resulting from establishing such a boundary. May be done. In certain embodiments, a relatively large number of particles within the catalyst may have a maximum cross-sectional dimension that is within a predetermined range. For example, in certain embodiments, at least about 50%, at least about 75%, at least about 90%, at least about 95%, or at least about 99% of the particles in the catalyst are less than about 5 microns, less than about 1 micron, about 500 nm. Less than, less than about 100 nm, from about 100 nm to about 5 microns, from about 500 nm to about 5 microns, from about 100 nm to about 1 micron, or from about 500 nm to about 1 micron.

  A relatively large proportion of the catalyst volume may be occupied by particles having a maximum cross-sectional dimension within a certain range in some cases. For example, in some embodiments, at least about 50%, at least about 75%, at least about 90%, at least about 95%, or at least about 99% of the total of all catalyst volumes used is less than about 5 microns, about 1 Occupied by particles having a maximum cross-sectional dimension of less than micron, less than about 500 nm, less than about 100 nm, about 100 nm to about 5 microns, about 500 nm to about 5 microns, about 100 nm to about 1 micron, or about 500 nm to about 1 micron .

Ｄ_ｉは粒子ｉの最大断面寸法であり、Ｄ_ａｖｇは全ての粒子の最大断面寸法の平均であり、ｎは触媒内の粒子の数である。標準偏差と上記で述べた粒子の平均最大断面寸法との間の比率の比較は、標準偏差を平均最大断面寸法で除し、１００％を掛けることによって求めることができる。 In some embodiments, the particles within the catalyst may be substantially the same size. For example, the catalyst has a standard deviation of the maximum cross-sectional dimension of the particles of about 50%, about 25%, about 10%, about 5%, about 2%, or about 1% of the average maximum cross-sectional dimension of the particles. It may comprise particles having a size distribution such that: The standard deviation (lowercase sigma) may be calculated as follows:

D _i is the maximum cross-sectional dimension of particle i, D _avg is the average of the maximum cross-sectional dimensions of all particles, and n is the number of particles in the catalyst. A comparison of the ratio between the standard deviation and the average maximum cross-sectional dimension of the particles described above can be determined by dividing the standard deviation by the average maximum cross-sectional dimension and multiplying by 100%.

  The use of a catalyst comprising particles within the selected size distribution shown above results in an increase in the yield and / or selectivity of the aromatics produced by the reaction of the hydrocarbonaceous material. obtain. For example, in some cases, using a catalyst comprising particles having a desired size range (eg, any of the size distributions shown above) has a size distribution outside the desired range (eg, 1 At least about 5%, at least about 10%, or at least about 20%, based on the amount of aromatic compound produced using a catalyst comprising particles) An increase in the amount of aromatics in the reaction product.

  Or, the catalyst has a pore size (eg, mesoporous and pore size typically associated with zeolites), eg, less than about 100 angstroms, less than about 50 angstroms, less than about 20 angstroms, less than about 10 angstroms, less than about 5 angstroms. Or an average pore size smaller than these may be selected. In some embodiments, a catalyst having an average pore size of about 5 angstroms to about 100 angstroms may be used. In some embodiments, a catalyst having an average pore size of about 5.5 angstroms to about 6.5 angstroms, or about 5.9 angstroms to about 6.3 angstroms may be used. In some cases, a catalyst having an average pore size of about 7 angstroms to about 8 angstroms, or about 7.2 angstroms to about 7.8 angstroms may be used.

  As used herein, the term “pore size” is used to refer to the minimum cross-sectional diameter of a pore. The minimum cross-sectional diameter of the pores may correspond to the minimum cross-sectional dimension (eg, cross-sectional diameter) as measured perpendicular to the pore length. In some embodiments, a catalyst having an “average pore size” or “pore size distribution” of X refers to a catalyst where the average of the smallest cross-sectional diameters of the pores in the catalyst is about X. As used herein, “pore size” or “minimum cross-sectional diameter” of a pore should be understood to refer to a Norman radius adjusted pore size well known to those skilled in the art. Norman radius adjustment pore size is determined by Cook, M; Conner, WC “How big are the pores in zeolite”, Minutes of the 12th International Zeolite Conference, Baltimore, July 1998- 10 (1999), 1, 409-414, which is incorporated herein by reference in its entirety. As a specific exemplary calculation, the atomic radius of the ZSM-5 pore is about 5.5 to 5.6 angstroms as measured by X-ray diffraction. In order to adjust the repulsive effect between oxygen atoms in the catalyst, Cook and Conner have shown that the Norman tuning radius is 0.7 Angstroms, which is larger than the atomic radius (about 6.2-6.3 Angstroms). .

  Those skilled in the art understand how to determine the pore size (eg, minimum pore size, average of minimum pore size) in the catalyst. For example, X-ray diffraction (XRD) can be used to determine atomic coordination (or coordinates). XRD techniques for determining pore size are described, for example, in Pecharsky, V. K. et al., “Fundamental Powder Diffraction and Structural Properties of Materials”, Springer Science Business Media, New York, 2005, which is entirely described. Which is incorporated herein by reference. Other techniques that may be beneficial in determining pore size (eg, zeolite pore size) include, for example, helium pycnometry or low pressure argon adsorption techniques. These and other techniques are described in Maggie JS et al., “Fluid Catalytic Catalytic Cracking: Science and Technology,” Elsevier Publishing Company, July 1, 1993, pages 185-195, which is incorporated by reference in its entirety. Incorporated herein. The pore size of the mesoporous catalyst is described in Greg, S. J et al., “Adsorption, Surface Area and Pore”, 2nd edition, Academic Press, New York, 1982, and incorporated herein by reference in its entirety. Rouquerol, F et al., "Principles of adsorption and methodologies with powders and porous materials, methodology and applications" can be determined using, for example, nitrogen adsorption techniques as described in Academic Press, New York, 1998. Unless otherwise indicated, the pore sizes shown herein were determined by X-ray diffraction modified as described above to reflect the Norman radius adjusted pore size.

  The screening method can be used to select a catalyst having an appropriate pore size for converting a particular pyrolysis product molecule. The screening method may include determining the molecular size of the desired pyrolysis product for catalysis (eg, the molecular dynamics diameter of the pyrolysis product molecule). One skilled in the art can calculate a predetermined molecular dynamics diameter, for example. The catalyst type may be selected such that the pores of the catalyst (eg, Norman tuning minimum radius) are large enough to allow the pyrolysis product molecules to diffuse into and / or react with the catalyst. . In some embodiments, the catalyst is selected such that the pore size of the catalyst is small enough to inhibit the entry and / or reaction of undesirable pyrolysis products of the reaction.

  The catalyst may be selected from naturally occurring zeolites, synthetic zeolites, and combinations thereof. The catalyst may be an inverted mordenite structure (MFI) type such as a ZSM-5 zeolite catalyst. Catalysts comprising ZSM-5 that may be used with or without modification are commercially available. The catalyst provided herein may comprise an acid or a catalytically active site. Without being bound by theory, the various acid sites in ZSM-5 and other zeolites can be dehydrated, decarbonylated, decarboxylated, isomerized, oligomerized and / or dehydrogenated hydrocarbon-based materials. It is considered catalytically active due to the reaction. Thus, the terms “acid site” and “catalytically active site” may be used interchangeably. Other types of useful zeolite catalysts are ferrierite, zeolite Y, zeolite beta, mordenite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S) AIPO- 31, SSZ-23, a mixture of two or more thereof, and the like.

  The catalyst may comprise one or more additional metals and / or metal oxides in addition to alumina and silica. Suitable metals and / or oxides may comprise, for example, nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper, chromium, gallium, and / or any of these oxides, among others. . In some embodiments, the metal and / or metal oxide can be impregnated into the catalyst (eg, in the interstices of the catalyst lattice structure). The metal or metal oxide can be added to the zeolite by any of a number of techniques known to those skilled in the art including but not limited to impregnation, ion exchange, vapor deposition and the like. Zeolites may comprise small amounts of structural stabilizing elements such as phosphorus, lanthanum, rare earths, typically at a level of less than about 1% by weight of the zeolite. The catalyst may be adjusted prior to operation of the process by a wide range of techniques known to those skilled in the art such as, but not limited to, oxidation, calcination, reduction, cyclic oxidation and reduction, steaming, hydrolysis, etc. . Metals and / or metal oxides may be incorporated into the catalyst lattice structure. For example, metals and / or metal oxides may be included during the preparation of the catalyst, and the metals and / or metal oxides may occupy lattice sites of the resulting catalyst (eg, a zeolite catalyst). As another example, the metal and / or metal oxide may react or interact with the zeolite such that the metal and / or metal oxide replaces atoms in the lattice structure of the zeolite.

  In certain embodiments, an inverted mordenite structure (MFI) zeolite catalyst comprising gallium may be used. For example, a galloaluminosilicate MFI (GaAlMFI) zeolite catalyst may be used. Those skilled in the art are familiar with GaAlMFI zeolites which may be considered aluminosilicate MFI zeolites in which some of the Al atoms are replaced with Ga atoms. In certain examples, the zeolite catalyst may be in the hydrogen form (eg, H-GaAlMFI). The galloaluminosilicate MFI catalyst may in one aspect be a ZSM-5 zeolite catalyst in which some of the aluminum atoms are replaced with gallium atoms.

  In one example, the ratio of Si moles in the galloaluminosilicate zeolite catalyst to the sum of the moles of Ga and Al in the galloaluminosilicate zeolite catalyst (ie, the mole ratio expressed as Si: (Ga + Al)) is at least about The ratio may be 15: 1, at least about 20: 1, at least about 25: 1, at least about 35: 1, at least about 50: 1, at least about 75: 1 or higher. In some embodiments, the ratio of Si moles in the zeolite catalyst to the sum of moles of Ga and Al in the zeolite catalyst is from about 15: 1 to about 100: 1, from about 15: 1 to about 75: 1, from about 25: It may be advantageous to use a catalyst that is 1 to about 80: 1, or about 50: 1 to about 75: 1. In one example, the ratio of Si moles in the galloaluminosilicate zeolite catalyst to Ga moles in the galloaluminosilicate zeolite catalyst is at least about 30: 1, at least about 60: 1, at least about 120: 1, at least about 200: 1, about 30: 1 to about 300: 1, about 30: 1 to about 200: 1, about 30: 1 to about 120: 1, or about 30: 1 to about 75: 1. The ratio of Si moles in the galloaluminosilicate zeolite catalyst to Al moles in the galloaluminosilicate zeolite catalyst is at least about 10: 1, at least about 20: 1, at least about 30: 1, at least about 40: 1, at least about 50: 1, at least about 75: 1, about 10: 1 to about 100: 1, about 10: 1 to about 75: 1, about 10: 1 to about 50: 1, about 10: 1 to about 40: 1, Or from about 10: 1 to about 30: 1.

  Further, in some cases, the properties of the catalyst (eg, pore structure, type and / or number of acid sites, etc.) may be selected to selectively produce the desired product.

  In certain embodiments, it may be desirable to use one or more catalysts to establish a bimodal distribution of pore sizes. In some cases, a single catalyst having a bimodal distribution of pore sizes (eg, a single catalyst comprising mainly 5.9-6.3 angstrom pores and 7-8 angstrom pores). Can be used. In other cases, a mixture of two or more catalysts may be used to establish a bimodal distribution (eg, in a mixture of two catalysts, each catalyst type includes a different range of average pore sizes. Become). In some embodiments, one of the one or more catalysts comprises a zeolite catalyst and the other of the one or more catalysts comprises a non-zeolite catalyst (eg, a mesoporous catalyst, a metal oxide catalyst, etc.).

  For example, in certain embodiments, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 98% of the pores of one or more catalysts (eg, zeolite catalyst, mesoporous catalyst, etc.), Or at least about 99% has a minimum cross-sectional diameter that is within the first size distribution or the second size distribution. In some cases, at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have a minimum cross-sectional diameter within the first size distribution, and the one or more catalysts At least about 2%, at least about 5%, or at least about 10% have a minimum cross-sectional diameter within the second size distribution. In some cases, the first and second size distributions are selected from the ranges provided above. In certain aspects, the first and second size distributions are different from each other and do not overlap. Examples of non-overlapping ranges are 5.9 to 6.3 angstroms and 6.9 to 8.0 angstroms, and examples of overlapping ranges are 5.9 to 6.3 angstroms and 6.1 to 6 .5 Angstrom. The first and second size distributions may be selected such that the ranges are not immediately adjacent to one another, examples being pore sizes of 5.9 to 6.3 angstroms and 6.9 to 8.0 angstroms. Examples of ranges that are immediately adjacent to each other are pore sizes of 5.9 to 6.3 angstroms and 6.3 to 6.7 angstroms.

  As a specific example, in one embodiment, one or more catalysts are used to provide a bimodal pore size distribution for the simultaneous production of aromatic and olefinic compounds. That is, one pore size distribution can be advantageous to produce a relatively large amount of aromatic compound, while the other pore size distribution can produce a relatively large amount of olefinic compound. Can be advantageous. In some embodiments, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 98%, or at least about 99% of the pores of one or more catalysts are about 5.9 angstroms. Having a minimum cross-sectional diameter of from about 6.3 angstroms or from about 7 angstroms to about 8 angstroms. Further, at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have a minimum cross-sectional diameter of about 5.9 angstroms to about 6.3 angstroms, and At least about 2%, at least about 5%, or at least about 10% of the pores of the catalyst have a minimum cross-sectional diameter of about 7 angstroms to about 8 angstroms.

  In some embodiments, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 98%, or at least about 99% of the pores of one or more catalysts are about 5.9 angstroms. Having a minimum cross-sectional diameter of from about 6.3 angstroms or from about 7 angstroms to about 200 angstroms. Further, at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have a minimum cross-sectional diameter of about 5.9 angstroms to about 6.3 angstroms, and At least about 2%, at least about 5%, or at least about 10% of the pores of the catalyst have a minimum cross-sectional diameter of about 7 angstroms to about 200 angstroms.

  In some embodiments, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 98%, or at least about 99% of the pores of one or more catalysts are first and second Has the smallest cross-sectional diameter within the distribution. The first distribution is about 5.9 angstroms to about 6.3 angstroms, and the second distribution is different from the first distribution and does not overlap with the first distribution. In certain embodiments, the second pore size distribution may be from about 7 angstroms to about 200 angstroms, from about 7 angstroms to about 100 angstroms, from about 7 angstroms to about 50 angstroms, or from about 100 angstroms to about 200 angstroms. In certain embodiments, the second catalyst may be mesoporous, for example, having a pore size distribution of about 2 nm to about 50 nm.

  In certain embodiments, a bimodal distribution of pore sizes can be beneficial for reacting two or more hydrocarbonaceous feedstock components. For example, one embodiment comprises providing a solid hydrocarbonaceous material comprising a first component and a second component in a reactor. The first component and the second component are different. Examples of compounds that may be used as the first or second component are any of the hydrocarbonaceous materials described herein (eg, sugarcane bagasse, glucose, wood, corn stover, cellulose, hemicellulose, lignin, or others). Comprising. For example, the first component may comprise one of cellulose, hemicellulose and lignin and the second component comprises one of cellulose, hemicellulose and lignin. The method may further include providing first and second catalysts in the reactor. In some embodiments, the first catalyst may have a first pore size distribution, the second catalyst may have a second pore size distribution, and the first and second pore size distributions are different. And do not overlap. The first pore size distribution may be, for example, from about 5.9 angstroms to about 6.3 angstroms. The second pore size distribution may be, for example, from about 7 angstroms to about 200 angstroms, from about 7 angstroms to about 100 angstroms, from about 7 angstroms to about 50 angstroms, or from about 100 angstroms to about 200 angstroms. In some cases, the second catalyst may not be mesoporous or porous.

  The first catalyst can be selected to catalytically react the first component or derivative thereof to produce a flowable hydrocarbon product. Further, the second catalyst can be selected to catalytically react the second component or derivative thereof to produce a flowable hydrocarbon product. The method includes pyrolyzing at least a portion of the hydrocarbonaceous material in a reactor under reaction conditions sufficient to produce one or more pyrolysis products, and one or more hydrocarbon products. The method may further comprise the step of catalytically reacting at least a portion of the pyrolysis product with the first and second catalysts for production. Also, in some examples, an at least partially deactivated (or deactivated) catalyst may be used.

  In certain embodiments, a method used in combination with the embodiments described herein can increase the mass ratio of the catalyst hydrocarbonaceous material in the composition to increase the production of identifiable aromatic compounds. Comprising. As shown herein, which shows certain differences from certain conventional catalytic pyrolysis methods, the products and methods described herein are benzene, toluene, propylbenzene, ethylbenzene, methylbenzene, methyl Produces separately identifiable aromatic and biofuel compounds selected from ethylbenzene, trimethylbenzene, xylene, indane, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene, dimethylhydrindene and combinations thereof Can be used to

  In certain embodiments, the reaction chemistry of the catalyst can be affected by adding one or more additional compounds. For example, the addition of metal to the catalyst can result in a shift in the selective formation of a particular compound (eg, the addition of alumina silicate catalyst metal can result in more CO production). Furthermore, when the flowable fluid comprises hydrogen, the amount of coke formed on the catalyst can be reduced.

The catalyst may comprise both silica and alumina. Silica (SiO ₂ ) and alumina (A1 ₂ O ₃ ) in the catalyst may be present in any suitable molar ratio. For example, in some cases, the catalyst in the feedstock has a molar ratio of silica (SiO ₂ ) to alumina (A 1 ₂ O ₃ ) of about 10: 1 to about 50: 1, about 10: 1 to about 40: 1, It may comprise silica (SiO ₂ ) and alumina (A 1 ₂ O ₃ ) that is about 10: 1 to about 20: 1, or about 15: 1.

  In some embodiments, for example, the feed composition is combined with a catalyst (eg, via one or more feed streams comprising a catalyst and a hydrocarbon-based material, or via separate catalyst and hydrocarbon-based material feed streams). In the case of comprising a hydrocarbonaceous material (eg, a circulating fluidized bed reactor), the catalyst and hydrocarbonaceous material may be present in any suitable ratio. As another example, if the reactor is initially loaded with a mixture of catalyst and hydrocarbonaceous material (eg, a batch reactor), the catalyst and hydrocarbonaceous material may be present in any suitable mass ratio. . In certain embodiments comprising a circulating fluidized bed reactor, the mass ratio of catalyst to hydrocarbonaceous material in the feed stream, i.e., in the composition comprising the catalyst and hydrocarbonaceous material fed to the reactor, is: May be at least about 0.5: 1, at least about 1: 1, at least about 2: 1, at least about 5: 1, at least about 10: 1 at least about 15: 1, at least about 20: 1, or more. . In some embodiments comprising an annular fluidized bed reactor, the mass ratio of catalyst to hydrocarbonaceous material in the feed stream is less than about 0.5: 1, less than about 1: 1, less than about 2: 1, about 5 : 1, less than about 10: 1, less than about 15: 1, less than about 20: 1, about 0.5: 1 to about 20: 1, about 1: 1 to about 20: 1, or about 5: 1. It may be about 20: 1. By using a catalyst with a relatively high mass ratio to the hydrocarbon-based material, the volatile organic compound formed from the pyrolysis of the feedstock in the catalyst before the volatile organic compound is pyrolyzed into coke. It can be easy to introduce. While not wishing to be bound by any theory, this effect is due, at least in part, to the presence of stoichiometric excess catalyst sites in the reactor.

  In certain embodiments, the articles and methods described herein may be configured to selectively produce aromatic compounds (eg, p-xylene) in a single stage pyrolysis apparatus or a multistage pyrolysis apparatus. For example, in some embodiments, the mass yield of aromatics in the flowable hydrocarbon product is at least about 18%, at least about 20%, at least about 25%, at least about 30%, at least about 35%. %, At least about 39%, about 18% to about 40%, about 18% to about 35%, about 20% to about 40%, about 20% to about 35%, It may be from 25% to about 40%, from about 25% to about 35%, from about 30% to about 40%, or from about 30% to about 35%. The mass yield of p-xylene may be at least about 1.5 wt%, at least about 2 wt%, at least about 2.5 wt%, or at least about 3 wt%.

  As used herein, the “mass yield” of aromatics or p-xylene in a given product is the total weight of aromatics or p-xylene present in the flowable hydrocarbon product. Calculated as 100% multiplied by the weight of the solid hydrocarbonaceous material used to form the reaction product.

As used herein, the term “aromatic compound” refers to, for example, a single aromatic ring system (eg, benzyl, phenyl, etc.), and fused polycyclic aromatic ring systems (eg, naphthyl, 1,2, Used to refer to a hydrocarbon compound comprising one or more aromatic groups such as 3,4-tetrahydronaphthyl and the like. Examples of the aromatic compound include, but are not limited to, benzene, toluene, indane, indene, 2-ethyltoluene, 3-ethyltoluene, 4-ethyltoluene, trimethylbenzene (for example, 1,3,5- Trimethylbenzene, 1,2,4-trimethylbenzene, 1,2,3-trimethylbenzene, etc.), ethylbenzene, methylbenzene, propylbenzene, xylene (eg, p-xylene, m-xylene, o-xylene, etc.), naphthalene Methylnaphthalene (eg, 1-methylnaphthalene, anthracene, 9,10-dimethylanthracene, pyrene, phenanthrene, dimethylnaphthalene (eg, 1,5-dimethylnaphthalene, 1,6-dimethylnaphthalene, 2,5-dimethylnaphthalene, etc.) ), Ethyl naphthalene, hi Linden, methyl hydrate Linden, and dimethyl hydrate Linden. May be generated by the single ring and / or more ring aromatics some embodiments. Aromatic compound, for _{example, C 5 ~C 14, C 6} ~ C ₈ , C ₆ -C ₁₂ , C ₈ -C ₁₂ , C ₁₀ -C ₁₄ may have carbon numbers.

In certain embodiments, the normalized mass space velocity of the solid hydrocarbonaceous material fed to the reactor is up to about 3 hours ⁻¹ (or less), up to about 2 hours ^−1, up to about 1.5 hours ⁻¹ , about 0. until .9 hr ^-1, or about 0.01 hr ^-1 to about 3 hours ^-1, about 0.01 hr ^-1 to about 2 hr ^-1, about 0.01 hr ^-1 to about 1.5 hr ^{- 1} , about 0.01 hour ⁻¹ to about 0.9 hour ⁻¹ , about 0.01 hour ⁻¹ to about 0.5 hour ⁻¹ , about 0.1 hour ⁻¹ to about 0.9 hour ⁻¹ , Alternatively, aromatic compounds (particularly p-xylene) may be selectively produced when from about 0.1 hour- ¹ to about 0.5 hour- ¹ . In one example, when the reactor is operated at a temperature of about 400 ° C. to about 600 ° C. (about 425 ° C. to about 500 ° C., or about 440 ° C. to about 460 ° C.), aromatic compounds (particularly p-xylene) ) May be selectively generated. Further, certain heating rates (eg, at least about 50 ° C./second, or at least about 400 ° C./second), high mass ratios of catalyst to feedstock (eg, at least about 5: 1), and / or silica in the catalyst A high molar ratio of to alumina (eg, at least about 30: 1) can be used to facilitate the selective production of aromatics (particularly p-xylene). Several such and other process conditions are combined with specific reactor types such as fluidized bed reactors (eg, circulating fluidized bed reactors) to selectively produce aromatic and / or olefinic compounds. It's okay.

  Further, in certain embodiments, the catalyst can be selected to facilitate the selective production of aromatic products, particularly p-xylene. For example, in some cases, ZSM-5 may preferentially produce a relatively large amount of aromatic compound. In some cases, a catalyst comprising a Bronsted acid moiety can facilitate the selective production of aromatic compounds. Further, a catalyst having an ordered pore structure can facilitate the selective production of aromatic compounds. For example, in certain embodiments, a catalyst having an average pore size of about 5.9 angstroms to about 6.3 angstroms may be particularly beneficial for the production of aromatic compounds. In addition, a catalyst having an average pore size of about 7 angstroms to about 8 angstroms can be beneficial in producing olefins. In certain embodiments, a combination of one or more of the above process parameters can be used to facilitate the selective production of aromatic and / or olefinic compounds. The ratio of aromatic to olefin produced is, for example, from about 0.1: 1 to about 10: 1, from about 0.2: 1 to about 5: 1, from about 0.5: 1 to about 2: 1, It may be about 0.1: 1 to about 0.5: 1, about 0.5: 1 to about 1: 1, about 1: 1 to about 5: 1, or about 5: 1 to about 10: 1. .

  In some embodiments, the mass ratio of catalyst to hydrocarbonaceous material in the feed is adjusted to produce the desired product and / or good yield. The weight ratio of catalyst to hydrocarbonaceous material can be, for example, in some embodiments, at least about 0.5: 1, at least about 1: 1, at least about 2: 1, at least about 5: 1, at least about 10: 1, It may be at least about 15: 1, at least about 20: 1, or more, or in other embodiments, less than about 0.5: 1, less than about 1: 1, less than about 2: 1, about 5: 1. Less than about 10: 1, less than about 15: 1, or less than about 20: 1.

  Furthermore, the processes described herein can result in lower coke formation than certain existing methods. For example, in some embodiments, pyrolysis products wherein less than about 30%, less than about 25%, less than about 20%, less than about 15%, or less than about 10% by weight of the pyrolysis product are coke. Things are formed. The amount of coke formed is measured as the weight of coke formed in the system divided by the weight of the hydrocarbonaceous material used to form the pyrolysis product.

  The following non-limiting examples are intended to illustrate various aspects and features of the present invention.

  A series of zeolite catalysts were prepared and used in the CFP process to convert furan, 2-methylfuran (2MF) and pinewood into a fluid hydrocarbon product.

(catalyst)
Four catalysts identified as ZSM, GaZSM, SD and GASD were used. ZSM is ZSM-5, (Si / Al = 15). GaZSM was performed using ion exchange in which 1 g of ZSM was refluxed to 100 mL of an aqueous solution of Ga (NO ₃ ) ₃ (0.01 M) at 70 ° C. for 12 hours. After ion exchange, the solution was dried at 110 ° C. to form a dry powder. The dried powder was fired under air at 550 ° C. SD is spray dried HZSM-5. GaSD was made by initial wet impregnation of SD using Ga (NO ₃ ) ₃ solution (0.43M). The impregnated GaSD was dried at 110 ° C. and fired in air at 550 ° C. The Ga content of GaZSM and GaSD was determined by inductively coupled plasma (ICP) analysis.

Chemical liquid deposition (CLD) using tetraethylorthosilicate (TEOS) was used to modify the catalyst and thereby reduce the pore opening size of the catalyst. 1 g of the catalyst was dispersed in 25 mL of hexane. Thereafter, 0.15 mL of TEOS was added. The mixture was refluxed at 90 ° C. for 1 hour with stirring. The catalyst was recovered by centrifugation. The catalyst was dried for 2 hours at 100 ° C. and calcined in dry air for 4 hours at 500 ° C. The pore opening correction process was repeated two more times. A catalyst with a modified pore opening may be referred to as a “silylation” catalyst, and is identified by “*”, for example, ZSM ^* , GaZSM ^*, and GASD ^* . The modification process may be referred to as a TEOSCLD silylation process.

  Catalyst samples were analyzed by temperature programmed desorption of isopropylamine (IPA) or 2,4,6-collidine (2,4,6-trimethylpyridine). Prior to adsorption, the sample was degassed at 823K for 2 hours. After cooling the sample to 393 K, the sample was exposed to He saturated with isopropylamine or 2,4,6-collidine at room temperature for 1 hour by flowing pure He through a bubbler comprising amine. Samples were held at 393 K with He flow for 2 hours to remove physisorbed IPA or 2,4,6-collidine. The sample was heated to 973K at 10K / min. The total number of acid sites was calculated using the total amount of desorbed amine, and the number of Bronsted acid sites for each catalyst was calculated using the amount of amine desorbed between about 580K and 650K. Due to the size of 2,4,6-collidine, 2,4,6-collidine does not enter the pores of ZSM-5. Thus, 2,4,6-collidine desorption detected acid sites only on the outer surface of the catalyst or in the pores near the pore openings. IPA desorption detected acid sites in the zeolite pores and acid sites on the outer surface of the catalyst or in the pores near the pore openings. The decrease in 2,4,6-collidine adsorption that occurs during the silylation treatment showed a decrease in the number of acid sites in the outer surface of the catalyst and in or near the pore openings. These reductions in external sites reduce the formation of m-xylene and o-xylene and the re-equilibration of p-xylene formed in the pores to m-xylene o-xylene, thereby selecting p-xylene. It is thought to be a factor that improves sex.

Table 1 Thermal desorption analysis of acid concentration obtained from 2,4,6-collidine (kinetic diameter = 7.4 angstrom) and isopropylamine (IPA, kinetic diameter = 5.2 angstrom)

(Catalytic conversion of furan and 2MF)
The catalytic reaction was 0.5 inch (1.27 cm) O.D. Performed in D fixed bed quartz reactor. A catalyst in the form of a fixed bed of solid particles was held in the reactor by a quartz frit. The catalyst bed was calcined at a temperature of 600 ° C. with air flowing at a rate of 60 mL / min. After calcination, the reactor was purged with helium at 408 mL / min for 5 minutes. The furan was pumped into the helium stream using a syringe pump. Prior to testing, the furan bypassed the reactor for 30 minutes. A helium stream comprising furan was switched to pass through the reactor. An air bath condenser was used to trap heavy product. The gas phase product was collected by an airbag. All experiments were performed at atmospheric pressure. No pressure drop was detected through the catalyst bed. After the reaction process was complete, the reactor was purged with helium at a flow rate of 408 mL / min for 45 seconds at the reaction temperature. The effluent was collected using an airbag. After completion of each reaction, the catalyst used was regenerated at a temperature of 600 ° C. with air at a flow rate of 60 mL / min. It was converted to C0 ₂ by copper converter (copper oxide) at a temperature of 240 ° C. The formed CO during playback. CO ₂ was trapped by a CO ₂ trap. Coke yield was determined by measuring the change in weight of the CO ₂ trap. Gas products were identified by GC-MS (Shimadzu Corporation-2010) and quantified by GC-FID / TCD (Shimadzu 2014 for gas samples and HP-7890 for liquid samples). All hydrocarbons in the gas phase product were quantified by GC-FID. CO and C0 ₂ in the gas phase product was quantified by GC-TCD. C ₂ -C ₆ standard olefin standards (Scott special gas, 1000 ppm for each olefin), furan, benzene, toluene, xylene (gas phase standards are prepared for these aromatics that can be evaporated at room temperature), ethylbenzene GC-FID was calibrated (or calibrated) with styrene, indene, naphthalene, and benzofuran. The sensitivity of the hydrocarbon is assumed to be proportional to the number of carbon molecules having a similar structure (eg, styrene vs. methylstyrene, indene vs. methylindene). The GC-TCD was calibrated with CO and CO ₂ standards (air gas, air gas, 6% CO ₂ and 14% CO kept in equilibrium with helium). Less than 0.05% carbon or product was collected in the condenser. Most of the product was either in the gas phase or in the coke deposited on the catalyst. The carbon balance was close to greater than 90% of the total run.

The reaction conditions for furan conversion were a temperature of 550 ° C., a space velocity (WHSV) of 10.2 h ⁻¹ , and a partial pressure of 6 Torr. The furan was pumped at a pump speed of 0.58 mL / h and the carrier gas was maintained at 408 mL / min. The amount of catalyst charged in the reactor was 53 mg.

The 2MF reaction process was the same as the furan process except that pure helium was used as the carrier gas instead of 2% propylene (1.986% propylene equilibrated with helium). Reaction conditions for 2MF conversion were 600 ° C., WHSV of 5.7 h ⁻¹ , and partial pressure of 4.9 torr. 2MF was pumped at a pump speed of 0.57 mL / min and the carrier gas flow rate was maintained at 408 mL / min. The amount of catalyst charged into the reactor was 92 mg.

  The results are shown in Table 2.

Figures A and B show the selectivity and xylene distribution of total p-xylene obtained from the conversion of 2MF and furan, respectively. In FIG. A, the carbon selectivity of all p-xylene obtained from ZSM was 5%. This value was increased to 15% using ZSM ^* . Silylation significantly increased paraselectivity from 32% to 92%. Similarly, silylated GaZSM ^* significantly increased para selectivity from 58% to 96%. However, the selectivity of GaZSM ^* for total p-xylene was lower than GaZSM. Table 2 shows that the 2MF conversion of GaZSM ^* (74%) is lower than the 2MF conversion of GaZSM (98%). The ZSM ^* 2MF conversion (89%) was lower than the ZSM 2MF conversion (99%). The decrease in activity is believed to be due to the removal of some active sites on the outer surface and the surface near the pore opening by silica deposition. Ga deposited on ZSM can increase total aromatic selectivity. This is also shown in Table 2 (2MF + propylene) where the aromatic selectivity is 60% for ZSM and 69% for GaZSM. However, the deactivation caused by silylation using GaZSM ^* is more severe than that using ZSM ^* due to 2MF conversion, and some active Ga species are located on these surfaces and the active Ga species is silylated. Indicates that it was deactivated. These Ga species provide space confinement that causes an increase in paraselectivity (Figure A, 58% for GaZSM, 32% for ZSM). The silylation of GaZSM further provides a lot of space confinement, thereby providing better paraselectivity (96%).

In the case of furan conversion (Figure B), an increase in total p-xylene selectivity and para to xylene species was observed from the silylation catalyst. Para-selection increased from 53% for ZSM to 89% for ZSM ^* and from 57% for GaZSM to 96% for GaZSM ^* . Silylation reduces activity, as shown in Table 1, where furan conversion is low with silylation catalysts. The increase in activity of GaZSM (41%) compared to ZSM (33%) is due to Ga species. However, the furan conversion observed with GaZSM ^* was the lowest (24%) due to the catalytic sites deactivated by silica deposition. The significant decrease in total xylene selectivity in the 2MF + propylene reaction is believed to be due to the fact that furan itself is not a good Diels-Alder reagent for xylene production. The results of furan conversion using GaSD ^* are shown in Table 2. This table shows that the para-selectivity of this catalyst is 87% and that silylation can be used for FCC catalysts to increase the yield of p-xylene from biomass conversion in a fluidized bed reactor. Is shown. This is shown below where GaSD ^* is used for the conversion of pine wood in a bubbling fluidized bed reactor. Table 3 shows that the para-selectivity increased from 40% for GaSD to 72% for GaSD ^* .

(Catalytic conversion of pine wood in fluidized bed reactor)
Pine wood CFP was performed in a fluidized bed reactor. The fluidized bed reactor has a diameter of 2 inches (5.08 cm), a height of 10 inches (25.4 cm) and is made from 316 stainless steel. In the reactor, the catalyst bed was supported by a distribution plate consisting of laminated 316 stainless mesh (300 mesh). Solid pine wood was introduced into the reactor from a sealed feed hopper. Test runs were performed using GaSD and GaSD ^* . Prior to the trial run, the pine wood was crushed and sieved to particles with a size in the range of 0.25 to 1 mm. During the reaction, the catalyst was fluidized with helium gas flowing at 800 standard cubic centimeters (sccm), allowing the reactor to operate in a bubbling flow mode. The hopper and feed chamber were continuously purged with helium at 200 sccm to maintain an inert environment. The total gas flow flowing through the reactor was 1000 sccm helium. Both the reactor and the incoming gas stream were heated to the reaction temperature (550 ° C.). The reactor took 2 hours to reach this temperature before starting the reaction. The effluent gas exiting the reactor was flushed through a cyclone to remove entrained particles. The effluent was then flowed through 7 condensers in series to separate the liquid and gas phase products. The first three condensers with ethanol as a solvent in each condenser were placed in an ice-water bath, and the other four condensers without any solvent inside the condenser were placed in a dry ice and acetone bath (−55 ° C). After the biomass first entered the reactor, the uncondensed gas phase product was collected in the airbag at 5, 10, 20 and 30 minutes. The reaction time was 30 minutes. After 30 minutes, the reactor was purged with helium at a flow rate of 1000 sccm for 30 minutes to remove any CFP product other than coke in the catalyst. The liquid product was extracted from the condenser using ethanol. In addition to 200 sccm of helium from the feed chamber purge, the catalyst was regenerated with air at 800 sccm for 3 hours. During playback, the effluent gas CO is passed through a copper converter changes to CO _2, and the CO ₂ is trapped by CO ₂ trap. The gas phase product was analyzed by GC-FID / TCD (Shimadzu 2014). The liquid sample was analyzed by GC-FID (HP7890). The yield of coke was obtained by analyzing the change in weight of the CO ₂ trap. The results are shown in Table 3.

  While the invention has been described in connection with various aspects, it will be understood that various modifications will become apparent to those skilled in the art upon reading this specification. Accordingly, the invention disclosed herein is to be understood as comprising any modification as falling within the scope of the appended claims.

Claims (40)

A method for producing a flowable hydrocarbon product comprising p-xylene from a hydrocarbon-based material comprising:
Supplying a hydrocarbon-based material to the reactor;
Pyrolyzing at least a portion of the hydrocarbonaceous material in the reactor under reaction conditions sufficient to produce a pyrolysis product; and at least a portion of the pyrolysis product under reaction conditions in which a zeolite catalyst was present Catalytically reacting to produce a fluid hydrocarbon product,
The zeolite catalyst comprises pores having pore openings, catalyst sites on the outer surface of the catalyst, and a treatment layer;
The method wherein the treatment layer is an effective amount of treatment layer derived from a silicone compound to reduce the size of the pore openings and render at least some of the catalyst sites on the outer surface of the catalyst inaccessible to pyrolysis products .   The method of claim 1, wherein the catalyst sites are located in pores near the pore openings and the treatment layer renders at least some of the catalyst sites in the pores near the pore openings inaccessible to pyrolysis products.   The selectivity for p-xylene of xylene is at least about 40%, at least about 45%, at least about 50%, at least about 55%, at least about 60%, at least about 65%, at least about 70%, at least about 75%. 3. The process of claim 1 or 2, wherein the flowable hydrocarbon product comprises xylene that is at least about 80%, at least about 85%, or at least about 90%.   At least about 15%, at least about 25%, at least about 35%, at least about 45%, at least about 55%, at least about 65%, at least about 75%, at least about 85%, at least of the catalyst sites on the outer surface of the catalyst 4. The method of any of claims 1-3, wherein about 90%, at least about 95%, at least about 98%, or at least about 99% are inaccessible to pyrolysis products.   At least 20%, at least about 25%, at least about 30%, at least about 33%, at least about 35%, at least about 40%, at least about 45%, at least about 49% of the catalytic sites of the pores near the pore openings, At least about 50%, at least about 55%, at least about 60%, at least about 65%, at least about 70%, at least about 75%, at least about 80%, at least about 85%, at least about 90%, at least about 95%, 3. The method of claim 2, wherein at least about 99% is inaccessible to pyrolysis products.   The zeolite catalyst comprises silica and alumina, wherein the molar ratio of silica to alumina is about 10: 1 to about 50: 1, about 10: 1 to about 40: 1, about 10: 1 to about 20: 1, Or a method according to any one of claims 1 to 5, which is about 15: 1.   The zeolite catalyst further comprises nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper, chromium, gallium, one or more oxides thereof, or a mixture of two or more thereof. The method according to claim 6. 8. A method according to any one of the preceding claims, wherein the silicone compound comprises at least one group represented by the following formula:

The method according to claim 1, wherein the silicone compound is represented by the following formula:

(R ₁ and R ₂ independently comprise hydrogen, halogen, hydroxyl, alkyl, alkoxyl, halogenated alkyl, aryl, aryl halide, aralkyl, halogenated aralkyl, alkaryl or alkaryl, and n is at least 2 Is a number.) R ₁ and / or R ₂ comprises methyl, ethyl, or phenyl, A method according to claim 9.   11. The method of claim 9 or 10, wherein n is a number from about 3 to about 1000.   12. The method of any one of claims 1-11, wherein the silicone compound has a number average molecular weight of about 80 to about 20,000 or about 150 to 10,000.   Silicone compounds are dimethyl silicone, diethyl silicone, phenyl methyl silicone, methyl hydrogen silicone, ethyl hydrogen silicone, phenyl hydrogen silicone, methyl ethyl silicone, phenyl ethyl silicone, diphenyl silicone, methyl trifluoropropyl silicone, ethyl trifluoropropyl silicone, poly Dimethyl silicone, tetrachlorophenyl methyl silicone, tetrachlorophenyl ethyl silicone, tetrachlorophenyl hydrogen silicone, tetrachlorophenyl phenyl silicone, methyl vinyl silicone, hexamethylcyclotrisiloxane, octamethylcyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenylcyclotetra Rokisan or comprising mixtures of two or more of these A method according to any one of claims 1 to 12,.   9. A method according to any one of claims 1 to 8, wherein the silicone compound comprises tetraorthosilicate.   15. The method of claim 14, wherein the silicone compound comprises tetramethylorthosilicate, tetraethylorthosilicate, or a mixture thereof.   The method according to any of claims 1 to 15, wherein the reactor comprises a continuous stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor, or a fluidized bed reactor.   The process according to any of the preceding claims, wherein the reactor comprises a fluidized bed reactor.   18. The hydrocarbon-based material comprises a solid hydrocarbon-based material, a semi-solid hydrocarbon-based material, a liquid hydrocarbon-based material, or a mixture of two or more thereof. the method of.   The method according to any of claims 1 to 18, wherein the hydrocarbonaceous material comprises biomass.   Hydrocarbon materials include plastic waste, recycled plastic, agricultural solid waste, municipal solid waste, food waste, animal waste, carbohydrates, lignocellulosic material, xylitol, glucose, cellobiose, hemicellulose, lignin, sugarcane bagasse 20. A method according to any of claims 1 to 19, comprising glucose, wood, corn stover, or a mixture of two or more thereof.   The hydrocarbon-based material may be a pyrolysis oil derived from biomass, a carbohydrate derived from biomass, an alcohol derived from biomass, a biomass extract, a pretreated biomass, a digested biomass product, or two or more of these 21. A method according to any of claims 1 to 20, comprising a mixture.   The method according to any of claims 1 to 18, wherein the hydrocarbonaceous material comprises furan and / or 2-methylfuran.   The method according to claim 1, wherein the hydrocarbonaceous material comprises pine wood.   24. The method of any one of claims 1-23, wherein the reactor is at a temperature of about 400C to about 600C, about 425C to about 500C, or about 440C to about 460C. A hydrocarbon-based material is fed to the reactor at a normalized mass space velocity of up to about 3 hours ⁻¹ , about 2 hours ⁻¹ , about 1.5 hours ⁻¹ , or about 0.9 hours ⁻¹ , or hydrocarbon based The material is about 0.01 hours- ¹ to about 3 hours- ¹ , about 0.01 hours- ¹ to about 2 hours- ¹ , about 0.01 hours- ¹ to about 1.5 hours- ¹ , about 0.01. Time- ¹ to about 0.9 hours- ¹ , about 0.01 hours- ¹ to about 0.5 hours- ¹ , about 0.1 hours- ¹ to about 0.9 hours- ¹ , or about 0.1 hours 25. The method of any of claims 1-24, wherein the reactor is fed to the reactor at a normalized mass space velocity of ^-1 to about 0.5 hours- ¹ .   26. A process according to any of claims 1 to 25, wherein the flowable hydrocarbon product comprises an aromatic compound and the aromatic carbon molar yield is at least about 22%.   27. The flowable hydrocarbon product comprises an olefin compound and has an olefin carbon molar yield of at least about 3%, at least about 6%, at least about 9%, or at least about 12%. The method of crab.   28. A method according to any preceding claim, further comprising the step of recovering the flowable hydrocarbon product.   29. A process according to any of claims 1 to 28, wherein the flowable hydrocarbon product further comprises an aromatic compound and / or an olefinic compound.   Fluid hydrocarbon products are benzene, toluene, ethylbenzene, methylethylbenzene, trimethylbenzene, o-xylene, m-xylene, indane, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrinden, dimethylhydrindene. 30. A method according to any of claims 1 to 29, further comprising a mixture of two or more thereof.   The mass yield of p-xylene in the flowable hydrocarbon product is at least about 1.5 wt%, at least about 2 wt%, at least about 2.5 wt%, or at least about 3 wt%. The method in any one of 1-30.   32. The method of any of claims 1-31, wherein the reactor is operated at a pressure of at least about 100 kPa, at least about 200 kPa, at least about 300 kPa, or at least about 400 kPa.   32. The method of any of claims 1-31, wherein the reactor is operated at a pressure of less than about 600 kPa, less than about 400 kPa, or less than about 200 kPa.   32. The method of any of claims 1-31, wherein the reactor is operated at a pressure of about 100 to about 600 kPa, about 100 to about 400 kPa, or about 100 to about 200 kPa.   35. A process according to any of claims 1 to 34, wherein dehydration, decarbonylation, decarboxylation, isomerization, oligomerization and / or dehydrogenation reactions are carried out during the catalytic reaction step.   A pyrolysis product is formed wherein less than about 30%, less than about 25%, less than about 20%, less than about 15%, or less than about 10% of the pyrolysis product is coke. Item 34. The method according to any one of Items 1 to 35.   The method according to any one of claims 1 to 36, wherein the pyrolysis step and the catalytic reaction step are performed in a single vessel.   The method according to any one of claims 1 to 36, wherein the pyrolysis step and the catalytic reaction step are performed in separate vessels.   The residence time of the hydrocarbonaceous material in the reactor is at least about 1 second, at least about 2 seconds, at least about 5 seconds, at least about 7 seconds, at least about 10 seconds, at least about 15 seconds, at least about 20 seconds, at least about 25 39. The method of any of claims 1-38, wherein the method is seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds.   The contact time between the pyrolysis product and the catalyst is at least about 1 second, at least about 2 seconds, at least about 5 seconds, at least about 7 seconds, at least about 10 seconds, at least about 15 seconds, at least about 20 seconds, at least about 25 40. The method of any of claims 1-39, wherein the method is seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds.

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