US20150166899A1

US20150166899A1 - Catalysts and Processes for Producing p-xylene from Biomass

Info

Publication number: US20150166899A1
Application number: US14/569,284
Authority: US
Inventors: Jian Shi; Yu-Ting Cheng; Ruozhi Song, Jr.; Terry Mazanec
Original assignee: Anellotech Inc
Current assignee: Anellotech Inc
Priority date: 2013-12-14
Filing date: 2014-12-12
Publication date: 2015-06-18
Also published as: WO2015089442A9; WO2015089442A1

Abstract

Biomass is converted to a fluid hydrocarbon product comprising p-xylene by reaction over a zeolite catalyst. An iron-modified zeolite catalyst having a siliceous coating and methods of making the catalyst are also described.

Description

RELATED APPLICATIONS

This application claims the priority benefit of U.S. Provisional Patent Application Ser. No. 61/916,180, filed Dec. 14, 2013.

FIELD OF INVENTION

This invention relates to a method for converting biomass to a fluid hydrocarbon product comprising p-xylene by reaction over a zeolite catalyst. An iron-modified zeolite catalyst having a siliceous coating and methods of making the catalyst are also described.

BACKGROUND

p-Xylene is used as a starting material for plasticizers and polyester fibers. The oxidation of p-xylene is used to commercially synthesize terephthalic acid. Further esterification of the acid with methanol forms dimethyl terephthalate. Both monomers may be used in the production of polyethylene terephthalate (PET) plastic bottles and polyester clothing.
p-Xylene may be the most valuable of the xylenes (i.e., o-, m- and p-xylenes). However, during the catalytic pyrolysis of various hydrocarbonaceous materials, the xylenes may be formed with the m-xylene selectivity and/or o-xylene selectivity being the same as or higher than the p-xylene selectivity. The p-xylene that is produced may also isomerize to m-xylene and/or o-xylene. As a result, xylenes with undesirably high selectivities to m-xylene and/or o-xylene may be formed. Thus, there is a well-known need for the production of p-xylene or xylenes with a relatively high selectivity to p-xylene.

SUMMARY

In a first aspect, the invention provides a process for converting biomass to liquid hydrocarbons, comprising: feeding biomass into a reactor; heating the biomass in the presence of an aluminosilicate zeolite catalyst; and, wherein the aluminosilicate zeolite catalyst further comprises at least 0.2 wt % Fe wherein the Fe is not derived from biomass or reactor walls, or wherein the aluminosilicate zeolite catalyst has been manufactured to contain at least 0.2 wt % Fe or pretreated to contain at least 0.2 wt % Fe; and converting the biomass to a gaseous product stream comprising p-xylene.
In the present invention, an aluminosilicate zeolite means a zeolite having a Si/Al molar ratio of 200:1 to 1:1, more preferably 150:1 to 1.5:1, in some embodiments 100:1 to 1:1, in some embodiments 120:1 to 5:1. One preferred zeolite is ZSM-5. The phrase “at least 0.2 wt % Fe” is determined by elemental analysis of catalyst separated from biomass and separated from any ash, to the extent practicable, where the elemental analysis is preferably conducted by ICP. The term “pretreated” means treated prior to use in a catalytic pyrolysis process. The phrase “wherein the Fe is not derived from biomass or reactor walls” is to distinguish Fe made in catalyst preparation or catalyst pretreatment from Fe that may be deposited on the catalyst or may co-occur in ash as a result of the pyrolysis process. It is believed that Fe that is not derived from biomass or reactor walls will not be as effective as Fe added during catalyst manufacture or pretreatment, and, in any case, such adventitious Fe would be uncontrolled, variable, and unavailable for initial processing. Such adventitious Fe will be distinguishable from Fe added during catalyst preparation or catalyst pretreatment by characterization techniques such as SEM and Mossbauer spectroscopy.
In some preferred embodiments, the aluminosilicate zeolite catalyst further comprises at least 0.5, or at least 1.0, or at least 1.5 wt % Fe, and in some embodiments up to 10 wt % Fe, up to 5 wt % Fe, or up to 3 wt % Fe. The Fe is not derived, from biomass or reactor walls; or the aluminosilicate zeolite catalyst has been manufactured to contain or pretreated to contain the stated amount or ranges of Fe.
Preferably, the gaseous product stream comprises at least 10 wt % of aromatic compounds or at least 15 wt % of aromatic compounds and in some embodiments up to 30 wt % aromatics, in some embodiments up to 25 wt %; and/or wherein at least 85% of the xylenes in the gaseous product stream is p-xylene. Note that wt % is identical to mass %.
In some preferred embodiments, the aluminosilicate zeolite catalyst has a siliceous coating. A siliceous coating can be applied to a zeolite surface by reaction with silicones or siloxanes as described elsewhere herein and in the literature. A siliceous coating can be identified can be identified by a higher (at least 10% higher or at least 30% higher or at least 50% higher or at least 100% higher Si/Al ratio) in the exterior 50 A (or exterior 100 A) as compared to the Si/Al ratio at greater depths in the catalyst. The preferred technique for analyzing the Si/Al ratio is SEM/XPS before and after sputtering off 50 or 100 A.
Some embodiments further comprise removing the catalyst from the reactor after it has been used to pyrolyze the catalysis, heating the used catalyst in the presence of an oxygen containing gas (preferably O2) to form a regenerated catalyst, and returning the regenerated catalyst to the reactor, and again using the catalyst to catalyze the conversion of the biomass to a gaseous product stream comprising p-xylene. Surprisingly, the regenerated catalyst was found to have superior selectivity to p-xylene as compared to the freshly prepared catalyst.
In another aspect, the invention provides a method of making a catalyst, comprising: providing a zeolite catalyst; treating the catalyst to increase the iron content and applying a siliceous coating to the catalyst. This results in an iron-modified, zeolite catalyst having a siliceous coating. The step of applying a siliceous coating can be conducted before, during or after step of a treating the catalyst to increase the iron content; in some embodiments the invention can be characterized by applying the siliceous coating prior to treatment to increase Fe content. Preferably the zeolite catalyst is an aluminosilicate catalyst having a Si/Al ratio of 100 or less.
In any of the inventive catalysts, the Fe-modified, zeolite catalyst having a siliceous coating is used to catalyze the pyrolysis of biomass; and subsequent to the pyrolysis of biomass, the invention can include a step of regenerating the catalyst by heating in the presence of an oxygen containing gas.
In another aspect, the invention provides a hydrocarbon mixture, comprising: a biomass-derived (i.e., ¹⁴C-containing) mixture of hydrocarbons comprising at least 10 mass % of xylenes; wherein the xylenes are made up of 85 to about 91% p-xylene; or comprising at least 1.5 mass % of xylenes; wherein the xylenes are made up of at least 85% p-xylene. The invention includes a hydrocarbon mixture made by any of the processes described herein. The invention also includes a hydrocarbon mixture comprising catalyst particles of the type described herein.
In a further aspect, the invention provides a chemical system, comprising: a reactor, comprising: an iron-modified zeolite catalyst, biomass, and a hydrocarbon product stream comprising at least 10 mass % xylenes wherein at least 80% of the xylenes are p-xylene; or comprising at least 1.5 mass % of xylenes; wherein the xylenes are made up of at least 85% p-xylene. Preferably, the Fe-modified zeolite catalyst comprises a siliceous coating. Preferably, the Fe-modified zeolite catalyst comprises ZSM-5. The hydrocarbon product stream may be a stream condensed from the gaseous product stream (where the gaseous stream may include suspended liquid droplets and solid particulates).
In another aspect, the invention provides an aluminosilicate zeolite catalyst having a Si/Al molar ratio of 100 or less, comprising: at least 0.2 wt % Fe wherein the Fe is not derived from biomass or reactor walls, or wherein the aluminosilicate zeolite catalyst has been manufactured to contain at least 0.2 wt % Fe or pretreated to contain at least 0.2 wt % Fe; and a siliceous coating. In some embodiments, the catalyst comprises one or more of the following characteristics: Fe evenly distributed over the surface as measured by SEM-EDS; the catalyst comprises ZSM-5; wherein the catalyst has a surface ratio of Si/Fe in the ratio of 50:1 to 4:1; preferably 30:1 to 5:1; in some embodiments 20:1 to 7:1; wherein the Fe is concentrated in clusters on the surface of the catalyst; wherein the catalyst has a Brønsted acidity of greater than 0.01, 0.05 or greater, or a Brønsted acidity in the range of 0.01 to 0.2, preferably in the range of 0.04 to 0.15, preferably in the range of 0.05 to 0.1 μmol/mg; wherein the moles of Brønsted acid sites deactivated by the siliceous coating is at least 0.015 or at least 0.03 per mole of Si added as measured by desorption of isopropyl amine (IPA) in a temperature programmed desorption experiment.
The invention includes any combination of the inventive aspects. For example, the catalyst as defined in the descriptions above can be present in any of the inventive processes or systems.
Other advantages and novel features of the present invention will become apparent from the following detailed description of various non-limiting embodiments of the invention when considered in conjunction with the accompanying figures. In cases where the present specification and a document incorporated by reference include conflicting and/or inconsistent disclosure, the present specification shall control.

GLOSSARY

All ranges and ratio limits disclosed in the specification and claims may be combined in any manner. It is to be understood that unless specifically stated otherwise, references to “a,” “an,” and/or “the” may include one or more than one, and that reference to an item in the singular may also include the item in the plural.
The phrase “and/or” should be understood to mean “either or both” of the elements so conjoined, i.e., elements that are conjunctively present in some cases and disjunctively present in other cases. Other elements may optionally be present other than the elements specifically identified by the “and/or” clause, whether related or unrelated to those elements specifically identified unless clearly indicated to the contrary. Thus, as a non-limiting example, a reference to “A and/or B,” when used in conjunction with open-ended language such as “comprising” can refer, in one embodiment, to A without B (optionally including elements other than B); in another embodiment, to B without A (optionally including elements other than A); in yet another embodiment, to both A and B (optionally including other elements); etc.
The word “or” should be understood to have the same meaning as “and/or” as defined above. For example, when separating items in a list, “or” or “and/or” shall be interpreted as being inclusive, i.e., the inclusion of at least one, but also including more than one, of a number or list of elements, and, optionally, additional unlisted items.
The term “aromatic compound” is used to refer to a hydrocarbon compound comprising one or more aromatic groups such as, for example, single aromatic ring systems (e.g., benzyl, phenyl, etc.) and fused polycyclic aromatic ring systems (e.g. naphthyl, 1,2,3,4-tetrahydronaphthyl, etc.). Examples of aromatic compounds include, but are not limited to, benzene, toluene, indane, indene, 2-ethyl toluene, 3-ethyl toluene, 4-ethyl toluene, trimethyl benzene (e.g., 1,3,5-trimethyl benzene, 1,2,4-trimethyl benzene, 1,2,3-trimethyl benzene, etc.), ethylbenzene, methylbenzene, propylbenzene, xylenes (e.g., p-xylene, m-xylene, o-xylene, etc.), naphthalene, methyl-naphthalene (e.g., 1-methyl naphthalene, anthracene, 9.10-dimethylanthracene, pyrene, phenanthrene, dimethyl-naphthalene (e.g., 1,5-dimethylnaphthalene, 1,6-dimethylnaphthalene, 2,5-dimethylnaphthalene, etc.), ethyl-naphthalene, hydrindene, methyl-hydrindene, and dymethyl-hydrindene. Single ring and/or higher ring aromatics may be produced in some embodiments. The aromatic compounds may have carbon numbers from, for example, C5-C14, C6-C8, C6-C12, C8-C12, C10-C14.
The term “biomass” refers to living and recently dead biological material. In accordance with the inventive method, biomass may be converted, for example, to liquid fuel (e.g., biofuel or biodiesel) or to other fluid hydrocarbon products. Biomass may include trees (e.g., wood) as well as other vegetation; agricultural products and agricultural waste (e.g., corn stover, bagasse, fruit, garbage, silage, etc.); energy crops (e.g. switchgrass, miscanthus); algae and other marine plants; metabolic wastes (e.g., manure, sewage); and cellulosic urban waste. Biomass may be considered as comprising material that recently participated in the carbon cycle so that the release of carbon in a combustion process may result in no net increase averaged over a reasonably short period of time. For this reason, peat, lignite, coal, shale oil or petroleum may not be considered as being biomass as they contain carbon that may not have participated in the carbon cycle for a long time and, as such, their combustion may result in a net increase in atmospheric carbon dioxide. The term biomass may refer to plant matter grown for use as biofuel, but may also include plant or animal matter used for production of fibers, chemicals, heat, and the like. Biomass may also include biodegradable waste or byproducts that can be burnt as fuel or converted to chemicals. These may include municipal waste, green waste (the biodegradable waste comprised of garden or park waste such as grass or flower cuttings, hedge trimmings, and the like), byproducts of farming including animal manures, food processing wastes, sewage sludge, black liquor from wood pulp or algae, and the like. Biomass may be derived from plants, including miscanthus, spurge, sunflower, switchgrass, hemp, corn (maize), poplar, willow, sugarcane, and oil palm (palm oil), and the like. Biomass may be derived from roots, stems, leaves, seed husks, fruits, and the like. The particular plant or other biomass source used may not be important to the fluid hydrocarbon product produced in accordance with the inventive method, although the processing of the biomass may vary according to the needs of the reactor and the form of the biomass.
The term “catalytic pyrolysis” refers to pyrolysis performed in the presence of a catalyst. Catalytic fast pyrolysis (CFP), is a process that may be used to convert a hydrocarbonaceous material (e.g., biomass) into a fluid hydrocarbon product using rapid heating rates in the presence of a catalyst. With the inventive method, the fluid hydrocarbon product comprises p-xylene, and may further comprise additional aromatics, olefins, and the like.
Contact time is the residence time of a material in a reactor or other device, when measured or calculated under standard conditions of temperature and pressure (i.e., 0° C. and 100 kPa absolute pressure). In some cases contact time can be expressed for an individual component while in other cases contact time can be expressed based on more than one component and in other cases contact time can be expressed based on the entire reaction mixture of all components. For example, a 2 liter reactor to which is fed 3 standard liters per minute of gas A has a contact time of ⅔ minute, or 40 seconds for gas A. For a chemical reaction, contact time or residence time is based on the volume of the reactor where substantial reaction is occurring; and would exclude volume where substantially no reaction is occurring such as an inlet or an exhaust conduit. For catalyzed reactions, the volume of a reaction chamber is the volume where catalyst is present.
A hydrocarbonaceous feed material may comprise a solid, a semi-solid, a liquid, or a mixture of two or more thereof. The solids content of the hydrocarbonaceous feed may be up to about 100% by weight, or from about 30% to about 100% by weight, or from about 50% to about 100%, or from about 70% to about 100%, or from 90% to about 100%, or from about 95% to about 100%, or from about 98% to about 100%, or from about 30% to about 95%, or from about 50% to about 95%, or from about 70% to about 95%, or from about 80% to about 95%, or from about 85% to about 95%, or from about 90% to about 95% by weight. The carbon content of the hydrocarbonaceous feed may be up to about 90% by weight, or from 20% to 90% by weight, or from 25% to 75%, or from 30% to 65%, or from 35% to 60%, or from 40 to 50% by weight. The hydrocarbonaceous material may comprise biomass. The hydrocarbonaceous material may comprise plastic waste, recycled plastics, agricultural solid waste, municipal solid waste, food waste, animal waste, carbohydrates, lignocellulosic materials, xylitol, glucose, cellobiose, cellulose, hemi-cellulose, lignin, sugar cane bagasse, glucose, wood, corn stover, or a mixture of two or more thereof. The hydrocarbonaceous material may comprise furan, 2-methylfuran, furfural, ethylene glycol, glycerine, or any combination of these. The hydrocarbonaceous material may comprise pinewood. The hydrocarbonaceous material may comprise 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 a mixture of two or more thereof. Mixtures of two or more of any of the foregoing may be used.
The term “conversion of a reactant” may refer to the reactant mole or mass change between a material flowing into a reactor and a material flowing out of the reactor divided by the moles or mass of reactant in the material flowing into the reactor. For example, if 100 g of ethylene are fed to a reactor and 30 g of ethylene are flowing out of the reactor, the conversion is [(100−30)/100]=70% conversion of ethylene.
The term “fluid” may refer to a gas, a liquid, a mixture of a gas and a liquid, or a gas or a liquid containing dispersed solids, liquid droplets and/or gaseous bubbles. The terms “gas” and “vapor” have the same meaning and are sometimes used interchangeably. In some embodiments, it may be advantageous to control the residence time of the fluidization fluid in the reactor. The fluidization residence time of the fluidization fluid is defined as the volume of the reactor divided by the volumetric flow rate of the fluidization fluid under process conditions of temperature and pressure.
The term “fluidized bed reactor” may be used to refer to reactors comprising a vessel that contains a granular solid material (e.g., silica particles, catalyst particles, etc.), in which a fluid (e.g., a gas or a liquid) is passed through the granular solid material at velocities sufficiently high as to suspend the solid material and cause it to behave as though it were a fluid. The term “circulating fluidized bed reactor” may be used to refer to fluidized bed reactors in which the granular solid material is passed out of the reactor, circulated through a line in fluid communication with the reactor, and recycled back into the reactor. Bubbling fluidized bed reactors, circulating fluidized bed reactors, or turbulent fluidized bed reactors may be used. Examples of fluidized bed reactors, circulating fluidized bed reactors, bubbling and turbulent fluidized bed reactors are described in Fluidization Engineering, 2nd Edition, D. Kunii and O. Levenspiel, Butterworth-Heinemann, 1991, Chapter 3, pages 61-94, these pages being incorporated herein by reference.
The terms “olefin” or “olefin compound” (a.k.a. “alkenes”) may be used to refer to any unsaturated hydrocarbon containing one or more pairs of carbon atoms linked by a double bond. Olefins may include both cyclic and acyclic (aliphatic) olefins, in which the double bond is located between carbon atoms forming part of a cyclic (closed-ring) or of an open-chain grouping, respectively. In addition, olefins may include any suitable number of double bonds (e.g., monoolefins, diolefins, triolefins, etc.). Examples of olefin compounds may include ethene, propene, allene (propadiene), 1-butene, 2-butene, isobutene (2 methyl propene), 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 that include aromatic moieties are considered olefins, for example, benzyl acrylate or styrene.
The term “overall residence time” refers to the volume of a reactor or device or specific portion of a reactor or device divided by the exit flow of all gases out of the reactor or device including fluidization gas, products, and impurities, measured or calculated at the average temperature of the reactor or device and the exit pressure of the reactor or device.
As used herein, the term “pore size” is used to refer to the smallest cross-sectional diameter of a pore. The smallest cross-sectional diameter of a pore may correspond to the smallest cross-sectional dimension (e.g., a cross-sectional diameter) as measured perpendicularly to the length of the pore. In some embodiments, a catalyst with an “average pore size” or a “pore size distribution” of X refers to a catalyst in which the average of the smallest cross-sectional diameters of the pores within the catalyst is about X. It should be understood that “pore size” or “smallest cross sectional diameter” of a pore as used herein refers to the Norman radii adjusted pore size well known to those skilled in the art. Determination of Norman radii adjusted pore size is described, for example, in Cook, M.; Conner, W. C., “How big are the pores of zeolites?” Proceedings of the International Zeolite Conference, 12th, Baltimore, Jul. 5-10, 1998; (1999), 1, pp 409-414, which is incorporated herein by reference in its entirety. As a specific exemplary calculation, the atomic radii for ZSM-5 pores are about 5.5-5.6 Å, as measured by x-ray diffraction. In order to adjust for the repulsive effects between the oxygen atoms in the catalyst, Cook and Conner have shown that the Norman adjusted radii are 0.7 Å larger than the atomic radii (about 6.2-6.3 Å).
One of ordinary skill in the art will understand how to determine the pore size (e.g., minimum pore size, average of minimum pore sizes) in a catalyst. For example, x-ray diffraction (XRD) can be used to determine atomic coordinates. XRD techniques for the determination of pore size are described, for example, in Pecharsky, V. K. et al, “Fundamentals of Powder Diffraction and Structural Characterization of Materials,” Springer Science+Business Media, Inc., New York, 2005, incorporated herein by reference in its entirety. Other techniques that may be useful in determining pore sizes (e.g., zeolite pore sizes) include, for example, helium pycnometry or low pressure argon adsorption techniques. These and other techniques are described in Magee, J. S. et al, “Fluid Catalytic Cracking: Science and Technology,” Elsevier Publishing Company, Jul. 1, 1993, pp. 185-195, which is incorporated herein by reference. Pore sizes of mesoporous catalysts may be determined using, for example, nitrogen adsorption techniques, as described in Gregg, S. J. at al, “Adsorption, Surface Area and Porosity,” 2nd Ed., Academic Press Inc., New York, 1982 and Rouquerol, F. et al, “Adsorption by powders and porous materials. Principles, Methodology and Applications,” Academic Press Inc., New York, 1998, both incorporated herein by reference in their entirety. Unless otherwise indicated, pore sizes referred to herein are those determined by x-ray diffraction corrected as described above to reflect their Norman radii adjusted pore sizes.
The terms “pyrolysis” and “pyrolyzing” refer to the transformation of a material (e.g., a solid hydrocarbonaceous material) into one or more other materials (e.g., volatile organic compounds, gases, coke, etc.) by heat, without oxygen or other oxidants or without significant amounts of oxygen or other oxidants, and with or without the use of a catalyst.
The term “reactant residence time” of a reactant in the reactor is defined as the amount of time the reactant spends in the reactor. Residence time may be based on the feed rate of reactant and is independent of rate of reaction. The reactant residence time of the reactants in a reactor may be calculated using different methods depending upon the type of reactor being used. For gaseous reactants, where flow rate into the reactor is known, this is typically a simple calculation. In the case of solid reactants in which the reactor comprises a packed bed reactor into which only reactants are continuously fed (i.e. no carrier or fluidizing flow is utilized), the reactant residence time in the reactor may be calculated by dividing the volume of the reactor by the volumetric flow rate of the hydrocarbonaceous material and fluid hydrocarbon product exiting the reactor. In cases where the reaction takes place in a reactor that is closed to the flow of mass during operation (e.g., a batch reactor), the batch residence time of the reactants in such may be reactor is defined as the amount of time elapsing between the time at which the temperature in the reactor containing the reactants reaches a level sufficient to commence a pyrolysis reaction (e.g., for CFP, typically about 300° C. to about 1000° C. for many typical hydrocarbonaceous feedstock materials) and the time at which the reactor is quenched (e.g., cooled to a temperature below that sufficient to support further pyrolysis—e.g. typically about 300° C. to about 1000° C. for many hydrocarbonaceous feedstock materials).
The residence time of the catalyst in a fluidized bed reactor may be defined as the volume of the reactor filled with catalyst divided by the volumetric flow rate of the catalyst through the reactor. For example if a 3 liter reactor contains 2 liters of catalyst and a flow of 0.4 liters per minute of catalyst is fed through the reactor, i.e., both fed and removed, the catalyst residence time will be 2/0.4 minutes, or 5 minutes.
The term “silicon-containing compound” is used herein to refer to any compound that contains one or more Si—O groups. The silicon-containing compound may be a silicate containing one or more of SiO44-, Si2O76- or Si6O1812-groups. These may include one or more tetraorthosilicates. The silicon-containing compound may include one or more siloxanes containing one or more silicon-oxygen backbones (—Si—O—Si—O—) with organic (e.g., hydrocarbon) side groups attached to the silicon atoms. These may include one or more siloxane polymers (e.g., polydimethyl siloxane). The silicon-containing compound may be a straight chain, branched chain or cyclical compound. The silicon-containing compound may be monomeric, oligomeric or polymeric. The silicon-containing compound may comprise a compound containing at least one group represented by the formula
The silicon-containing compound may be represented by the formula:
wherein R₁and R₂independently comprise hydrogen, halogen, hydroxyl, alkyl, alkoxyl, halogenated alkyl, aryl, halogenated aryl, aralkyl, halogenated aralkyl, alkaryl or halogenated alkaryl; and n is a number that is at least 2. R₁and/or R₂may comprise methyl, ethyl or phenyl. n may be a number in the range from about 3 to about 1000.
The term “selectivity” refers to the amount of production of a particular product in comparison to a selection of products. Selectivity to a product may be calculated by dividing the amount of a particular product by the amount of a number of products produced. For example, if 75 grams of aromatics are produced in a reaction and 20 grams of benzene are found in these aromatics, on a mass basis the selectivity to benzene amongst aromatic products is 20/75=26.7%. Selectivity may be calculated on a mass basis, as in the aforementioned example, or it may be calculated on a carbon basis where the selectivity is calculated by dividing the amount of carbon that is found in a particular product by the amount of carbon that is found in a selection of products. Unless specified otherwise, for reactions involving biomass as a reactant, selectivity is on a mass basis. The carbon selectivities for various materials can be determined using the following equations:
$Overall selectivity = \frac{moles of carbon in a product}{moles of carbon in all products} \times 100 %$ $Aromatic selectivity = \frac{moles of carbon in an aromatic product}{moles of carbon in all aromatic products} \times 100 %$ $Olefin selectivity = \frac{moles of carbon in an olefinic product}{moles of carbon in all olefins products} \times 100 %$ $p - Xylene selectivity in xylenes = \frac{moles of p - xylene isomer}{moles of all xylene isomers} \times 100 %$
The term “yield” is used herein to refer to the amount of a product flowing out of a reactor divided by the amount of reactant flowing into the reactor, usually expressed as a percentage or fraction. Yields are often calculated on a mass basis, carbon basis, or on the basis of a particular feed component. Mass yield is the mass of a particular product divided by the weight of feed used to prepare that product. For example, if 500 grams of biomass is fed to a reactor and 45 grams of p-xylene is produced, the mass yield of p-xylene would be 45/500=9% p-xylene. 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 grams of biomass that contains 40% carbon is reacted to produce 45 g of p-xylene that contains 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, if 500 grams of biomass containing 40% carbon and 100 grams of CO2 are reacted to produce 40 g of p-xylene (containing 90.6% carbon), the carbon yield on biomass is [(40*0.906)/(500*0.40)]=18.1%; note that the mass of CO2 does not enter into the calculation.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1. is a schematic illustration of a CFP process for converting a solid hydrocarbonaceous material to a fluid hydrocarbon product

FIG. 2 is a graph of the acidity of silicone treated zeolite vs the amount of SiO₂added to the zeolite for two different zeolites and two different coating compounds.

DETAILED DESCRIPTION

Referring to FIG. 1, feed stream 10 includes a solid hydrocarbonaceous material (typically biomass) that can be fed to reactor 20. Certain solid hydrocarbonaceous materials may also comprise relatively minor proportions of other elements such as nitrogen and sulfur. The feed streams to the reactor may be free of olefins, or may contain olefins in an insignificant amount (e.g., such that olefins make up less than about 1 wt %, less than about 0.1 wt %, or less than about 0.01 wt % of the total weight of reactant fed to the reactor). In other embodiments, however, olefins may be present in one or more reactant feed streams.
The solid hydrocarbonaceous material feed composition (e.g., in feed stream 10 of FIG. 1) may comprise a mixture of solid hydrocarbonaceous material and a catalyst. The mixture may comprise, for example, a solid catalyst and a solid hydrocarbonaceous material. In other embodiments, a catalyst may be provided separately from the solid hydrocarbonaceous material (e.g., by co-feeding the catalyst via an independent catalyst inlet). A variety of catalysts may be used. For example, in some instances, zeolite catalysts with varying molar ratios of silica to alumina, and/or varying pore sizes and/or pore opening sizes, and/or varying catalytically active metals and/or metal oxides, may be used.
A hydrocarbonaceous material may be fed to a reactor (e.g., a fluidized-bed reactor) where the hydrocarbonaceous material may first thermally decompose to form one or more pyrolysis products. The pyrolysis products may comprise one or more pyrolysis vapors. These pyrolysis products may react in the presence of a modified zeolite catalyst to form one or more aromatic compounds as well as olefin compounds, water, CO, and CO₂. The modified zeolite catalyst may comprise acid sites on its external surface and in its pores. The modified zeolite may be modified to change the number of acid sites on its external surface and may be modified to reduce the size or number of the pore mouth openings. The modified zeolite may be modified by the incorporation of one or more promoter elements to improve catalyst stability, activity, or selectivity. The pyrolysis products may enter the pores in the modified zeolite catalyst where they may undergo reaction or react on the surface of the modified zeolite. The products formed in the catalyst pores may then diffuse out of the pores. The aromatic compounds may comprise p-xylene or xylenes with a relatively high selectivity towards p-xylene. Advantages of this process may include one or more of the following: 1) all the desired chemistry may occur in a single-step process, 2) the process may use a relatively inexpensive zeolite catalyst, 3) p-xylene may be produced at a relatively high level of production and selectivity amongst the xylenes, and 4) the activity, selectivity, or activity and selectivity of the catalyst may be maintained or improved with repeated cycling.
The hydrocarbonaceous material may comprise solids of any suitable size. In some cases, it may be advantageous to use hydrocarbonaceous solids with relatively small particle sizes. Small-particle solids may, in some instances, react more quickly than larger solids due to their relatively higher surface area to volume ratios compared to larger solids. In addition, small particle sizes may allow for more efficient heat transfer within each particle and/or within the reactor volume. This may prevent or reduce the formation of undesired reaction products. Moreover, small particle sizes may provide for increased solid-gas and solid-solid contact, leading to improved heat and mass transfer. In some embodiments, the mass 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 250 microns (60 mesh), less than about 149 microns (100 mesh), less than about 105 microns (140 mesh), less than about 88 microns (170 mesh), less than about 74 microns (200 mesh), less than about 53 microns (270 mesh), or less than about 37 microns (400 mesh), or smaller.
It may be desirable to employ a feed material with an average particle size above a minimum amount in order to reduce the pressure required to pass the solid hydrocarbonaceous feed material through the reactor. For example, it may be desirable to use a solid hydrocarbonaceous feed material with an average particle size of at least about 37 microns (400 mesh), at least about 53 microns (270 mesh), at least about 74 microns (200 mesh), at least about 88 microns (170 mesh), at least about 105 microns (140 mesh), at least about 149 microns (100 mesh), at least about 250 microns (60 mesh), at least about 0.5 mm, a least about 1 mm, at least about 2 mm, at least about 5 mm, or higher.
In some embodiments, catalyst and hydrocarbonaceous material may be present in any suitable ratio. For example, the catalyst and hydrocarbonaceous material may be present in any suitable mass ratio in cases where the feed composition (e.g., through one or more feed streams comprising catalyst and hydrocarbonaceous material or through separate catalyst and hydrocarbonaceous material feed streams), comprises catalyst and hydrocarbonaceous material (e.g., circulating fluidized bed reactors). As another example, in cases where the reactor is initially loaded with a mixture of catalyst and hydrocarbonaceous material (e.g., a batch reactor), the catalyst and hydrocarbonaceous material may be present in any suitable mass ratio. In some embodiments involving circulating fluidized bed reactors, the mass ratio of the catalyst to hydrocarbonaceous material in the feed stream—i.e., in a composition comprising a catalyst and a hydrocarbonaceous material provided to a reactor—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 higher. In some embodiments involving circulating fluidized bed reactors, the mass ratio of the catalyst to hydrocarbonaceous material in the feed stream may be less than about 0.5:1, less than about 1:1, less than about 2:1, less than about 5:1, less than about 10:1, less than about 15:1, or less than about 20:1; or from about 0.5:1 to about 20:1, from about 1:1 to about 20:1, or from about 5:1 to about 20:1. Employing a relatively high catalyst to hydrocarbonaceous material mass ratio may facilitate introduction of the volatile organic compounds, formed from the pyrolysis of the feed material, into the catalyst before they thermally decompose to coke. Not wishing to be bound by any theory, this effect may be at least partially due to the presence of a stoichiometric excess of catalyst sites within the reactor.
The reactor may comprise a continuously stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor, or a fluidized bed reactor. Advantageously, the reactor may comprise a fluidized bed reactor. The catalytic reaction step may be achieved by co-feeding the catalyst with the hydrocarbonaceous material. The catalyst may be fed separately. Part of the catalyst may be fed with the hydrocarbonaceous feed material and part of the catalyst may be fed separately.
The inventive method preferably comprises a catalytic fast pyrolysis (CFP) process. Aspects of a CFP process have been described in U.S. Pat. No. 8,277,643, U.S. Pat. No. 8,864,984, US Patent Application 2012/0203042 A1, US Patent Application 2013/0060070 A1, US Patent Application 2014/0027265 A1 and US Patent Application 2014/0303414 A1 all incorporated herein in full by reference.
For CFP processes, particularly advantageous catalysts may include those containing internal porosity selected according to pore size (e.g., mesoporous and pore sizes typically associated with zeolites), e.g., average pore sizes of less than about 100 Angstroms (Å), less than about 50 Å, less than about 20 Å, less than about 10 Å, less than about 5 Å, or smaller. In some embodiments, catalysts with average pore sizes of from about 5 Å to about 100 Å may be used. In some embodiments, catalysts with average pore sizes of between about 5.5 Å and about 6.5 Å, or between about 5.9 Å and about 6.3 Å may be used. In some cases, catalysts with average pore sizes of between about 7 Å and about 8 Å, or between about 7.2 Å and about 7.8 Å may be used.
The reactor may be operated at a temperature in the range from about 400° C. to about 650° C., or from about 500° C. to about 600° C., or from about 525° C. to about 575° C.
The hydrocarbonaceous material may be fed to the reactor at a mass normalized space velocity of up to about 12 hour⁻¹, or up to about 6 hour⁻¹, or up to about 3 hour⁻¹% or up to about 1.5 hour⁻¹, or up to about 0.9 hour⁻¹, or in the range from about 0.01 hour⁻ to about 12 hour⁻, or in the range from about 0.01 to about 2 hour⁻¹, or in the range from about 0.01 to about 1.5 hour⁻¹, or in the range from about 0.01 to about 0.9 hour⁻¹, or in the range from about 0.01 hour-1 to about 0.5 hour⁻¹, or in the range from about 0.1 hour⁻ to about 0.9 hour⁻¹, or in the range from about 0.1 hour⁻ to about 0.5 hour⁻.
The reactor may be operated at a pressure of at least about 100 kPa, or at least about 200 kPa, or at least about 300 kPa, or at least about 400 kPa. The reactor may be operated at a pressure below about 600 kPa, or below about 400 kPa, or below about 200 kPa. The reactor may be operated at a pressure in the range from about 100 to about 600 kPa, or in the range from about 100 to about 400 kPa, or in the range from about 100 to about 200 kPa. The method may be conducted under reaction conditions that minimize coke production. The pyrolysis product may be formed with less than about 30 wt %, or less than about 25 wt %, or less than about 20 wt %, or less than about 10 wt %, of the pyrolysis product being coke.
The catalyst may comprise any catalyst suitable for conducting the catalytically reacting step of the inventive method. The catalyst may be used to lower the activation energy (increase the rate) of the reaction conducted in the catalytically reacting step and/or improve the distribution of products or intermediates during the reaction (for example, a shape selective catalyst). Examples of reactions that can be catalyzed include: dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization, decarbonylation, decarboxylation, aldol condensation, and combinations thereof. The catalyst components may be acidic, neutral or basic.
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 acid sites. These acid sites may also be referred to as catalytically active sites. Other zeolite catalysts that may be used may include ferrierite, zeolite Y, zeolite beta, mordenite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)AlPO-31, SSZ-23, and the like. The catalyst may comprise silica and alumina, and further comprise one or more additional promoter elements such as metals and/or metal oxides. Suitable metals and/or oxides may include, for example, nickel, palladium, silver, platinum, palladium, titanium, vanadium; chromium, manganese, iron, cobalt, zinc, copper, gallium, sodium, potassium, magnesium, calcium, the rare earth elements, i.e., elements 57-71, i.e., La, Ce, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm, Yb, Lu, or zirconium, hafnium, tantalum, phosphorus, and/or any of their oxides, among others. In addition, in some cases, properties of the catalysts (e.g., pore structure, type and/or number of catalytic sites, etc.) may be chosen to selectively produce a desired product. Preferable promoter elements are Ga, La and Fe. A particularly preferable promoter element is Fe.
The promoter element(s) may be added in an amount sufficient to provide an effect on the yield, selectivity, or both yield and selectivity of the biomass conversion process. The amount of promoter elements can be from 0.1% by weight to 30% by weight, or 0.2% to 20%, or 0.5% to 10%, or 1.0% to 5%, or at least 0.2%, or at least 0.5%, or at least 1.0%. The amount of promoter element(s) added to a zeolite catalyst can be adjusted with respect to the alumina content of the zeolite or number of acid sites of the zeolite. The ratio of the added promoter element(s) to the number of acid sites on a mole to mole basis (moles of promoter element(s) to moles of acid sites on the unpromoted catalyst) can be from 0.05 to 20.0, or from 0.1 to 10, or from 0.5 to 6, or at least 0.1, or at least 0.5, or at least 2, or at least 6.0.
In some embodiments, the mass average diameter (as measured by conventional sieve analysis) of catalyst objects, which may in certain instances each comprise a single catalyst particle or in other instances comprise an agglomerate of a plurality of particles, may be less than about 5 mm, less than about 2 mm, less than about 1 mm, less than about 0.5 mm, less than about 250 microns (60 mesh), less than about 149 microns (100 mesh), less than about 105 microns (140 mesh), less than about 88 microns (170 mesh), less than about 74 microns (200 mesh), less than about 53 microns (270 mesh), or less than about 37 microns (400 mesh), or smaller.
The catalyst may comprise particles having a maximum cross-sectional dimension of less than about 5 microns, less than about 1 micron, less than about 500 nm, less than about 100 nm, between about 100 nm and about 5 microns, between about 500 nm and about 5 microns, between about 100 nm and about 1 micron, or between about 500 nm and about 1 micron. Catalyst particles having the dimensions within the ranges noted immediately above may be agglomerated to form discrete catalyst objects having dimensions within the ranges noted above. As used here, the “maximum cross-sectional dimension” of a particle refers to the largest dimension between two boundaries of a particle. One of ordinary skill in the art would be capable of measuring the maximum cross-sectional dimension of a particle by, for example, analyzing a scanning electron micrograph (SEM) of a catalyst preparation. In embodiments comprising agglomerated particles, the particles should be considered separately when determining the maximum cross-sectional dimensions. In such a case, the measurement may be performed by establishing imaginary boundaries between each of the agglomerated particles, and measuring the maximum cross-sectional dimension of the hypothetical, individuated particles that result from establishing such boundaries. In some embodiments, a relatively large number of the particles within a catalyst may have maximum cross-sectional dimensions that lie within a given range. 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 particles within a catalyst have maximum cross-sectional dimensions of less than about 5 microns, less than about 1 micron, less than about 500 nm, less than about 100 nm, between about 100 nm and about 5 microns, between about 500 nm and about 5 microns, between about 100 nm and about 1 micron, or between about 500 nm and about 1 micron.
A relatively large percentage of the volume of the catalyst can be occupied by particles with maximum cross-sectional dimensions within a specific 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 sum of the volumes of all the catalyst used is occupied by particles having maximum cross-sectional dimensions of less than about 5 microns, less than about 1 micron, less than about 500 nm, less than about 100 nm, between about 100 nm and about 5 microns, between about 500 nm and about 5 microns, between about 100 nm and about 1 micron, or between about 500 nm and about 1 micron.
Using catalysts including particles within a chosen size distribution indicated above can lead to an increase in the yield and/or selectivity of aromatic compounds produced by the reaction of the hydrocarbonaceous material. For example, in some cases, using catalysts containing particles with a desired size range (e.g., any of the size distributions outlined above) can result in an increase in the amount of aromatic compounds in the reaction product of at least about 5%, at least about 10%, or at least about 20%, relative to an amount of aromatic compounds that would be produced using catalysts containing particles with a size distribution outside the desired range (e.g., with a large percentage of particles larger than 1 micron, larger than 5 microns. etc. as above).
Alternatively, catalysts may be selected according to pore size (e.g., mesoporous and pore sizes typically associated with zeolites), e.g., average pore sizes of 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 smaller. In some embodiments, catalysts with average pore sizes of from about 5 Angstroms to about 100 Angstroms may be used. In some embodiments, catalysts with average pore sizes of between about 5.5 Angstroms and about 6.5 Angstroms, or between about 5.9 Angstroms and about 6.3 Angstroms may be used. In some cases, catalysts with average pore sizes of between about 7 Angstroms and about 8 Angstroms, or between about 7.2 Angstroms and about 7.8 Angstroms may be used.
The zeolite catalyst typically comprises silicon and aluminum. The silicon to aluminum molar ratio may be at least 10:1 or at least 15:1, or at least 25:1, in the range from about 10:1 to about 240:1, or in the range from about 10:1 to about 40:1, or in the range from about 20:1 to about 50:1, or about 30:1, or at least about 30:1. The zeolite catalyst may further comprise nickel, palladium, silver, platinum, palladium, titanium, vanadium, chromium, manganese, iron, cobalt, zinc, copper, gallium, sodium, potassium, magnesium, calcium, zirconium, lanthanum, cerium, phosphorus, an oxide of one or more thereof, or a mixture of two or more thereof.
A screening method may be used to select catalysts with appropriate pore sizes for the conversion of specific pyrolysis product molecules. The screening method may comprise determining the size of pyrolysis product molecules desired to be catalytically reacted (e.g., the molecular kinetic diameters of the pyrolysis product molecules). One of ordinary skill in the art may calculate, for example, the kinetic diameter of a given molecule. The type of catalyst may then be chosen such that the pores of the catalyst (e.g., Norman adjusted minimum radii) are sufficiently large to allow the pyrolysis product molecules to diffuse into and/or react with the catalyst. In some embodiments, the catalysts may be chosen such that their pore sizes are sufficiently small to prevent entry and/or reaction of pyrolysis products whose reaction would be undesirable.
The catalyst may be treated or impregnated one or more times with a coating compound such as a silicone compound to reduce the size of the pore mouth openings in the catalyst as well as cover or obscure catalytic sites on the external surface of the catalyst and inside the pores of the catalyst near the pore mouth openings. The covering of the catalytic sites with the treatment layer may inhibit and/or extinguish their catalytic activity. In order to facilitate a more controlled application of the coating compound, the coating compound may be dispersed in a carrier, for example, an aqueous or organic liquid carrier.
In each phase of the catalyst treatment process, the coating compound may be deposited on the external surface of the catalyst by any suitable method. For example, the coating compound may be dissolved in an organic carrier, mixed with the catalyst, and then dried by evaporation or vacuum distillation. The catalyst may be contacted with the coating compound at a catalyst to coating compound weight ratio in the range from about 1000:1 to about 1:10. The content of the coating compound in the coated catalyst may be from 0.1% to 30% by weight, or 0.5% to 20% by weight, or 1% to 12% by weight, or 4% to 8% by weight, or at least 2%, or at least 4% or at least 6%, or at least 8% by weight of the final catalyst weight.
The zeolite catalyst may be treated with the silicon-containing compound to reduce the size of the pore openings, and cover or obscure catalytic sites on the external surface of the catalyst. This treatment process may also be used to cover or obscure catalytic sites in the pores near the pore mouths openings. The covered or obscured catalytic sites may be referred to as deactivated catalytic sites. A silicone coating compound preferably has a molecular size that is incapable of entering the pores of the catalyst. During the catalyst treatment process, the silicon-containing compound may be applied to the catalyst and subsequently calcined. This process may be repeated until the desired level of treatment is provided. The fraction of catalytic sites on the external surface of the catalyst that may be deactivated by treatment with the coating compound may be at least about 5%, or at least about 10%, or at least about 15%, or at least about 25%, or at least about 35%, or at least about 45%, or at least about 55%, or at least about 65%, or at least about 75%, or at least about 85%, or at least about 90%, or at least about 95%, or at least about 98%, or at least about 99%, of the available catalytic sites on the external surface of the catalyst. The fraction of catalyst sites on the external surface of the catalyst that remain as acid sites after coating with the coating compound may be at least about 95%, at least about 90%, at least about 80%, at least about 70%, at least about 60%, at least about 50%, at least about 40%, at least about 30%, at least about 20%, at least about 10%, at least about 5%, at least about 1%, or between 1% and 95%, or between 5% and 90%, or between 50% and 90%, or between 60% and 90% of the active sites on the external surface of the non-coated material. The fraction of catalytic sites on the external surface of the catalyst may be measured by a temperature programmed desorption experiment using 2,4,6-collidine (2,4,6-trimethylpyridine). A sample of the material to be measured is degassed for 2 h at 823 K. After cooling the sample to 393 K, it is exposed for 1 h to He (helium) that had been saturated with 2,4,6-collidine at room temperature by flowing pure He through a bubbler containing the amine. Then the sample is held at 393 K with He flow for 2 h to remove physisorbed 2,4,6-collidine. The sample is heated to 973 K at 10 K./min. The total amount of 2,4,6-collidine desorbed is used to calculate the total number of acid sites on the external surface of the catalyst, and the amount of amine that desorbs between about 580 K and 650 K is used to calculate the number of Brønsted acid sites for each catalyst. Due to the size of 2,4,6 collidine, it does not enter ZSM-5 pores. Therefore, desorption of 2,4,6-collidine only detects acid sites on the external surface of the catalyst or in the pores near the pore mouth openings. The decrease in the 2,4,6-collidine adsorption that occurs with the catalyst treatment shows the decrease in the number of acid sites on the external surface and in or near the pore mouth openings. The decrease in these external sites is believed to be a factor in the production of m- and o-xylene and in reducing the re-equilibration of p-xylene formed in the pores to m- or o-xylene, and thus improving the selectivity to p-xylene.
The catalyst may be treated with a tetraorthosilicate or other coating compound using a chemical liquid deposition (CLD) process or any of a number of other coating processes known to those skilled in the art.
The coating compound may be provided in the form of a solution or an emulsion under the conditions of contact with the catalyst. The deposited coating compound may cover, and reside substantially exclusively on, the external surface of the catalyst, blocking external sites and partially blocking pore mouths and sites in or near the pore mouths openings. Examples of methods of depositing silicone compounds on the surface of zeolites may be found in U.S. Pat. Nos. 4,090,981; 5,243,117; 5,403,800, and 5,659,098, which are incorporated by reference herein.
The coated catalyst may comprise silica and alumina. The silicon to aluminum molar ratio may be in the range from about 10:1 to about 240:1, or in the range from about 10:1 to about 40:1, or in the range from about 20:1 to about 50:1, or about 30:1, or at least 30:1.
The catalyst may be treated in-situ with a silicon-based coating by flowing a hydrocarbon solution containing the silicone compound over the catalyst prior to introduction of biomass feed. If the hydrocarbon solution contains toluene, then the p-xylene selectivity increase during the coating process can be monitored and adjusted to the desired selectivity level.
During the addition of a metal or metal oxide compound, the metal compound may be applied to the catalyst and subsequently calcined. This process may be repeated until the desired level of metal is provided. The fraction of acid sites on the catalyst that may be deactivated by treatment with the metal compound may be at least about 5%, or at least about 10%, or at least about 15%, or at least about 25%, or at least about 35%, or at least about 45%, or at least about 55%, or at least about 65%, or at least about 75%, or at least about 85%, or at least about 90%, or at least about 95%, or at least about 98%, or at least about 99%, of the available catalytic sites. The fraction of acid sites on the catalyst that remain as acid sites after treatment with the metal may be at least about 95%, at least about 90%, at least about 80%, at least about 70%, at least about 60%, at least about 50%, at least about 40%, at least about 30%, at least about 20%, at least about 10%, at least about 5%, at least about 1%, or between 1% and 95%, or between 5% and 90%, or between 50% and 90%, or between 60% and 90% of the active sites of the material before metal addition.
The coating compound may have a number average molecular weight in the range from about 80 to about 20,000, or from about 150 to 10,000. The coating compound may be a Dynasylan from Evonik Industries AG, such as Hydrosil 2909, Hydrosil 2627, or any of the other Dynasylan products, or a silicone emulsifier product such as Dow Corning® 5329 Silicone Emulsifier or any similar product obtainable from Dow Corning, or a modified polysiloxane such as KF 6015 or any of those obtainable from Shin Etsu, or the like.
The coating compound may comprise dimethylsilicone, diethylsilicone, phenylmethylsilicone, methylhydrogensilicone, ethylhydrogen silicone, phenylhydrogen silicone, methylethyl silicone, phenylethyl silicone, diphenyl silicone, methyltrifluoropropyl silicone, ethyltrifluoropropyl silicone, polydimethyl silicone, tetrachloro-phenylmethyl silicone, tetrachlorophenylethyl silicone, tetrachlorophenylhydrogen silicone, tetrachlorophenylphenyl silicone, methylvinyl silicone, hexamethyl cyclotrisiloxane, octamethyl cyclotetrasiloxane, hexaphenyl cyclotrisiloxane, octaphenyl cyclotetrasiloxane, or a mixture of two or more thereof. The coating compound may comprise a tetraorthosilicate. The coating compound may comprise tetramethylorthosilicate, tetraethylorthosilicate, or a mixture thereof.
The kinetic diameter of the coating compound may be larger than the pore diameter of the catalyst in order to avoid entry of the silicone compound into the pore and any concomitant reduction in the internal activity of the catalyst.
The organic carrier for the coating compound may comprise hydrocarbons such as linear, branched, and cyclic alkanes having five or more carbons. The carrier may comprise a linear, branched or cyclic alkane having a boiling point greater than about 70° C., and containing about 6 or more carbons. Optionally, mixtures of low volatility organic compounds, such as hydrocracker recycle oil, may also be employed as carriers. Low volatility hydrocarbon carriers for the coating compound may comprise decane, dodecane, mixtures thereof, and the like. A preferred carrier is water.
Following each deposition of the coating compound, the catalyst may be calcined to decompose the molecular or polymeric species to a solid state species. The catalyst may be calcined at a rate of from about 0.2° C./minute to about 5° C./minute to a temperature greater than about 200° C., but below a temperature at which the crystallinity of the zeolite may be adversely affected. Generally, such temperature will be below about 600° C. The temperature of calcination may be in the range from about 350° C. to about 550° C. The catalyst may be maintained at the calcination temperature for about 1 to about 24 hours, or about 2 to about 6 hours, or longer.
The catalyst may be treated with a water soluble silicone compound or polysiloxane in a similar fashion. The soluble silicone compound, silane, or polysiloxane may be added to the catalyst in any manner such as incipient wetness, adsorption, spray coating, or wet coating, or any other process known to those skilled in the art. In a preferred procedure, the soluble silicone compound or polysiloxane is dispersed in a solution of water at from 1 to 95% by weight concentration of coating compound, or from 10 to 80% by weight, or from 20 to 65% by weight, or from 40 to 60% by weight, or at least 20% b/w, or at least 35% b/w, or at least 50% b/w, or employed neat (i.e., with no water). The mixture is impregnated into a dried catalyst by incipient wetness, i.e., just enough is added to nearly fill the pores, but not enough to cause observable wetness as is known to those skilled in the art. The resulting impregnated material is dried in air, preferably for at least 30 minutes at greater than 100 C, heated to calcination temperature at a heating rate of at least 1 C/minute, or at least 2 C/min, or at least 5 C/min, or at least 10 C/min, or more rapidly, and calcined in air at a temperature sufficient to decompose the organic functionality of the polysiloxane, preferable at least 500 C, or at least 550 C, or at least 600 C, for 1 hour, or 2 hours, or 3 hours, or 5 hours, or longer.
The silicone or other coating compound can be largely deposited on the surface of the catalyst and will be found to be enriched on the surface of the catalyst relative to the whole catalyst. Detection of the coating material can be performed by Low-Energy Ion Scattering spectroscopy (LEIS), sometimes referred to simply as ion scattering spectroscopy (ISS), which is a surface-sensitive analytical technique used to characterize the chemical and structural makeup of materials. In some cases the surface coating may be detected by X-ray Photoelectron Spectroscopy (XPS, sometimes called Electron Spectroscopy for Chemical Analysis, ESCA) or by Scanning Electron Microscopy/Energy Dispersive X-ray spectroscopy (SEM/EDX), or Transmission Electron Microscopy (TEM). Another useful technique is Secondary Ion Mass Spectroscopy (SIMS) in which a sample is bombarded with an ionic species (eg Ar+) and the resulting ejected ions are detected. With SIMS a depth profile of the various components can be obtained. For a silicone coating LEIS, SEM/EDX, TEM, or SIMS can be used to show that the silica to alumina ratio (SAR) of the surface atomic layers is higher than the SAR of the bulk zeolite particles. For example, if the SAR (SiO2/Al2O3) ratio of the bulk zeolite is 30:1, the SAR of the surface will be at least 40:1 or at least 50:1 or at least 100:1.
For catalysts that are both coated with a coating compound and promoted by addition of metal or metal oxide promoters, either the catalyst can be coated first and then promoted or the catalyst can be promoted first and then coated with the coating compound. Catalysts that are coated first with the coating compound and promoted with a metal or metal oxide promoter in a later step are preferred.
The zeolite catalyst may be treated with the coating compound to reduce the size of the pore openings, block pore openings, and cover or obscure catalytic sites on the external surface of the catalyst. This treatment process may also be used to cover or obscure catalytic sites in the pores near the pore mouths openings. The covered or obscured catalytic sites may be referred to as deactivated catalytic sites. The coating compound may have a molecular size that is incapable of entering the pores of the catalyst or only partially entering the pores. During the catalyst treatment process, the coating compound may be applied to the catalyst and subsequently calcined. This process may be repeated until the desired level of treatment is provided. The fraction of catalytic sites on the external surface of the catalyst that may be deactivated by treatment with the coating compound may be at least about 15%, or at least about 25%, or at least about 35%, or at least about 45%, or at least about 55%, or at least about 65%, or at least about 75%, or at least about 85%, or at least about 90%, or at least about 95%, or at least about 98%, or at least about 99%, of the available catalytic sites.
The coating compound may be a silicon-containing compound having a number average molecular weight in the range from about 80 to about 20,000, or from about 150 to 10,000.
The coating compound may comprise oxides of Ti, V, Zr, Cr, Mo, or Mn or mixtures of these. The coating may comprise carbon or coke from prior exposure of the catalyst to CFP or other coke forming conditions.
The pores with pore mouth openings that have been reduced in size or number may allow for an increase in para selectivity for the xylenes. This may be due to the fact that the reduced pore mouth openings may allow p-xylene to diffuse out of the pores while the diffusion of m-xylene and o-xylene may be restricted.
Reduction of the number of pore mouth openings may have a similar effect. Due to the more rapid diffusion of p-xylene within the zeolite when compared to isomers m- and o-xylene, closing pore mouths increases the diffusion length of molecules entering and leaving the zeolite; ie a linear pore that is closed on one end has twice the diffusion path length compared to a pore that is open on both ends. Increased pore length permits the faster diffusing p-xylene to diffuse out of the pore more quickly than the o- and m-xylenes.
Promoters can be added to the zeolite that impact the pore and channel sizes as well. Promoter elements situated in or near the pore mouth opening are capable of reducing the effective diameter of the pore. Pores with reduced effective diameters have slower rates of diffusion. For the xylene isomers, the rates of diffusion of the o- and m-xylene isomers are more reduced than for the p-xylene isomer. Again the faster relative rate of diffusion of the p-xylene isomer permits it to diffuse out of the pores more rapidly than the m- and o-xylene isomers, thus increasing the selectivity of p-xylene among the xylenes in the product. Promoters on the zeolite external surface can likewise inhibit the activity of the external sites that are non-shape-selective, hence increasing p-xylene selectivity.
The selectivity to p-xylene in the xylenes may be up to 100%, or, p-xylene may be produced in preference to o-xylene and/or m-xylene, but some o-xylene and/or m-xylene may nevertheless be produced. The fluid hydrocarbon product produced using the foregoing method may comprise xylenes and may be characterized by a p-xylene selectivity in the xylenes of at least about 50%, or at least about 55%, or at least about 60%, or at least about 65%, or at least about 70%, or at least about 75%, or at least about 80%, or at least about 85%, or at least about 90%, or at least about 95%.
The method may further comprise the step of recovering the fluid hydrocarbon product. The fluid hydrocarbon product may further comprise, in addition to p-xylene, other aromatic compounds and/or olefin compounds. The fluid hydrocarbon product may further comprise benzene, toluene, ethylbenzene, styrene, methylethylbenzene, trimethylbenzene, o-xylene, m-xylene, indanes, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene, dimethylhydrindene, or a mixture of two or more thereof.
The carbon yield of aromatics in the fluid hydrocarbon product may be at least about 13%, or at least about 17%, or at least about 20%. The carbon yield of olefins in the fluid hydrocarbon product may be at least about 7%, or at least about 9%, or at least about 11%. The mass yield of p-xylene may be at least about 1.5 wt %, or at least about 2 wt %, or at least about 2.5 wt %, or at least about 3 wt %.
The inventive method may comprise a single-stage method for the pyrolysis of the hydrocarbonaceous material. This method may comprise providing or using a single-stage pyrolysis apparatus. A single-stage pyrolysis apparatus may be one in which pyrolysis and subsequent catalytic reactions are carried out in a single vessel. The single-stage pyrolysis apparatus may comprise a continuously stirred tank reactor, a batch reactor, a semi-batch reactor, a fixed bed reactor or a fluidized bed reactor. Multi-stage apparatuses may also be used for the production of fluid hydrocarbon products in accordance with the invention.
The catalysts provided for herein may be particularly suited for producing xylenes with a relatively high selectivity to p-xylene in the xylenes of at least about 40%, or at least about 45%, or at least about 50%, or at least about 55%, or at least about 60%, or at least about 65%, or at least about 70%, or at least about 75%, or at least about 80%, or at least about 85%, or at least about 90%.
Referring to FIG. 1, feed stream 10 includes a solid hydrocarbonaceous material that can be fed to reactor 20. Moisture 12 may optionally be removed from the solid hydrocarbonaceous feed composition prior to being fed to the reactor, e.g., by an optional dryer 14. Removal of moisture from the solid hydrocarbonaceous material feed stream may be advantageous for several reasons. For example, the moisture in the feed stream may require additional energy input in order to heat the solid hydrocarbonaceous material to a temperature sufficiently high to achieve pyrolysis. Variations in the moisture content of the solid hydrocarbonaceous feed may lead to difficulties in controlling the temperature of the reactor. In addition, removal of moisture from the solid hydrocarbonaceous feed can reduce or eliminate the need to process the water during later processing steps.
The particle size of the solid hydrocarbonaceous feed composition may be reduced in an optional grinding system 16 prior to passing the solid hydrocarbonaceous feed to the reactor. The hydrocarbonaceous material may be transferred to reactor 20. The reactor may be used, in some instances, to perform catalytic pyrolysis of at least a portion of the first reactant comprising the hydrocarbonaceous material under reaction conditions sufficient to produce one or more pyrolysis products. In the illustrative embodiment of FIG. 1, the reactor comprises any suitable reactor known to those skilled in the art. For example, in some instances, the reactor may comprise a continuously stirred tank reactor (CSTR), a batch reactor, a semi-batch reactor, or a fixed bed catalytic reactor, among others. Preferably, the reactor comprises a fluidized bed reactor, e.g., a circulating fluidized bed reactor, bubbling bed reactor, or riser reactor. Fluidized bed reactors may, in some cases, provide improved mixing of the catalyst, solid biomass during pyrolysis and/or subsequent reactions, which may lead to enhanced control over the reaction products formed. The use of fluidized bed reactors may also lead to improved heat transfer within the reactor. In addition, improved mixing in a fluidized bed reactor may lead to a reduction of the amount of coke adhered to the catalyst, resulting in reduced deactivation of the catalyst in some cases.
Higher yields of desired product formation, lower yields of coke formation, and/or more controlled product formation (e.g., higher production of p-xylene relative to other products) may be achieved when particular combinations of reaction conditions and system components are implemented in methods and systems described herein. For example, conditions such as the mass normalized space velocity(ies) (e.g., of the solid hydrocarbonaceous material and/or the fluidization fluid), the temperature of the reactor and/or solids separator, the reactor pressure, the heating rate of the feed stream(s), the catalyst to solid hydrocarbonaceous material mass ratio, the residence time of the hydrocarbonaceous material in the reactor, the residence time of the reaction products in the solids separator, and/or the catalyst type (as well as silica to alumina molar ratio and pore mouth opening size) may be controlled to achieve beneficial results.
The reactor(s) may be operated at any suitable temperature. In some instances, it may be desirable to operate the reactor(s) at intermediate temperatures, compared to temperatures typically used in many previous catalytic pyrolysis systems. For example, the reactor may be operated at temperatures of between about 400° C. and about 650° C., between about 425° C. and about 600° C., or between about 525° C. and about 575° C. Operating the reactor(s) at these intermediate temperatures may allow one to maximize the amount of desirable products. The invention may not be limited to the use of such intermediate temperatures, however, and in other embodiments, lower and/or higher temperatures can be used.
The reactor(s) 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, or at least about 300 kPa, or at least about 400 kPa. The reactor may be operated at a pressure below about 600 kPa, or below about 400 kPa, or below about 200 kPa. The reactor may be operated at a pressure in the range from about 100 to about 600 kPa, or in the range from about 100 to about 400 kPa, or in the range from about 100 to about 200 kPa. The invention may not be limited to the use of such pressures, however, and in other embodiments, lower and/or higher pressures may be employed.
The mass-normalized space velocity of the hydrocarbonaceous material may be selected to selectively produce a desired array of fluid hydrocarbon products. As used herein, the term “mass-normalized space velocity” of a component is defined as the mass flow rate of the component into the reactor (e.g., as measured in g/hr) divided by the mass of catalyst in the reactor (e.g., as measured in g) and has units of inverse time. For example, the mass-normalized space velocity of solid hydrocarbonaceous material fed to the reactor may be calculated as the mass flow rate of the solid hydrocarbonaceous material into the reactor divided by the mass of catalyst in the reactor. The mass-normalized space velocity of a component (e.g., the hydrocarbonaceous material) in the reactor may be calculated using different methods depending upon the type of reactor being used. For example, in systems employing batch or semi-batch reactors, wherein the solid hydrocarbonaceous material is not fed continuously to the reactor, the solid hydrocarbonaceous material does not have a mass-normalized space velocity. For systems in which catalyst is fed to and/or extracted from the reactor during reaction (e.g., circulating fluidized bed reactors), the mass-normalized space velocity may be determined by calculating the average amount of catalyst within the volume of the reactor over a period of operation (e.g., steady-state operation).
The mass-normalized space velocity of the hydrocarbonaceous material fed to the reactor may be at a mass normalized space velocity of up to about 3 hour-1, or up to about 2 hour-1, or up to about 1.5 hour-1, or up to about 0.9 hour-1, or in the range from about 0.01 hour-1 to about 3 hour-1, or in the range from about 0.01 to about 2 hour-1, or in the range from about 0.01 to about 1.5 hour-1, or in the range from about 0.01 to about 0.9 hour-1, or in the range from about 0.01 hour-1 to about 0.5 hour-1, or in the range from about 0.1 hour-1 to about 0.9 hour-1, or in the range from about 0.1 hour-1 to about 0.5 hour-1. The invention may not be limited to the use of such mass-normalized space velocities, however, and in other embodiments, lower and/or higher mass-normalized space velocities can be used.
The residence time of a reactant (e.g., the hydrocarbonaceous material) in the reactor (i.e., the reactant residence time) may be 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 a reactant (e.g., the hydrocarbonaceous material) in the reactor may be less than about 5 minutes, or from about 1 second and about 4 minutes, or from about 2 seconds to about 4 minutes, or from about 5 seconds to about 4 minutes, or from about 7 seconds to about 4 minutes, or from about 10 seconds to about 4 minutes, or from about 12 seconds to about 4 minutes, or from about 15 seconds to about 4 minutes, or from about 20 seconds to about 4 minutes, or from about 30 seconds to about 4 minutes, or from about 60 seconds to about 4 minutes. Previous “fast pyrolysis” studies have, in many cases, employed systems with very short reactant residence times (e.g., less than 2 seconds). In some cases, however, the use of relatively longer residence times may allow for additional chemical reactions to form desirable products. Long residence times may be achieved by, for example, increasing the volume of the reactor and/or reducing the volumetric flow rate of the hydrocarbonaceous materials. It should be understood, however, that in some embodiments described herein, the residence time of the reactant (e.g., hydrocarbonaceous material) may be relatively shorter, e.g., less than about 2 seconds, or less than about 1 second.
The contact time of the pyrolysis product (e.g., pyrolysis vapor) with the catalyst in the reactor may be 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. The contact time may be less than about 5 minutes, or from about 1 second and about 4 minutes, or from about 2 seconds to about 4 minutes, or from about 5 seconds to about 4 minutes, or from about 7 seconds to about 4 minutes, or from about 10 seconds to about 4 minutes, or from about 12 seconds to about 4 minutes, or from about 15 seconds to about 4 minutes, or from about 20 seconds to about 4 minutes, or from about 30 seconds to about 4 minutes, or from about 60 seconds to about 4 minutes.
In certain cases where fluidized bed reactors are used, the feed material (e.g., a solid hydrocarbonaceous material) in the reactor may be fluidized by flowing a fluid stream through the reactor. In the exemplary embodiment of FIG. 1, a fluid stream 44 is used to fluidize the feed material in reactor 20. Fluid may be supplied to the fluid stream from a fluid source 24 and/or from the product streams of the reactor via a compressor 26. As used herein, the term “fluid” means a material generally in a liquid, supercritical, or gaseous state. Fluids, however, may also contain solids such as, for example, suspended or colloidal particles. In some embodiments, it may be advantageous to control the residence time of the fluidization fluid in the reactor. The residence time of the fluidization fluid may be defined as the volume of the reactor divided by the volumetric flow rate of the fluidization fluid. The residence time of the fluidization fluid may be at least about 0.1 second, at least about 0.2 second, at least about 0.5 second, at least about 1 second, at least about 2 seconds, at least about 3 seconds, at least about 4 seconds, at least 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 residence time of the fluidization fluid may be from about 0.1 second to about 48 seconds, from about 0.2 second to about 48 seconds, from about 0.5 second to about 480 seconds, from about 1 second to about 48 seconds, from about 3 seconds to about 48 seconds, from about 5 seconds to about 48 seconds, from about 6 seconds to about 48 seconds, from about 8 seconds to about 48 seconds, from about 10 seconds to about 48 seconds, from about 12 seconds to about 48 seconds, or from about 24 seconds to about 48 seconds.
Suitable fluidization fluids that may be used in this invention include, for example, inert gases (e.g., helium, argon, neon, etc.), hydrogen, nitrogen, steam, carbon monoxide, and carbon dioxide, among others.
As shown in the illustrative embodiment of FIG. 1, the products (e.g., fluid hydrocarbon products) formed during the reaction of the reactants (e.g., the solid hydrocarbonaceous material) exit the reactor via a product stream 30. In addition to the reaction products, the product stream may, in some cases, comprise unreacted reactant(s), fluidization fluid, char, ash, and/or catalyst. In one set of embodiments, the desired reaction product(s) (e.g., liquid aromatic hydrocarbons, olefin hydrocarbons, gaseous products, etc.) may be recovered from an effluent stream of the reactor.
As shown in the illustrative embodiment of FIG. 1, product stream 30 may be fed to an optional solids separator 32. The solids separator may be used, in some cases, to separate the reaction products from catalyst (e.g., at least partially deactivated catalyst) present in the product stream. In addition, the solids separator may be used, in some instances, to remove coke and/or ash from the catalyst. In some embodiments, the solids separator may comprise optional purge stream 33, which may be used to purge coke, ash, and/or catalyst from the solids separator.
The solids separator may not be required in all embodiments. For example, for situations in which catalytic fixed bed reactors are employed, the catalyst may be retained within the reactor, and the reaction products may exit the reactor substantially free of catalyst, thus negating 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, not shown in FIG. 1. The catalyst exiting the separator may be at least partially deactivated. The separated catalyst may be fed to a regenerator 36 in which any catalyst that was at least partially deactivated may be re-activated. The regenerator may comprise an optional purge stream 37, which may be used to purge coke, ash, and/or catalyst from the regenerator. Methods for activating catalyst are well-known to those skilled in the art, for example, as described in Kirk-Othmer Encyclopedia of Chemical Technology (Online), Vol. 5, Hoboken, N.J.: Wiley-Interscience, c2001-, pages 255-322, which are incorporated herein by reference.
A portion of the catalyst may be removed from the reactor through a catalyst exit port (not shown in FIG. 1.). The catalyst removed from the reactor may be partially deactivated and passed via a conduit into regenerator 36, or into a separate regenerator (not shown in FIG. 1). Removed catalyst that has been regenerated may be returned to the reactor via stream 47, or may be returned to the reactor separately from the fluidization gas via a separate stream (not shown in FIG. 1.).
An oxidizing agent may be fed to the regenerator via a stream 38, e.g., as shown in FIG. 1. The oxidizing agent may originate from any source including, for example, a tank of oxygen, atmospheric air, steam, among others. In the regenerator, the catalyst may be re-activated by reacting the catalyst with the oxidizing agent. The deactivated catalyst may comprise residual carbon and/or coke, which may be removed via reaction with the oxidizing agent in the regenerator. The regenerator in FIG. 1 comprises a vent stream 40 which may include regeneration reaction products, residual oxidizing agent, etc.
The regenerator may be of any suitable size mentioned above in connection with the reactor or the solids separator. In addition, the regenerator may be operated at elevated temperatures in some cases (e.g., at least about 300° C., 400° C., 500° C., 600° C., 700° C., 800° C., or higher). The residence time of the catalyst in the regenerator may also be controlled using methods known by those skilled in the art, including those outlined above. The mass flow rate of the catalyst through the regenerator may be coupled to the flow rate(s) in the reactor and/or solids separator in order to preserve the mass balance in the system.
The regenerated catalyst may exit the regenerator via stream 42. The regenerated catalyst may be recycled back to the reactor via recycle stream 47. In some cases, catalyst may be lost from the system or removed intentionally during operation. Additional “makeup” catalyst may be added to the system via a makeup stream 46. The regenerated and makeup catalyst may be fed to the reactor with the fluidization fluid via recycle stream 47. Alternatively, the catalyst and fluidization fluid may be fed to the reactor via separate streams.
Referring to solids separator 32 in FIG. 1, the reaction products (e.g., fluid hydrocarbon products) may exit the solids separator via stream 48. In some cases, a fraction of stream 48 may be purged via purge stream 60. The contents of the purge stream may be fed to a combustor or a water-gas shift reactor, for example, to recuperate energy that would otherwise be lost from the system. In some cases, the reaction products in stream 48 may be fed to an optional condenser 50. The condenser may comprise a heat exchanger which condenses at least a portion of the reaction product from a gaseous to a liquid state. The condenser may be used to separate the reaction products into gaseous, liquid, and solid fractions. The operation of condensers is well known to those skilled in the art. The condenser may also make use of pressure change to condense portions of the product stream. In FIG. 1, stream 54 may comprise the liquid fraction of the reaction products (e.g., water, aromatic compounds, olefin compounds, etc.), and stream 74 may comprise the gaseous fraction of the reaction products (e.g., CO, CO₂, H₂, etc.). In some embodiments, the gaseous fraction may be fed to a vapor recovery system 70. The vapor recovery system may be used, for example, to recover any desirable vapors within stream 74 and transport them via stream 72. In addition, stream 76 may be used to transport CO, CO₂, and/or other gases from the vapor recovery system. The optional vapor recovery system may be placed in other locations. For example, in some embodiments, a vapor recovery system may be positioned downstream of purge stream 54. One skilled in the art can select an appropriate placement for a vapor recovery system.
Other products (e.g., excess gas) may be transported to optional compressor 26 via stream 56, where they may be compressed and used as fluidization gas in the reactor (stream 22) and/or where they may assist in transporting the hydrocarbonaceous material to the reactor (streams 58) or may be used to transport catalyst to the reactor (not shown), or may be used to transport additional non-solid feeds to the reactor. In some instances, the liquid fraction may be further processed, for example, to separate the water phase from the organic phase, to separate individual compounds, etc.
It should be understood that, while the set of embodiments described by FIG. 1 includes a reactor, solids separator, regenerator, condenser, etc., not all embodiments will involve the use of these elements. For example, in some embodiments, the feed stream(s) may be fed to a catalytic fixed bed reactor, reacted, and the reaction products may be collected directly from the reactor and cooled without the use of a dedicated condenser. In some instances, while a dryer, grinding system, solids separator, regenerator, condenser, and/or compressor may be used as part of the process, one or more of these elements may comprise separate units not fluidically and/or integrally connected to the reactor. In other embodiments one or more of the dryer, grinding system, solids separator, regenerator, condenser, and/or compressor may be absent. In some embodiments, the desired reaction product(s) may be recovered at any point in the production process (e.g., after passage through the reactor, after separation, after condensation, etc.).
Catalyst components useful in the context of this invention can be selected from any catalyst known in the art, or as would be understood by those skilled in the art made aware of this invention. Functionally, catalysts may be limited only by the capability of any such material to promote and/or effect dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization, decarbonylation, decarboxylation, aldol condensation and/or any other reaction or process associated with or related to the pyrolysis of a hydrocarbonaceous material. Catalyst components can be considered acidic, neutral or basic, as would be understood by those skilled in the art.
A screening method may be used to select catalysts with appropriate pore sizes for the conversion of specific pyrolysis product molecules. The screening method may comprise determining the size of pyrolysis product molecules desired to be catalytically reacted (e.g., the molecule kinetic diameters of the pyrolysis product molecules). One of ordinary skill in the art can calculate, for example, the kinetic diameter of a given molecule. The type of catalyst may then be chosen such that the pores of the catalyst (e.g., Norman adjusted minimum radii) are sufficiently large to allow the pyrolysis product molecules to diffuse into and/or react with the catalyst. In some embodiments, the catalysts are chosen such that their pore sizes are sufficiently small to prevent entry and/or reaction of pyrolysis products whose reaction would be undesirable.
The catalyst may be selected from naturally-occurring zeolites, synthetic zeolites and combinations thereof. The catalyst may be a Mordenite Framework Inverted (MFI) type zeolite catalyst, such as a ZSM-5 zeolite catalyst. Catalysts comprising ZSM-5 that may be used with or without modification are available commercially. The catalysts that are provided for herein may comprise acid or catalytically active sites. While not wishing to be bound by theory, it is believed that various acid sites in ZSM-5 and other zeolites are catalytically active for reactions of the hydrocarbonaceous materials including dehydration, decarbonylation, decarboxylation, isomerization, oligomerization and/or dehydrogenation, hence the terms “acid sites” and “catalytically active sites” may be used interchangeably. Other types of useful zeolite catalysts may include ferrierite, zeolite Y, zeolite beta, modernite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)AlPO-31, ZSM-11, SSZ-23, mixtures of two or more thereof, and the like.
The zeolite catalyst may contain binders and fillers in addition to the zeolite. Typical binders or fillers include amorphous materials such as alumina, silica, silica-alumina, titania, aluminum phosphate, kaolin, attapulgite clay and various types of clays. Preferred zeolite catalysts comprise at least 25%, or at least 30%, or at least 35%, or at least 40%, or at least 45%, or at least 50%, or at least 60% by weight zeolite in the catalyst particle.
The catalyst may comprise, in addition to alumina and silica, one or more additional metals and/or a metal oxides. Suitable metals and/or oxides may include, for example, nickel, palladium, silver, platinum, palladium, titanium, vanadium, chromium, manganese, iron, cobalt, zinc, copper, gallium, sodium, potassium, magnesium, calcium, zirconium, lanthanum, cerium, phosphorus, an oxide of one or more thereof, or a mixture of two or more thereof, among others. The metal and/or metal oxide can be impregnated into the catalyst (e.g., in the interstices of the lattice structure of the catalyst), in some embodiments. 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, such as, but not limited to, impregnation, ion exchange, vapor deposition, and the like. The zeolite may comprise small amounts of structure stabilizing elements such as phosphorus, lanthanum, rare earths, and the like, typically at levels that are less than about 4% by weight of the zeolite. The catalyst may be conditioned before operation in 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, and the like. The metal and/or metal oxide may be incorporated into the lattice structure of the catalyst. For example, the metal and/or metal oxide may be included during the preparation of the catalyst, and the metal and/or metal oxide may occupy a lattice site of the resulting catalyst (e.g., a zeolite catalyst). As another example, the metal and/or metal oxide may react or otherwise interact with a zeolite such that the metal and/or metal oxide displaces an atom within the lattice structure of the zeolite.
In addition, in some cases, properties of the catalysts (e.g., pore structure, type and/or number of acid sites, etc.) may be chosen to selectively produce a desired product. It may be desirable, in some embodiments, to employ one or more catalysts to establish a bimodal distribution of pore sizes. In some cases, a single catalyst with a bimodal distribution of pore sizes may be used (e.g., a single catalyst that contains predominantly 5.9-6.3 Å pores and 7-8 Å pores). In other cases, a mixture of two or more catalysts may be employed to establish the bimodal distribution (e.g., a mixture of two catalysts, each catalyst type including a distinct range of average pore sizes). In some embodiments, one of the one or more catalysts comprises a zeolite catalyst and another of the one or more catalysts comprises a non-zeolite catalyst (e.g., a mesoporous catalyst, a metal oxide catalyst, etc.).
For example, 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 the one or more catalysts (e.g., a zeolite catalyst, a mesoporous catalyst, etc.) have smallest cross-sectional diameters that lie within a first size distribution or a 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 smallest cross-sectional diameters that lie within the first size distribution; and at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have smallest cross-sectional diameters that lie within the second size distribution. In some cases, the first and second size distributions are selected from the ranges provided above. In certain embodiments, the first and second size distributions are different from each other and do not overlap. An example of a non-overlapping range is 5.9-6.3 Å and 6.9-8.0 Å, and an example of an overlapping range is 5.9-6.3 Å and 6.1-6.5 Å. The first and second size distributions may be selected such that the ranges are not immediately adjacent one another, an example being pore sizes of 5.9-6.3 Å and 6.9-8.0 Å. An example of a range that is immediately adjacent one another is pore sizes of 5.9-6.3 Å and 6.3-6.7 Å. As a specific example, in some embodiments one or more catalysts are used to provide a bimodal pore size distribution for the simultaneous production of aromatic and olefin compounds. That is, one pore size distribution may advantageously produce a relatively high amount of aromatic compounds, particularly p-xylene, and the other pore size distribution may advantageously produce a relatively high amount of olefin compounds. 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 the one or more catalysts have smallest cross-sectional diameters between about 5.9 Å and about 6.3 Å or between about 7 Å and about 8 Å. In addition, at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have smallest cross-sectional diameters between about 5.9 Å and about 6.3 Å; and at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have smallest cross-sectional diameters between about 7 Å and about 8 Å.
In some embodiments, it 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 the one or more catalysts have smallest cross-sectional diameters between about 5.9 Å and about 6.3 Å or between about 7 Å and about 200 Å. In addition, at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have smallest cross-sectional diameters between about 5.9 Å and about 6.3 Å; and at least about 2%, at least about 5%, or at least about 10% of the pores of the one or more catalysts have smallest cross-sectional diameters between about 7 Å and about 200 Å. 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 the one or more catalysts have smallest cross-sectional diameters that lie within a first distribution and a second distribution, wherein the first distribution is between about 5.9 Å and about 6.3 Å and the second distribution is different from and does not overlap with the first distribution. In some embodiments, the second pore size distribution may be between about 7 Å and about 200 Å, between about 7 Å and about 100 Å, between about 7 Å and about 50 Å, or between about 100 Å and about 200 Å. In some embodiments, the second catalyst may be mesoporous (e.g., have a pore size distribution of between about 2 nm and about 50 nm).
In some embodiments, a bimodal distribution of pore sizes may be beneficial in reacting two or more hydrocarbonaceous feed material components. For example, some embodiments comprise providing a solid hydrocarbonaceous material comprising a first component and a second component in a reactor, wherein the first and second components are different. Examples of compounds that may be used as first or second components include any of the hydrocarbonaceous materials. For example, the first component may comprise one of cellulose, hemi-cellulose and lignin, and the second component comprises one of cellulose, hemicellulose and lignin. The method may further comprise providing first and second catalysts in the reactor. In some embodiments, the first catalyst may have a first pore size distribution and the second catalyst may have a second pore size distribution, wherein the first and second pore size distributions are different and do not overlap. The first pore size distribution may be, for example, between about 5.9 Å and about 6.3 Å. The second pore size distribution may be, for example, between about 7 Å and about 200 Å, between about 7 Å and about 100 Å, between about 7 Å and about 50 Å, or between about 100 Å and about 200 Å. In some cases, the second catalyst may be mesoporous or non-porous. The first catalyst may be selective for catalytically reacting the first component or a derivative thereof to produce a fluid hydrocarbon product. In addition, the second catalyst may be selective for catalytically reacting the second component or a derivative thereof to produce a fluid hydrocarbon product. The method may further comprise pyrolyzing within the reactor at least a portion of the hydrocarbonaceous material under reaction conditions sufficient to produce one or more pyrolysis products and catalytically reacting at least a portion of the pyrolysis products with the first and second catalysts to produce the one or more hydrocarbon products. In some instances, at least partially deactivated catalyst may also be used.
In certain embodiments, a method used in combination with embodiments described herein includes increasing the catalyst to hydrocarbonaceous material mass ratio of a composition to increase production of identifiable aromatic compounds. As illustrated herein, representing but one distinction over certain prior catalytic pyrolysis methods, articles and methods described herein can be used to produce discrete, identifiable aromatic, biofuel compounds selected from but not limited to benzene, toluene, propylbenzene, ethylbenzene, styrene, methylbenzene, methylethylbenzene, trimethylbenzene, xylenes, indanes, naphthalene, methylnaphthelene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene, and dimethylhydrindene and combinations thereof.
In some embodiments, the reaction chemistry of a catalyst may be affected by adding one or more additional compounds. For example, the addition of a metal to a catalyst may result in a shift in selective formation of specific compounds (e.g., addition of metal to alumina-silicate catalysts may result in the production of more CO or CO₂). In addition, when the fluidization fluid comprises hydrogen, the amount of coke formed on the catalyst may be decreased.
The catalyst may comprise both silica and alumina. The silica (SiO₂) and alumina (Al₂O₃) in the catalyst may be present in any suitable molar ratio. For example, in some cases, the catalyst in the feed may comprise a silica (SiO₂) to alumina (Al₂O₃) molar ratio (SAR) of between about 10:1 to about 240:1, or in the range from about 10:1 to about 40:1, or in the range from about 20:1 to about 50:1, or about 30:1, or greater than 30:1.
In some embodiments, the methods described herein may be configured to selectively produce aromatic compounds (e.g., p-xylene) in a single-stage, or alternatively, a multi-stage pyrolysis apparatus. For example, in some embodiments, the mass yield of the aromatic compounds in the fluid hydrocarbon product may be at least about 13 wt %, at least about 17 wt %, at least about 20 wt %, at least about 30 wt %, between about 13 wt % and about 40 wt %, between about 13 wt % and about 35 wt %, between about 17 wt % and about 40 wt %, between about 17 wt % and about 35 wt %, between about 20 wt % and about 40 wt %, between about 20 wt % and about 35 wt %, between about 30 wt % and about 40 wt %, or between about 30 wt % and about 35 wt %. The mass yield of p-xylene may be at least about 1.5% by weight, or at least about 2% by weight, or at least about 2.5% by weight, or at least about 3% by weight. The “mass yield” of aromatic compounds or p-xylene in a given product is calculated as the total weight of the aromatic compounds or p-xylene present in the fluid hydrocarbon product divided by the weight of the solid hydrocarbonaceous material used in forming the reaction product, multiplied by 100%.
In some embodiments, aromatic compounds (especially p-xylene) may be selectively produced when the mass-normalized space velocity of the solid hydrocarbonaceous material fed to the reactor is up to about 3 hour-1, or up to about 2 hour-1, or up to about 1.5 hour-1, or up to about 0.9 hour-1, or in the range from about 0.01 hour-1 to about 3 hour-1, or in the range from about 0.01 to about 2 hour-1, or in the range from about 0.01 to about 1.5 hour-1, or in the range from about 0.01 to about 0.9 hour-1, or in the range from about 0.01 hour-1 to about 0.5 hour-1, or in the range from about 0.1 hour-1 to about 0.9 hour-1, or in the range from about 0.1 hour-1 to about 0.5 hour-1. In some instances, aromatic compounds (especially p-xylene) may be selectively produced when the reactor is operated at a temperature of between about 400° C. and about 650° C. (or between about 425° C. and about 600° C., or between about 500° C. and about 575° C.). In addition, certain heating rates (e.g., at least about 50° C./s, or at least about 400° C./s), high catalyst-to-feed mass ratios (e.g., at least about 5:1), and/or high silica to alumina molar ratios in the catalyst (e.g., at least about 30:1) may be used to facilitate selective production of aromatic compounds (especially p-xylene). Some such and other process conditions may be combined with a particular reactor type, such as a fluidized bed reactor (e.g., a bubbling fluidized bed, a turbulent fluidized bed, a fast fluidized bed, circulating fluidized bed reactor), to selectively produce aromatic and/or olefin compounds.
Furthermore, in some embodiments, the catalyst may be chosen to facilitate selective production of aromatic products (especially p-xylene). For example, ZSM-5 may, in some cases, preferentially produce relatively higher amounts of aromatic compounds. In some cases, catalysts that include Brønsted acid sites may facilitate selective production of aromatic compounds. In addition, catalysts with well-ordered pore structures may facilitate selective production of aromatic compounds. For example, in some embodiments, catalysts with average pore diameters between about 5.9 Å and about 6.3 Å may be particularly useful in producing aromatic compounds. In addition, catalysts with average pore diameters between about 7 Å and about 8 Å may be useful in producing olefins. In some embodiments, a combination of one or more of the above process parameters may be employed to facilitate selective production of aromatic and/or olefin compounds. The ratio of aromatics to olefins produced on a carbon basis may be, for example, between about 0.1:1 and about 10:1, between about 0.2:1 and about 5:1, between about 0.5:1 and about 4:1, between about 0.1:1 and about 0.5:1, between about 0.5:1 and about 1:1, between about 1:1 and about 5:1, or between about 5:1 and about 10:1.
Furthermore, processes described herein may result in lower coke formation than certain existing methods. For example, in some embodiments, a pyrolysis product can be formed with less than about 30 wt %, less than about 25 wt %, less than about 20 wt %, than about 15 wt %, or less than about 10 wt % of the pyrolysis product being coke. The amount of coke formed is measured as the weight of coke formed in the system divided by the weight of hydrocarbonaceous material used in forming the pyrolysis product.
The catalyst of the present invention may comprise at least 0.2 wt % Fe, or at least 0.3% Fe, or at least 0.5% Fe, or at least 1.0% Fe, or at least 1.5% Fe, or at least 1.7% Fe, or at least 2.0% Fe, or from 0.2% to 10% Fe, or from 0.3% to 2.0% Fe, or from 0.5% to 1.7% Fe. The phrase “at least 0.2 wt % Fe” is determined by elemental analysis of catalyst separated from biomass and separated from any ash, to the extent practicable, where the elemental analysis is preferably conducted by ICP. The term “pretreated” means treated prior to use in a catalytic pyrolysis process.
The catalyst of the present invention may comprise an aluminosilicate zeolite wherein the catalyst has been treated with sufficient iron such that the Fe:acid site ratio, defined as the number of moles of Fe in the catalyst divided by the number of moles of acid sites in the catalyst that has not been treated with Fe, is at least 0.1, or at least 0.5, or at least 1, or at least 5, or at least 6, or less than 20, or less than 15, or less than 10, or less than 7, or from 0.1 to 20, or from 0.5 to 10 wherein the Fe is determined by elemental analysis and the acid site concentration is determined by desorption of isopropylamine, where the elemental analysis for Fe is preferably conducted by ICP.
A siliceous coating can be applied to a zeolite surface by reaction with silicones or siloxanes as described elsewhere herein and in the literature. A siliceous coating can be identified can be identified by a higher (at least 10% higher or at least 30% higher or at least 50% higher or at least 100% higher) in the exterior 50 A (or exterior 100 A) as compared to the Si/Al ratio at greater depths in the catalyst. The preferred technique for analyzing the Si/Al ratio is SEM/XPS before and after sputtering off 50 or 100 A, or by cross-sectional analysis by SEM/XPS.
The following non-limiting examples are intended to illustrate various aspects and features of the invention.

EXAMPLES

Catalyst A

A sample of a spray-dried catalyst containing 50% NH₄-ZSM-5 with a silica binder was obtained commercially (Zeolite A) and calcined at 550 C for 2 hours in air to form Catalyst A.

Catalyst B

A 200 g portion of the calcined Catalyst A was impregnated with 105 g of aqueous solution of 14.3 wt % Fe(NO₃)₃in deionized water to achieve a loading of 1.7% Fe by weight. The impregnated sample was dried at 120 C for two hours in air and then oven calcined five hours at 600 C in air with a temperature ramp rate of 10 C/min. The catalyst thus obtained was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst C

A second batch of Fe-impregnated catalyst was prepared by incipient wetness using the same procedure as Catalyst B. The catalyst thus obtained was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst D

A 200 g portion of calcined Catalyst A was mixed with 500 ml of dried hexane in a 1 liter round-bottom flask. The mixture was heated to reflux under N₂with stirring. While refluxing, 30 ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h through a syringe pump. Upon completion of adding TEOS, the mixture was refluxed for additional 30 minutes, and then was cooled to ambient temperature. The solid material was collected by vacuum filtration, dried one hour at 120 C and then calcined five hours at 600 C in air. The calcined material has a nominal composition of 4.0% SiO₂/ZSM-5 and was tested in a fluidized bed reactor for biomass conversion to aromatics and olefins. Another 200 g batch was prepared following the same procedure and was used for the preparation of Catalyst E.

Catalyst E

A 200 g portion of Catalyst D was impregnated with 105 g of aqueous solution of 14.3 wt % Fe(NO₃)₃in deionized water to achieve an iron loading of 1.7% Fe by weight. The impregnated sample was dried at 120 C for two hours and then oven calcined five hours at 600 C in air with a temperature ramp rate of 10 C/min. The catalyst thus obtained has a nominal composition of 1.7% Fe, 4.0% SiO₂/ZSM-5 and was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst F

A 200 g portion of calcined Catalyst A was mixed with 500 ml of dried hexane in a 1 liter round-bottom flask. The mixture was heated to reflux under N2 with stirring. While refluxing, 45 ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h through a syringe pump. Upon completion of adding TEOS, the mixture was refluxed for additional 30 minutes, and then was cooled to ambient temperature. The solid material was collected by vacuum filtration, dried one hour at 120 C and then calcined five hours at 600 C in air. The calcined material has a nominal composition of 5.7% SiO₂/ZSM-5 and was tested in fluidized bed reactor for biomass conversion to aromatics and olefins. Another 200 g batch was prepared following the same procedure and was used for the preparation of Catalyst G.

Catalyst G

A 200 g portion of Catalyst F was impregnated with 105 g of aqueous solution of 14.3 wt % Fe(NO₃)₃in deionized water to achieve an iron loading of 1.7% Fe by weight. The impregnated sample was dried at 120 C for two hours and then oven calcined five hours at 600 C in air with a temperature ramp rate of 10 C/min. The catalyst thus obtained has a nominal composition of 1.7% Fe, 5.7% SiO₂/ZSM-5 and was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst H

An aminofunctional oligomeric siloxane (Hydrosil 2627) was diluted with deionized water at 1:1 weight ratio. A 400 g portion of spray-dried NH₄ZSM-5 (50% ZSM-5) obtained commercially was impregnated with 194 g of the silane polymer/water solution. The impregnated sample was dried two hours at 120 C and was then calcined at 600 C for five hours in air with a temperature ramp rate of 10 C/min. The catalyst thus obtained has a nominal composition of 4% SiO₂/ZSM-5. A portion of the calcined material was tested in fluidized bed reactor for the conversion of biomass to aromatics and olefins and the other portion was used for preparing Catalyst I.

Catalyst I

A 200 g portion of Catalyst H was impregnated with 105 g of aqueous solution of 14.3 wt % Fe(NO3)3 in deionized water to achieve a loading of 1.7% Fe by weight. The impregnated sample was dried at 120 C for two hours and then oven calcined five hours at 600 C with a temperature ramp rate of 10 C/min. The catalyst thus obtained has a nominal composition of 1.7% Fe, 4% SiO₂/ZSM-5 and it was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst J

A sample of a second commercially obtained spray-dried zeolite containing H-ZSM-5 with clay, an alumina binder, and some residual carbon (Zeolite B) was calcined at 550 C for 2 hours in air to form Catalyst J.

Catalyst K

A 200 g portion of calcined Catalyst J was mixed with 500 ml of dried hexane in a 1 liter round-bottom flask. The mixture was heated to reflux under N₂with stirring. While refluxing, 30 ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h through a syringe pump. Upon completion of adding TEOS, the mixture was refluxed for an additional 60 minutes, and then was cooled to ambient temperature. The solid material was collected by vacuum filtration, dried one hour at 120 C and then calcined five hours at 600 C in air. The procedures were repeated to double the amount of SiO2 loadings. The resulting material was impregnated with 105 g of aqueous solution of 14.3 wt % Fe(NO3)3 in deionized water to achieve a loading of 1.7% Fe by weight. The impregnated sample was dried at 120 C for two hours and then oven calcined five hours at 600 C with a temperature ramp rate of 10 C/min. The catalyst thus obtained has a nominal composition of 1.7% Fe, 7.48% SiO₂/ZSM-5 and it was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst L

A 200 g portion of calcined Catalyst J was mixed with 500 ml of dried hexane in a 1 liter round-bottom flask. The mixture was heated to reflux under N2 with stirring. While refluxing, 45 ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h through a syringe pump. Upon completion of adding TEOS, the mixture was refluxed for an additional 30 minutes, and then was cooled to ambient temperature. The solid material was collected by vacuum filtration, dried one hour at 120 C and then calcined five hours at 600 C in air. The catalyst thus obtained has a nominal composition of 5.7% SiO₂/ZSM-5 and it was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalyst M

A 200 g portion of calcined Catalyst J was mixed with 500 ml of dried hexane in a 1 liter round-bottom flask. The mixture was heated to reflux under N₂with stirring. While refluxing, 30 ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h through a syringe pump. Upon completion of adding TEOS, the mixture was refluxed for an additional 60 minutes, and then was cooled to ambient temperature. The solid material was collected by vacuum filtration, dried one hour at 120 C and then calcined five hours at 600 C in air. The procedures were repeated to double the amount of SiO₂loadings. The catalyst thus obtained has a nominal composition of 7.48% SiO₂/ZSM-5 and it was tested in a fluidized bed reactor for the conversion of biomass to aromatics and olefins.

Catalysts N, O, P, Q, and R

Catalysts N, O, P, Q, and R were prepared using the procedure of Catalyst D, except on a smaller scale with only 10 g of Catalyst A and the appropriate amount of TEOS.

Catalysts S and T

Catalysts S and T were prepared using the procedure of Catalyst H except on a smaller scale with only 10 g of Catalyst A and the appropriate amount of Hydrosil 2627.

Catalyst U

Catalyst U was prepared by calcining a sample of Catalyst G retrieved from Experiment 22 and then ion-exchanging it with a solution of NH4NO3 in water for 2 hours. The catalyst was dried at 120 C to give Catalyst U.

Catalyst V

A sample of a third commercially obtained spray-dried zeolite containing H-ZSM-5 with clay, an alumina binder, and some residual carbon (Zeolite C) was calcined at 550 C for 2 hours in air to form Catalyst V.

Catalysts W, X, Y, and Z

For each of these catalysts, a sample of Catalyst V was impregnated with a measured amount of TEOS in accord with the procedure used to make Catalyst L, except at a smaller scale. The catalysts were calcined 5 hours at 600 C in air to give catalysts W, X, Y, and Z.

Catalyst AA

A sample of a fourth commercially obtained spray-dried zeolite containing —H-ZSM-5 was calcined at 600 C for 2 hours in air to form Catalyst AA.

Catalyst AB—4% SiO2/ZSM5

A 10 g sample of the catalyst AA was suspended in a 200 mL hexane in a round-bottom flask. The solution was heated under stirring until reflux. A de-humidified N2 (60 mL/min) was used to purge the reflux flask and the condenser throughout the entire process. The mixture was allowed to reflux for 1 hour at 90-110° C. and the system was filled with dry N2. A 1.5 mL portion of TEOS, corresponding to 1.85 wt % Si (4.0 wt % SiO₂), was added into the mixture. Another 60 min was given under stirring and refluxing after the addition of TEOS to complete the silylation. The mixture was filtered to recover the catalyst. The catalyst was dried at 120° C. for 1 h and calcined at 600° C. for 5 h.

Catalyst AC—4% La/4% SiO2/ZSM-5

Lanthanum was loaded onto the catalyst by the incipient wetness method. A weighed portion of La(NO₃)₃(hydrated form) was dissolved in de-ionized water to obtain the desired concentration. A sample of Catalyst AB was impregnated by the lanthanum nitrate solution until the catalyst pores were filled with the solution (45 mL solution per 100 g of catalyst). The La-loaded catalyst was dried at 120° C. for 2 h and calcined at 600° C. for 5 h. In this example, 4 wt % La/was deposited on the catalyst.

Catalyst AD—4% La/ZSM-5

Lanthanum was loaded onto the catalyst by the incipient wetness method. A weighed portion of La(NO3)3 (hydrated form) was dissolved in de-ionized water to obtain the desired concentration. A sample of Catalyst AA was impregnated by the lanthanum nitrate solution until the catalyst pores were filled with the solution (45 mL solution per 100 g of catalyst). The La-loaded catalyst was dried at 120° C. for 2 h and calcined at 600° C. for 5 h. In this example, 4 wt % La/was deposited on the catalyst.

Catalyst AE—4% SiO2/ZSM-5

A 15 g sample of the catalyst AA was suspended in a 300 mL hexane in a round-bottom flask. The solution was heated under stirring until reflux. A de-humidified N2 (60 mL/min) was used to purge the reflux flask and the condenser throughout the entire process. The mixture was allowed to reflux for 1 hour at 90-110° C. and the system was filled with dry N2. A 1.58 mL portion of tetramethoxysilane (MEOS), corresponding to 1.85 wt % Si (4.0 wt % SiO₂), was added into the mixture. Another 60 min was given under stirring and refluxing after the addition of MEOS to complete the silylation. The mixture was filtered to recover the catalyst. The catalyst was dried at 120° C. for 1 h and calcined at 600° C. for 5 h.

Catalyst AF—4% SiO2/ZSM-5

The procedure of Catalyst AE was repeated.

Catalyst AG—4% SiO2/4% La/ZSM-5

A 15 g sample of the catalyst AH was suspended in a 300 mL hexane in a round-bottom flask. The solution was heated under stirring until reflux. A de-humidified N₂(60 mL/min) was used to purge the reflux flask and the condenser throughout the entire process. The mixture was allowed to reflux for 1 hour at 90-110° C. and the system was filled with dry N₂. A 1.58 mL portion of tetramethoxysilane (MEOS), corresponding to 1.85 wt % Si (4.0 wt % SiO₂), was added into the mixture. Another 60 min was given under stirring and refluxing after the addition of MEOS to complete the silylation. The mixture was filtered to recover the catalyst. The catalyst was dried at 120° C. for 1 h and calcined at 600° C. for 5 h.

Catalyst AH—4% La/4% SiO2/ZSM-5

Lanthanum was loaded onto the catalyst by the incipient wetness method. A weighed portion of La(NO3)3 (hydrated form) was dissolved in de-ionized water to obtain the desired concentration. A sample of Catalyst AE was impregnated by the lanthanum nitrate solution until the catalyst pores were filled with the solution (45 mL solution per 100 g of catalyst). The La-loaded catalyst was dried at 120° C. for 2 h and calcined at 600° C. for 5 h. In this example, 4 wt % La/was deposited on the catalyst.

Catalyst AI 4.6% SiO2/ZSM-5

A sample of Catalyst A was silylated by Chemical Liquid Deposition (CLD) using tetraethoxysilane (TEOS). A 314 g sample of the parent catalyst was suspended in a 700 mL dried hexane in a round-bottom flask. The solution was heated under stirring until reflux. A de-humidified N2 (60 mL/min) was used to purge the reflux flask and the condenser throughout the entire process. After reflux, another 1 hour was given to ensure the temperature was stabilized at 90-110° C. and the system was filled with dry N₂. A 56.25 mL portion of TEOS, corresponding to 2.2 wt % Si (4.6 wt % SiO₂), was introduced into the mixture at 12 mL/h using a syringe pump (SyringePump, NE-300). The mixture was stirred under reflux an additional 60 minutes to complete the silylation. The mixture was filtered to recover solid catalyst. The catalyst was dried at 120° C. for 1 h and calcined at 600° C. for 5 h.

Catalyst AJ. 0.5% Fe/4.6% SiO2/ZSM-5

A sample of Catalyst AI was impregnated with Fe by incipient wetness. A portion of Fe(NO3)3 (hydrated form) was dissolved in de-ionized water to obtain a solution of 4.2 wt % Fe(NO3)3. Catalyst F was then impregnated by the iron nitrate solution until the catalyst pores were filled with the solution (45 mL solution per 100 g of catalyst). The catalyst was dried at 120° C. for 2 h and calcined at 600° C. for 5 h.

Catalyst AK. 5% Fe/4.6% SiO2/ZSM-5

A sample of Catalyst AI was impregnated with Fe by incipient wetness. A portion of Fe(NO₃)₃(hydrated form) was dissolved in de-ionized water to obtain a solution of 42 wt % Fe(NO₃)₃. Catalyst F was then impregnated by the iron nitrate solution until the catalyst pores were filled with the solution (45 mL solution per 100 g of catalyst). The catalyst was dried at 120° C. for 2 h and calcined at 600° C. for 5 h.

Example 1

Catalytic fast pyrolysis (CFP) experiments were conducted in a fluidized bed reactor. The fluidized bed reactor was 1.94 inches in diameter (ID) and 24 inches in height and was made of 316 stainless steel. Inside the reactor, the catalyst bed was supported by a distributor plate made of 316 stainless steel plate with 1/16 inch circular openings.
Biomass was charged to the biomass hopper and its flow rate was controlled by an augur inside the hopper that delivers the biomass to the feed tube. A curved ¼-inch OD 316-stainless steel tube extended from the feed hopper to the biomass inlet port. A series of sieve trays made of perforated 316 stainless steel with ⅛ inch openings and 42% open area were installed inside the reactor. There were six sieve trays attached to a central, threaded rod with a 1-inch spacing between the sieve trays. The reactor was loaded with 172.97 g of catalyst A prior to the experiment and the catalyst was calcined in-situ in air at the flow rate of 1.5 SLPM and 3 SLPM of N₂for 2 hours at 600° C. The biomass feed was sieved to 20-40 mesh particle size. A 383.9 gram portion of hardwood pellets (46.67% C) was weighed and loaded into the hopper system. The reactor was purged with a flow of N2 at 3.0 SLPM for 30 minutes prior to starting the biomass conversion.
The reactor was heated to 575° C. and the fluidization gas feeding tube was heated to approximately 500° C. The solid biomass was introduced into the reactor from a side feeding tube with N2 flow. Gas flow rate through the biomass screw auger feed tube was 3.0 SLPM. The biomass flow rate was adjusted to approximately 1.4 g/min and 41.49 g of biomass was fed during the 30 minute experiment. During reaction, 1.5 SLPM of N₂was passed into the reactor through the distributor plate to fluidize the catalyst in addition to the feeding tube N₂flow.
The reactor effluent exited the reactor from the top through a heated cyclone (350° C.) to remove solid particles, including small catalyst and char particles. The effluent exiting the cyclone flowed into a product collection system that included two bubblers and three condensers. The bubblers were placed in an ice water bath and charged with 150 ml of isopropanol inside as solvent; the three condensers contained no solvent and were placed inside a Dry Ice/isopropanol bath. The uncondensed gas phase products that exited the last condenser were collected in gas bags. The reaction time was typically 30 min and two gas bag samples were taken at 15 and 30 minutes time on stream after initiating the feed of biomass. After each bag was taken, the total gas flow rate was measured with a bubble flow meter; at least 4 measurements were made and the average was used for performance calculations. The gas bags samples were analyzed by injection into a Shimadzu GC 2014 equipped with TCD and FID detectors that had been calibrated with analytical standard gas mixtures.
The contents of the two bubblers were combined into a first sample for analysis. The three condensers were rinsed with isopropanol to collect the contents, and the emptied bubblers were rinsed with isopropanol; the condenser rinse and bubbler rinse were combined into a second sample. The volumes of the two liquid samples were measured and weights determined. Liquid samples were analyzed by injection into a Shimadzu GC 2010-plus equipped with an FID detector, and 60 meter Rtx-1 capillary column from Restek
After the experiment the reactor was flushed an additional 15 minutes with N2 to ensure that the condensable products were swept into the product collection train, and then allowed to cool. The solid catalyst and char were removed from the reactor and sieved through a 60 mesh sieve. The 169.26 g portion passing through the sieve contained most of the catalyst. A small portion of this material was analyzed for carbon content on a Total Organic Carbon analyzer (TOC). The 4.19 g of larger particles were mostly char derived from biomass. This fraction was also analyzed for carbon content on a TOC. The carbon yield of aromatics was determined to be 26.47% of the carbon fed.

Example 2

The experiment in Example 1 was repeated with 173.46 g of Catalyst B and 46.59 g of hardwood.

Example 3

The experiment in Example 1 was repeated with 173 g of Catalyst C and 46.4 g of hardwood and a temperature of 525 C.
A comparison of Examples 2 and 3 with Example 1 shows the effect of iron promotion on the selectivity of the xylenes. A measurable increase in the selectivity to p-xylene was observed, from 49.0% selectivity to p-xylene without iron to 52.3 or 54.2% selectivity to p-xylene for two separate catalyst preparations containing iron in Examples 2 and 3, respectively.
Example 3 shows that the increase in the selectivity to p-xylene with the iron promoted catalyst extends to lower temperature (525 C) as well as at 575 C.

Example 4

The experiment in Example 1 was repeated with 178.3 g of Catalyst D and 50.2 g of newsprint (41.70% Carbon) and a temperature of 525 C.

Example 5

The catalyst in Example 4 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 4 was repeated with regenerated Catalyst D and 63.72 g of newsprint.
Examples 4 and 5 show the impact of coating the catalyst with SiO₂on the selectivity to p-xylene. With 4% SiO₂coating the selectivity to p-xylene increased from 49.0% from Example 1 to 71.5 and 75.5% selectivity to p-xylene in Examples 4 and 5, respectively.

Example 6

Cycle 1

The experiment in Example 1 was repeated with 173.0 g of Catalyst E and 52.29 g of hardwood and a temperature of 550 C.
Example 6 shows the surprising effect of the addition of iron to a catalyst that is coated with 4% SiO2 on the selectivity to p-xylene. Without the added iron the selectivity to p-xylene was 71.5 and 75.5 (Examples 4 and 5, respectively), whereas with the addition of 1.7% Fe the selectivity to p-xylene increased to 84.1% among the xylenes.

Example 7

Cycle 2

The catalyst in Example 6 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 6 was repeated with regenerated Catalyst E and 48.1 g of hardwood at a temperature of 550 C.
Example 7 shows the surprising effect on the selectivity to p-xylene of cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene increased from the initial 84.1% selectivity to p-xylene in the first cycle (Example 6) to 84.9% in the second cycle (Example 7).

Example 8

Cycle 3

The catalyst in Example 7 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 6 was repeated with regenerated Catalyst E and 50.24 g of hardwood and a temperature of 525 C.
The results of Example 8 show that even at lower temperature (525 C) the selectivity to p-xylene of a catalyst that has been cycled (85.2%) is greater than the selectivity to p-xylene of the fresh catalyst (84.1%).

Example 9

Cycle 4

The catalyst in Example 8 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 8 was repeated with regenerated Catalyst E and 52.76 g of hardwood at a temperature of 525 C.

Example 10

Cycle 5

The catalyst in Example 9 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 6 was repeated with regenerated Catalyst E and 50.19 g of newsprint and a temperature of 600 C.
Experiment 10 further shows that the improved selectivity to p-xylene is maintained with different feedstocks, in this case newsprint, as well as with hardwood.

Example 11

Cycle 6

The catalyst in Example 10 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 6 was repeated with regenerated Catalyst E and 45.1 g of hardwood and a temperature of 568 C.
Examples 7 through 11 show the surprising effect on the selectivity to p-xylene of repeated cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene increased from the initial 84.1% selectivity to p-xylene in the first cycle (Example 6) to 84.9%, 85.2%, 86.7%, 85.9%, and 86.0% p-xylene in successive cycles (Examples 7 through 11). These surprising results show that cycling the catalyst improves the p-xylene selectivity compared to the 84.1% selectivity to p-xylene in the fresh catalyst, and that the improved selectivity to p-xylene is maintained for multiple cycles of biomass conversion and catalyst regeneration.
Example 11 shows that the yield of aromatics of a catalyst that has been cycled 6 times and operated at 568 C (18.43% aromatics, Example 11) is higher than the fresh catalyst when operated at 550 C (17.93% aromatics, Example 6). Moreover, the yield of most desirable products, ie aromatics plus olefins, is greater with the catalyst that has been cycled 6 times (31.06% carbon yield of aromatics plus olefins, Example 11) than the fresh catalyst operated at 550 C (25.75% carbon yield of aromatics plus olefins, Example 6).

Example 12

The Experiment in Example 1 was repeated with 173.0 g of Catalyst F and 49.45 g of newsprint and a temperature of 525 C.

Example 13

The catalyst in Example 12 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 12 was repeated with regenerated Catalyst F and 48.2 g of newsprint and a temperature of 525 C.
Examples 12 and 13 show that a catalyst coated with 5.7% by weight SiO₂has a greater selectivity to p-xylene (79.4% and 83.8% selectivity to p-xylene in Examples 12 and 13, respectively) than a catalyst coated with 4% by weight SiO₂(71.5% and 75.5% selectivity to p-xylene in Examples 4 and 5, respectively). Moreover, for both the first cycle of biomass conversion with the catalyst (79.4% selectivity to p-xylene in Example 12 vs 71.5% selectivity to p-xylene in Example 4) and for a catalyst in the second cycle (83.8% selectivity to p-xylene in Example 13 vs 75.5% selectivity to p-xylene in Example 5), the selectivity to p-xylene is greater for the catalyst with 5.7% by weight SiO₂than for the catalyst with 4% by weight SiO₂.

Example 14

Cycle 1

The Experiment in Example 1 was repeated with 173.0 g of Catalyst G and 49.97 g of hardwood and a temperature of 525 C. The results show that for the catalyst with 5.7% by weight SiO₂and 1.7% by weight Fe, the selectivity to p-xylene (86.5%) was greater than for a catalyst with 1.7% Fe (52.3% and 54.2% selectivity to p-xylene in Examples 2 and 3, respectively). The results show that for the catalyst with 5.7% by weight SiO2 and 1.7% by weight Fe, the selectivity to p-xylene (86.5%) was greater than for a catalyst with 5.7% by weight SiO₂(79.4% and 83.8% selectivity to p-xylene in Examples 12 and 13, respectively). This surprising result shows the synergistic effect of a catalyst that is promoted with Fe in addition to being coated with SiO₂.

Example 15

Cycle 2

The catalyst in Example 14 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 14 was repeated with regenerated Catalyst G and 51.76 g of newsprint and a temperature of 525 C.

Example 16

Cycle 3

The catalyst in Example 15 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 15 was repeated with regenerated Catalyst G and 59.25 g of hardwood.
Example 16 shows the surprising effect on the selectivity to p-xylene of cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene increased from the initial 86.5% selectivity to p-xylene in the first cycle (Example 14) to 87.9% in the third cycle (Example 16).

Example 17

Cycle 4

The catalyst in Example 16 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 16 was repeated with regenerated Catalyst G and 48.25 g of hardwood.
Example 17 shows the surprising effect on the selectivity to p-xylene of cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene increased from the 87.9% selectivity to p-xylene in the third cycle (Example 16) to 89.3% in the fourth cycle (Example 17).

Example 18

Cycle 5

The catalyst in Example 17 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 17 was repeated with regenerated Catalyst G and 49.05 g of hardwood.

Example 19

Cycles 6 through 12

The catalyst in Example 18 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 18 was repeated with regenerated Catalyst G and 45 to 60 g of hardwood 6 additional times, regenerating the catalyst with air and N2 for 2 hours at 600 C after each exposure to hardwood. The Experiment in Example 18 was repeated a seventh time with regenerated Catalyst G and 46.46 g of hardwood. At this juncture the catalyst has participated in 12 cycles of biomass conversion.

Example 20

Cycle 13

The catalyst in Example 19 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 18 was repeated with regenerated Catalyst G and 48.88 g of hardwood.

Example 21

Cycle 14

The catalyst in Example 20 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 20 was repeated with regenerated Catalyst G and 47.62 g of hardwood.

Example 22

Cycle 15

The catalyst in Example 21 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 21 was repeated with regenerated Catalyst G and 46.87 g of hardwood.
Examples 16 through 22 show the surprising effect on the selectivity to p-xylene of repeated cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene increased from the initial 84.1% selectivity to p-xylene in the first cycle (Example 14) to 87.9%, 89.3%, 89.3%, 89.1%, 89.3%, 88.8%, 89.6% and 89.3% selectivity to p-xylene in 3, 4, 5, 12, 13, 14, and 15 cycles (Examples 16 through 22). These surprising results show that cycling the catalyst improves the p-xylene selectivity compared to the 86.5% selectivity to p-xylene in the first cycle with the catalyst, and that the improved selectivity to p-xylene is maintained for multiple cycles of biomass conversion and catalyst regeneration.
Examples 21 and 22 show that the yields of aromatics of a catalyst that has been cycled 14 and 15 times and operated at 525 C (20.79% aromatics, Example 21, 19.17% aromatics, Example 22) are higher than the fresh catalyst when operated at 525 C (18.82% aromatics, Example 14). Moreover, the yield of most desirable products, ie aromatics plus olefins, is greater with the catalyst that has been cycled 14 or 15 times (31.77% and 30.42% carbon yield of aromatics plus olefins, Example 21 and Example 22, respectively) than the first cycle with the catalyst operated at 525 C (28.91% carbon yield of aromatics plus olefins, Example 14).

Example 23

The Experiment in Example 1 was repeated with 135.8 g of Catalyst H and 48 g of hardwood and a temperature of 525 C.
Example 23 shows that the catalyst can be coated with any of a variety of sources of silica and still provide increased selectivity of p-xylene. Example 23 shows that a silicone polymer can be used as the SiO₂source and provide high selectivity to p-xylene (84.0% in Example 23) compared to the non-coated material (49.0% in Example 1).

Example 24

The catalyst in Example 23 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 23 was repeated with regenerated Catalyst H and 48 g of hardwood.
Example 24 shows that a catalyst coated with polymer-derived SiO₂that has been cycled through a biomass conversion and catalyst regeneration cycle retains a high selectivity to p-xylene (84.4%), and that the selectivity to p-xylene increases with cycling (84.4% vs 84.0% for the first cycle, Example 23).

Example 25

The Experiment in Example 23 was repeated with 173.02 g of Catalyst I and 24 g of hardwood and a temperature of 525 C.
Example 25 shows that addition of iron to a catalyst coated with SiO₂, wherein the SiO₂was derived from a silicone polymer, provides an increased selectivity to p-xylene (88.1% selectivity to p-xylene in Example 25) compared to the coated catalyst without Fe addition (84.0% selectivity to p-xylene in Example 23).

Example 26

The catalyst in Example 25 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 25 was repeated with regenerated Catalyst I and 26 g of hardwood.
Example 26 shows that a catalyst coated with polymer-derived SiO₂that has been cycled through a biomass conversion and catalyst regeneration cycle retains a higher selectivity to p-xylene (87.8%) compared to the catalyst without iron addition that has been cycled through a biomass conversion and catalyst regeneration cycle (84.4% in Example 24).
Moreover Examples 6-11, 14-22, and 25-26 show that iron addition to a SiO₂coated catalyst increases the selectivity to p-xylene for different SiO2 loadings and sources of the SiO₂.

Example 27

The Experiment in Example 23 was repeated with 170.87 g of Catalyst J and 51 g of hardwood and a temperature of 571 C. The results show that with Catalyst J the selectivity to p-xylene is 46.6%.

Example 28

The catalyst in Example 27 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 27 was repeated with regenerated Catalyst J and 27 g of hardwood and a temperature of 572 C.
The results of Example 28 show that the selectivity to p-xylene with Catalyst J that has no coating or iron addition remains low and declines relative to the first cycle (45.0% selectivity to p-xylene in the second cycle, Example 28 vs 46.6% selectivity to p-xylene in the first cycle).

Example 29

The Experiment in Example 27 was repeated with 173.02 g of catalyst K and 57 g of hardwood and a temperature of 523 C.
Results of Example 29 show that a catalyst that contains clay and is coated with 7.48% SiO₂and has added iron provides a higher selectivity to p-xylene (90.7% in Example 29) than the uncoated, unpromoted catalyst (Examples 27-28 where the selectivities to p-xylene were 46.4 and 45.0, respectively) or than the coated, unpromoted catalyst (Examples 34-35, where the selectivities to p-xylene were 87.5 and 88.3, respectively).

Example 30

The catalyst in Example 29 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 29 was repeated with regenerated Catalyst K and 60 g of hardwood and a temperature of 522 C.
Example 30 shows that a catalyst coated with 7.48% SiO2 and promoted with Fe that has been cycled through a biomass conversion and catalyst regeneration cycle retains a high selectivity to p-xylene (91.3%), and that the selectivity to p-xylene unexpectedly increases with cycling (91.3% vs 90.7%% for the first cycle, Example 29). Example 30 shows that with an appropriate combination of catalyst components, catalyst binder, SiO₂coating content, and promoter elements, the selectivity to p-xylene is increased with cycling in biomass conversion and catalyst regeneration cycles.

Example 31

The Experiment in Example 23 was repeated with 170.87 g of Catalyst L and 51 g of hardwood and a temperature of 526 C.
The results of Example 31 show that coating a clay-containing catalyst with 5.7% by weight SiO2 can increase the selectivity to p-xylene (85.3% in Example 31) compared to the uncoated catalyst (46.6% in Example 27). The results of Example 31 show that the increase in selectivity to p-xylene produced by SiO₂coating is general for ZSM-5 containing catalysts with different binders and from different sources.

Example 32

The catalyst in Example 31 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 31 was repeated with regenerated Catalyst L and 27 g of hardwood and a temperature of 527 C.
Example 32 shows that a catalyst coated with SiO₂that has been cycled through a biomass conversion and catalyst regeneration cycle retains a high selectivity to p-xylene (76.4%). Example 32 shows that the selectivity to p-xylene with a catalyst containing ZSM-5 and comprising clay and coated with 5.7% SiO₂does not increase with cycling (76.4%% vs 85.3%% for the first cycle, Example 31).

Example 33

The catalyst in Example 32 was regenerated in the reactor by reaction with air and N2 over the course of 2 hours at 600 C. The Experiment in Example 31 was repeated with regenerated Catalyst L and 24 g of hardwood and a temperature of 553 C.
Example 33 shows that a catalyst coated with SiO₂that has been cycled through a biomass conversion and catalyst regeneration cycle retains a high selectivity to p-xylene (74.2%). Example 33 shows that the selectivity to p-xylene with a ZSM-5 containing catalyst comprising clay and coated with 5.7% SiO₂does not increase with cycling (74.2% for the third cycle vs 85.3%% for the first cycle, Example 31, and 76.4% for the second cycle, Example 32).
Examples 32 and 33 show that the increase in selectivity to p-xylene is limited to particular combinations of catalyst components, catalyst binder, SiO₂coating content, and presence of promoter elements.

Example 34

The Experiment in Example 23 was repeated with 162.9 g of Catalyst M and 45 g of hardwood and a temperature of 522 C.
The results of Example 34 show that coating a clay-containing catalyst with 7.48% by weight SiO2 can increase the selectivity to p-xylene (87.5% in Example 34) compared to the uncoated catalyst (46.6% in Example 27). The results of Example 34 show that the increase in selectivity to p-xylene produced by SiO₂coating is general for ZSM-5 containing catalysts with different binders and from different sources. The results of Example 34 show that with some catalysts a greater amount of SiO₂coating increases the selectivity to p-xylene compared to a lesser amount of SiO₂; Example 34 with a catalyst having 7.48% SiO₂coating provided 87.5% selectivity to p-xylene compared to Example 31 with a catalyst having 5.7% SiO₂coating that provided 85.3% selectivity to p-xylene.

Example 35

The catalyst in Example 34 was regenerated in the reactor by reaction with air and N2 over the course of 3 hours at 600 C. The Experiment in Example 34 was repeated with regenerated Catalyst M and 45 g of hardwood and a temperature of 525 C.
Example 35 shows the surprising effect on the selectivity to p-xylene of cycling of the catalyst. As the catalyst was cycled through biomass upgrading and regeneration cycles, the selectivity to p-xylene unexpectedly increased from the initial 87.5% in the first cycle (Example 34) to 88.3% in the second cycle (Example 35). Example 35 shows that with an appropriate combination of catalyst components, catalyst binder, SiO₂coating content, and promoter elements, the selectivity to p-xylene is increased with cycling in biomass conversion and catalyst regeneration cycles.

Examples 36-51

In order to determine the acid site concentration on various materials the desorption of isopropylamine from the materials is measured by means of Isopropyl-amine Temperature-Programmed Desorption (IPA-TPD). For the IPA-TPD experiments, a TGA instrument (Shimadzu TGA-50) is adjusted to read zero with an empty platinum sample cell. The sample cell is then filled with a sample of catalyst powder (30-50 mg). The catalyst is pre-treated at 500 C under 50 mL/min N2. It is then cooled to 120 C under a 50 mL/min flow of N2. Isopropylamine (IPA) is fed into the TGA chamber at this temperature by flowing a 2nd portion of N2 gas (<10 mL/min) through a bubbler filled with liquid IPA while monitoring the weight of the sample. The feed of IPA is stopped when the catalyst is saturated as indicated by no more weight increase. The flows of N2 are maintained through the chamber, but bypassing the IPA bubbler, for an additional 120 min to remove weakly adsorbed′IPA. The TGA chamber is then heated up to 700° C. at a ramping rate of 10° C./min to obtain desorption curves, and the weight is monitored as a function of temperature.
In the desorption curve, the sharp desorption at 270-380 C is assigned to IPA decomposition into propylene and NH3 occurring on the Brønsted acid sites. The peak area under the desorption curve measured from 270 to 380 C is used for quantifying the number of Brønsted acid sites for a particular sample.
Examples 36-41 and FIG. 2 show that the addition of a coating, e.g., silicone coating derived from TEOS, reduces the concentration of acid sites on the catalyst and that more coating reduces the concentration further.
Example 41 and FIG. 2 show that at some level of SiO₂coating (ie 11.46% SiO₂in Example 41 compared to 7.80 wt % in Examples 39 and 40 and lower wt % in Examples 37-38), the concentration of acid sites is no longer significantly reduced by addition of more SiO₂.
Examples 42 and 43 show that a polysiloxane coating reduces the concentration of acid sites compared to the concentration of acid sites in the parent catalyst in Example 36. Example 43 shows that increasing the amount of coating with a polysiloxane from 3.96% to a larger amount does not further reduce the concentration of acid sites.
Example 42 shows that the use of a polysiloxane reduces the concentration of acid sites more than an equivalent amount of SiO₂derived from TEOS (compare to Example 37). Example 43 and FIG. 2 shows that addition of polysiloxane beyond 3.96% by weight does not significantly reduce the concentration of acid sites in the catalyst.
Example 44 shows that the addition of Fe to a silicone coated catalyst further reduces the concentration of acid sites compared to the catalyst without Fe in Example 39.
Example 45 shows that cycling the coated and Fe-promoted catalyst multiple times through biomass upgrading and catalyst regeneration reduces the concentration of acid sites and with Example 22 shows that the carbon yields of aromatics (19.17%) and olefins (11.25%) remain high and the selectivity to p-xylene remains high (89.3%), and is higher than in the first cycle. Example 45 shows that a catalyst with reduced concentration of acid sites (5.9% of the parent zeolite) can be effectively used to convert biomass to useful products.
Example 46 shows that the concentration of acid sites in a silicone coated and Fe-promoted catalyst that has been cycled multiple times can be increased by an ion exchange with acid solution (NH₄NO₃).
Example 47 through 50 and FIG. 2 shows that a commercially obtained spray-dried zeolite containing H-ZSM-5 with clay, an alumina binder, and some residual carbon (Zeolite C) shows a reduction in acid concentration with added coating of SiO₂. Examples 48 through 50 show that addition of SiO₂beyond 5.89% by weight SiO₂does not significantly further reduce the concentration of acid sites
Examples 36-43 and examples 47-51 show that the moles of Brønsted acid sites deactivated by the siliceous coating is at least 0.015 or at least 0.03, or at least 0.04, or at least 0.06 per mole of Si added as measured by desorption of isopropyl amine (IPA) in a temperature programmed desorption experiment.

Example 52

The 2-methylfuran conversion was carried out in ½″ O.D. stainless steel tubular reactor. A 1-gram sample of Catalyst AA was held by a piece of quartz wool inside the reactor. N2 was used as the carrier gas. 2-methylfuran (b.p.=65° C.) was dripped into the top of the reactor by an HPLC pump and it was carried with a down-flow of N₂into the catalyst bed. The product stream was collected in gas bags at 4, 7, 10, 15, and 20 minutes time on stream. Coke deposited on the catalyst was quantified by using a TOC instrument (Shimadzu SSM-5000A). The results in Table 4 are averages of the samples collected after 7 and 10 minutes have elapsed since the introduction of 2-methylfuran.

Example 53

The experiment from Example 52 was repeated with Catalyst AB in place of Catalyst AA. The experiment was repeated after regenerating the catalyst by exposure to air at 650 C. The cycle was repeated for a total of 3 experiments with Catalyst AB. The results in Table 4 are the average of the three experiments.
Example 53 shows that coating a zeolite with 4 wt % SiO₂using TEOS as the source of the coating results in high selectivity to p-xylene (92.3%) compared to the uncoated catalyst (75.3%).

Example 54

The experiment in Example 52 was repeated with Catalyst AC in place of Catalyst AA. Example 54 shows that addition of 4% by weight of La to a catalyst coated with 4% by weight SiO₂increases the aromatic yield (16.6% vs 12.3%) compared to the catalyst that has no La added and that the high selectivity to p-xylene is maintained (92.1% vs 93.3%).

Example 55

The experiment in Example 52 was repeated with Catalyst AD in place of Catalyst AA. Example 55 shows that the addition of 4% by weight of La to an uncoated catalyst improves the aromatics yield (16.1% vs 11.7%) and the p-xylene selectivity (80.2% vs 75.3%).

Example 56

The experiment in Example 52 was repeated with Catalyst AE in place of Catalyst AA. Example 56 shows that a catalyst coated with SiO₂using MEOS as the source of SiO₂shows increase p-xylene selectivity compared to the uncoated catalyst (95.5% vs 75.3%).

Example 57

The experiment in Example 52 was repeated with Catalyst AF in place of Catalyst AA. Example 57 shows that a catalyst prepared at larger scale provides similar improvement in p-xylene selectivity as the catalyst prepared at the smaller scale.

Example 58

The experiment in Example 52 was repeated with Catalyst AG in place of Catalyst AA. Example 58 shows that when the catalyst that is promoted first with 4% by weight La and then silicone coated the yield of aromatics is increased compared to the coated catalysts (14.7% vs 9.6%) but the p-xylene selectivity is lower than the catalyst that has only been coated (80.2% vs 75.3%). Example 58 shows that a La-promoted and then silicone coated catalyst has higher yield of aromatics (14/7% vs 11.7%) and p-xylene selectivity (86.2% vs 75.3%) compared to an unpromoted, uncoated catalyst.

Example 59

The experiment in Example 52 was repeated with Catalyst AH in place of Catalyst AA. Example 59 shows that when the catalyst that is first coated with 4% by weight SiO₂and then promoted with 4% La the yield of aromatics (14.8 vs 14.7%) and the p-xylene selectivity (95.3% vs 86.2%) are higher than when the catalyst is first impregnated with La and then coated with SiO₂; Example 59 shows that the sequence of catalyst promotion and coating is critical to obtaining an improved catalyst and that the sequence in which the catalyst is first coated with silicone and then promoted with metal addition provides high yield of aromatics and higher selectivity to p-xylene compared to the catalyst prepared by promotion followed by coating.

Example 60

The furfural conversion was carried out in ½″ O.D. stainless steel tubular reactor. A 1-gram sample of Catalyst F was held by a piece of quartz wool inside the reactor. N₂was used as the carrier gas. Furfural (b.p.=162° C.) was dripped into the top of the reactor by an HPLC pump and it was carried with a down-flow of N₂into the catalyst bed. The product stream was collected in gas bags at 4, 8, 12, and 16 minutes time on stream. Coke deposited on the catalyst was quantified by using a TOC instrument (Shimadzu SSM-5000A).
The results in Table 4 are averages of the samples collected after 8 and 12 minutes have elapsed since the introduction of furfural.

Example 61

The experiment of Example 60 was repeated with Catalyst AI in place of Catalyst F.
The results in Table 5 show that addition of 0.5% Fe to the coated catalyst increased the selectivity of p-xylene (83.2% vs 74.9%) and yield of p-xylene (1.70% vs 1.36%) by the addition of Fe.

Example 62

The experiment in Example 60 was repeated with Catalyst AJ in place of Catalyst F.
The results in Table 5 show that addition of 5.0% Fe to the coated catalyst increased the selectivity of p-xylene (84.6% vs 74.9%) and yield of p-xylene (2.04% vs 1.36%) by the addition of 5% Fe compared to the Fe-free catalyst, and that the addition of 5.0% Fe increased the selectivity of p-xylene (84.6% vs 83.2%) and yield of p-xylene (2.04% vs 1.70%) by the addition of 5.0% Fe compared to the 0.5% Fe catalyst. The results in Table 5 show that the addition of Fe to the catalyst decreased the ratio of CO/CO₂with increasing Fe content indicating a more efficient rejection of oxygen from the product mixture with added Fe.

TABLE 1

Catalyst Preparations. The Wt % SiO2 is the fraction of SiO2 added in
the treatment step.

	Run		Metal
Catalyst	Numbers	Zeolite	Additive	Wt % SiO2	SiO2 Source

A	1	A	none	none	none
B	2	A	1.7% Fe	none	none
C	3	A	1.7% Fe	None	none
D	4-5	A	none	4	TEOS
E	6-11	A	1.7% Fe	4	TEOS
F	12-13	A	none	5.7	TEOS
G	14-22	A	1.7% Fe	5.7	TEOS
H	23-24	A	none	4	H2O soluble
					polymer
I	25-26	A	1.7% Fe	4	H2O soluble
					polymer
J	27-28	B	none	none	none
K	29-30	B	1.7% Fe	7.48	TEOS
L	31-33	B	none	5.7	TEOS
M	34-35	B	none	7.48	TEOS

TABLE 6

Brønsted acid concentration obtained from IPA-TPD experiments

		Brønsted acid	% of Acid			Brønsted acid
	Fe	concentration	Sites in	Fe added	Fe:	sites deactivated
	content	(mmol/mg) ×	Parent	(mmol/mg) ×	Acid site	per Fe added
Catalyst	(wt %)	1000	Zeolite	1000	ratio (1)	(mol/mol) (2)

A	0	0.2000	100	—	—
AI	0	0.1445	72	—	—
AJ	0.5	0.0901	45	0.089	0.62	0.61
AK	5.0	0.0414	21	0.89	6.2	0.12

(1) Moles of Fe per mole of acid sites in the Fe-free catalyst AI.
(2) Moles of Brønsted acid sites deactivated per moles of Fe added as determined by IPA-TPD.

TABLE 1

Experimental testing results of catalysts in fluidized bed reactor tests.

						Xylene
Run		Biomass			Yield, % Carbon	Selectivity, %

No

Catalyst

Feed

T, C.

Cycle

Aromatics

Olefins

CO

CH₄

CO₂

pX

mX

oX

1	A	Hardwood	575	1	26.47	7.5	18.19	2.50	7.39	49.0	40.8	10.2
2	B	Hardwood	575	1	22.32	9.2	12.91	2.67	7.24	52.3	36.7	11.1
3	C	Hardwood	525	1	20.49	8.1	15.85	2.36	7.88	54.2	35.8	10
4	D	Newsprint	525	1	21.98	6.4	11.41	1.57	4.25	71.5	22.6	5.9
5	D	Newsprint	525	2	17.90	7.5	13.24	1.82	4.97	75.5	19.5	5.0
6	E	Hardwood	550	1	17.93	7.82	11.01	2.43	5.64	84.1	13.0	2.9
7	E	Hardwood	550	2	16.83	7.73	10.88	2.41	5.57	84.9	12.1	3.0
8	E	Hardwood	525	3	16.07	9.39	13.22	2.61	6.69	85.2	12.1	2.7
9	E	Hardwood	525	4	15.19	8.12	11.59	2.18	5.69	86.7	10.8	2.5
10	E	Newsprint	600	5	17.98	11.33	18.23	4.71	8.21	85.9	11.3	2.8
11	E	Hardwood	568	6	18.43	12.63	18.74	4.72	8.38	86.0	10.8	3.2
12	F	Newsprint	525	1	24.69	8.33	14.27	2.07	5.57	79.4	16.7	3.9
13	F	Newsprint	525	2	24.37	8.41	14.48	2.11	5.65	83.8	13.1	3.2
14	G	Hardwood	525	1	18.82	10.09	13.22	2.70	7.03	86.5	11.2	2.4
15	G	Newsprint	525	2	19.79	9.60	13.31	2.12	6.10	86.2	11.2	2.6
16	G	Hardwood	525	3	16.67	10.10	13.79	2.77	6.92	87.9	9.8	2.4
17	G	Hardwood	525	4	19.20	11.21	14.46	2.88	7.23	89.3	8.6	2.1
18	G	Hardwood	525	5	19.25	11.35	15.29	3.00	7.58	89.1	8.8	2.1
19	G	Hardwood	525	12	17.45	9.68	12.17	2.49	6.15	89.3	—	—
20	G	Hardwood	525	13	17.01	11.61	14.90	2.63	7.92	88.8	—	—
21	G	Hardwood	525	14	20.79	10.98	13.91	2.81	7.06	89.6	—	—
22	G	Hardwood	525	15	19.17	11.25	14.26	2.87	7.24	89.3	—	—
23	H	Hardwood	525	1	22.94	10.28	14.05	2.17	6.62	84.0	11.0	5.1
24	H	Hardwood	525	2	24.43	10.43	14.27	2.21	6.84	84.4	10.6	5.0
25	I	Hardwood	525	1	22.02	10.41	13.37	2.40	7.36	88.1	9.3	2.5
26	I	Hardwood	525	2	22.07	10.72	13.49	2.35	7.03	87.8	9.7	2.4
27	J	Hardwood	571	1	16.08	8.96	18.93	3.68	4.90	46.6	38.8	14.5
28	J	Hardwood	572	2	16.16	7.91	16.12	3.37	4.18	45.0	40.2	14.8
29	K	Hardwood	523	1	14.15	9.91	13.47	2.43	5.50	90.7	7.0	2.6
30	K	Hardwood	522	2	13.48	19.43	14.29	5.90	5.69	91.3	7.0	2.2
31	L	Hardwood	526	1	19.11	10.48	15.97	2.69	5.30	85.3	11.6	3.1
32	L	Hardwood	527	2	21.95	11.41	17.22	2.42	5.85	76.4	18.7	4.9
33	L	Hardwood	553	3	21.03	12.17	19.61	2.94	6.20	74.2	20.4	5.4
34	M	Hardwood	522	1	14.03	9.21	16.72	2.97	5.27	87.5	10.0	2.6
35	M	Hardwood	525	2	13.23	9.24	16.79	2.98	5.29	88.3	9.2	2.2

TABLE 2

IPA-TPD experimental results for catalysts.

							Brønsted	Acidity -
					Wt %	Moles Si loading	acidity	% of	Brønsted acid sites
				Coating	SiO₂	[mmol/mg]	[mmol/mg]	Parent	deactivated per Si
Example	Catalyst	Zeolite	Promoter	Compound	Added	(×1000)	(×1000)	Zeolite	added [mol/mol]

36	A	A	none	na	—	—	0.2000	100	—
37	N	A	Si	TEOS	3.96	0.661	0.1749	87	0.038
38	O	A	Si	TEOS	5.89	0.982	0.1623	81	0.038
39	P	A	Si	TEOS	7.80	1.300	0.1448	72	0.042
40	Q	A	Si	TEOS	7.80	1.300	0.1424	71	0.044
41	R	A	Si	TEOS	11.46	1.911	0.1419	71	0.030
42	S	A	Si	Polysiloxane	3.96	0.661	0.1382	69	0.094
43	T	A	Si	Polysiloxane	5.89	0.982	0.1364	68	0.065
44	G	A	Si, Fe	TEOS	5.7	1.000	0.1177	58.8	na
45	G* (1)	A	Si, Fe	TEOS	5.7	1.000	0.0118	5.9	na
46	U	A	Si, Fe	TEOS	5.7	1.000	0.0349	17.5	na
47	V	C	none	na	—	—	0.0854	100	—
48	W	C	Si	TEOS	3.96	0.661	0.0628	73	0.034
49	X	C	Si	TEOS	5.89	0.982	0.0523	61	0.034
50	Y	C	Si	TEOS	7.80	1.300	0.0504	59	0.027
51	Z	C	Si	TEOS	11.46	1.911	0.0530	62	0.017

(1). Sample G* is a sample from Experiment 22 that has been subjected to 15 cycles of biomass conversion and regeneration in air.

TABLE 3

Results of 2-methyl-furan (2MF) conversion experiments.

2MF

Yield, Carbon %

Coating

SiO₂

conversion

Aromatics +

Selectivity %

Example	Catalyst	Sequence	Source	(%)	Aromatics	Olefins	Olefins	pXylene	pX	mX	oX

52	AA	None	na	86.2	11.7	12.6	24.34	0.97	75.3	21.7	3.1
53	AB	Si only	TEOS	90.4	12.3	13.4	17.63	1.27	93.3	5.4	1.6
54	AC	La on Si	TEOS	98.8	16.6	18.1	34.71	1.26	92.1	6.6	1.3
55	AD	La only	TEOS	96.1	16.1	16.1	32.18	1.23	80.2	17.2	2.6
56	AE	Si only	MEOS	82.0	9.4	10.8	20.15	0.98	95.5	3.5	0.9
57	AF	Si only	MEOS	80.8	9.6	11.6	21.25	1.02	96.0	3.0	1.1
58	AG	Si on La	MEOS	98.1	14.7	15.0	29.74	1.16	86.2	12.2	1.6
59	AH	La on Si	MEOS	96.4	14.8	15.8	30.61	1.37	95.3	4.2	0.5

TABLE 4

Results of Furfural Conversion Experiments.

Furfural

Yield, % Carbon

Wt %

conversion

CO/CO₂

Aromatics +

Selectivity, %

Example	Catalyst	Fe	(%)	ratio	Aromatics	Olefins	Olefins	pXylene	pX	mX	oX

60	AI	0.0	100	9.77	24.85	7.58	32.43	1.36	74.9	21.2	3.9
61	AJ	0.5	100	7.00	23.00	8.95	31.95	1.70	83.2	14.1	2.7
62	AK	5.0	100	5.67	24.44	8.78	33.22	2.04	84.6	13.1	2.3

Claims

What is claimed:

1. A process for converting biomass to liquid hydrocarbons, comprising:

feeding biomass into a reactor;

heating the biomass in the presence of an aluminosilicate zeolite catalyst; and

wherein the aluminosilicate zeolite catalyst further comprises at least 0.2 wt % Fe wherein the Fe is not derived from biomass or reactor walls, or

wherein the aluminosilicate zeolite catalyst has been manufactured to contain at least 0.2 wt % Fe or pretreated to contain at least 0.2 wt % Fe; and

converting the biomass to a gaseous product stream comprising p-xylene.

2. The process of claim 1 wherein the aluminosilicate zeolite catalyst further comprises at least 0.5, or at least 1.0, or at least 1.5 wt % Fe, and in some embodiments up to 10 wt % Fe, up to 5 wt % Fe, or up to 3 wt % Fe.

3. The process of claim 1 wherein the gaseous product stream comprises at least 10 wt % of aromatic compounds or at least 15 wt % of aromatic compounds and in some embodiments up to 30 wt % aromatics, in some embodiments up to 25 wt %; and/or wherein at least 85% of the xylenes in the gaseous product stream is p-xylene.

4. The process of claim 1 wherein the aluminosilicate zeolite catalyst has been pretreated with a siliceous coating, or wherein the aluminosilicate zeolite catalyst has a siliceous coating.

5. The process of any of claim 1 further comprising:

removing the catalyst from the reactor after it has been used to pyrolyze the catalysis,

heating the used catalyst in the presence of an oxygen containing gas (preferably O2) to form a regenerated catalyst, and returning the regenerated catalyst to the reactor, and again using the catalyst to catalyze the conversion of the biomass to a gaseous product stream comprising p-xylene.

6. A method of making a catalyst, comprising:

providing a zeolite catalyst;

treating the catalyst to increase the iron content; and

applying a siliceous coating to the catalyst;

resulting in an Fe-modified, zeolite catalyst having a siliceous coating.

7. The method of claim 2, comprising:

using the Fe-modified, zeolite catalyst having a siliceous coating to catalyze the pyrolysis of biomass; and

subsequent to the pyrolysis of biomass, regenerating the catalyst by heating in the presence of an oxygen containing gas.

8. A hydrocarbon mixture, comprising: a biomass-derived (i.e., ¹⁴C-containing) mixture of hydrocarbons comprising at least 10 mass % of xylenes; wherein the xylenes are made up of 85 to about 91% p-xylene.

9. The hydrocarbon mixture of claim 8 made by the process of any of claims 1 to 5.

10. The hydrocarbon mixture of claim 8 comprising catalyst particles of the type described herein.

11. A chemical system, comprising:

a reactor, comprising

an Fe-modified zeolite catalyst;

biomass; and

a hydrocarbon product stream comprising at least 10 mass % xylenes wherein at least 80% of the xylenes are p-xylene.

12. The chemical system of claim 11 wherein the Fe-modified zeolite catalyst comprises a siliceous coating.

13. The chemical system of claim 11 wherein the zeolite catalyst comprises ZSM-5.

14. An aluminosilicate zeolite catalyst having a Si/Al molar ratio of 100 or less, comprising:

at least 0.2 wt % Fe wherein the Fe is not derived from biomass or reactor walls, or

a siliceous coating.

15. The aluminosilicate zeolite catalyst of claim 14 wherein the catalyst has Fe evenly distributed over the surface as measured by SEM-EDS.

16. The aluminosilicate zeolite catalyst of claim 14 wherein the catalyst comprises ZSM-5.

17. The aluminosilicate zeolite catalyst of claim 14 wherein the catalyst has a surface ratio of Si/Fe in the ratio of 50:1 to 4:1; preferably 30:1 to 5:1; in some embodiments 20:1 to 7:1.

18. The aluminosilicate zeolite catalyst of claim 14 wherein the Fe is concentrated in clusters on the surface of the catalyst.

19. The aluminosilicate zeolite catalyst of claim 14 having a Brønsted acidity of greater than 0.01, 0.05 or greater, preferably a Brønsted acidity in the range of 0.01 to 0.2, preferably in the range of 0.04 to 0.15, preferably in the range of 0.05 to 0.1 μmol/mg.