SYSTEMS AND PROCESSES FOR CATALYTIC BIOMASS PYROLYSIS AND HYDROCARBONIFIC MATERIALS FOR PRODUCTION O

专利摘要:
systems and processes for catalytic pyrolysis of biomass and hydrocarbon materials for optional olefin recycling aromatics production and catalysts having particle size selected for catalytic pyrolysis This invention relates to compositions and methods for the fluid hydrocarbon product, and more specifically to compositions and methods for the fluid hydrocarbon product by catalytic pyrolysis. Some embodiments refer to methods for the production of specific aromatic products (e.g. benzene, toluene, naphthalene, xylene, etc.) by catalytic pyrolysis. some of these methods may involve the use of a composition comprising a mixture of a solid hydrocarbonaceous material and a heterogeneous pyrolytic catalyst component. In some embodiments, an olefin compound may be co-fed to the reactor and / or separated from a product stream and recycled to the reactor to improve yield and / or selectivity of certain products. The methods described herein may also involve the use of specialized catalysts. For example, in some cases, zeolite catalysts may be used. In some cases, catalysts are characterized by particle sizes in certain identified ranges that may lead to improved yield and / or selectivity of certain products.
公开号:BR112012005379B1
申请号:R112012005379-5
申请日:2010-09-09
公开日:2019-03-06
发明作者:Anne Mae Gaffney；Jungho Jae；Yu-Ting Cheng；George W. Huber
申请人:University Of Massachusetts；Anellotech, Inc；
IPC主号:

专利说明:

SYSTEMS AND PROCESSES FOR CATALYTIC PYROLYSIS OF BIOMASS AND HYDROCARBONIFERING MATERIALS FOR THE PRODUCTION OF AROMATICS WITH OPTIONAL OLEPHINE RECYCLING AND CATALYSTS HAVING PARTICLE SIZE SELECTED FOR CATALYTIC PYROLYSIS
RELATED REQUESTS
This application claims priority under 35 USC §119 (e) for Provisional Patent Application No. 61 / 241,018, filed September 9, 2009, and entitled Systems and Processes for Catalytic Pyrolysis of Biomass and Hydrocarbon Materials for the Production of Aromatic Compounds with Optional Olefine Recycling and Catalysts having Selected Particle Size for Catalytic Pyrolysis, which is incorporated by reference in its entirety for all purposes.
DECLARATION ON FEDERAL RESEARCH OR SPONSORED DEVELOPMENT
The US Government has a license terminated on this invention and the right in limited circumstances to require the patent owner to license other interested parties on reasonable terms, as provided for under Grant No. CBET-0747996 granted by the National Science Foundation.
FIELD OF THE INVENTION
This invention relates to compositions and methods for the production of biochemical products, such as biofuel, aromatic compounds and olefin, and more specifically, to compositions and methods for biochemical production through catalytic pyrolysis.
BACKGROUND
With its low cost and high availability, lignocellulosic biomass has been the object of study worldwide as a raw material for renewable liquid biofuels. One impulse, in particular, is that
2/119 fuels derived from biomass have zero net CO2 emissions if produced without the use of fossil fuels. However, currently, lignocellulosic biomass is not commonly used as a source of liquid fuel because typical current conversion processes are not considered to be economically viable. Several routes are being examined to convert solid biomass to liquid fuel. At low temperatures (eg 200-260 ° C) alkanes of diesel variation can be produced by multi-stage aqueous phase processing (APP) of aqueous carbohydrate solutions involving dehydration, ---- aldolic condensation è dehydration / hydrogenation (GW Huber, JA Dumesic, Catalysis Today 2006, 111, 119-132.). However, APP requires that solid lignocellulosic biomass is first converted to aqueous carbohydrates. At higher temperatures (~ 800 ° C), solid biomass can be reformed to produce synthesis gas through partial oxidation on catalysts in an auto-thermal packed bed reactor. (P.J. Dauenhauer, J.D. Dreyer, N.J. Degenstein, L.D. Schnudt, Angew. Chem. Int. Ed. 2007, 46, 5864-5867.). The synthesis gas produced from the reaction can be fed to a secondary process to make fuels and chemicals. For certain applications, an ideal process for converting solid biomass may involve the production of a liquid fuel that fits within the existing infrastructure from solid biomass in a single step, in short residence times. Unfortunately, neither APP nor the gas synthesis process meets these criteria.
Another approach to biofuel production is rapid pyrolysis, which can involve, for example, rapidly heating biomass (eg ~ 500 ° C / sec) to
3/119 intermediate temperatures (eg ~ 400-500 ° C) followed by rapid cooling (eg residence times 1-2 s). (See, AV Bridgwater, Fast Pyrolysis of Biomass: A Handbook Volume 2, CPL Press, Newbury, UK, 2002.) Conventional rapid pyrolysis often produces a mixture of the thermally unstable liquid product called bio-oils, a more acidic liquid fuel mixture. of 300 compounds that degrades over time. However, bio-oils are not compatible with existing liquid transport fuels, such as gasoline and diesel, and yields are low.
---- By — COTrs ^ gxrinTtey-permarriza a search in the course in the technique for an efficient economic route for the production of useful biofuels and related compounds from solid biomass.
SUMMARY OF THE INVENTION
This invention relates, generally, to compositions and methods for the production of biochemical products, such as biofuel, aromatic compounds and olefins. The object of the present invention involves, in some cases, interrelated products, alternative solutions to a particular problem, and / or a plurality of different uses of one or more systems and / or articles.
In one aspect, a method for producing one or more fluid hydrocarbon products from a hydrocarbonaceous material is provided. The method may comprise feeding a hydrocarbonaceous material to a reactor, and pyrolyzing within the reactor at least a portion of the hydrocarbonaceous material under sufficient reaction conditions to produce one or more pyrolysis products.
In some embodiments, the method may comprise catalytically reacting within the reactor at least one
4/119 portion of one or more pyrolysis products under sufficient reaction conditions to produce one or more fluid hydrocarbon products comprising olefins and aromatics, separating at least a portion of the olefins in the fluid hydrocarbon products to produce a recycling stream comprising at least the separate olefins, and a product stream, and feed at least a portion of the recycle stream to the reactor.
In some embodiments, the method may comprise feeding a solid hydrocarbon material to a reactor, pyrolyzing at least a portion of the hydrocarbon material within the reactor under sufficient reaction conditions to produce one or more pyrolysis products, and catalytically reacting at least a portion of one or more pyrolysis products, in sufficient reaction conditions to produce one or more fluid hydrocarbon products comprising olefins and aromatics. In some cases, the method may further comprise separating at least a portion of the olefins into one or more fluid hydrocarbon products to produce a recycling stream comprising at least the separate olefins and a product stream, and feeding at least a portion of the recycling stream for the reactor.
In some cases, the method may comprise providing a hydrocarbonaceous material and one or more catalysts comprising a plurality of particles having maximum cross-sectional dimensions of less than about 1 micron, and pyrolyzing at least a portion of the hydrocarbonaceous material under sufficient reaction conditions to produce one or more pyrolysis products. In some cases, the method may further catalytically react at least a portion of the products
5/119 pyrolysis with the catalysts to produce one or more hydrocarbon products.
The method may comprise, in some embodiments, providing a hydrocarbonaceous material and a zeolite catalyst comprising gallium, pyrolyzing at least a portion of the hydrocarbonaceous material under sufficient reaction conditions to produce one or more pyrolysis products, and catalytically reacting at least a portion of the pyrolysis products with the catalysts to produce one or more hydrocarbon products.
Other new advantages and features of the present invention will be evident from the following detailed description of the various non-limiting modalities of the invention when considered in conjunction with the accompanying figures. In cases where the present report and a document incorporated by reference include conflicting and / or inconsistent disclosure, the present report shall prevail. If two or more documents incorporated by reference include conflicting and / or inconsistent disclosure with respect to each other, then the document with the effective later date must prevail.
BRIEF DESCRIPTION OF THE DRAWINGS
Non-limiting modalities of the present invention will be described by way of example, with reference to the accompanying figures, which are schematic and are not intended to be launched to scale. In the figures, each identical or nearly identical component illustrated is typically represented by a single numeral. For the sake of clarity, not every component is marked on each figure, nor is every component of each embodiment of the invention shown where illustration is not necessary to allow a person skilled in the art to understand the invention. In the figures:
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FIG. 1 is a schematic diagram of a catalytic pyrolysis process, according to a set of modalities;
Figs. 2A-2B are graphs of (A) carbon yield for various raw materials derived from biomass (aromatics: horizontal lines, CO2: white, CO: diagonal lines, coke: black, and unidentified: gray) and (B) selectivity aromatic for benzene (Ben.), toluene (Tol.), ethyl-benzene and xylenes (E-Ben., Xil.), methyl-ethyl-benzene and trimethyl-benzene (m, e-Ben., tmBen. ), Indians (Ind.), and Naphthalenes (Naf.) according to a fashion group;
FIG. 3 is a graph of carbon yield of CO (), aromatics (±), CO ₂ (Δ), and coke (·) as a function of the nominal heating rate for a catalytic glucose pyrolysis with ZSM5, according to a set of modalities;
Figs. 4A-4B are graphs of (A) carbon yield of CO (), aromatics (^), CO ₂ (Δ), partially deoxygenated species (□), and coke (·) as a catalyst function for mass ratio of glucose and (B) a distribution of partially deoxygenated species of hydroxyacetylaldehyde (HA), acetic acid (AA), furan (Fur.), furfural (Furf), methyl furan (M-Fur), 4-methylfurfural (4-M- Furf) and furan-2-methanol (fur-2-MeOH), according to a set of modalities;
FIG. 5 is a graph of carbon yields
after performing a catalytic glucose pyrolysis with several catalysts (Aromatic: horizontal lines, CO _2: white, CO : diagonal lines, species partially
deoxygenated: gray and coke: black) according to a set of modalities;
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Figs. 6A-6B are graphs of (A) carbon yield for various molar ratios from silica to alumina in the catalyst and (B) aromatic selectivity for benzene (Ben.), Toluene (Tol.), Ethyl5 benzene and xylenes (E- Ben., Xyl.), Methyl-ethyl-benzene and trimethyl-benzene (m, e-Ben., TmBen.), Indanes (Ind.), And naphthalenes (Naf.) For various molar ratios from silica to alumina in the catalyst according to a set of modalities;
FIG. 7 is a schematic diagram of a catalytic pyrolysis process with two reactors, according to a --c on j-un-te-de- ^ ned-aLídades -; ------------ ----------------- Figs. 8A-8B are graphs of (A) carbon yield for raw material material of various hydrocarbons and (B) aromatic selectivity for feeds of benzene (Ben.), Toluene (Tol.), Ethylbenzene and xylenes (E-Ben. , Xyl.), Methyl-ethyl-benzene and trimethyl-benzene (m, e-Ben, tmBen) and indanes (Ind.), And naphthalenes (Naf.) For various 20 hydrocarbon raw materials, according to a set of modalities;
FIG. 9 includes a graph of aromatics yield and the amount of energy per unit mass as a function of theoretical yield, according to a set of modalities;
FIG. 10 is a graph of carbon yield of
CO (), aromatics (^), CO ₂ (Δ), and coke (·) as a function of the reactor temperature for a catalytic glucose pyrolysis with ZSM5, according to a set of modalities;
FIG. 11 is a graph of carbon yield of
CO (), aromatics (^), and CO ₂ (Δ) as a function of silica for molar ratio of alumina to a pyrolysis
8/119 glucose catalytic with ZSM-5, according to a set of modalities;
FIG. 12 is a graph plotting the carbon yields of olefins and aromatics as a function of space velocity for a set of modalities;
FIG. 13 is a graph illustrating the carbon yields of various compounds according to a set of modalities;
FIG. 14 is a schematic diagram of a set of modalities in which a fluidized bed reactor is used;
FIG. T5 is a graph plotting the yield of aromatics and olefins for a range of modalities;
Figs. 16A-16B include graphs of olefin and yield and aromatic selectivity, respectively, as a function of spatial speed, according to a set of modalities;
Fig. 17 includes an exemplary table of adjusted pore sizes of Norman rays of zeolite catalysts, according to a set of modalities;
FIG. 18 includes a graph of carbon yield of various reaction products for various experimental operations, according to a set of modalities;
FIG. 19 includes an exemplary schematic diagram of a reactor system configuration;
Figs. 20A-20B include graphs of carbon yield for various reaction products and reaction feed compositions, according to a set of modalities;
Figs. 21A-21B include graphs of (A) carbon yield and (B) aromatic selectivity for various
9/119 reaction products and reaction feed compositions, according to a set of modalities;
Figs. 22A-22F include images WITHOUT copies in catalysts;Figs. 23A-23B include graphics in (A) income carbon and (B) aromatic selectivity for several products of reaction as a occupation of size of particles, of according to a set of modalities;Figs. 24A-24D include images WITHOUT copies in catalysts;FIG. 25 includes a graphic in Yield in
carbon from various reaction products for different catalysts, according to a set of modalities;
Figs. 26A-26C include powder X-ray diffraction (PXRD) standards for various ZSM-5 catalysts, according to a set of modalities;
FIG. 27 includes an exemplary schematic illustration of a catalytic pyrolysis process, according to a set of modalities, and
FIG. 28 includes, according to some modalities, an exemplary aroma yield graph as a function of the recycling ratio.
DETAILED DESCRIPTION
The report describes compositions and methods of the invention for the production of biochemical products, such as biofuel, aromatic compounds and olefins and, more specifically, compositions and methods for the production of biochemical products through catalytic pyrolysis. Some modalities refer to methods for the production of fluid hydrocarbon products (for example, a liquid, a supercritical fluid, and / or a gas), such as aromatic compounds (for example, benzene, toluene, naphthalene, xylene, etc.) and olefins (e.g. ethylene,
10/119 propene, butene, etc.) through catalytic pyrolysis processes (for example, rapid catalytic pyrolysis). In certain embodiments, hydrocarbon products or a portion thereof are liquid at room temperature and standard pressure (SATP - for example, 25 ° C and 100 KPa of absolute pressure). Some of these methods may involve the use of a composition comprising a hydrocarbonaceous mixture, a component hydrocarbonaceous material, for example, a liquid, gaseous and / or solid material, and a heterogeneous pyrolytic catalyst.
In some embodiments, the hydrocarbonaceous material can be fed by a reactor, under catalytic polio, and a portion of the product stream can be recycled to the feedstream comprising the hydrocarbonaceous material. In a particular embodiment, a portion of the olefins in the product stream is selectively recycled to the feed stream. Such modalities can be useful, for example, in increasing the amount of aromatic compounds present in the product stream, in relation to the amount of aromatic compounds that would be present in the product stream in the absence of recycling (for example, olefin recycling), but under essentially identical conditions.
In some embodiments, the mixture can be pyrolyzed at elevated temperatures (for example, between 500 ° C and 1000 ° C). Pyrolysis can be conducted for an amount of time, at least partially sufficient for the discrete production of identifiable fluid hydrocarbon products. Some embodiments involve heating the hydrocarbonaceous mixture at relatively high rates of catalyst and heating material (for example, about 400 ° C per second and about 1000 ° C per second). The methods described here
11/119 may also involve the use of specialized catalysts. For example, in some cases, zeolite catalysts are used; optionally, the catalysts used here can have high molar ratios from silica to alumina. The catalyst may, in some cases, be formed of, or comprise relatively small particles, which can be agglomerated. In some cases, the composition fed to the pyrolysis reactor has a relatively high catalyst for mass ratio of hydrocarbonaceous material (for example, from about 5: 1 to about 20: 1).
--A + gumas — fashion ± dars — can be assigned to a single-stage method for biomass pyrolysis. Such a method may comprise providing or using a single stage pyrolysis apparatus. A single stage pyrolysis apparatus is one in which pyrolysis and subsequent catalytic reactions are carried out in a single vessel. In some embodiments, the single stage pyrolysis apparatus comprises a fluidized bed reactor. Multistage devices can also be used for the production of fluid hydrocarbon products, as described in more detail below.
As used herein, the terms pyrolysis and pyrolize are given their conventional meanings in the art and are used to refer to the transformation of a compound, for example, hydrocarbonaceous material, into one or more other substances, for example, volatile organic compounds, gases and coke, by heating alone without oxidation, which can occur with or without the use of a catalyst. Catalytic pyrolysis refers to pyrolysis performed in the presence of a catalyst, and may involve steps, as described in greater detail below. Examples of catalytic pyrolysis processes are described,
12/119 for example, in Huber, G.W. et al, Synthesis of Transportation Fuels from Biomass: Chemistry, Catalysts, and Engineering, Chem. Rev. 106, (2006), pp 4044-4098, which is incorporated by reference in its entirety.
As used here, the term biomass is given its conventional meaning in the art and is used to refer to any organic source of energy or chemicals that are renewable. Its main components can be (1) trees (wood) and any other vegetation; (2) agricultural products and waste (corn, fruit, silage, etc.); (3) algae and other marine plants; (4) wastes — me t abo 1 i co s ------- (e trume, sewage); and (5) urban cellulosic waste. Examples of biomass materials are described, for example, in Huber, G.W. et al, Synthesis of
Transportation Fuels from Biomass: Chemistry, Catalysts, and Engineering, Chem.
The inventors have discovered within the context of the present invention that for some reactions, some changes in the conditions and combinations of the reaction of such changes can produce favorable products and / or yields, lower yields of coke formation and / or product formation with greater production controlled compounds (by aromatics and / or, for example, olefins related to other fuels) that cannot be obtained otherwise, but for changes in reaction conditions. For example, incorporating a recycling stream comprising olefins can produce a relatively high and / or low amount of aromatic coke compounds in the product stream, relative to an amount of aromatic compounds and / or coke that would be produced in the absence of the recycling chain. As another example, the use of elevated temperatures (for example, in the reactor and / or solid separator) can produce
13/119 products favorable to and / or yields from reactions that may not occur at temperatures inventors have also discovered within lower. Those in the context of the invention that it may be advantageous in some cases to heat the feed stream (for example, a gaseous or liquid hydrocarbonaceous material, a solid hydrocarbonaceous material, a mixture of a solid hydrocarbonaceous material and a solid catalyst, a relatively low As it enters the reactor, the inventors have also found that providing a feed with a high mass ratio of catalyst to hydrocarbonaceous material can produce desirable yields of aromatic products and / or olefins, for example, without wishing to be limited by theory , the inventors currently believe that high heating rates and high feed mass ratios for catalysts can facilitate the introduction of volatile organic compounds, formed from the pyrolysis of the hydrocarbon feed, into the catalyst before they thermally decompose, thus leading high incomes aromatic compounds and / or olefins. Spatial velocities normalized at relatively low mass have also been shown to produce desirable yields of aromatic compounds and / or olefins. In addition, the inventors have found that the relatively long residence times of hydrocarbon material in high temperature components of the reactor and / or the solids separator) can allow adequate time for additional chemical reactions to form desirable products.
The inventors have also discovered within the context of the present invention that with specific properties the use of catalysts can be useful in the formation of
14/119 a relatively large amount of aromatic products and / or olefins. For example, in some cases, the use of catalysts comprising particles of relatively small sizes can result in the production of a relatively high amount of aromatic compounds and / or a relatively low amount of coke. As another example, in certain embodiments, ZSM-5, in combination with certain reaction conditions, has been found to produce, preferably, aromatic compounds and / or olefins. In addition, they found that certain catalysts, which include sites of Bronstead acids and / or well-ordered pore structures, selectively produce aromatic compounds and / or olefins in some cases. Pore size of the catalyst can also be used, in some cases, to affect the quantities and types of product compounds produced.
The modalities described here also involve chemical process designs used to perform catalytic pyrolysis. In some cases, processes may involve the use of one or more fluidized bed reactors (for example, a circulating fluidized bed reactor, turbulent fluidized bed reactor, bubbling fluidized bed reactor, etc.). The process designs described here they can optionally involve specialized handling of the material fed to one or more reactors. For example, in some embodiments, the feed material can be dried, cooled, and / or crushed, before supplying the material to a reactor. Other aspects of the invention relate to product compositions produced using the process designs described herein.
Without being linked to a particular model of action or order of steps in the global thermal / catalytic conversion process, it is believed that catalytic pyrolysis is
15/119 to involve at least partial thermal pyrolysis of the hydrocarbon material (for example, solid biomass, such as cellulose) to produce one or more pyrolysis products (for example, volatile organic compounds, gases, solid coke, etc.) and catalytic reaction of at least a portion of one or more pyrolysis products using a catalyst under sufficient reaction conditions to produce fluid hydrocarbon products. The catalytic reaction may involve volatile organics entering a catalyst (for example, a zeolite catalyst), where they are converted to, for example, hydrocarbons, such as aromatic compounds and carbon monoxide, carbon dioxide, water , and coke. Within or in contact with the catalyst, species derived from biomass can undergo a series of dehydration, decarbonylation, decarboxylation, isomerization, oligomerization, and the dehydrogenation reactions that lead to aromatic products, olefins, CO, CO2 and water. A challenge in the production of selective aromatics and / or olefin is to minimize the formation of coke. For example, global stoichiometries for converting xylitol and glucose to toluene, CO, and H ₂ O are shown in Equations 1 and 2, respectively.
C5O5H12 -> 12/22 CvHg (76%
carbon) + 26/22 CO (24% in income from carbon) + 84/22 H ₂ O (D C ₆ O ₆ H ₁₂ -► 12/22 C ₇ H ₈ (63% of income from carbon) + 48/22 CO (36% in income from carbon) + 84/22 H ₂ O (2)
As shown in these equations, oxygen must be removed from species derived from biomass, such as a combination of CO (or CO ₂ ), and H ₂ O when aromatics are produced. The maximum theoretical yields of toluene
16/119 for xylitol and glucose are 76% and 63%, respectively. FIG.
includes a graph of aromatic yield and the amount of energy per unit mass as a function of theoretical yield, according to a set of modalities. In FIG. 9, the yield axis corresponds to the gallons of aromatic compounds produced by the process per metric ton of biomass feed for the process. The Energy axis corresponds to the amount of energy (calculated using heat of combustion) in aromatic products per metric ton of biomass fed to the process. The figure includes a curve showing the amount of aromatic pnod / utos produced due to the amount of aromatic products calculated from Equation 2, assuming that the biomass fed contains 75% by weight of carbohydrates.
Other challenges associated with the conversion of biomass are the removal of oxygen and the enrichment of the hydrogen content of the hydrocarbon product. A factor commonly referred to as effective hydrogen to molar carbon ratio, H / C _and f f, illustrates the chemistry required for efficient conversion of oxygenates derived from biomass.
Η H-2O
Ç
The molar ratios H / C _eff of glucose, sorbitol and glycerol (all compounds derived from biomass) are 0, 1/3 and 2/3, respectively. By comparison, the H / Ceff molar ratio of petroleum-derived feeds ranges from slightly greater than 2 (for liquid alkanes) to 1 (for benzene). In this respect, biomass can be seen as deficient hydrogen compared to petroleum based raw materials.
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Some of these and other fuel production issues can be addressed using the methods and processes described here. For example, aromatics and / or olefins can be controlled produced from feeds of hydrocarbon material by controlling a variety of process parameters, including, for example: catalyst selection (including type and physical properties (for example, pore size, particle size, existence and degree of agglomeration, particle / agglomerate shape, etc.)), selection of hydrocarbon material, recycling ratios, composition of the recycling stream, heating rates, ~ reaction temperature, catalyst for hydrocarbon mass ratios (eg in the feed stream, reactor, etc.), molar ratios from silica to alumina catalyst, standardized spatial mass velocities, residence times in different processing components, among others . In some embodiments, process parameters can be selected in such a way that coke formation rates are relatively low.
In one aspect, the chemical processes for the reaction of the hydrocarbonaceous material are described. The process may involve, in some embodiments, pyrolizing within a reactor (for example, a fluidized bed reactor) at least a portion of a hydrocarbonaceous material under sufficient reaction conditions to produce one or more pyrolysis products. In addition, the process may catalytically involve reacting at least a portion of one or more pyrolysis products using a catalyst under sufficient reaction conditions to produce one or more fluid hydrocarbon products. In some embodiments, one or more fluid hydrocarbon products can be produced from such
18/119 pyrolysis products by reactions of dehydration, decarbonylation, decarboxylation, isomerization, oligomerization, and dehydrogenation. The pyrolysis and catalytic reaction processes can occur, in some cases, in a single reactor. Chemical processes can be used, in some cases, for the specific production of the hydrocarbon fluid product (for example, aromatic products and / or olefins). In some cases, a portion of the olefins produced by the chemical process can be recycled into a feed stream through which the hydrocarbonaceous material is fed to the reactor (for example, the pyrolysis reactor).
FIG. 1 includes a schematic illustration of an exemplary design of a chemical process used to perform catalytic pyrolysis, according to a set of modalities. In some embodiments, such a process can be used to perform catalytic pyrolysis. As shown in the illustrative embodiment of FIG. 1, a feed stream 10 includes a feed composition comprising hydrocarbonaceous material that will be fed to a reactor 20. The hydrocarbonaceous material can generally comprise carbon and hydrogen, in which carbon is the most abundant component in mass, as well as proportions smaller than other elements, such as oxygen, nitrogen and sulfur. The hydrocarbonaceous material in the feed composition can comprise a solid, liquid, and / or gas. Specific examples of hydrocarbon materials are provided below.
In some embodiments, the feed composition (for example, in feed stream 10 of FIG. 1) comprises a mixture of hydrocarbonaceous material and a catalyst. The mixture can comprise, for example, solids, liquids and / or gases. In certain modalities, the
19/119 The mixture comprises a composition of a solid catalyst and a solid hydrocarbon material. In other embodiments, a catalyst can be provided separately from the feed composition. A variety of catalysts can be used, as described in more detail below. For example, in some cases, zeolite catalysts with molar ratios ranging from silica to alumina and / or varying pore sizes can be used.
One or more olefin compounds can be fed with the hydrocarbonaceous material and / or the cacalisador, —was some modalities. The olef2s can be of any suitable phase (for example, solid, liquid, or gas). Examples of suitable olefin compounds that can be fed with the hydrocarbon material include, but are not limited to, ethylene, propene, butene, ethylene, butene, pentene, hexane, and the like. The olefins and the hydrocarbonaceous material can be fed as part of the same stream as the hydrocarbonaceous material and / or a catalyst, or the olefins can be fed separately from the hydrocarbonaceous material and / or the catalyst. In some embodiments, the olefins may originate from one or more product streams from the reactor and / or a downstream process, such as a separation process (ie, the olefins may be part of a recycling stream ). Olefins can be reacted with the hydrocarbonaceous material in any suitable quantity and proportion. A desired ratio of hydrocarbonaceous material to olefins can be achieved according to the invention, for example, by adjusting the flow rates of hydrocarbonaceous material and olefins, or by pre-mixing the appropriate amounts of hydrocarbonaceous materials and olefins. In some modalities, for
20/119 example, when solid hydrocarbonaceous materials are used, moisture 12 can optionally be removed from the feed composition before being fed to the reactor, for example, by an optional dryer 14. Removal of moisture from the feed stream feeding can be advantageous for several reasons. For example, moisture in the feed stream may require additional energy to enter to heat the feed to a temperature high enough to achieve pyrolysis. Variations in the moisture content of the food can lead to difficulties in controlling the reactor temperature. In addition, removing an umfade from the feed can reduce the need to process water during later processing steps.
In some embodiments, the feed composition may be dried until the feed composition comprises less than about 10%, less than about 5%, less than about 2%, or less than about 1% water by weight . Suitable equipment capable of removing water from the feed composition is known to a person skilled in the art. As an example, in a set of embodiments, the dryer comprises an oven heated to a particular temperature (for example, at least about 80 ° C, at least about 100 ° C, at least about 150 ° C, or higher ) through which the feed composition is continuously, semi-continuously or periodically passed. In some cases, the dryer may comprise a vacuum chamber in which the composition of the feed is processed as a batch. Other modes of the dryer may combine high temperatures with vacuum operation. The dryer can be integrally connected to the reactor or it can be supplied as a separate unit from the reactor.
11/21
In some cases, the particle size of the feed composition can be reduced in an optional milling system 16, before passing the feed to the reactor. In some embodiments, the average diameter of the crushed feed composition exiting the milling system can comprise no more than about 50%, no more than about 25%, no more than about 10%, no more than about 5%, no more than about 2% of the average diameter of the feed composition fed into the grinding system. The large particle feed material can be more easily transported and less-cluttered-than-small particle-adding material. On the other hand, in some cases it may be advantageous to feed small particles to the reactor (as discussed below). The use of a grinding system allows the transport of large particles feed between the source and the process, while allowing the feeding of small particles to the reactor.
Suitable equipment capable of crushing the feed composition is known to those skilled in the art. For example, the grinding system can comprise an industrial mill (for example, hammer mill, ball mill, etc.), a unit with blades (for example, chipper, crusher, etc.), or any other suitable type of system crushing process. In some embodiments, the shredding system may comprise a cooling system (for example, an active cooling system, such as a pumped fluid heat exchanger, a passive cooling system, such as one including fins, etc.), which can be used to maintain the feed composition at relatively low temperatures (e.g., room temperature) before introducing the feed composition to the reactor. O
22/119 grinding system can be integrally connected to the reactor or can be supplied as a separate reactor unit. While the grinding step is shown following the drying step in FIG. 1, the order of these operations can be reversed in some modalities. In still other modalities, the drying and crushing steps can be achieved using an integrated unit.
In some cases, crushing and cooling of the hydrocarbonaceous material can be achieved using separate units. Cooling of the hydrocarbonaceous material may be desirable, for example, to redirect undesirable distribution of the feed material before passing it through the reactor. In a number of embodiments, the hydrocarbonaceous material can be passed to a crushing system to produce a crushed hydrocarbonaceous material. The crushed hydrocarbon material can then be passed from the crushing system to a cooling and cooling system. The hydrocarbonaceous material can be cooled to a temperature less than about 300 ° C, less than about 200 ° C, less than about 100 ° C, less than about 75 ° C, less than about 50 ° C, less than about 35 ° C, or less than about 20 ° C before introducing the hydrocarbonaceous material to the reactor. In modalities that include the use of a cooling system, the cooling system includes an active cooling unit (for example, a heat exchanger) capable of lowering the temperature of the biomass. In some embodiments, two or more of the dryers, grinding system, and cooling system can be combined into a single unit. The cooling system can, in some modalities, be integrated directly with one or more reactors.
11/23
As illustrated in FIG. 1, the feed composition can be transferred to reactor 20. The reactor can be used, in some cases, to perform the catalytic pyrolysis of hydrocarbonaceous material. In the illustrative embodiment of FIG. 1, the reactor comprises any suitable reactor known to those skilled in the art. For example, in some cases, the reactor may comprise a continuously agitated tank reactor (CSTR), a batch reactor, a semi-batch reactor, or a fixed bed catalytic reactor, among others. In some cases, the reactor comprises a fluidized bed reactor, for example, a r and a r o r — de — 1e ± to — fluidized circulant e. Fluidized 1 μl reactions can, in some cases, provide improved mixing of the catalyst and / or hydrocarbon material during pyrolysis and / or subsequent reactions, which can lead to improved control over the formed reaction products. The use of fluidized bed reactors can also lead to improved heat transfer within the reactor. In addition, improved mixing in a fluidized bed reactor can lead to a reduction in the amount of coke adhered to the catalyst, resulting in deactivation of the reduced catalyst in some cases.
As used herein, the term fluidized bed reactor is given its conventional meaning in the art and is used to refer to reactors comprising a vessel that may contain a solid granular material (e.g., silica particles, catalyst particles, etc.). ), in which a fluid (for example, a gas or liquid) is passed through the granular solid material at speeds high enough to suspend the solid material and make it behave as if it were a fluid. Examples of fluidized bed reactors are described in Kirk-Othmer Encyclopedia
24/119 of Chemical Technology (online), vol. 11, Hoboken, N.J .: Wiley-Interscience, c2001-, pages 791-825, incorporated herein by reference. The term circulating fluidized bed reactor is also given its conventional sense in the art and is 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 to the reactor. Examples of circulating fluidized bed are described in Kirk-Othmer Encyclopedia of Chemical Technology (online), vol. 11, Hoboken, N.J .: WileyTrrterscienceç — c2001-7 — panges “^ 79 ± = ít2IT7 -------- Bubble fluid bed reactors and turbulent fluid bed reactors are also known to those skilled in the art. In bubble fluidized bed reactors, the fluid stream used to fluidize the granular solid material is operated at a sufficiently low flow rate such that bubbles and gaps are observed within the fluidized bed volume during operation. In turbulent fluidized bed reactors, the flow rate of the fluidizing stream is higher than that used in a bubble fluidized bed reactor, and therefore bubbles and gaps are not observed within the fluidized bed volume during operation. Examples of bubble and turbulent fluidized bed reactors are described in KirkOthmer Encyclopedia of Chemical Technology (online), vol. 11, Hoboken, N.J .: Wiley-Interscience, c2001-, pages 791825, incorporated herein by reference.
The reactor (s) can be of any size suitable for carrying out the processes described herein. For example, the reactor can have a volume between 0.1-1 L, 150 L, 50-100 L, 100-250 L, 250-500 L, 500-1000 L, 1000-5000
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L, 5000-10,000 L, or 10,000-50,000 L. In some cases, the reactor has a volume greater than about 1 L, or in other cases, greater than about 10 L, 50 L, 100 L, 250 L , 500 L, 1000 L, or 10,000 L. Reactor volumes greater than 50,000 L are also possible. The reactor can be cylindrical, spherical, or any other suitable shape.
The inventors found that the highest yields of production of the desired product, the lowest yields of coke formation, and / or formation of the most controlled product (for example, greater production of aromatics and / or olefins in relation to other products) they ask — to be — achieved — when — particular combinations of reaction conditions and system components are implemented in methods and systems described here. For example, reaction conditions, such as reactor temperature and / or solids separator, reactor pressure, heating rate of the feed stream, catalyst for mass ratio of the hydrocarbonaceous material, hydrocarbonaceous material to olefin (for example , through a recycling stream) feed rate, standardized spatial velocities of mass, residence time of hydrocarbonaceous material in the reactor, residence time of reaction products in the solids separator, and / or type of catalyst (as well as molar silica to alumina for zeolite catalysts) can be controlled to achieve beneficial results, as described below.
In some embodiments, olefins can be fed (for example, through a recycling stream), in addition to the hydrocarbonaceous material, to a vessel in which the hydrocarbonaceous material is to be reacted (for example, through catalytic pyrolysis). In some cases, co-feeding of olefins with the hydrocarbonaceous material can lead to an increase in the amount of compounds
26/119 aromatics produced by the reaction of the hydrocarbonaceous material. In some embodiments, co-feeding olefins into the reactor can result in an increase in 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 in the absence of the olefin co-feed. Olefins can be reacted with the hydrocarbonaceous material in any suitable ratio. In some embodiments, the ratio of the mass of carbon within the hydrocarbonaceous material to the mass of carbon in the olefins in a mixture of hydrocarbonaceous material and olefins, which must be reacted, is between about 2: 1 and about 20: 1, between about 3: 1 and about 10: 1, or between about 4: 1 and about 5: 1.
The reactor (s) can be operated at any suitable temperature. In some cases, it may be desirable to operate the reactor at relatively high temperatures. For example, the reactor can be operated at temperatures of at least about 300 ° C, at least about 400 ° C, at least about 500 ° C, at least about 600 ° C, at least about 700 ° C at least about 800 ° C, at least about 900 ° C, or at least about 1000 ° C. In some embodiments, the reactor can be operated at temperatures between about 500 ° C and about 1000 ° C, between about 525 ° C and about 800 ° C, between about 550 ° C and about 700 ° C, or between about 575 ° C and about 650 ° C. In other embodiments, the reactor can be operated between about 500 ° C and about 600 ° C. Not wishing to be bound by any theory, relatively high operating temperatures can affect the kinetics of reactions, in such a way that the desired reaction products are formed and / or the formation of the unwanted product is inhibited or reduced. FIG. 10 includes
27/119 a graph of the carbon yield of various products as a function of the reactor temperature for catalytic glucose pyrolysis with ZSM-5 catalyst, in a particular embodiment. Note that in the exemplary embodiment of FIG. 10, the yield of aromatic compounds (indicated by solid triangles) increases with an increase in temperature from 400 ° C to 800 ° C. In addition, the relative amount of coke produced decreases as the temperature is increased from 400 ° C to 800 ° C. In other embodiments, however, lower temperatures can be used.
The reactor (s) can also be operated at an appropriate pressure. In some embodiments, the reactor can be operated at pressures between about 1-4 atm. In some embodiments, the reactor can be operated at a pressure of at least about 1 atm, at least about 2 atm, at least about 3 atm, or at least about 4 atm.
The inventors have found that, in certain embodiments, it is advantageous to heat the feed stream (for example, a gaseous hydrocarbonaceous material, a solid hydrocarbonaceous material, a mixture of a solid hydrocarbonaceous material and any additional olefins and / or a solid catalyst, etc. .), at a relatively quick rate as it enters the reactor. High heating rates can be advantageous for a number of reasons. For example, high heating rates can increase the rate of mass transfer of reagents from bulky solid biomass to catalytic reagent sites. This can, for example, facilitate the introduction of volatile organic compounds formed during the pyrolysis of the hydrocarbonaceous material into the catalyst before completely decomposing the hydrocarbonaceous material into products,
28/119 generally undesirable (for example, coke). In addition, high heating rates can reduce the amount of time that reagents are exposed to temperatures (temperatures between feed temperature and the desired reaction temperature). Prolonged exposure of reagents to intermediate temperatures can lead to the formation of undesirable products through undesirable decomposition and / or reaction paths. Examples of appropriate heating rates to heat the supply stream when entering the reactor from the supply stream include, for example,
example, —Ma ± or what close in Srd ^ C / s, ma ± Or oue about 100 ° C / s, bigger what fence in 200 ° C / s, bigger what fence in 300 ° C / s, bigger what fence in 400 ° C / s, bigger what fence in 500 ° C / s, bigger what fence in 600 ° C / s, bigger what fence in 700 ° C / s, bigger what fence in 800 ° C / s, bigger what fence in 900 ° C / s, bigger what fence < 1000 ° C / s, or greater. In some
embodiments, the supply stream can be heated to a heating rate of between about 500 ° C / s, and about 1000 ° C / s. In some embodiments, the rate of heating to heat the supply current upon entering the reactor can be between about 50 ° C / s, and about 1000 ° C / s, or between about 50 ° C / s and about 400 ° C / s.
In some embodiments, the normalized spatial velocity of the mass of the hydrocarbon material can be selected to selectively produce a desired matrix of fluid hydrocarbon products. As used here, the term normalized space velocity of mass is defined as the mass flow rate of the hydrocarbonaceous material into the reactor (for example, as measured in g / h) divided by the mass of catalyst in the reactor (for example, as measured in g) and has units of inverse time. The mass normalized spatial velocity of the material
29/119 hydrocarbonaceous in a reactor can be calculated using different methods depending on the type of reactor being used. For example, systems employing batch or semi-batch reactors, the hydrocarbonaceous material does not have a normalized spatial velocity of mass. For systems where catalyst is fed and / or extracted from the reactor during the reaction (for example, circulating fluidized bed reactors), the normalized space velocity of mass can be determined by calculating the average amount of catalyst within the reactor volume over a period of operation (for example, steady state operation).
Any suitable normalized space velocity of mass can be used in the modalities described here. In some cases, a normalized spatial speed of mass less than about 10 hours ” ¹ , less than about 5 hours” ¹ , less than about 1 hour ' ¹ , less than about 0.5 hour' ¹ , less than about 0.1 hour ' ¹ , less than about 0.05 hour' ¹ , or less than about 0.01 hour ' ¹ can be employed. In some embodiments, a normalized spatial velocity of mass between about 0.01 hour ' ¹ and about 10 hours' ¹ , between about 0.01 hour ' ¹ and about 5 hours' ¹ , between about 0, 01 hour ' ¹ and about 0.1 hour' ¹ , between about 0.1 hour ” ¹ and about 1 hour” ¹ , or between about 1 hour ' ¹ and about 10 hours' ¹ can be employed. It may also be advantageous, in some modalities, to employ standardized spatial velocities of mass less than about 1 hour ” ¹ , less than about 0.5 hours” ¹ , less than about 0.1 hours' ¹ , less than about 0.05 hour ' ¹ , less than about 0.01 hour' ¹ , between about 0.01 hour ' ¹ and 0.1 hour' ¹ , or between about 0.1 hour ' ¹ and 1 hour' ¹ using a fluidized bed reactor.
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Some modalities include varying the normalized spatial velocity of the mass of the hydrocarbon material to selectively produce different fluid hydrocarbon products. For example, in some embodiments, varying the normalized spatial velocity of mass of the hydrocarbon material can control the relative amounts of aromatic compounds in the reaction product and olefin. For example, relatively low mass standardized spatial velocities can be used to produce a relatively greater amount of aromatics than olefins. Relatively high normalized space velocities of mass can be used to produce a relatively greater amount of olefins than aromatics. In some embodiments, the solid hydrocarbon material is supplied in a fluidized bed reactor at a standardized spatial mass speed of between about 0.1 hour ^-1 and about 10 hours ^-1 to selectively produce olefin compounds, or between about 0.01 hour ^-1 and about 0.1 hour ^-1 to selectively produce aromatic compounds.
In some cases, it is beneficial to control the residence time of the hydrocarbonaceous material (for example, a solid hydrocarbonaceous material) in the reactor and / or under a defined set of reaction conditions (that is, conditions under which the hydrocarbonaceous material can undergo pyrolysis in a supplied reactor system). In continuous flow systems, the residence time of the hydrocarbonaceous material in the reactor is defined as the amount of time that the hydrocarbonaceous material and any reaction products formed therein (excluding products that accumulate in the reactor such as, for example, coke deposited in the catalyst) spent in the reactor. The residence time of the hydrocarbonaceous material in a reactor can
11/31 be calculated using different methods depending on the type of reactor being used.
For example, in modalities where the reactor comprises a packaged bed reactor in which only the hydrocarbonaceous material is continuously fed, no vehicle or fluidization flow of the hydrocarbonaceous material in the reactor as used here can be determined by the volume of the reactor divided by the rate of volumetric flow of product gases leaving the reactor.
In cases where the reaction takes place in a reactor that is closed to the mass flow during the operation (for example, a flask reactor), the ma t ex ± a ± hydrocarbonaceous reaction time in such a reactor is defined as the amount of time that elapses between reactor containing the material the time that the temperature in the hydrocarbon reaches a level sufficient to initiate a pyrolysis reaction (for example, typically about
300 ° C at about
1000 ° C for many materials in typical hydrocarbon feed stock) the time the reactor is cooled instantly (for example, cooled to a temperature below that sufficient to withstand more pyrolysis for example, typically from about 300 ° C to about 1000 ° C for many materials in typical hydrocarbon feed stock).
In some cases, for example, for certain fluidized bed reactors, the supply stream (s) will include supply stream (s) comprising auxiliary materials (i.e. materials other than hydrocarbonaceous and / or olefins).
is,
For example, in certain cases where fluidized beds are used as reactors, the feed stream may comprise cases where circulating fluidized beds are used, fluidizing fluid and
32/119 catalyst can either be fed / recycled to the reactor. In some cases, auxiliary materials may comprise contaminants entrained in the hydrocarbonaceous material. In such cases, the residence time of the hydrocarbonaceous material in the reactor can be determined as the volume of the reactor divided by the volumetric flow rate of the hydrocarbonaceous material and gases from the reaction products exiting the reactor as with the packaged bed situation described above; however, since the flow rate of the hydrocarbonaceous material and the gases of reaction products exiting the reactor may not be convenient to determine directly, the volumetric flow rate of the ~ hydrocarbonaceous material and the gases of reaction products leaving the reactor can be estimated by subtracting the volumetric flow rate of the auxiliary materials (for example, the fluidization fluid, catalyst, contaminants, etc.) in the reactor from the total volumetric flow rate of the gas stream (s) exiting the reactor.
In some embodiments, the residence time of a material (for example, a hydrocarbonaceous material or any other suitable feed material) in the reactor is at least about 2 seconds, at least about 5 seconds, at least about 10 seconds at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds. In some cases, the residence time of a material (for example, a hydrocarbonaceous material or any other suitable feed material) in the reactor is less than about 5 minutes, between about 1 minute and about 4 minutes, or from about 2 seconds to about 480 seconds. Rapid pyrolysis studies have in many cases employed systems with very short feed material
33/119 (for example, hydrocarbonaceous material) residence times (for example, less than 2 seconds). The inventors have found, however, that in some cases, the use of relatively longer residence times allows adequate time for additional chemical reactions to form desirable products. Long residence times can be achieved by, for example, increasing the volume of the reactor and / or reducing the volumetric flow rate of hydrocarbon materials. It should be understood, however, that, in some embodiments described here, the residence time of the feed material (for example, hydrocarbon material may be relatively short, for example, less than about 2 seconds or less than about 1 second.
In certain cases where fluidized bed reactors are used, the feed material (for example, a solid hydrocarbonaceous material) in the reactor can be fluidized by fluidizing a stream of fluid through the reactor. In the exemplary embodiment of FIG. 1, a fluid stream 44 is used to flow the feed material into the reactor 20. The fluid can be supplied to the fluid stream from a fluid source 24 and / or from the reactor product streams through of a compressor 26 (which will be described in more detail below). As used herein, the term fluid means a material, generally in a liquid, supercritical, or gaseous state. Fluids, however, can also contain solids, such as, for example, in suspended or colloidal particles. In some embodiments, it may be advantageous to control the residence time of the fluidizing fluid in the reactor. The residence time of the fluidization fluid is defined as the volume of the reactor divided by the volumetric flow rate of the fluidization fluid. In some cases, the
34/119 fluidization fluid residence can be at least about 5 seconds, at least about 10 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, or at least about 480 seconds. In some cases, the fluidization fluid residence time can be from about 2 seconds to about 480 seconds, from about 5 seconds to about 480 seconds, from about 10 seconds to about 480 seconds , from about 30 seconds to about 480 seconds, from about 60 seconds to about 4ΤΓ0 seconds, about ^- ~ ft2ô seconds - air about 480 seconds, or from about 240 seconds about 480 seconds.
Suitable fluidizing fluids that can be used in the present invention include, for example, inert gases (for example, helium, argon, neon, etc.), hydrogen, nitrogen, carbon monoxide and carbon dioxide, among others.
As shown in the illustrative embodiment of FIG.
1, the products (for example, fluid hydrocarbon products) formed during the reaction of the hydrocarbon material come out of the reactor through a product stream 30. In addition to the reaction products, the product stream may, in some cases, comprise the unreacted hydrocarbon material, fluidization fluid, and / or catalyst. In a set of modalities, the desired reaction product (s) (for example, aromatic liquid hydrocarbons, olefin hydrocarbons, gaseous products, etc.) can be recovered from of a stream of effluent from the reactor.
In some embodiments, at least a portion of the olefins in the fluid hydrocarbon product stream 30
35/119 is separated from the rest of the product stream to produce a recycle stream 100, comprising at least a portion of the separate olefins, and product stream 31A. The separation of olefins from liquid hydrocarbon products can be performed by an olefin recycler 102. While the olefin recycler is shown to be positioned directly downstream of reactor 20 in FIG. 1, it should be understood that the olefin recycler can be positioned at any point downstream of the reactor, and the separation of olefins from other fluid hydrocarbon products can potentially be carried out at any one of a number of points after that fluid hydrocarbon products are produced. In addition, while the recycling stream 100 is illustrated in FIG. 1 as being combined with the feed stream 10 upstream of the dryer 14, it should be understood that the recycling stream 100 can alternatively be combined with the feed stream 10 downstream of the dryer 14 and / or crusher 16, fed directly to the reactor 20, and / or combined with any of the catalyst streams (e.g., 34, 42, 44, 46, and / or 47) described in more detail below.
Suitable methods for separating olefins from other fluid hydrocarbon products are known to those skilled in the art. For example, olefins can be separated from other fluid hydrocarbon products by cooling product stream 30 to a temperature that is between the boiling points of the olefins and the other fluid hydrocarbon products.
Optionally, the olefin recycler 102 may comprise a multi-stage separator. For example, the olefin recycler can comprise a first
36/119 separator that directly separates gaseous products (including olefins) from liquid products (for example, aromatic products with high boiling points, such as benzene, toluene, xylene, etc.), and a second separator that separates at least a portion of the olefins from other gaseous products (for example, aromatic gases, CO ₂ , CO, etc.). The methods and / or conditions used to carry out the separation may depend on the relative amounts and the types of compounds present in the stream of fluid hydrocarbon product, and one of ordinary skill in the art will be able to select a — method — and — the — suitable conditions for achieving a separation provided the guidance provided herein.
In some embodiments, the optional product reactor 104 can be incorporated into the process. The product reactor can be used, for example, to convert one or more of the fluid hydrocarbon products (e.g., olefins, aromatics, etc.) in the product stream 31A to one or more other products (output as stream 31B in the FIG. 1). In some cases, the product reactor may contain a catalyst (for example, a zeolite catalyst), which can be used to perform one or more catalytic reactions. For example, in some embodiments, the product reactor can be used to oligomerize (for example, using a catalyst) one or more olefin products to produce one or more aromatic products. As another example, the product reactor can be used to carry out a carbonylation reaction involving aromatic compounds (for example, carbonylation from ethylene to acrylene through the addition of carbon monoxide and water). One of ordinary skill in the art is able to select the appropriate types of reactors and / or conditions for carrying out such reactions. As a
37/119 optional product 104 is shown located directly downstream of olefin recycler 102 in FIG. 1, it should be understood that the olefin recycler can be positioned at any point downstream of the reactor (for example, a reactor 20 in FIG. 1), and the reaction of olefins, aromatics, or other fluid hydrocarbon products can potentially be performed at any of a variety of points after fluid hydrocarbon products are produced.
As shown in the illustrative embodiment of FIG.
1, product stream 31B (or 31A) can be fed to solids, in some cases, it can be used to separate reaction products from the catalyst (e.g., at least one partially deactivated catalyst) present in the product stream. In addition, the solids separator can be used, in some and / or ashes from the cases, to remove catalyst. In coke some modalities, the solids separator can comprise optional purge current
33, which can be used to purge coke, ash, and / or catalyst from the solids separator.
of solids readily to achieve required equipment and / or decoking steps designed by an ordinary expert the separation can be in the art.
For example, solids can mesh that permeate in a set of modalities, the separator comprise a vessel that comprises a material defines a retaining portion and a portion of
The mesh can serve to retain the retention portion while allowing vessel.
catalyst within which the reaction product passes to the permeate portion. The catalyst can escape from the solids separator through an exhaust on the mesh retaining side,
38/119 while the reaction product may escape into the permeate side of the mesh. Other examples of solids separators and / or decockers are described in more detail in Kirk-Othmer Encyclopedía of Chemical
Technology (Online), Vol. 11, Hoboken, N.J .: WileyInterscience, c2001-, pages 700-734; and C. D. Cooper and F.
C. Alley. Air Pollution Control, A Design Approach. Second Ed. Prospect Heights, Illinois: Waveland Press, Inc. cl 994, pages 127-149, incorporated herein by reference.
The solids separator can be operated at any suitable temperature. In some embodiments, the solids separator can be operated at high temperatures. The inventors found that by certain reactions, the use of elevated temperatures in the solids separator can allow for further reform and / or reaction of the compounds from the reactor. This can allow for increased formation of desirable products. Not wishing to be bound by any theory, the high temperatures in the solids separator can provide enough energy to conduct 20 endothermic reform reactions. The solids separator can be operated at a temperature of, for example, between about 25 ° C and about 200 ° C, between about 200 ° C and about 500 ° C, between about 500 ° C and about 600 ° C, or between about 600 ° C and about 800 ° C. In some cases, the solids separator can be operated at temperatures of at least about 500 ° C, at least about 600 ° C, at least 700 ° C, at least 800 ° C, or greater.
In some cases, it can be beneficial to control the residence time of the catalyst in the 30 solids separator. The residence time of the catalyst in the solids separator is defined as the volume of the solids separator divided by the volumetric flow rate of the catalyst through the solids separator. In some cases, times
39/119 relatively long residues of the catalyst in the solids separator may be desired in order to facilitate removal of sufficient amounts of ash, coke, and / or other undesirable products from the catalyst. In addition, the inventors have found that by employing relatively long residence times of the catalyst in the solids separator, the pyrolysis products can be further reacted to produce desirable products. In some embodiments, the residence time and temperature in the solids separator are selected together in such a way that a product stream of catalyst residues in the solids separator is at least about 1 second, at least about 5 seconds, at least about 10 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, at least about 300 seconds, at least about 600 seconds, or at least at least 1200 seconds. Methods for controlling the residence time of the catalyst in the solids separator are known to those skilled in the art. For example, in some cases, the interior wall of the solids separator may comprise deflectors which serve to restrict the flow of catalyst through the solids separator and / or increase the length of the fluid flow path in the solids separator. Additionally or alternatively, the residence time of the catalyst in the solids separator can be controlled by controlling the flow rate of the catalyst through the solids separator (for example, by controlling the flow rate of the fluidization fluid through the reactor).
solids separator can be any suitable size. For example, the solids separator can have a
40/119 volume between 0.1-lL, 1-50L, 50-100 L, 100-250 L, 250-500 L, 500-1000 L, 1000-5000 L, 5000-10,000L, or 10,000-50,000 L In some cases, the solids separator has a volume greater than about 1 L, or in other cases, greater than about 10 L, 50 L, 100 L, 250 L, 500 L, 1000 L, or 10,000 L Separator volumes of solids greater than 50,000 L are also possible. The solids separator can be cylindrical, spherical, or any other shape and can be circulating or non-circulating. In some embodiments, the solids separator may comprise a vessel or other operation of the unit similar to that used for one or more of the reactors used in the purcxres. The catalyst flow in the solids separator may comprise any suitable geometry. For example, the flow path can be substantially straight. In some cases, the solids separator may comprise a flow channel with a coil, winding, helical, or any other suitable shape. The ratio of the length of the flow path of the solids separator (or, in certain embodiments, the length of the catalyst path through the solids separator) to the average diameter of the solids separator channel can comprise any suitable proportion. In some cases, the ratio can be at least 2: 1, at least 5: 1, at least 10: 1, at least 50: 1, at least 100: 1, or greater.
The parameters outlined above can be used in any suitable combination to produce desirable reaction products (for example, aromatic compounds and / or olefin) and / or favorable yields or particular components. For example, the use of long residence times can be combined with the use of a circulating or turbulent fluidized bed reactor to process solid hydrocarbonaceous material. In some modalities,
41/119 relatively high temperatures (for example, at least 500 ° C) and long residence times (for example, at least about 1 second, at least about 5 seconds, at least about 10 seconds, at least about 30 seconds, at least about 60 seconds, at least about 120 seconds, at least about 240 seconds, at least about 300 seconds, at least about 600 seconds, or at least about 1200 seconds, etc.) can be used in the solids separator after pyrolyzing a solid hydrocarbon material in the reactor. In other embodiments, normalized spatial velocities of relatively — low— mass (per — example ·, —lower — than about 0.1 hour ' ¹ , less than about 0.05 hour' ¹ , less than about 0.01 hour ' ¹ , etc.) can be used to produce a relatively greater amount of aromatics than olefins in a fluidized bed reactor, for example, at least about 6% of aromatics or more. Relatively high standardized space velocities of mass (for example, at least about 0.1 hour ' ¹ , at least about 0.5 hour' ¹ ) can be used to produce a relatively greater amount of olefins than aromatics in a reactor. fluidized bed, for example, at least about 3% by weight, at least about 6% by weight, at least about 10% by weight, at least about 15% by weight, or at least about 20% by weight olefins weight). In another set of embodiments, a solid hydrocarbon material and a zeolite catalyst comprising a large molar ratio of silica to alumina (for example, at least about 30) can be heated in a reactor at a high rate (for example, greater than about 500 ° C / s). In some cases, a catalyst and a solid hydrocarbon material can be fed to a reactor in a mass ratio of at least about 0.5: 1 and heated
42/119 at a temperature of, for example, between 500 ° C and 1000 ° C. In some cases, a catalyst and a solid hydrocarbon material can be fed to a reactor in a mass ratio of at least about 0.5: 1 such that the mixture has a relatively long residence time (for example, at least about seconds). In yet another set of embodiments, a relatively high fluidization residence time (for example, at least about 5 seconds) and a relatively high reactor temperature (for example, between about 500 ° C and about 1000 ° C) can be used.
As mentioned earlier, the solids separator may not be required in all modes. For example, for situations in which fixed bed catalytic reactors are employed, the catalyst can be retained within the reactor, and the reaction products can leave the reactor substantially free of catalyst, thus negating the need to separate the separation step.
In the set of modalities illustrated in FIG. 1, separate catalyst can exit the solids separator via stream 34. In some cases, the catalyst leaving the separator can be at least partially deactivated. The separate catalyst can be fed, in some embodiments, to a regenerator 36, in which any catalyst that was at least partially deactivated can be reactivated. In some embodiments, the regenerator may comprise an optional purge stream 37, which can be used to purge coke, ash, and / or catalyst from the regenerator. Methods for activating the catalyst are well known to those skilled in the art, for example, as described in Kirk-Othmer Encyclopedia of Chemical Technology (Online), Vol. 5,
43/119
Hoboken, N.J.: Wiley-Interscience, c2001-, pages 255-322 incorporated herein by reference.
In a set of embodiments, an oxidizing agent is fed to the regenerator via a stream 38, for example, as shown in FIG. 1. The oxidizing agent can originate from any source, including, for example, an oxygen tank, atmospheric air, steam, among others. In the regenerator, the catalyst is reactivated by reacting the catalyst with the oxidizing agent. In some cases, the deactivated catalyst may comprise residual carbon and / or coke, which can be removed by reacting with the non-regenerating oxidizing agent. The regenerate in FIG. 1 comprises a ventilation stream 40, which may include reaction regeneration products, residual oxidizing agent, etc.
The regenerator can be of any suitable size mentioned above in connection with the reactor or the solids separator. In addition, the regenerator can be operated at elevated temperatures in some cases (for example, 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 can also be controlled using methods known to those skilled in the art, including those described above. In some cases, the mass flow rate of the catalyst through the regenerator will be coupled with the flow rate (s) in the reactor and / or solids separator in order to preserve the mass balance in the system.
As shown in the illustrative embodiment of FIG.
1, the regenerated catalyst can exit the regenerator via stream 42. The regenerated catalyst can be recycled back to the reactor via recycle stream 47. In some cases, the catalyst can be
44/119 lost from the system during operation. In some of these and other cases, additional constitution catalyst can be added to the system via a constitution stream 46. As shown illustratively in FIG. 1, the regenerated and constituting catalyst can be fed to the reactor with the fluidizing fluid through a recycling stream 47, although in other embodiments, the catalyst and fluidizing fluid can be fed to the reactor via separate streams.
Referring again to the separation solids 32 in FIG. 1, the reaction products (for example, fluid hydrocarbon productsf leave the solids separator via stream 48. In some cases, a fraction of stream 48 can be purged via purge stream 60. The contents of the purge stream they can be fed to a combustor or a water-gas displacement reactor, for example, to recover energy that would otherwise be lost from the system. In some cases, the reaction products in current 48 can be fed to a condenser optional 50. The condenser can comprise a heat exchanger, which condenses at least a portion of the reaction product from a gaseous state to a liquid.The condenser can be used to separate reaction products into gaseous, liquid and Condenser operation is well known to those of skill in the art Examples of condensers are described in more detail in Perry's Chemical Engineers' Handbook, Sect ion 11: Heat Transfer Equipment. 8th ed. New York: McGraw-Híll, c2008, incorporated herein by reference.
The condenser may also, in some embodiments, make use of the pressure change to condense portions of the product stream. In FIG. 1, the current 54
45/119 can comprise the liquid fraction of the reaction products (for example, water, aromatics, olefin compounds, etc.), and stream 74 can comprise the gas fraction of the reaction products (for example, CO, C0 ₂ , H ₂ , etc.). In some embodiments, the gas fraction can be fed to a vapor recovery system 70. The vapor recovery system can be used, for example, to recover any desirable vapors within stream 74 and transport them through stream 72.
In addition, stream 76 can be used to transport CO, CO ₂ , and / or other non-recoverable gases from the vapor recovery system. Should you choose ^- some modalities, the optional vapor recovery system can be placed in other locations. For example, in some embodiments, a vapor recovery system can be positioned downstream of the purge stream 54. A person skilled in the art can select a suitable placement for a vapor recovery system.
Other products (for example, excess gas) 20 can be transported to the optional compressor 26 via chain 56, where they can be compressed and used as fluidizing gas in the reactor (chain 22) and / or where they can assist in the transport of material hydrocarbonaceous reactor (currents 58). In some 25 cases, the liquid fraction can be further processed, for example, to separate the aqueous phase from the organic phase, to separate the individual compounds, etc.
It should be understood that, while the set of modalities described by FIG. 1 includes a reactor, 30 solids separator, regenerator, condenser, etc., not all modalities will involve the use of these elements. For example, in some embodiments, the supply current can be fed to a bed reactor
46/119 catalytic fixed, reacted, and reaction products can be collected directly from the reactor and cooled without the use of a dedicated condenser. In some cases, while a dryer, grinding system, solids separator, regenerator, condenser and / or compressor can be used as part of the process, one or more of these elements may comprise separate units not fluidly 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 missing. In some modalities,
--- product (s) --- da --- reaction - desired (s) --- (for - exempiov liquid aromatic hydrocarbons, olefin hydrocarbons, gaseous products, etc.) can be recovered at any time point in the production process (for example, after passing through the reactor, after separation, after condensation, etc.).
In some embodiments, a process of the invention may involve the use of more than one reactor. For example, multiple reactors can be connected in fluid communication with each other, for example, to operate in series and / or in parallel, as shown in the exemplary embodiment of FIG. 7. In some embodiments, the process may comprise providing a hydrocarbonaceous material in a first reactor and pyrolyzing, from the first reactor, at least a portion of the hydrocarbonaceous material under sufficient reaction conditions to produce one or more pyrolysis products. In some embodiments, a catalyst can be supplied to the first reactor, and at least a portion of one or more
pyrolysis in first reactor are catalytically reacted using the catalyst under conditions in reaction enough to product one or more products of
47/119 fluid hydrocarbons. The process may further comprise catalytically reacting at least a portion of one or more pyrolysis products in a second reactor using a catalyst under reaction conditions sufficient to produce one or more fluid hydrocarbon products. In some cases, after catalytically reacting at least a portion of one or more pyrolysis products in the second reactor, the process may further comprise a step of reacting within the second reactor at least a portion of one or more fluid hydrocarbon products from the first reactor to produce one or more other hydrouarbonetry products ^
In FIG. 7, the reaction product of reactor 20 is transported to a second reactor 20 '. Those skilled in the art are familiar with the use of multiple reactor systems for the pyrolysis of organic material to produce organic products, and such systems are known in the art. While FIG. 7 illustrates a set of modalities, in which the reactors are in fluid communication with each other, in some cases, the two reactors cannot be in fluid communication. For example, a first reactor can be used for the first reaction which can be used to produce a product from a transported to a separate location for the reaction in a second reactor.
In some cases, a composition comprising the hydrocarbonaceous material (with or without a catalyst) can be heated in a first reactor, and the hydrocarbonaceous material can be a portion of at least one pyrolyzed to produce a pyrolysis product (and optionally at least a partially catalyst) disabled). The first pyrolysis product can be in the form of a liquid and / or a gas. The composition comprising the first pyrolysis product can then be heated in a second reactor, which can be
48/119 or cannot be in fluid communication with the first reactor. After the heating step in the second reactor, a second pyrolysis product from the second reactor can be collected. The second pyrolysis product can be in the form of a liquid and / or a gas. In some cases, the composition comprising hydrocarbonaceous material that is fed into the first reactor can comprise, for example, a mixture of a solid hydrocarbonaceous material and a solid catalyst. The first product produced by pyrolysis from the first reactor can be different in its chemical composition, quantity, state (for example: fTu ± do ~ vst a ~ gáraj do ~ that ~ o ^ second ~ product ^- ^ of pyrolysis. For example, the first pyrolysis product can substantially include a liquid, while the second pyrolysis product can substantially include a gas. In another example, the first pyrolysis product includes a fluid product (for example, a bio-oil, sugar) , and the second pyrolysis product comprises a relatively greater amount of aromatics than the first pyrolysis product. In some cases, the first pyrolysis product includes a fluid product (for example, including aromatics), and the second pyrolysis product comprises a relatively greater amount of olefins than the first pyrolysis product. In yet another example, the first pyrolysis product includes a fluid product (for example, a bio-oil, sugar), and the second pyrolysis product comprises a relatively greater amount of oxygenated aromatic compounds than the first pyrolysis product.
One or more of the reactors in a multiple reactor configuration may comprise a fluidized bed reactor (for example, a circulating fluidized bed reactor, a turbulent fluidized bed reactor, etc.)
49/119 or, in other cases, any other type of reactor (for example, any of the above mentioned reactors). For example, in a set of embodiments, the first reactor comprises a circulating fluidized bed reactor or a turbulent fluidized bed reactor, and the second reactor comprises a circulating fluidized bed reactor or a turbulent fluidized bed reactor in fluid communication with the first reactor. In addition, the multiple reactor configuration can include any of the additional processing steps and / or equipment mentioned here (for example, a solids separator, a regenerator, a uondeiTHadryrç etcrj The reactors and% or additional processing equipment can be operated using any of the processing parameters (for example, temperatures, residence times, etc.) mentioned here.
Hydrocarbonaceous material useful in the context of the present invention may comprise, for example, a component such as xylitol, glucose (for example, α-D-glucose, β-D-glucose), cellobiosis, cellulose, hemicellulose, lignin, sugarcane bagasse, glucose, wood, and corn straw together with their pyrolysis products and combinations of such components and / or their pyrolysis products. Other examples of hydrocarbon materials include, for example, plastic waste, recycled plastics, municipal and agricultural solid waste, food waste, animal waste, carbohydrates, Iignocellulosic materials (for example, wood chips or pellets, Iignocellulosic biomass, etc. .), or their combinations, among others. Figs. 8A and 8B include product distribution charts for various hydrocarbon feeds including cane bagasse, glucose, wood and corn straw. In the modalities illustrated in Fig. 8A, all raw materials for the
50/119 tests produced with relatively high aromatic yields (for example, greater than 20% carbon yields (equivalent to yields greater than about 8%)). Carbon yields greater than 40% 5 (equivalent to yields greater than about 18.5%) were produced using a glucose feed in the set of modalities. FIG. 8B includes a graph of aromatic selectivity for various hydrocarbonaceous materials as a raw material. The aromatic species 10 included in FIG. 8B are benzene, toluene, ethyl-benzene and xylenes, methyl-ethyl-benzene and trimethyl-benzene, indans, and naphtalenes.
As demonstrated here, the choice of hydrocarbon materials and catalyst can be used to vary the resulting composition of the fluid hydrocarbon product. For example, a wide range of hydrocarbon materials (eg, without limitation, glucose, cellulose, cellobiose, xylitol, etc.) can be used for the production of naphthalenes. In another example, certain 20 hydrocarbon materials (for example, cellulose) can be used for the selective production of toluene. Alternatively, without limitation, where a hydrocarbonaceous material comprises glucose, adjustment of a mass ratio of the catalyst to glucose of the feed composition can be used to vary the production of identifiable oxygenated compounds (e.g., oxygenated aromatic compounds). The catalyst for mass ratio of glucose in the feed composition can be adjusted by increasing or decreasing the amount of catalyst fed to the reactor in relation to the amount of glucose fed to the reactor. Some of these and other compounds can be isolated as special chemicals for further reaction or incorporated into
51/119 subsequent biofuel. In certain other embodiments, a hydrocarbonaceous material may comprise a lignin pyrolysis product such as, for example, benzyl phenyl ether. Pyrolysis compounds of this and others can be used to produce a range of aromatic compounds for use as fuel or basic chemical additives. Regardless of the initial hydrocarbon materials or resulting pyrolysis products, the processes described herein can optionally include hydrogenation of various unsaturated or aromatic compounds to produce tri-hydrogenation products that can be used as incorporated in the production of biofuel.
As described above, the hydrocarbonaceous material in the feed composition can comprise a solid, liquid and / or gas. In cases where the hydrocarbonaceous material includes solids, the solids can be of any suitable size. In some cases, it may be advantageous to use hydrocarbon solids with relatively small particle sizes. Small particles of solids can, in some cases, react more quickly than larger solids due to their relatively larger surface area for volume rates compared to larger solids. In addition, small particle sizes can allow for more efficient heat transfer within each particle and / or within the reactor volume. This can prevent or reduce the formation of unwanted reaction products. In addition, small particle sizes can predict increased solid-gas and increased solid-solid contact, leading to improved heating and mass transfer. In some embodiments, the average size of the solid hydrocarbonaceous material is less than about 5 mm, less than about 2 mm,
52/119 less than about 1 mm, less than about 500 microns, less than about 60 mesh (250 microns), less than about 100 mesh (149 microns), less than about 140 mesh (105 microns), less than about 170 mesh (88 microns), less than about 200 mesh (74 microns), less than about 270 mesh (53 microns) or less than about 400 mesh (37 microns) or less.
In some cases, it may be desirable to employ feed material with an average particle size above a minimum value in order to reduce the pressure required to pass the hydrocarbonaceous material through the reactor. For example, in some cases, it is desirable to use solid hydrocarbon material with an average particle size of at least about at least about 270 mesh (53 microns), at least about at least about 170 microns), at least about less than about
100 mesh (149 microns), at least about 60 mesh (250 microns), at least about 500 microns, at least about 1 mm, at least about 2 mm, at least about 5 mm, or greater.
Useful components of the catalyst in the context of the present invention can be selected from any catalyst known in the art, or as would be understood by those skilled in the art aware of the present invention. Functionally, catalysts may be limited only by the ability of any material such to promote and / or perform dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization, decarbonylation, decarboxylation, aldolic condensation and / or any other reaction or process associated with or related to with the pyrolysis of a material of
53/119 hydrocarbon. Components of the catalyst can be considered acidic, neutral or basic, as would be understood by those skilled in the art.
The catalyst particles described herein can comprise polycrystalline solids (for example, polycrystalline particles) in some cases. The catalyst particles can also comprise a single crystal, in some embodiments. In certain cases, particles can be separate and distinct physical objects that are autonomous. In other cases, the particles may, at least at certain points in their preparation and / or use, comprise an ag + cmermer — of — a pturaiztidade — of p ^ nrt ± raw ± ars ÍHd ± v ± ~ dua ± s ~ em intimate contact with each other.
A catalyst used in embodiments described herein (for example, in the supply stream, in the reactor, etc.) can be of any suitable size. In some cases, it may be advantageous to use catalysts comprising relatively small particles of catalyst, which may, as mentioned above, in certain embodiments, be in the form of larger catalyst objects which can be comprised of a plurality of agglomerated catalyst particles. In some embodiments, for example, the use of small catalyst particles can increase the extent to which the hydrocarbonaceous material can contact the catalyst surface sites due to, for example, the increased catalytic outer surface area and decreased diffusion distances across of the catalyst. In some cases, the size of the catalyst and / or particle size of the catalyst can be chosen based at least in part, for example, the type of fluid flow desired and the life span of the catalyst.
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In some embodiments, the average diameter (as measured by conventional sieve analysis) of catalyst objects, which may, in certain cases, each comprise a single particle of catalyst or, in other cases, comprise a cluster of a plurality of particles , can be less than about 5 mm, less than about 2 mm, less than about 1 mm, less than about 500 microns, less than about 60 mesh (250 microns), less than about 100 mesh (149 microns), less than about 140 mesh (105 microns), less than about 170 mesh (88 microns), less than about 200 mesh (74 microns), —lower — than about 270 me ~ sh (“53 nm ± CTO “ns”) u or less than about 400 mesh (37 microns), or less.
In some embodiments, the catalyst objects may be or be formed of particles having a maximum transverse dimension 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. As noted earlier, in certain cases, the catalyst particles having the dimensions within the ranges noted immediately above can be agglomerated to form discrete catalyst objects with dimensions within the ranges noted in the previous paragraph. As used here, the maximum cross-sectional dimension of a particle refers to the largest dimension between two boundaries of a particle. A person skilled in the art would be able to measure the particle cross-sectional dimension, for example, by analyzing an electron scanning micrograph (SEM) of a catalyst preparation. In modalities comprising agglomerated particles, the
55/119 particles should be considered separately when determining maximum cross-sectional dimensions.
In such a case, the measurement would be carried out by establishing imaginary limits between each of the agglomerated particles, and measuring the maximum cross-sectional dimension of the hypothetical, individualized particles that result from the establishment of such limits. In some embodiments, a relatively large number of particles within a catalyst have, at most, cross-sectional dimensions that are within a certain range.
For example, in some fashion, 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 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 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 catalyst volume can be occupied by particles with maximum cross-sectional dimensions within a specific range, in some cases. For example, in some modalities, 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 all catalyst volumes used is occupied by particles having maximum cross-sectional dimensions less than about 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
56/119 microns, between about 100 nm and about 1 micron, or between about 500 nm and about 1 micron.
In some embodiments, the particles within a catalyst can be substantially the same size. For example, the catalyst can comprise particles with a size distribution, such that the standard deviation of the maximum particle cross-sectional dimensions is not greater than about 50%, not greater than about 25%, not greater than about 10%, is not greater than about 5%, is not greater than about 2%, or is not greater than about 1% of the maximum mean cross-sectional dimensions of the particles. Standard deviation (sigma lower case) and claniclot are their normal meaning in the technique, and can be calculated as:
where Di is the maximum cross-sectional dimension of particle i, D _avg is the average of the maximum cross-sectional dimensions of all particles, and n is the number of particles within the catalyst. The percentage comparisons between the standard deviation and the maximum mean cross-sectional dimensions of the particles described above can be obtained by dividing the standard deviation by the mean and multiplying by 100%.
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 (for example, any of the
57/119 size distributions described 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%, over a 5 amount of aromatic compounds that would be produced using catalysts containing particles with a size distribution outside the desired range (for example, with a large percentage of particles larger than 1 micron, larger than 5 microns, etc.)
Alternatively, alone or in conjunction with the considerations mentioned above, the catalysts can be selected according to pore size (poir-exemplcr; pore sizes and mesoporous typically associated with zeolites), for example, average pore sizes less than 15 about 100 Angstroms, less than about 50 Angstroms, less than about 20 Angstroms, less than about 10 Angstroms, less than about 5 Angstroms, or less. In some embodiments, catalysts with average pore sizes of about 5 Angstroms to about 100 Angstroms 20 can be used. In some embodiments, catalysts with average pore sizes between about 5.5 Angstroms and about 6.5 Angstroms, or between about 5.9 Angstroms and about 6.3 Angstroms can be used. In some cases, catalysts with 25 average pore sizes between about 7 Angstroms and about 8
Angstroms, or between about 7.2 Angstroms and about 7.8
Angstroms can be used.
As used here, the term used for transverse pore size is to refer to that of a pore. The smallest cross-sectional diameter The smallest cross-sectional diameter of a pore can correspond to the smallest cross-sectional dimension (for example, a cross-sectional diameter), as measured perpendicular to the
58/119 pore length. In some embodiments, a catalyst with an average pore size or pore size distribution of X refers to a catalyst, where the average of the smallest pore cross-sectional diameters within the catalyst is about X. It should be It is understood that pore size or smallest cross-sectional diameter of a pore as used herein refers to the adjusted pore sizes of Norman rays well known to those skilled in the art. Determination of the adjusted pore sizes of Norman rays is described, for example, in Cook, M .; Conner, W. C, How big are the pores
Conference, 12th, Baltimore, July 5-10, 1998; (1999), 1, pp 409-414, which is incorporated herein by reference in its entirety. A list of adjusted pore sizes of exemplary Norman rays are shown, for example, in Fig. 17. As an exemplary specific calculation, the atomic rays for ZSM-5 pores are about 5.5-5.6 Angstroms, measured by X-ray diffraction. In order to adjust for the repulsion effects between the oxygen atoms in the catalyst, Cook and Conner showed that the adjusted Norman radius is 0.7 Angstroms greater than the atomic radius (about 6.2-6.3 Angstroms).
One of ordinary skill in the art will understand how to determine the pore size (e.g., minimum pore size, average minimum pore sizes) in a catalyst. For example, X-ray diffraction (XRD) can be used to determine atomic coordinates. XRD techniques for determining 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
59/119 reference in its entirety. Other techniques that may be useful in determining pore sizes (for example, 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, July lrst, 1993, pp. 185-195, which is incorporated herein by reference in its entirety. Pore sizes of mesoporous catalysts can be determined using, for example, gas adsorption techniques, as described in Gregg, SJ et al, Adsorpttonç Surface Area and Porosity, 2τπ3 “^ ± γ, Academic Press Inc., New York, 1982 and Rouquerol, F. et al, Adsorption by powders and porous materials. Principies, Methodology and Applications, Academic Press Inc., New York, 1998, both of which are 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 your adjusted Norman radius pore sizes.
In some embodiments, a screening method is used to select catalysts with appropriate pore sizes for the conversion of specific molecules from pyrolysis products. The screening method may comprise determining the size of the desired pyrolysis product molecules to be catalytically reacted (for example, the kinetic molecule 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 can then be chosen in such a way that the pores of the catalyst (for example, adjusted minimum Norman radii) are large enough to allow the molecules of
60/119 pyrolysis products can diffuse in and / or react with the catalyst. In some embodiments, catalysts are chosen in such a way that their pore sizes are small enough to prevent the entry and / or reaction of pyrolysis products, the reaction of which would be undesirable.
Without limitation, some of these other catalysts can be selected from naturally occurring zeolites, synthetic zeolites and their combinations. In certain embodiments, the catalyst may be a zeolite catalyst of the Mordenite Framework type
Inverted
Modernite Inverted Structure (MFI), such as a ZSM-5 zeolite talisher, as would be understood by those skilled in the art. Optionally, such a catalyst can comprise acidic sites. Other types of zeolite catalysts include ferrierite, zeolite Y, beta zeolite, modernite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S) A1PO-31, SSZ- 23, among others. In other embodiments, non-zeolite catalysts can be used. For example, WO _x / ZrO ₂ , aluminum phosphates, etc.
In some embodiments, the catalyst may comprise a metal and / or a metal oxide.
Suitable metals and / or oxides include, for example, nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper, chromium, gallium, and / or any of its oxides, among others. The metal and / or metal oxide can be impregnated in the catalyst (for example, in the interstices of the crystalline structure of the catalyst), in some embodiments. The metal and / or metal oxide can be incorporated into the lattice structure of the catalyst. For example, metal and / or metal oxide may be included during the preparation of the catalyst, and the metal and / or metal oxide may occupy a latex site of the resulting catalyst (for example, a zeolite catalyst). As another example, metal and / or oxide
61/119 metal can react or otherwise interact with a zeolite so that the metal and / or metal oxide displaces an atom within the lattice structure of the zeolite.
In certain embodiments, Mordenite Framework Inverted - Modernite Inverted Structure (MFI) zeolite catalyst comprising gallium can be used. For example, a MFI cock-aluminosilicate zeolite catalyst (GaAlMFI) can be used. One of ordinary skill in the art would be familiar with GaAlMFI zeolites, which can be thought of as MFI aluminosilicate zeolites in which some of the Al atoms have been replaced with Gat atoms In some cases, the zeotitre catalyst may be in the form of hydrogen ( for example, Η-GaAlMFI). The MFI galloaluminosilicate catalyst can be a ZSM-5 zeolite catalyst in which some of the aluminum atoms have been replaced by gallium atoms, in some embodiments.
In some cases, the proportion of moles of Si in the gallaluminosilicate zeolite catalyst to the sum of moles of Ga and Al (that is, the molar ratio expressed as Si: (Ga + Al)) in the gallaluminosilicate zeolite catalyst can be at least about 15: 1, at least about 20: 1, at least about 25: 1, at least about 35: 1, at least about 50: 1, at least about 75: 1, or greater. In some embodiments, it may be advantageous to employ a catalyst with a mole ratio of Si in the zeolite to the sum of moles of Ga and Al of between about 15: 1 and about 100: 1, from about 15: 1 to about 75: 1, between about 25: 1 and about 80: 1, or between about 50: 1 and about 75: 1. In some cases, the ratio of Si moles in the gallaluminosilicate zeolite catalyst to the Ga moles in the gallaluminosilicate zeolite catalyst can be at least about 30: 1, at least about 60: 1, at least about 120: 1 at least about 200: 1, between about
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30: 1 and about 300: 1, between about 30: 1 and about 200: 1, between about 30: 1 and about 120: 1, or between about 30: 1 and about 75: 1. The ratio of Si moles in the gallaluminosilicate zeolite catalyst to Al moles in the gallaluminosilicate zeolite catalyst can be at least about 10: 1, at least about 20: 1, at least about 30: 1, at least about 40: 1, at least about 50: 1, at least about 75: 1, between about 10: 1 and about 100: 1, between about 10: 1 and about 75: 1, between about 10: 1 to about 50: 1, between about 10: 1 to about 40: 1, or between about 10: 1 to about 30: 1.
In addition, in some cases, the properties of the catalysts (for example, the pore structure, type and / or the number of acidic sites, etc.) can 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 pore size distribution. In some cases, a single catalyst with a bimodal pore size distribution may be used (for example, a single catalyst that contains predominantly 5.9-6.3 Angstrom pores and 7-8 Angstrom pores). In other cases, a mixture of two or more catalysts can be used to establish the bimodal distribution (for example, a mixture of two catalysts, each type of catalyst including a distinct range of the 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 (for example, a mesoporous catalyst, a metal oxide catalyst, etc.).
For example, in some modalities, at least about 70%, at least about 80%, at least about
63/119
90%, at least about 95%, at least about 98%, or at least about 99% of the pores of one or more catalysts (for example, a zeolite catalyst, a mesoporous catalyst, etc.) have smaller diameters of cross sections that are 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 one or more catalysts have smaller cross-section diameters that are within the first size distribution, and at least about 2%, at least about 55 or at least about 10% of the pores of one or more catalysts have smaller cross-section diameters that are 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 an overlapping range is 5.9-6.3 Angstroms and 6.9-8.0 Angstroms, and an example of an overlapping range is 5.96.3 Angstroms and 6.1 -6.5 Angstroms . The first and second size distributions can be selected in such a way that the strips are not immediately adjacent to each other, an example being the pore sizes of 5.9-6.3 Angstroms and 6.9-8.0 Angstroms. An example of a strip that is immediately adjacent to each other is pore sizes of 5.9-6.3 Angstroms and 6.3-6.7 Angstroms.
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 aromatics and olefins. That is, a pore size distribution can
64/119 advantageously produce a relatively high amount of aromatic compounds, and the other pore size distribution can 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 a or more catalysts have smaller cross-sectional diameters between about 5.9 Angstroms and about 6.3 Angstroms, or between about 7 Angstroms and about 8 Angstroms. In addition, at least about 2%, at least about — of — 5 ~% ~, —or at least about 10 ^ ”dOs ^__ p <Tros ~ dre” unrσα plus catalysts have smaller cross-section diameters between about 5.9 Angstroms and about 6.3 Angstroms, and at least about 2%, at least about 5%, or at least about 10% of the pores of one or more catalysts have smaller cross-sectional diameters between about 7 Angstroms and about 8 Angstroms.
In some embodiments, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 98%, or at least about 99% of the pores of a or more catalysts have smaller cross-sectional diameters between about 5.9 Angstroms and about 6.3 Angstroms, or between about 7 Angstroms and about 200 Angstroms. In addition, at least about 2%, at least about 5%, or at least about 10% of the pores of one or more catalysts have smaller cross-section diameters between about 5.9 Angstroms and about 6.3 Angstroms ; and at least about 2%, at least about 5%, or at least about 10% of the pores of one or more catalysts have smaller cross-section diameters between about 7 Angstroms and about 200 Angstroms.
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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 a or more catalysts have smaller cross-section diameters that are within a first distribution and a second distribution, where the first distribution is between about 5.9 Angstroms and about 6.3 Angstroms and the second distribution is different and not overlaps with the first distribution. In some embodiments, the second pore size distribution may be comprised of about 7 Angstroms and about 200 Angstroms, between about 7 Angstroms and about 100 Angstroms, between about 7 Angstroms and about 50 Angstroms, or between about 100 Angstroms and about 200 Angstroms. In some embodiments, the second catalyst can be mesoporous (for example, having a pore size distribution of between about 2 nm and about 50 nm).
In some embodiments, the bimodal pore size distribution may be beneficial in reacting two or more components of hydrocarbonaceous material in the feed. For example, some embodiments comprise providing a solid hydrocarbon material comprising a first component and a second component in a reactor, where the first and second components are different. Examples of compounds that can be used as the first or second components include any of the hydrocarbonaceous materials described herein (for example, sugarcane bagasse, glucose, wood, corn straw, cellulose, hemicellulose, lignin, or any others). For example, the first component may comprise cellulose, hemicellulose and lignin, and the second component comprises cellulose, hemicellulose and lignin. The method can also
66/119 understand to supply first and second catalysts in the reactor. In some embodiments, the first catalyst may have a first pore size distribution and a second catalyst may have a second pore size distribution, in which the first and second pore size distributions are different and do not overlap. The first pore size distribution can be, for example, between about 5.9 Angstroms and about 6.3 Angstroms. The second pore size distribution can be, for example, between about 7 Angstroms and about 200 Angstroms, between about 7 Angstroms and about 100 angstroms, between about 7 Angsüwms and about ~ 50 Angstroms, or between about 100 Angstroms and about 200 Angstroms. In some cases, the second catalyst can be mesoporous or non-porous.
The first catalyst can be selective to catalytically react the first component or a derivative thereof to produce a fluid hydrocarbon product. In addition, the second catalyst can be selective to catalytically react 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 hydrocarbon 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 one or more more hydrocarbon products. In some cases, at least a partially deactivated catalyst can also be produced.
In certain embodiments, a method used in combination with embodiments described herein includes increasing the mass ratio of the catalyst to the hydrocarbonaceous material of a composition to increase production
67/119 of identifiable aromatic compounds. As illustrated here, representing, but with a distinction over certain prior catalytic pyrolysis methods, articles and methods described herein can be used to produce aromatic, discrete and identifiable biofuel compounds selected from, but not limited to, benzene, toluene, propylbenzene , ethylbenzene, methylbenzene methylethylbenzene, trimethylbenzene, xylenes, indans, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene, and dimethylhydrindene and their combinations.
In some embodiments, the chemistry of a catalyst reaction can be affected by the addition of one or more additional compounds.
For example, adding a metal to a catalyst can result in a change in the selective formation of specific compounds (for example, adding metal to alumina-silicate catalysts can result in the production of more CO). In addition, when the fluidization fluid comprises hydrogen, the amount of coke formed on the catalyst can be decreased.
In some embodiments, the catalyst can comprise both silica and alumina (for example, a zeolite catalyst). The silica and alumina in the catalyst can be present in any suitable molar ratio. In some embodiments, it may be advantageous to employ catalysts with a greater number of moles of silica than the number of moles of alumina (i.e., a molar ratio of silica to high alumina). The inventors have unexpectedly discovered that the molar ratios of silica to high alumina, for example, in combination with modalities described herein, can result in the formation of a relatively large amount of aromatic product. For example, in some cases, the feed composition
68/119 can comprise a molar ratio of silica to alumina of at least about
30: 1, at least about 40: 1, at least about 50: 1, at least about 75: 1, at least about 100: 1, at least about 150: 1, or greater.
In some embodiments, it may be advantageous to employ a silica alumina for catalyst with a molar ratio
in between about 30: 1 and fence in 200: 1, of fence fence 150: 1 , in between fence in 50 1 e about in between about 100: 1 and fence in 150: 1. In some modalities , O
30: 1 material
160: 1, or hydrocarbonaceous and catalyst can be present in any suitable ratio. For example, the hydrocarbonaceous material and catalyst may be present in any appropriate mass ratio in cases where the composition of the feed (for example, through one or more feed streams comprising catalyst and the hydrocarbonaceous material or through feed streams separated from hydrocarbonaceous material catalyst), comprises hydrocarbonaceous material catalyst (for example, circulating fluidized bed reactors). As another example, in cases where the hydrocarbon (for example, a batch reactor), the catalyst and the hydrocarbonaceous material can be present in any appropriate mass ratio. In some embodiments involving circulating fluidized bed reactors, the mass ratio of the catalyst to the hydrocarbonaceous material in the feed stream ie in a composition comprising a solid catalyst and a solid hydrocarbonaceous material supplied to a reactor can 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 greater. In some modalities involving
69/119 circulating fluidized bed reactors, the mass ratio of the catalyst to the hydrocarbonaceous material in the feed stream can 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. By employing a relatively high mass ratio of the catalyst to the hydrocarbonaceous material, it can facilitate the introduction of volatile organic compounds, formed from the pyrolysis of the additive material; in the catalyst before its thermal decomposition 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.
In another aspect, a process product is described. In a set of embodiments, a product (e.g., a pyrolysis product) comprises a fluid composition comprising a portion of a reaction product of a solid hydrocarbonaceous material. Such products can be isolated for use as especially chemical products (for example, used as a fuel directly or as additives for high octane fuels) or, alternatively, hydrogenated for use as a biofuel. Products can also be further processed to make other useful compounds.
In some embodiments, the articles and methods described herein are configured to selectively produce aromatic compounds, for example, in a single stage, or alternatively, a multi-stage pyrolysis apparatus.
A fluid hydrocarbon product can comprise, for example, an amount of compounds
Aromatic 70/119, comprising at least about 10% by weight, at least about 15% by weight, at least about 20% by weight, at least about 25% by weight, at least about 30% by weight at least about 35% by weight, at least about 39% by weight, between about 10% by weight and about 40% by weight, between about 10% by weight and about 35% by weight, between about 15% by weight and about 40% by weight, between about 15% by weight and about 35% by weight, between about 20% by weight and about 40% by weight, between about 20% by weight and about 35% by weight, between about 25% by weight and about 40% by weight, between about 25% by weight and about — 35% —in — pestr; —between — about — 3O ~ % —In ~ peSO — and about
40% by weight, or between about 30% by weight and about 35% by weight of the total reaction product of the solid hydrocarbon material. In some cases, such amounts of aromatic compounds have an octane rating greater than or equal to about 90, for example at least 92, 95 or 98. The amount of aromatic compounds comprising a weight percentage of the reaction product Total solid hydrocarbon material is calculated as the weight of aromatic compounds present in the fluid hydrocarbon product, divided by the weight of the hydrocarbon material used in the formation of the pyrolysis products. As used herein, 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 (for example, benzyl, phenyl, etc.) and fused polycyclic aromatic ring (for example, naphthyl,
1,2,3,4-tetrahydro-fertile, etc.). Examples of aromatics include, but are not limited to, benzene, toluene, indane, indene, 2-ethyl toluene, 3-ethyl toluene, 4 ethyl toluene, trimethyl benzene (eg 1,3,5-trimethyl
71/119 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, 1,6dimethylnaphthalene, 2, 5-dimethylnaphthalene, etc.), ethylnaphthalene, hydrindene, methylhydrindene, and dimethylhydrindene Single-ring and / or upper ring aromatic compounds can be produced in some embodiments. Aromatic compounds may have carbon numbers from, for example , C ₅ -Ci4, C ₆ -C ₈ , C ₆ -ei27 ~~ er ^ Ú ^ r ~ Ciõ ^ Ci47
In some embodiments, the articles and methods described herein are configured to selectively produce olefin compounds, for example, in a single-stage, or alternatively, multi-stage pyrolysis apparatus. A fluid composition (for example, liquid and / or gaseous pyrolysis product) can comprise, for example, an amount of olefin compounds that includes at least about 3% by weight, at least about 7% by weight, at least less than about 10% by weight, at least about 12.5% by weight, at least about 15% by weight, at least about 20% by weight or more of the total reaction product of the solid hydrocarbon material. The amount of olefin compounds that comprise a percentage by weight of the total reaction product of the solid hydrocarbon material is calculated as the weight of the olefin compounds present in the fluid hydrocarbon product, divided by the weight of the hydrocarbon material used in the formation of the pyrolysis products . As used here, the terms olefins or olefin compound (known as alkenes) are given their common meaning in the art, and are used to refer to
72/119 any unsaturated hydrocarbon containing one or more pairs of carbon atoms linked by a double bond. Olefins include both cyclic and acyclic (aliphatic) olefins, where the double bond is located between carbon atoms that form part of a cyclical (closed ring) or open chain group, respectively. In addition, olefins can include any suitable number of double bonds (e.g., mono-olefins, diolefins, triolefins, etc.). Examples of olefin compounds include, but are not limited to, ethylene, propene, butene, butadiene, and isoprene, among
carbon from, for example, C ₂ -C ₄ , C ₂ -Cs, C ₄ -Cs, or C ₂ Process conditions can be chosen, in some cases, in such a way that aromatics and / or olefins they are selectively produced, for example, in a single-stage, or alternatively, multi-stage pyrolysis apparatus. For example, in some embodiments, aromatics and / or olefins can be selectively produced when the reactor is operated at a temperature of around 600 ° C (or higher, in some cases). In addition, certain heating rates (for example, at least about 50 ° C / s, or at least about 400 ° C / s), high feed mass ratios for catalyst (for example, at least about 5 : 1), and / or high molar ratios of silica to alumina in the catalyst (for example, at least about 30: 1) can be used to facilitate the selective production of aromatics and / or olefins. Some of these and other process conditions can be combined with a particular type of reactor, such as a fluidized bed reactor (for example,
73/119 example, a circulating fluidized bed reactor), to selectively produce aromatic compounds and / or olefins.
In addition, in some embodiments, the catalyst can be chosen to facilitate the selective production of aromatic products and / or olefins. For example, ZSM-5 may, in some cases, preferably produce relatively higher amounts of aromatic compounds and / or olefins.
In some cases, catalysts that include Bronstead acid sites can facilitate selective production aromatics.
In addition, catalysts with well-ordered pore structures can facilitate the sequential production of aromatic compounds. ~ For example, in some embodiments, catalysts with an average pore diameter between about 5.9 Angstroms and about 6.3 Angstroms they can be particularly useful in the production of aromatic compounds. In addition, catalysts with an average pore diameter between about 7 Angstroms and about 8 Angstroms can be useful in the production of olefins. In some embodiments, a combination of one or more of the above process parameters can be employed to facilitate the selective production of aromatic and / or olefin compounds. The ratio of aromatics to olefins produced can be, for example, between
fence in 0.1: 1 and about 10: 1, between fence in 0.2: 1 and fence in 5: 1, between about 0.5: 1 and about 2: 1, enter fence in 0.1: 1 and about 0.5: 1, between fence in 0.5: 1 and fence in 1: 1, between about 1: 1 and about out of 5: 1, or between fence in 5: 1 and about 10: 1.
In some embodiments, the mass ratio of catalyst to hydrocarbonaceous material is not adjusted to produce favorable yields.
food is desirable products and / or
In some embodiments, oxygenated compounds can be produced, such as, for example,
74/119 example, acetic acid, formic acid, hydroxyacetylaldehyde, furfural, 2-methyl furan, furan, 4-methyl furfural, furan-2-methanol, and levoglucosan, among others. For example, in some cases, increasing the mass ratio of catalyst to hydrocarbonaceous material may result in an increase in the production of non-cyclic oxygenated carbonyl compounds. As a specific example, when the food catalyst (eg glucose) mass ratio in the feed is increased, but kept below a mass ratio of about 9, the relative amount of non-cyclic oxygenated carbonyl products (eg , hydroxyacetaldehyde, acetic acid, etc.) can be increased. In some cases, decreasing the mass ratio of the catalyst to hydrocarbon material can result in an increase in the production of cyclic oxygenated compounds. For example, in some cases, when the catalyst to feed (for example, glucose) mass ratio in the feed is decreased (for example, from about 19 to about 1), the relative amount of furan products, furfural , methyl furan, and / or 4-methyl furfural is increased. In still other embodiments, when the catalyst to feed (eg, glucose) mass ratio in the feed is decreased (for example, from about 19 to about 2.3), the amount of furan-2- product methanol can be increased; and when the catalyst to feed (eg, glucose) mass ratio in the feed is further decreased (for example, from about 2.3 to about 1.5), the amount of the furan-2- product methanol can be decreased. As such, the mass ratio of the catalyst to the hydrocarbon material can be, for example, 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,
75/119 at least about 20: 1, or greater, in some embodiments, or 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 in other modes.
In some embodiments, the process product may also comprise a high octane rating biofuel composition that comprises a pyrolysis product from a biomass hydrocarbon material. The pyrolysis product can be made using a single stage pyrolysis device or, alternatively, a tte-p ± in some cases, σ hydrocarbonaceous material can be mixed with a catalyst (eg a zeolite catalyst) during the reaction of pyrolysis. The composition may include, for example, discrete and identifiable aromatic compounds, one, more than one or each of those compounds characterized by an octane number greater than or equal to about 90, for example at least 92, 95, or 98. As distinguishable over some viscous niches and sludge from the prior art, such a biofuel composition can be characterized as soluble in petroleum-derived gasolines, diesel fuels and / or heating fuels. Such compounds may include, but are not limited to, benzene, toluene, ethylbenzene, methylethylbenzene, trimethylbenzene, xylenes, indanes, naphthalene, methylnaphthalene, dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene, and dimethylhydrindene and their combinations and their quantities and / or their combinations, and / or their combinations, and / or combinations thereof, and / or combinations thereof, and / or combinations thereof, and / or combinations thereof. vary, depending on the choice of biomass composition, type of catalyst, and / or any of the process parameters described here.
76/119
In some embodiments, the product may comprise a non-acid biofuel process compatible with existing gasoline and diesel fuel lines.
In addition, the processes described herein may result in the formation of a lower coke than certain existing methods. For example, in some embodiments, a pyrolysis product can be formed with less than about 30% by weight, less than about 25% by weight, less than about 20% by weight, less than about 15% by weight , or less than about 10% by weight of the pyrolysis product being coquex — The — quantity — of — coke — formed — is^-nred ± da — how — the pose of coke formed in the system divided by the weight of the hydrocarbonaceous material used in the formation of the pyrolysis product.
The following documents are hereby incorporated by reference in their entirety for all purposes: US Provisional Patent Application No. 61/068, 001, filed March 4, 2008, entitled Catalytic Fast Pyrolysis of Solid Biomass and Related Biofuels and Aromatic Compounds , by Huber, et al .; US Provisional Patent Application No. 61/098, 284, filed September 19, 2008, entitled Catalytic Pyrolysis of Solid Biomass and Related Biofuels and Aromatic Compounds, by Huber, et al .; and US Provisional Patent Application No. 12 / 397,303. filed on March 3, 2009, entitled Catalytic Pyrolysis of Solid Biomass and Related Biofuels, Aromatic, and Olefin Compounds, by Huber, et al.
EXAMPLES
The following non-limiting examples and data are intended to illustrate various aspects and relative characteristics of the methods and / or compositions of the present invention, including the selective production of various compounds
77/119 aromatics and / or oxygenates (for example, oxygenated hydrocarbons), as available through the pyrolytic methodologies described here, but does not exemplify the full scope of the invention. In comparison with the prior art, the methods and compositions of the present invention provide results and data that are surprising, unexpected, and contrary to them. Although the utility of the present invention is illustrated through the use of various catalyst materials and hydrocarbon sources it will be understood by those skilled in the art that if comparable results are obtained with various other catalyst materials and sources of carbon dioxide; as being compatible with the scope of the present invention.
EXAMPLE 1
Representing various modalities, the catalytic pyrolysis experiments described in Examples 1-9 below were carried out in a Pyroprobe 2000 batch pyrolysis reactor (CDS Analytical Inc.) with powder and feed catalyst (size <140 mesh). Unless otherwise specified, in this example, the reaction conditions for the experiments were: catalyst to feed weight ratio, 19; catalyst, ZSM5 (SiO2 / AI2O3 = 30); nominal heating rate, 1000 ° C s ¹ ; reaction temperature, 600 ° C; reaction time (feed residence time), 240 s. Figs 2A-2B show the yields of carbon and aromatic selectivity, respectively, for the catalytic pyrolysis of xylitol, cellobiose, glucose and cellulose with HZSM-5 (i.e., protonated ZSM-5). Aromatic yields were calculated as carbon yields. Carbon yield was calculated by dividing the moles of carbon in the product by moles of carbon in the food. Selectivity was calculated as the carbon moles in a given product divided by the moles
78/119 carbon in all products (excluding CO, CO ₂ , and coke (for example, the solid coke remaining on the catalyst)). As can be seen from Fig. 2A, the main products included aromatics, CO, CO ₂ and coke. Xylitol had a higher yield of aromatics than other food products. Xylitol also had a molar ratio (2/5) H / C _eff higher than other food products (0 for cellulose, glucose and cellobiosis). Carbon monoxide and carbon dioxide were generally present as products when aromatics were desired. The aromatic yields of these reactions were about half of the theoretical yields given by equations 1 and 2. Coke yield was about 30% for all tested catalysts, and in an industrial reactor it could be flared to provide heating to the pyrolysis process. catalytic.
It should be noted that an ordinary person skilled in the art will be able to convert between percentages in weight and carbon yield. The amount of carbon in a feed of hydrocarbonaceous material can be determined, for example, through chemical analysis. In addition, the carbon percentage of each of the reaction products can be calculated using molecular formulas. For example, 1 mole of benzene (C ₆ H ₆ ) contains about 72 grams of carbon and about 6 grams of hydrogen, resulting in a weight percentage of about 92.3% carbon. Likewise, methyl benzene contains about 91.5% by weight of carbon, and ethyl benzene and xylene contain about 90.5% by weight of carbon, etc. By dividing the mass of carbon in a particular product stream by the mass of carbon in the feed, percentages of carbon can be determined from percentages by weight.
79/119
In a specific example, toluene can be produced from a wood feed. Chemical analysis can be used to determine that the wood that is fed into the system is 44% carbon by mass (ie 44% carbon in the feed). The toluene produced is 91.25% carbon by mass (ie 91.25% carbon in the product). For a carbon yield (% C) of 5%, the weight percentage can be calculated as:
% by weight = (5% C) * (44%) / (91.25%) = 2.41 percent yield by weight of toluene
For a product mix (for example, berTzerroç txrtuerroç xylene and naphthalene), the sum of the individual product yields gives the total yield.
A person skilled in the art will be able to determine the amount of carbon in the supply stream of a given commercial technology available. The composition of feed or a feed of hydrocarbonaceous material, in terms of percentage of carbon and hydrogen can be determined, for example, by combustion analysis.
In combustion analysis, a feed sample is weighed and subsequently burned in the air (excess air) producing measurable combustion products, such as carbon dioxide and water. The developed carbon dioxide and water can be measured, for example, by trapping and weighing the gas or by gas chromatography. In this example, the moles of carbon dioxide (CO ₂ ) measured would be equivalent to the moles of carbon (C) in the feed sample. In addition, the moles of water (H2O) measured would be equal to 1/2 times the moles of hydrogen sample feed.
When the reactor operates in steady state, the mass leaving the reactor is equal to the mass fed to the
80/119 reactor. In some cases, however, the steady state may not be achieved. For example, material may accumulate (for example, coke) inside the reactor. In order to perform mass balance calculations, the amount of material that accumulates in the reactor must be determined. This can be achieved, for example, by weighing the contents of the reactor before and after the operation.
The aromatic distribution from the catalytic pyrolysis of various representative oxygenated compounds derived from biomass is shown in FIG. 2B. Interestingly, using the methods of the present invention, the tested raw materials produced similar distributions of aromatic products. The engine octane rating (MON) of the aromatics was estimated at 111. (For a complete list of the octane numbers (RON and MON) and boiling points of all quantified aromatics, see the examples below.) Such and others Aromatic products can be used as a fuel directly, as high octane fuel additives, or it can be another additional process to make different compounds. However, the naphthalenes produced had weakly cold flow properties (ie, low volatility) and current regulatory limit levels of gasoline at 25% by volume. To alleviate these concerns, naphthalenes and other aromatic compounds can be hydrogenated to alkanes in a secondary process, to increase their use as fuel additives.
As can be seen from Fig. 3, the product yield for catalytic glucose pyrolysis was a function of the heating rate. The maximum aromatic yield and the lowest coke yield were obtained at a nominal heating rate of 1000 ° C s ^-1 . When the heating rate has been reduced by three orders of magnitude 1 ° C s
81/119, the yield of aromatic compounds has halved, and the yield of coke has increased from 35 to 40%. Thus, it was determined that high heating rates can be used to avoid undesirable reactions of thermal decomposition and coke formation.
In addition, high heating rates, high catalyst mass ratios for biomass can be used advantageously for the production of aromatics. Figs. 4A-4B show the product selectivity for catalytic glucose pyrolysis as a function of the catalyst for the glucose mass ratio. The coke yield increased and the aromatic yield decreased as the mass ratio of the catalyst to glucose decreased. CO and CO ₂ yields also decreased as the mass ratio of the catalyst to glucose decreased. In addition, at mass ratio of catalyst to glucose lower than 19 thermally stable oxygenated compounds were formed. The yield of these oxygenated compounds decreased as the mass ratio of the catalyst to glucose increased. The oxygenated compounds formed included furan, 2-methyl furan, furfural, 4-methylfurfaral, furan-2-methanol, hydroxyacetylaldehyde, and acetic acid, as shown in FIG. 4B. In the largest mass ratio of the catalyst to glucose, the main oxygenated products were hydroxyacetylaldehyde and acetic acid. However, selectivity for furans increased as the mass ratio of the catalyst to glucose decreased. These results indicated that, in addition to aromatic compounds, catalytic pyrolysis can be tuned to form oxygenated compounds, which could be used as special chemicals or fuel precursors.
82/119
Selection of the appropriate catalyst can also be used to selectively produce aromatics.
FIG. 5 compares the carbon yield from the catalytic glucose pyrolysis over several different catalysts. HZSM-5 had the highest aromatic yield of any catalyst tested. When no catalyst was used, the primary product observed was coke. Two catalytic parameters that appear to have an effect on product distribution were the 10-pore structure and the type of acidic sites. The role of acidic sites on catalytic activity, using ZSM-5 ----- representative ^ s ± ± za ± ± t e and SiO ₂ AI2 O3 amorphous caxairisers, was examined. Both silicalite and ZSM-5 have the same pore structure, but silicalite does not contain 15 Bronstead acid sites. Silica-alumina contains sites of Bronstead acids, but does not have a well-ordered pore structure. Silicalite mainly produced coke indicating that the Bronstead acid sites may be useful for the production of aromatics. Silica-alumina also produces mainly coke, indicating that the pore structure of the zeolite can be used to selectively produce aromatics. Also shown in FIG. 5 are β-zeolite and Y-zeolite catalysts, both of which also produce large amounts of coke. The results in FIG. 5 indicate that the method (s) of this invention can be varied by catalyst, the type of active site, and the shape of the pores.
The experiments were carried out using a 2000 pyroprobe analytic pyroliser model (CDS Analytical Inc.). The sample was a resistively heated computer controlled element, which kept an open-ended quartz tube. The powder samples were kept in the
83/119 tube with loose quartz wool packaging; during pyrolysis, fluid vapors from the open ends of the quartz tube into a large cavity (the pyrolysis interface) with a stream of helium vehicle gas. The vehicle gas stream was routed to a model 5890 gas chromatograph interfaced with a Hewlett Packard 5972A selective mass detector model. The pyrolysis interface was maintained at 100 ° C and the temperature of the GC injector used was 275 ° C. Helium was used as inert pyrolysis gas, as well as vehicle gas for the GCMS system. A constant 0.5 ml min ^-1 flow program is ± —used-for — for — the — column — cap ±: GC air. The GC was programmed with the following temperature regime: maintained at 50 ° C for 1 min, increased to 200 ° C at 10 ° C min ^-1 , maintained at 200 ° C for 15 min.
EXAMPLE 2
Powdered reagents were prepared by physically mixing the carbohydrate feed and the catalyst. Both the feed and the catalyst were sieved to <140mesh before mixing. The physical glucose mixtures tested were prepared with a mass ratio Dglucose (Fisher) to HZSM-5 (Si / Al = 30, WR Grace) of 19, 9, 4, 2.3, and 1.5. Xyliol (Fisher) / ZSM-5, cellobiose (Acros) / ZSM-5, and cellulose (Whatnam) / ZSM-5 with a catalyst: 19 feed weight ratio was also prepared. The HZSM-5 was calcined at 500 ° C in air for 5 hours before the reaction. Samples with a mass ratio of the catalyst: glucose of 19 were also prepared with the following catalysts: silicalite, O-zeolite, Yzeolite, and mesoporous SiO2 / AI2O3 (SiO2 / AI2O3 = 35). The reaction conditions, product yield, and product selectivities for all pyrolysis operations are summarized in Table 1. All operations
84/119 were carried out with a reaction temperature of 600 ° C. The yields are presented in terms of molar carbon yield where the moles of carbon in the product are divided by the moles of carbon in the reagent. The aromatic yields 5 were calculated by dividing the moles of carbon in the aromatic product (s) by mol of carbon in the feed.
Table 1. Summary of Pyrolysis Experiments food Catalyst Reason inMass of Cata- Rate ofHeating ° ç / s Reaction Time (s) YieldAro- YieldOxi- CO ₂ Yield 1%} CO yield(%) YieldCokeÍ% 1 Total Carbon Undesirable(%) pain forfood CO(%) of(%)Gli-sew ZSM-5 19 1000 240 31.4 <1 12.6 15.3 33.2 92.5 7.5 Cel-bio-if ZSM-5 19 1000 240 28.2 <1 10.4 13.0 30.0 81.6 18.4 Cell-lose ZSM-5 19 1000 240 31, 1 <1 86 15.2 28.6 83.5 16.5 Xili-tol ZSM-5 19 1000 240 47, 5 <1 7.2 12.8 37.5 105.0 0.0 Gli-sew ZSM-5 19 1 240 14.9 <1 12.0 13.3 38.9 79.1 20.9 Gli-sew ZSM-5 19 5 240 23, 6 <1 8.5 10.5 35, 6 78.2 21.8 Gli-sew ZSM-5 19 50 240 29, 4 <1 6, 6 9.0 34.1 79.1 20.9 Gli-sew ZSM-5 9 1000 240 27.2 2.8 11.0 13, 6 32.3 86, 9 13.1 Gli-sew ZSM-5 4 1000 240 22, 9 8.2 9.3 11.4 43, 6 95.4 4.6 Gli-sew ZSM-5 2.3 1000 240 16.5 13, 6 7.3 8.7 41.7 87.8 12.2 Glucose ZSM-5 1.5 1000 240 13.2 14.9 6, 3 7.1 43, 9 85.3 14.7 Gli- Sili- 19 1000 240 6.5 12.3 5.6 6.1 69.4 “ 30, 6
85/119
sew cali-to -η --------- Gli-sew Sio ₂ - A12O3 19 1000 240 0, 6 1.5 4.5 4.3 89, ^la ) 10, 9Glucose B-Zeó-lita 19 1000 240 4.3 1.1 10, 5 7.8 76, 3 ^{| a} ) 23, 7Gli-sew Y-Zeó-lita 19 1000 240 1.1 1.8 5.3 5.3 86.5 ^la ) 13.5[a] Coke yield was estimated by mass balance
EXAMPLE 3
According to the results summarized above, xylitol and xylose can be converted into thermally stable compounds by catalytic pyrolysis without significant coke formation (see Table 2). Addition of catalyst to the pyrolysis process significantly decreases the formation of coke and increases the conversion to thermally stable products. Five different catalysts were tested for xylitol catalytic pyrolysis including: silica alumina (SiO ₂ -Al ₂ O ₃ Grace-Davison 3125), zirconium tungstate (WO _x / ZrO ₂ MEI X201251), zirconium sulfate (SO _x / ZrO ₂ MEI X20880), Pt-silica-alumina (Pt / SiO ₂ -Al ₂ O3 prepared according to Huber et al.) And ZSM-5 (35 WR Grace silica to alumina molar ratio). The catalyst structure greatly changes the product's selectivity, and high yields (50%) of aromatic compounds (which could be used as gasoline fuel additives) can be produced with a ZSM-5 catalyst. The system used detects thermally stable products against thermally unstable compounds, which decompose under GC conditions. Notably, xylose produces furfural with a higher selectivity (55%) than when xylitol is the food.
Table 2 plots the results for xylitol catalytic pyrolysis in the GCMS pyro-sample system. At
86/119 reaction conditions for these experiments were as follows:
Temperature 600 ° C; Elevation Rate, 1000 ° C / s;
Reaction time, 60 s, Xylitol to catalyst weight ratio, 0.18; xylitol added to the catalyst, as a mixture physically ground in mesh between 60-120; inert gas, 1 atm of He, unless otherwise indicated.
Table 2. Xylitol catalytic pyrolysis in the pyro-sample system-GCMS Catalyst Conversion (% P) ¹⁴¹ Conversion%Carbon! 51 Carbon Selectivity (%) ¹⁶¹ Netdo + Solidof Composition Stick itat the MethylStick it Furfu-ral Acetaldehyde CO ₂ CO Arom.[7] None ¹¹¹ 65 35 1 10 2 20 41 8 18 <1 SiO2 “Al ₂ O ₃ 95 5 12 27 7 7 25 22 13 <1 site ₂ - A1 ₂ O ₃ ¹²¹ 2 1 55 5 13 24 <1 Pt / Sio ₂ - A12O3 67 33 25 12 3 1 15 18 50 <1 Pt / Sio ₂ - AI2O3 with h ₂ ¹³¹ 86 13, 32 27 13 2 <1 14 45 <1 WO _x / ZrO ₂ 99 1 13 30 13 12 19 16 10 <1 SO _x / ZrO ₂ 92 8 11 30 12 7 7 34 9 <1 ZSM-5 67 33<1 <1 <1 <1 11 19 70 [1] 1.08 mg of xylitol was used for this experiment. Xylitol was ground to a mesh size of between 60-120.[2] Xylose was fed for this experiment.[3] This experiment was done in a hydrogen atmosphere to see the hydrogen effect.[4] Conversion based on change in weight.[5] This conversion only reports thermally stable components defined as products that can be analyzed with a current GTCMS system.[6] Selectivity is on a carbon basis and only includes thermally stable compounds identified as GTCMS.[7] Aromatic products include: Benzene, Toluene, Ethyl Benzene, Xylenes, Naphthalene, Methyl-Naphthalene, Dimethyl-Naphthalene, Ethyl-naphthalene, Hydrindene, Methyl-hydrindene, and dimethyl-hydrindene.
EXAMPLE 4
Addition of metal to silica-alumina shifted selectivity towards CO, which shows that metals can influence the reaction chemistry. Such results
87/119 suggest that by adding different amounts of metal to the catalysts, the rate of hydrogen producing reactions and the hydrogen transfer reactions can be increased. Catalytic hydropyrolysis (catalytic pyrolysis with hydrogen instead of He) decreased the formation of coke on the catalyst, illustrating yet another modality. These preliminary positive results show that catalytic pyrolysis can produce a range of products, including aromatics, which can be used as a mixture of gasoline or jet fuel. This aromatic blend can be produced from a variety — of raw materials — on — the — eataüsader — ZSM-5 — with selectivity of the similar product. (See, for example, Figures 1A-1B).
EXAMPLE 5
Compounds derived from lignin and lignin can also be converted to aromatic fluid compounds by catalytic pyrolysis, in accordance with the present invention (Tables 3 and 4). Organosolvent lignin pyrolysis mainly produces benzyl phenyl ether (BPE), ethanol, methanol, CO and CO ₂ . Catalyst pyrolysis increases the conversion to thermally stable products 3 to 10 times (compared to pyrolysis without catalysts) with SiO2-AI2O3 and ZSM-5, respectively. Organosolvent lignin is a lignin product from the organosolvent pulp process, and similar results can be anticipated from other solid compounds comprising lignin. These experiments show how catalysts can significantly change the products and reactivity of the lignin-derived feed in the pyrolysis process. One of the main products formed from organosolvent lignin is BPE, and Table 3 shows the results for pyrolysis
88/119 ítica catalytic: Benzene, phenol, toluene and other aromatic compounds. The catalytic pyrolysis process can be modified to produce benzene, phenol, toluene and other aromatic compounds directly from 5 streams of solid lignin. Benzene and toluene can be added directly to gasoline, while phenol is a valuable chemical.
Table 3 traces the results from the catalytic pyrolysis of organosolvent lignin (Aldrich) in the GCMS pyro-sample system. The reaction conditions were as follows: Temperature, 600 ° C; Elevation rate, 1000 ° C / s; Reaction-time, - 60 s; - Ra-z-èe-in-Rese-de-Lignina-to
Catalyst, 0.18; Lignin added to the catalyst, as a mixture physically ground to 60-120 mesh;
inert gas, 1 atm of He.
Table 3. Catalytic pyrolysis of organosolvent lignin in a Piro-sample systemGCMS Catalyst Conv. (% C) ¹¹¹ Selectivity (%) ¹³¹ Ethanol Methanol CO CO ₂ Methyl guaiacol Benzyl phenyl ether Xil. . ^U1 Hydroxyphenyl none 3.61 5.27 9.81 9.20 5, 40 4.66 50.97 0.00 - SiO ₂ -Al ₂ O ₃ 9.04 10.73 13, 49 16.85 7.82 - 43, 48 0.00 3.03 ZSM-5 37.41 0.21 0.36 9.06 2.42 - 72, 91 6.06 6, 19 [1] Conversion based on thermally stable compounds only.[2] Xylene includes meta- and para-xylene.[3] Other products observed in low selectivity, but not reported here include diphenylmethane, 1,2-diphenylethane, guaiacol, acetic acid, furfuraldehydeand toluene
Table 4 traces the results from the catalytic pyrolysis of benzyl-phenyl ether (BPE) in the system
GCMS pyro-sample. The reaction conditions were as follows:
89/119
Temperature, 600 ° C; Elevation Rate, 1000 ° C / s; Reaction time, 60 s; Weight ratio of lignin to catalyst, 0.18; Lignin added to the catalyst, as a physically ground mixture between 60 - 120 mesh;
inert gas, 1 atm of He.
Table 3. Catalytic pyrolysis of benzyl-phenyl ether (BPE) in a pyro-sample system ~ GCMS Catalyst Conv. (% C) ¹¹¹ Selectivity Benzene Phenol Toluene Biphenyl-methane[21 1.2-Diphenylethane Methylphenol none 14.33 4.16 80, 02 8, 91 0.00 4.72 0.00 SY2-A12O3 63.57 37.88 36.19 10.60 4.85 6.08 3.77 ZSM-5 53, 34 38, 62 42.87 6, 48 4.88 3.79 2.51 [1] Conversion based on stable terraicairente compounds only -; -------------------[2] Methyl phenol includes (2,3 and 4) -methylphenols.[3] Other products with low selectivity include benzyl alcohol and benzaldehyde.
EXAMPLE 6
According to certain modalities and by comparison with the prior art, the biomass was pyrolysed to condensable vapors, which were converted over the catalyst (in situ) at the same temperature in the same reaction chamber. A second stage, as described in the
US Patent Nos. 7,241,323 and 5,504,259, was eliminated from the process. The benefit of a one-stage process is twofold: less energy is used when the fluid product is condensed and perfected later (all chemistry happens at the same temperature), and condensable vapors do not have a chance to polymerize or otherwise. degrade during transfer to an improved second stage.
0 EXAMPLE 7
In general, prior art fluid fuel product compositions are not specified. While, the literature refers to the reduced oxygen content, it does not disclose components of
90/119 specific molecular fuel. Illustrating a range of modalities of the type described here, a ZSM-5 catalyst in a fixed bed reactor produced a fluid consisting almost entirely of aromatic compounds. Oxygen was removed from the biomass in the form of water, CO, and
CO _2. Specifically, the aromatic compounds quantified in the fuel included: benzene, toluene, xylenes, ethylbenzene, ethyl-methyl-benzene, trimethyl-benzene, indane, methyl-indane, naphthalene, methyl-naphthalene, and dimethyl-10 naphthalene. Such an aromatic mixture could be used as a high octane fuel additive. All aromatic izompxrstas ~ s — are — above — 100 — of — ethane — with — the exception of naphthalene (90 of octane). See Table 5, below.
Table 5. Properties of quantified aromatic species Compound Point ofboiling (° C) Number of Number ofengine octane (MON) octane research (RON) Benzene 84.35 98 90 Toluene 112.29 124 112 Ethylbenzene 135.17 124 107 o-Xylene 140.15 120 102 m-Xylene 140.15 145 124 p-Xylene 140.15 146 126 Ethyl-methylbenzene 163.03 126-155 112-138 Tri-methylbenzene 168.01 118-170 104-136 Indano 174.44 161 140 Naphthalene 199.91 not reported 90 Methyl-naphthalene 227.77 123-127 114-116
Source: Pure hydrocarbon shock characteristics (Research project 45)
91/119
American Society of Material Testing (ASTM), Special Technical Publication Number 225.
Philadelphia, PA, 1958.
As shown above, high quality aromatic fuel / additives can be produced directly from solid biomass raw materials by catalytic pyrolysis in a single catalytic reactor in short residence times. Through an understanding of the reaction chemistry, catalyst, and the design of the apparatus / reactor, catalytic pyrolysis can be used _______________ to efficiently generate biofuel fluids from a range of lignocellulosic biomass sources.
EXAMPLE 8
The effect of varying the molar ratio of silica to alumina on the catalyst was also investigated. The conditions for these experiments were as follows: mass ratio of the catalyst for feeding, 19; catalyst, 15 ZSM5; nominal heating rate, 1000 ° C s ^-1 ; reaction temperature, 600 ° C; reaction time (feed residence time), 240 s. Glucose was used as a hydrocarbon feed for these experiments. Figs. 6A-6B show the product selectivity for catalytic glucose pyrolysis as a function of the molar ratio of silica to alumina. As shown in FIG. 6A, the use of catalyst with a silica to alumina molar ratio of 30 produced with a greater amount of aromatic product compared to the use of catalysts with 25 silica to alumina molar ratios of 23, 55, or 80. As shown in FIG .
6B, the use of catalysts with different molar ratios from silica to alumina can produce higher yields of selective compounds. For example, to produce
92/119 higher naphthalene yields, a molar ratio of silica to alumina of about 30 or 50 can be used.
EXAMPLE 9
This example illustrates the effect of metal impregnation on the catalyst on product yields. Impregnation of ZSM-5 pores (molar ratio of silica to alumina of 30, Zeolyst) with metals shifted the product's selectivity towards CO and CO2, showing that metals can influence the reaction chemistry. Not wishing to be limited by any theory, metals can increase decarbonylation and / or decarboxylation reaction rates. The following metals have been tested: —Cu, —Mn, —Feb — Co, —Ni, Zn, Ga, and Pt. Table 6 summarizes the results obtained for catalytic glucose pyrolysis in the incorporated metal ZSM-5 in the system pyro-sample GCMS. Two different methods were used for the addition of ZSM-5 metal: wet impregnation and ion exchange. Catalysts impregnated using the ion exchange method produced higher yields of aromatic compounds and lower yields of coke compared to catalysts impregnated using the wet impregnation method.
Table 6. Summary of metal addition in ZSM-5 Catalyst Metal loading(%P) Method ofPreparation YieldAromatic(%Ç) YieldOxygenated(%Ç) CO ₂ yield (% C) CO yield(%Ç) Coke Yield(%Ç) Cu-ZSM-5 6.2% p stateSolid ExchangeIonic 8.1 <1 24.1 33, 6 31.4 Mn-ZSM-5 5% p Wet Impregnation 11.9 <1 18.2 25.5 45.0 Mn-ZSM-5 5% p ExchangeIonic 23.0 <1 5.8 22.5 34.1 Fe-ZSM-5 5% p Wet Impregnation 20.9 <1 11.5 24.8 41.0 Co-ZSM-5 5% p Impregnate 12.7 <1 28.0 44.5 19, 6
93/119
Wet SectionNi-ZSM-5 5% p Wet Impregnation 7.2 <1 34.7 47.1 12.0 Zn-ZSM-5 5% p Wet Impregnation 23.7 <1 14.7 23.7 37.9 Zn-ZSM-5 5% p ExchangeIonic 32, 9 <1 9, 9 29.2 30, 7 Ga-ZSM-5 5% p Wet Impregnation 28.5 <1 6.0 22.2 48.0 Ga-ZSM-5 5% p ExchangeIonic 33.3 <1 7.6 30, 4 23.2 Pt-ZSM-5 5% p Wet Impregnation 17.0 <1 15, 6 35.1 25.2
Catalyst pore sizes also affected aromatic yield. Table 7 includes data on glucose yield over several different zeolite structures. Not wishing to be limited by any theory, it may be desirable to use zeolite catalysts with pore sizes large enough to allow the diffusion of oxygenated intermediate molecules (eg methyl furfural, which has a kinetic diameter of 5.9 Angstroms) in the structure of zeolite. It may also be desirable to use zeolite catalysts with pore sizes small enough to selectively produce aromatics (<6.3 Angstroms).
Table 7 shows that ZK-5 did not produce aromatics, while
Y-zeolite mainly produced coke.
Catalysts with pore sizes closer to that of ZSM-5 (5.6 Â) produced the most aromatic yield.
Table 7. Summary of Different Pyrolysis Structures CatalyticZeolite Co- Structured Di ” of OK- Surrendered Surrendered Surrendered Surrendered Surrendereddig ture mensen ring trick - - - - -O0in ment ment ment ment mentIZA Pore Aroma Oxige in of CO in -tico -nothing_CO2 (%Ç) Coke (%Ç) (%Ç) (%Ç)(%Ç) Frog-34 TEA Romboe- 3 8 3.8X3, 0.3 8.6 9.5 21.7 68.8 dral 8
94/119
ZK-5 KFI cubic 3 8.8 3.9X3,9 0.0 1.8 18.1 25.2 46.4 Ferrieri-OK FER Orthorom-spout 2 10.8 4.2X5,43.5X4,8 0.0 3.4 8.8 27.7 34.6 ZSM-23 MTT Ortorômbi-co 1 10 4.5X5,2 7.0 1.2 4.8 19, 8 43.4 SSZ-20 TON Ortorômbi-co 1 10 4.6X5,7 7.4 2.8 3.7 15, 9 34.0 SSZ-41 VET tetragon1 1 12 5.9X5,9 5.2 0, 6 2.6 7.7 65.5 ZSM-12 MTW Ortorômbi-co 1 12 5 _f 9x6, 0 2.5 2.8 2.5 8.2 79.4 SSZ-55 ATS Ortorômbi-co 1 12 6.5X7,5 3.4 <1 5.6 21.4 83.7 B- BEA Tetrago- 3 12.1 6, 6X6, 4.3 1.1 10.5 7.8 67.0 Zeolitenal2 75.6X5,6Y-Zeolite FAU cubic 3 12.12 7.4X7,4 1.6 <1 7.6 25.3 85.0
The density and strength of acidic sites in the catalysts also had an effect on the production of aromatics. FIG. 11 is a graph of carbon yield from glucose pyrolysis using ZSM5 5 catalysts with different silica molar ratios for alumina (SiO2 / AI2O3 = 23, 30, 50 and 80, Zeolyst). ZSM-5 (SiO ₂ / Al ₂ O ₃ = 30) produced the maximum aromatic yield at 600 ° C with an elevation rate of 1000 ° C / s.
EXAMPLE 10
Naturally occurring biomass was also used as a raw material in some experiments to produce aromatics from fluid by catalytic pyrolysis.
Table 8 presents the results for biomass catalytic pyrolysis occurring naturally in the 15 pyro-sample-GCMS system. Pyrolysis of wood, sugar cane (Brazil and Hawaii) and corn straw under ZSM-5 (Si / Al = 60 WR Grace) produced aromatics, CO and CO ₂ . The aromatic yields produced using these raw materials were
95/119 comparable to glucose and cellulose yields. Such results suggest that catalytic pyrolysis can be used with naturally occurring biomass raw materials.
Table 8. Biomass Catalytic Pyrolysis occurring naturally in the Piroamostra-GCMS system Food-tation Catali-sador Molar ratio of catalyst-to-feed-dog Rateheating (° C / s) Reaction time(s) Aromatic Yield (% C) YieldOxygenated (% C) CO ₂ yield (% C) Yield-to ofCO (% C) wood ZSM-5 19 1000 240 26.4 <1 4.9 11.7 Sugar caneICA ZSM-5 19 1000 240 28.3 <1 8.0 12.2 Paulo) Sugar cane (Leandro ZSM-5 19 1000 240 29.9 <1 6, 6 12.9 tBrazil) Sugar cane (Puunene, HI ZSM-5 19 1000 240 26, 6 <1 5, 6 11.9 Straw ofCorn ZSM-5 19 1000 240 21.4 <1 6.8 10.1
EXAMPLE 11
This example describes the use of a fixed bed flow reactor system. In this example, a tubular quartz reactor 12.7 mm (0.5 inches) in diameter (approximately 50.8 mm (2 inches) in length) was used. The reactor was loaded with 50 mg of 10 ZSM-5 catalyst (ZEOLYST, CBV 3024E, SiO2 / AI2O3 = 30) to produce a fixed bed, which was supported by quartz wool and quartz granules. The quartz reactor was supported in a temperature controlled oven (oven: Lindberg, 55035 A; temperature controller: Omega, CN96211TR) at 600 ° C. The reactor temperature was controlled by a thermocouple
96/119 inserted through an internal quartz tube to the top surface of the packaging bed.
The raw material for this example was furan (Sigma-AIdrich, 99%). During the operation, helium (ultra-high purity, Airgas) was used as a carrier gas, and its flow rate was controlled by a mass flow controller (controller: Brooks, SLA5850S1 BAB1-C2A 1; control box : Brooks, 0154CFD2B31A). Liquid raw material (furan) was introduced into the vehicle gas via a syringe pump (KD Scientific, KDS 100) and was quickly vaporized. The vaporized feed was transported to the reactor. —Reactor temperature- was 600 ° C, under an atmosphere of He with a partial pressure of 0.76 KPa (5.7 torr) furan.
Products flowed from the reactor to a condenser placed in an acetone bath on dry ice. The condenser was maintained at a temperature of about -55 ° C and was used to condense products with relatively high boiling points. Gas products were collected in gas sampling bags. Liquid and gas products were analyzed using GC-FID (Shimadzu 2010).
The amount of carbon in the catalyst was determined during the regeneration of the catalyst. During regeneration, CO was converted to CO2 by a copper converter. CO2 was trapped by ascarite (Sigma-AIdrich, 5-20 mesh). The amount of carbon in the partially deactivated catalyst was calculated by measuring the weight change of the CO2 trap.
As shown in FIG. 12, the olefin yield was affected by the normalized spatial mass velocity (h ^_1 ) of furan. The normalized mass space velocity was calculated as the mass flow rate of
97/119 furan (g / h) divided by the mass of catalyst (g). The carbon yield was calculated by dividing the moles of carbon in a product by the moles of carbon fed to the reactor. The types of olefins produced included ethylene and propene, and a trace amount of butene was also detected. For a space speed of 0.35 h ¹ , olefins were not detected. However, at a spatial speed of 2.79 h ^{_ 1} , the olefin yield rose to 15%. On the contrary, the yield of aromatic compounds, which were the second most abundant product, decreased with increasing spatial velocity (and as the time of res ± dence ± u «imT ± nxríüayn -------- ----------------------------------- FIG. 13 includes a graph comparing the carbon yields of each aromatic product and olefins obtained from the conversion of furan in the fixed bed flow reactor and the pyro-sample.The reaction conditions for the fixed bed reactor were: reaction temperature, 600 ° C;
partial pressure of furan, 0.76 space KPa 0.89 h ^_1 . The pyro-speed conditions were: Reaction temperature, mass of the catalyst for feeding, reaction, 4 min. Both pyro-sample reactors used olefin catalysts were not produced in experiments (5.7 torr); and reaction to
600 ° C;
19; and fixed bed ratio time as in de
ZSM-5. In the pyro-sample, a substantial degree. Those from the pyro-sample, however, resulted in a large amount of naphthalene, and a relatively large amount of aromatic compounds. The fixed bed reactor experiments yielded slightly higher amounts of valuable products (aromatics plus olefins). Benzene and toluene yields were also higher in the fixed bed reactor than in pyro-samples. Ethylbenzene and trimethylbenzene (which are not shown in FIG.
13) were also detected in the
98/119 pyro-sample experiments, but not to an appreciable extent in flow reactor experiments
EXAMPLE 12
In this example, a fluidized bed was used to convert the solid biomass to hydrocarbons. FIG. 14 includes a schematic diagram of the fluidized bed system. The fluidized bed reactor was constructed using a diameter of 50.8mm (2 inches), stainless steel tube 316. The tube was 254mm (10 inches) 10 length. A distribution plate made from a stacked 316 stainless steel mesh (300 mesh) was
---- mounted inside the reactor. —The plate — dTSbribuddora — served to support the catalyst bed. The reactor was loaded with 90 grams of ZSM-5 catalyst (Grace).
Before the operation, the catalyst was calcined in the reactor for 4 hours at 600 ° C in 1200 mL min ^-1 of flowing air.
The catalyst was fluidized during the operation of the reactor through a stream of helium gas controlled by a mass flow controller. The flow rate of gas fluidization was 1200 mL min ^-1 in the SATP. Raw materials from solid wood biomass were injected by a stainless steel bit into the side of the reactor from a sealed feed hopper. The wood feed rate was 6 gh ^-1 , yielding a normalized mass spatial speed of 0.07 h ^_1 . To maintain an inert environment in the reactor, the hopper was swept with helium at a rate of 200 mL min ^-1 . Both the reactor and the fluidizing gas were resistively heated to a reaction temperature of 600 ° C.
During the operation, the product gases left the top of the reactor and passed through a cyclone, operated at 450 ° C, where the entrained solids were removed and collected. The steam then passed through a
Condenser 99/119. The first three condensers were operated at 0 ° C in an ice bath, and the following three condensers were operated at -55 ° C in a dry ice / acetone bath. Non-condensed vapors from the condenser train were collected in a tedlar sampling gas bag for GC MS and GC / FID analysis. Liquids collected in the condensers were quantitatively removed after the reaction with ethanol and analyzed by GC / MS and GC / FID.
The aromatic yields for wood pyrolysis with ZSM-5 catalyst in the fluidized bed reactor system are shown in FIG. 15. The quantified products — na — Fig. — 15 — include: —calf - (- Berrrj-ç — toluene (Tol.), Xylene (Xyl.), Naphthalene (Naf.), Ethylene (Et.) And propene ( The yield of naphthalene was relatively low, and the primary products were benzene, toluene and xylene.In addition to aromatic products, olefins such as ethylene and propene were produced during catalytic pyrolysis in the fluidized bed reactor.
The selectivity for olefins and aromatic products was adjusted by varying the normalized spatial velocity of biomass mass. Two different spatial velocities were tested in the fluidized bed reactor at a reactor temperature of 600 ° C. Loads of 30 g and 90 g of catalyst were used for the high and low spatial velocity operations of biomass, respectively. Power pinewood dried (80- 120 mesh) was fed at 30 ¹ gh and gh ¹ 6 for high speed operation and low spatial biomass, respectively. The fluidization gas temperature was 600 ° C, with a fluidization gas flow rate of 12 mL min ^-1 in the SATP. The cyclone temperature was 450 ° C, and a ZSM-5 catalyst was used. As seen in Figs 16 ^to -16B, the reaction is selective for the production of olefin at speed
100/119 high space while low space speed, aromatic products are favored. The quantified olefin products included :. ethylene, propene, butane, and butadiene. The aromatic products quantified 5 included: benzene, toluene, xylene, ethyl-benzene, styrene, and naphthalene.
EXAMPLE 13
This example describes the use of a fixed bed reactor to test the effects of feed olefins on the types and quantities of products formed during the catalytic pyrolysis of a hydrocarbon feed material ----- si-stem -a — de — re-a-to / n — de — flux-ode fixed bed was built for biomass conversion. The vehicle gas flow rates (described below) 15 were controlled using the mass flow controllers (controller: Brooks, SLA5850S1BAB1-C2A 1; control box: Brooks, 0154CFD2B31A). Liquid raw material was introduced into the carrier gas using a syringe pump (KD Scientific, KDS 100) and was sprayed essentially 20 immediately. The steam raw material was then loaded into a tubular quartz reactor. The quartz reactor was kept in a temperature control oven (oven: Lindberg, 55035A, temperature controller: Omega,
CN96211TR). The catalyst bed fixed in the reactor was supported by quartz wool and quartz granules. The reactor temperature was monitored using a thermocouple inserted through an inner quartz tube to the top surface of the packed bed. Products from the reactor were passed through a condenser placed in an ice-water bath to condense heavy liquid products. Gas products were collected in gas sampling bags. The liquid and gaseous products were analyzed by gas chromatography (GC) using a
101/119 flame ionization detector (FID) using HP 7890 and Shimadzu 2014 instruments, respectively. Coke yield was determined by thermogravimetric analysis (TGA) (TA Instrument, SDT-Q600), assuming that the weight loss was caused by carbon removal.
Several experiments were carried out in which furan (Sigma-Aldrich, 99%) was reacted under He atmosphere in the flow reactor. ZSM-5 was used as the catalyst (ZEOLYST, CBV 3024E, SiO2 / AI2O3 = 30). Ethene (known as 10 ethylene) and propene (known as propylene) were used as sources of olefins, and were incorporated into ---- gas-for-feed, -when-teatured ^ —The — compositions — of — gas for following vehicles were tested: 100% helium, 98% helium / 2% ethylene, 99.8% helium / 0.2% ethylene, 98% helium / 2% propene, and 99.8% helium / 0.2% propene.
Table 9 shows the reaction conditions for each experimental test. During each test, the reaction temperature was maintained at 600 ° C, the space velocity of furan (WHSV) was 10.5 h ¹ , the partial pressure of furan was 0.8 KPa 20 (6 torr), and the vehicle gas flow rate was 200 ml / min.
Table 9. Reaction conditions test Feedstock Vehicle gas, 200mL / min Conversion ofFuran (%) 1 Furano He 43 2 Furano 2% Ethene 46 3 Furano 0.2% Ethene 49 4 Furano 2% Propene 49 5 Furano 0.2% ofPropene 62
FIG. 18 includes a graph of carbon yield (expressed as a percentage) for several
102/119 products obtained from the catalytic conversion of raw materials derived from biomass in the quartz tube flow reactor, using ZSM-5 catalyst and the reaction conditions shown in Table 9. The following abbreviations are used in FIG. El: Eene = Ethene, Pene = Propene, Bene = Butene, B = Benzene, T = Toluene, X = Xylene, C = Coke. As shown in FIG. 18, aromatics and olefins were produced, and their carbon yields were affected by the composition of the raw material. The carbon yield was defined as the moles of carbon in the product divided by the moles of carbon in the furan feed. were used as the vehicle gas) are not shown in FIG. 18 because of its high values. Furan conversion is illustrated in Table 9.
Smaller products (contributing less than 4% of carbon yield) included styrene, benzofuran, indene and naphthalene. With an increase in the concentration of propene in the vehicle gas, the reaction was more selective for aromatics and light olefins, especially for ethylene, butene, toluene, and xylene. Referring to Operations 2 and 3, the inclusion of ethylene in the vehicle gas also led to the production of increased light and aromatic olefins, although the effect was not as significant as observed when propylene was used. In addition, the co-feed of ethylene and propene increased the conversion of furan and decreased the amount of coke produced. Due to the greater furan conversion, greater aromatic selectivity, and lower coke yield, it is evident that the incorporation of olefins into the feed stream (for example, through recycling) can be beneficial to increase the yield and selectivity of
103/119 certain products in processes involving the conversion of biomass.
EXAMPLE 14
This example describes the use of a fluidized bed reactor to test the effects of feed olefins on the types and quantities of products formed during the catalytic pyrolysis of a feed hydrocarbon material. The configuration of the fluidized bed reactor is illustrated in FIG. 19. The fluid bed reactor is comprised of a 31.8 stainless steel tube 50.8 mm (2 inches) in diameter and 254 mm (10 inches) in height. Biomass - solids - fermes --- ferarm injected into the side of the reactor from a sealed biomass hopper using a stainless steel drill bit. In order to maintain an inert environment in the reactor, the hopper was swept with helium at a rate of 200 mL min ^-1 . The catalyst bed was supported by a distributor plate made from stacked 316 stainless steel mesh (300 mesh). During the reaction, the catalyst was fluidized through a stream of helium gas controlled by a mass flow controller. Both the reactor and the inlet gas stream were resistively heated to the reaction temperature.
During the operation, the product gases left the top of the reactor and passed through a cyclone, where the entrained solids were removed and collected. The steam was then passed through a condenser train. The first three condensers were operated at 0 ° C in an ice bath, and the following three condensers were operated at -55 ° C in a dry ice / acetone bath. Non-condensing vapors from the condenser train were collected in a Tedlar® gas sampling bag for gas chromatography / mass spectrometry analysis
104/119 (GC-MS) and GC / FID. Liquids collected in the condensers were quantitatively removed after the reaction with ethanol and analyzed by GC / MS and GC / FID. catalyst ZSM-5 (Grace) was calcined in the reactor for 4 hours at 600 ° C in 1200 mL min ' ¹ flowing air before the reaction.
For the olefin co-feed experiments, the secondary gas (T2 in FIG. 19) was exchanged for either ethylene or propylene controlled, at a desired flow rate. The helium fluidization gas flow rate was adjusted to hold constant total inlet gas flow rate at 1200 mL min ' ¹ .
After removal, the secondary gas — was — exchanged — with air to regenerate the catalyst. The effluent combustion during regeneration was passed under a copper catalyst maintained at 150 ° C to convert carbon monoxide to carbon dioxide. The stream of carbon dioxide was then passed through a dryrite trap to remove water vapor. The dry carbon dioxide was harvested by a pre-weighed ascarite trap. The total moles of carbon dioxide collected in the trap were assumed to be substantially equal to the moles of carbon in the coke on the catalyst bed.
Olefin co-feeding experiments were conducted using the reaction parameters described in Table 10. All reactor parameters were kept constant, except for the olefin concentration in the incoming fluidization gas. As shown in Table 10, the olefin moles exiting the reactor, in the case of propylene, were about half the amount of feed. This indicated that propylene was consumed during the reaction. When ethylene was used as the co-feed, there was a net production of ethylene during the reaction, which
105/119 suggested that ethylene was a stable product and was non-reactive.
Table 10. Olefin reaction and conversion parameters forcatalytic wood pyrolysisCo-feedingPropylene Co-feedingEthylene No Calibration Operation 1 2 3 4 5 Temperature (° C) 600 600 600 600 600 velocitySpatialwood (h ~~) 0.24 0.24 0.26 0.27 0.21 g propylene / g 0.16 0.04 0, 08 0.02 0.00 woodPropylene / Wood(quantity ofcarbon) 0.30 0.09 0.15 0.05 0.00 Olefin Molesoutput input 0.50 0.45 0.96 1.33 AT
FIGS. 20A-20B include product yield charts for catalytic pyrolysis (CP) of wood with A) a propylene co-feed, and B) an ethylene co-feed. The quantified aromatic compounds included: benzene, toluene, xylene, ethyl benzene, styrene, indene, phenol and naphthalene. The quantified olefins included: ethylene, propylene, butene and butadiene. Carbon yields were calculated as the amount of carbon in the given product divided by the total amount of carbon in the food (wood and olefin). As shown in FIG. 20A, the aromatic yield increased slightly, while the coke yield decreased significantly from 30 to 15 25% when propylene was used as a co-feed.
The yield of carbon dioxide and carbon monoxide also decreased at higher concentrations of propylene feed. For co-feeding of ethylene,
106/119 there was a decrease in aromatic yield with increasing feed concentration (FIG. 20B). Coke yield also decreased slightly with increasing ethylene concentration, however, this change observed was less pronounced than that seen when propylene was co-fed.
Figs. 21A-21B include graphs of product selectivity for various aromatic compounds for catalytic wood pyrolysis with a) propylene co-feed, and 10 b) ethylene co-feed. As shown in Figs. 21A21B, selectivity for benzene, toluene, xylene and ----- naphthalene — has changed - with - different - eona-limitation-of-olefins. Propylene had little effect on selectivity for benzene and xylene. However, propylene affected the selectivity for toluene and naphthalene. The selectivity for toluene increased from 36% to 40%, while for naphthalene it decreased from 17% to 13%. Ethylene exhibited an opposite trend, as it affected the selectivity for benzene and xylene, and did not have a significant effect on toluene or naphthalene. Selectivity for benzene increased from 30% to 35%, while for xylene it decreased from 17 to 12%.
EXAMPLE 15
This example describes the effects of varying particle sizes within catalyst objects (i.e., in the form of catalyst particle agglomerates) on the types and amounts of compounds produced during the catalytic pyrolysis of hydrocarbonaceous material. Catalytic pyrolysis experiments were performed using a pyro-sample 2000 analytical pyrolyzer model (CDS Analytical Inc.). The sample included a resistively heated computer-controlled element, which kept a quartz tube with
107/119 open end. Powder samples were kept in the tube with loose quartz wool packaging. During pyrolysis, vapors flowed from the open ends of the quartz tube into a large cavity (the pyrolysis interface) with a stream of helium vehicle gas. The vehicle gas stream was sent to a 58 90 gas chromatograph (GC) model interfaced to a Hewlett Packard 5972 mass spectrometer (MS) model. The pyrolysis interface was maintained at 100 ° C, and the injector GC was maintained at a temperature of 275 ° C. Helium was used as an inert pyrolysis gas, as well as the vehicle gas — for — the — system — GCMS. — A —flow — program — of ~ constant 0.5 mL min ' ¹ was used for GC capillary column (Restek Rtx-5sil MS) The GC oven was programmed with the following temperature regime: Maintained at 50 ° C for 1 min, raised to 270 ° C at 10 ° C min ' ¹ , maintained at 270 ° C for 15 min. The products were quantified by injecting the calibration standards into the GC / MS system. All yields were reported in terms of molar carbon yield, calculated as the moles of carbon in the product divided by the moles of carbon in the feed. The aromatic selectivity reported was defined as the carbon moles in an aromatic species divided by the total carbon moles of the aromatic species.
Powdered reagents were prepared by mechanically mixing a D-glucose feed (Fisher) and the catalyst. For a typical operation, about 8-15 mg of reagent and catalyst mixture was used. Both the feed and the catalyst were sieved to <140 mesh before mixing. Physical glucose mixtures were prepared with a ZSM-5 to glucose ratio of 19. ZSM-5 was calcined at 500 ° C in air for 5 hours before the reaction. In all experiments, a heating rate
108/119 of 1000 ° C / s, and a reaction temperature of 600 ° C were used.
In order to investigate the effect of particle size on catalytic activity, three different particle sizes of ZSM-5 were prepared. The first sample of the ZSM-5 catalyst, designated as AG101, was prepared as follows. 0.93 g of NaOH was dissolved in 9 ml of deionized water, and 12 g of Ludox AS-40 was added to the solution and stirred to form a synthetic mixture. A gel was formed, and stirring was continued for 15 min. 0.233 g of NaAlO ₂ was dissolved in 2 ml of deionized water — The NaA10 ₂ solution was added to the mixture of synthesis and stirred for 10 min. Then, 8.0 g of TPAOH solution (40%) was added dropwise to the synthesis mixture, and stirring was continued for 1 h. The final pH of the synthesis mixture was measured to be about 13. The synthesis mixture was then transferred to a Teflon-coated autoclave (internal volume 45 ml). The hydrothermal synthesis was carried out under static conditions at 170 ° C for 72 h under autogeneous pressure. The contents of the autoclave were centrifuged and washed 5 times, each time with 50 ml of deionized water, and dried overnight at 100 ° C. The resulting material was calcined in air at 300 ° C for 3 h, followed by calcination at 550 ° C for 6 h, with an elevation temperature of 1 ° C / min. FIG. 26A includes a PXRD standard for AG101 catalyst material.
The process for producing the second ZSM-5 catalyst sample (designated as AG102) was similar to the process for the production of AG101, except a solution of 1.8 g of acetic acid in 1.8 ml of deionized water was added dropwise to drop after the dropwise addition of TPAOH solution. After stirring for 10 minutes, the final pH of the synthesis mixture was measured to be about
109/119 of 10. This synthetic mixture was then transferred to a Teflon-coated autoclave and processed in a similar manner as the 101 AG sample. FIG. 26B includes a PXRD standard for the AG 102 catalyst material.
The process for producing the third ZSM-5 catalyst sample (referred to as AG103) was similar to the process for producing AG102 (ie, including the dropwise addition of the acetic acid solution to produce a final pH of about 10 ), except hydrothermal synthesis was performed under static conditions at 170 ° C for 168 hours (as opposed to 72 hours for AG101 and AG 102 cataFisadores), under autogenous pressure — FIG. — 26C includes a PXRD standard for the catalyst material AG 103.
The chemical composition of the catalysts was determined using X-ray fluorescence (XRF). As shown in Table 11, all three catalysts had Si to Al ratios of about 30: 1 (corresponding silica to alumina ratios of about 60: 1). The particle size was determined by scanning electron microscopy (SEM, JEOL X-Vision 6320FXV FESEM) coupled with the ImageJ software. About 10 to 20 particles in each SEM image were chosen to calculate an average particle size. As shown in Figs. 22A-22F, all three samples had different particle sizes. Figs. 22A-22B include SEM images of AG101 catalyst, Figs. 22C-22D include SEM images of catalyst from AG102, and Figs. 22E-22F include SEM images of AG103 catalyst. The AG101 catalyst included particle sizes of about 1 micron, the AG102 catalyst included particle sizes of about 10 microns, and the AG103 catalyst included particle sizes of about 20 microns. FIG. 23A included a
110/119 graph of carbon yields of aromatic compounds, CO and CO ₂ as a function of particle size for catalytic glucose pyrolysis. FIG. 23B describes the aromatic selectivity of various aromatic compounds, including benzene, toluene, xylene, naphthalene, and others (which include ethyl methyl benzene, trimethyl benzene, indane and indene) as a function of particle size for catalytic pyrolysis of glucose. Generally, the yield of aromatic compounds increased as the particle size decreased. For example, catalytic pyrolysis using AG101 yielded 35.2% aromatics, ----- while the use of AG103 yielded 31.9in- In addition, two commercial ZSM-5 catalysts from WR Grace and Zeolyst were tested. The 15 commercial catalysts had the same chemical composition (Si / Al = 15), but different particle sizes. At
Figs. 24A-24B include SEM images of WR Grace's ZSM-5 catalyst, showing a particle size of about 6 microns. Figs. 24C-24D includes SEM images of the ZSM-5 Zeolyst catalyst, including a wide range of catalyst particle sizes less than 1 micron to a maximum of 3 microns. FIG. 25 includes a graph of the carbon yields of various products for catalytic glucose pyrolysis using the ZSM-5 25 Zeolyst and ZSM-5 WR Grace catalyst (with the results of the ZSM-5 catalyst from sample AG101 included as a reference). As before, these experiments employed a feed catalyst ratio of 19, a heating rate of 1000 ° C / s, and a reaction temperature of 600 ° C. From Fig. 25, it is evident that the ZSM-5 Zeolyst produced about 10% more aromatic compounds than the ZSM-5WR Grace catalyst.
111/119
Table 11. Elemental analysis of catalyst for AG by XRF Sample Na (% p) Si (% p) Al (% p) AG 101 1,667 43.56 2.66 AG 102 0.938 45.26 1.46 AG 103 0.625 45.31 1.39
EXAMPLE 16
This example includes a description of a theoretical calculation describing the effects of feed olefins on the types and quantities of products formed during the catalytic pyrolysis of a hydrocarbon feed material. The feasibility of recycling olefins for a reactor can be assessed using a simple mass balance in the model system illustrated in FIG. 27. In the set of modalities illustrated in FIG. 27, the wood (biomass labeled in Stream 1) is mixed with a recycling stream (Stream 3) containing olefins, CO and CO2. The mixture is fed to a fluidized bed reactor. Within the reactor, the wood (dry base) can react to form aromatic compounds and olefins by Reactions (4) and (5), respectively.
C ₃ , 8 ^ 5,802.7 “► 0.36 C ₇ Hg + 1.45 H2O + 1.25 CO (4)
C3.8H ₅ , 8O ₂ , 7 -> 0.67 C ₃ H ₆ + 0.9 H ₂ 0 + 1.8 CO (5)
The olefins in the reactor can be converted into additional aromatic compounds by Reaction (6).
C ₃ H ₇ C ₆ H ₈ 3/7 + 9/7 H ₂ (6)
For this exemplary calculation, it was assumed that the balance of biomass not converted into aromatic compounds or olefins is converted into coke and gases. In the set of modalities illustrated in FIG. 27, the spent coking catalyst is then sent to a regenerator and regenerated by burning the coke in a secondary regeneration reactor. Most likely, recirculation
112/119 of the catalyst will be adjusted to control the temperature of the reactor and the regenerator. For this example, it was assumed that the coke yield would be relatively high, and removal of heat from the regenerator may be necessary to avoid high temperatures in the regenerator. Excess heat can be used in any other part of the process. In this example, the product stream from the reactor (Stream 3) is separated into the condensable aromatic product (Stream 4), water and water-soluble compounds and non-condensable olefins and gases. The separation system would include a condenser system that removes cuttable compounds — from — the — gases — ree ± ed-aviable ^ —Θ — liquefied product would contain a mixture of water, aromatics and water-soluble compounds. The aromatic product would be decanted and further refined. Water-soluble products and water would go to wastewater treatment. From the separation system, the olefins are then recycled to the reactor with a recycling molar ratio defined as olefin moles in the recycling (Stream 6) divided by olefin moles in the purge stream (Stream 5). The purge stream can be used to remove CO and CO ₂ and / or to prevent the accumulation of any other non-reactive species in the system.
FIG. 28 includes an exemplary graph of the yield of aromatic compounds as a function of the recycling ratio. The continuous lines are drawn for olefin reaction extensions (according to Equation 6 above) of 0.1, 0.2, 0.3, 0.4 and 0.5. Therefore, FIG. 28 illustrates the effect of adjusting the olefin conversion and recycling ratio on the yield of aromatic compounds. The extent of reactions for Equations 4 and 5 were both set at 0.17 to match the experimental yield for olefins and aromatic compounds in co
113/119 zero supply of olefin in the fluidized bed reactor. As shown in FIG. 28, the yield of aromatic compounds increases both with a proportion of recycling and with an increase in the increasing extent of the reaction for Reaction 6. The recycling ratio is defined as the mass of carbon within the hydrocarbonaceous material within the feed for the mass of carbon in the olefins in the recycling stream. Using recycling ratios in excess of 2: 1 (or in excess of 3: 1, or in excess of 4: 1), and having high reaction spans for Reaction 6 could produce a two-fold increase in compound yield— aromatic to — the — system —; - in some — cases, —in which fluidized bed reactors are employed, it may be advantageous to employ recycling ratios less than 20: 1, less than 10: 1, or less than 5: 1, which can be useful in maintaining a well fluidized bed.
EXAMPLE 17
This example describes the use of a synthesized zeolite catalyst comprising gallium (GaAlMFI) in the fast catalytic pyrolysis of furan in a fixed bed stream reactor.
Η-GaAlMFI was synthesized using the methods described by Choudhary et al. in H-Gallosilicate (MFI)
Propane Aromatization Catalyst: Influence of Si / Ga Ratio on Acidity, Activity and Deactivation Due to Coking, J. Catai. 158 (1996) pages 34-50. The precursor solutions for
Η-GaAlMFI were prepared using new N-silicate (SiO / NaO = 3.22, PQ Corp.), Ganitrate (BDH), Aldrich bromide), deionized water and used to adjust the pH). final reaction in nitrate molar ratios (Sigma-Aldrich), Altetrapropylammonium (TPA-Br, sulfuric acid (which is The composition of the oxide mixture was: 2.5 AI2O3:
114/119
12.5 TPA-Br: 5020 H ₂ O. The reasons for
Si / Al in the mixture were 15, 60 and
20, respectively.
at 180 ° C under pressure, the reaction mixture was crystallized autogenously for 72 hours in an autoclave. After synthesis, the zeolite samples were washed with water and dried
80 ° C. The samples were then calcined in air at 550 ° C for 6 hours to remove occluded organic molecules. The zeolite samples were ionically exchanged with the H form by treatment in 0.1 M NH4NO3 at 70 ° C for hours, followed by filtration, night, and calcination under air at 550 °
Gr
The GaAlMFI type catalyst was synthesized to have an MFI structure including a mixture of
SiO ₂ ,
A1 ₂ O ₃ , and Ga ₂ O ₃ . The catalyst was similar to a conventional ZSM-5 catalyst, with gallium atoms added to the MFI structure at the site of a portion of the aluminum sites.
This type of substitution is often referred to as an isomorphic substitution (for example, in this case, an isomorphic substitution of Al in
Ga).
Without being bound by any particular theory, both Ga and Al are believed to generate Bronsted acid sites, which can balance the negative charge on the catalyst. Acid sites are believed to
Bronsted Ga and Al will have different resistance to acid, and therefore the nature of acidic sites is similar, but the resistance of acidic sites is different. This mixture of types of acid sites can produce a relative increase in catalyst activity, compared to catalysts that do not include a mixture of acid sites.
For the purpose of this example, the catalyst has not been analyzed to verify that Ga is within the structure
115/119 zeolite or not. While it is expected that Ga species are preferably located inside the structure, as an integrated structure, it is possible that some of Ga may be located outside the structure as a form of gallium oxide. In such a case, the Ga outside the structure can also promote the reaction, but not as an acid site.
A fixed flow bed reactor system was built to convert biomass. During the reaction, 10 helium (Airgas CO.) Was used as a vehicle gas, and its flow rate was controlled by a tidal flow controller (controller: —Brooksq — SLA5O-56-S-1BA- B1 — G2Al · ^ control box: Brooks, 0154CFD2B31A). Liquid raw material was introduced into the vehicle gas using a syringe pump (KD Scientific, KDS 100), and was vaporized immediately. The steam raw material was then loaded into a tubular quartz reactor. The quartz reactor was kept in a temperature controlled oven (oven: Lindberg, 55035A, temperature controller: Omega,
CN96211TR). The catalyst bed fixed in the reactor was supported by a quartz frit. The reactor temperature was monitored by a thermocouple inserted through an internal quartz tube to the top surface of the packed bed. The reactor products flowed through an air-bathed condenser to condense heavy products (carbon yield <0.05%). Gas products were collected in gas collection bags. Liquid and gas products were analyzed by GC-FID (Shimadzu 2014). After the reaction, the spent catalyst was regenerated by heating in air at 600 ° C. During regeneration, the effluent was passed through a copper converter (CuO, Sigma Aldrich), followed by a CO ₂ trap (Ascarite, Sigma Aldrich). CO was oxidized to CO ₂ in the copper converter.
116/119
C0 ₂ was trapped by the C0 ₂ trap · Coke yield was obtained by measuring the change in the weight of the C0 ₂ trap.
ZSM-5 catalysts (ZEOLYST, CBV 3024E,
SiO ₂ / Al ₂ O3 = 30) and zeolite catalyst GaAlMFI were used to react furan (Sigma-Aldrich, 99%), under a helium atmosphere in the flow reactor. The reaction conditions were as follows: partial pressure of furan 0.80KPa (6 torr), space speed per hour by weight (WHSV) 10.4 h ¹ , and temperature of 600 ° C. Table 12 summarizes 10 the carbon yield and carbon selectivity of the main products, which contributed to the 90% carbon yield. Tabre / ta 13 shows — the carbon selectivity of each aromatic product over aromatic compounds. Carbon yield was defined as the moles of carbon in a given product type divided by the moles of a carbon fed to the reactor. Carbon selectivity was defined as moles of carbon in one product divided by the sum of moles of carbon in all products.
As shown in Table 12, the choice of catalyst impacted on the types of aromatic compounds produced and their carbon selectivity. The carbon yield of aromatics increased from 11.84
18.15% for GaAlMFI. The selectivity of aromatic compounds increased from
39.1% for GaAlMFI. In addition, income increased significantly,
O, the of
30.30% for ZSM-5 carbon from for ZSM-5 the conversion of furan
0.48 for carbon ZSM-5 but due to an increase in
0.54 for
GaAlMFI. The olefins selectivity amount of carbon did not alter the aromatics produced.
Coke yield also decreased when using coke
GaAlMFI as a catalyst.
117/119
Table 12. Carbon Yield (%) and Carbon Selectivity(%) for catalytic furan pyrolysisCarbon Yield(%) Carbon Selectivity(%) Products ZSM-5 GaAlMFl ZSM-5 GaAlMFl Benzene 3.43 7.38 8.77 15.91 Toluene 3.13 3.22 8.00 6, 94 Ethylbenzene 0.16 0.07 0.40 0.16 Xylene 0.58 0.51 1.47 1.10 Styrene 1.08 1.49 2.76 3.22 Benzofuran 0.79 0.93 2.01 1.99 Indeno 1.77 2.46 4.52 5.31 Naphthalene 0.92 2.08 2.36 4.49CO 5.94 9.15 15.21 19.73 CO ₂ 0.48 0.45 1.23 0.96 Ethylene 3.16 2.33 8.09 5, 03 Propylene 2.87 3.38 7.33 7.28 C ₄ Olefins 0.35 0.41 0.90 0.88 Coke 14.44 12.54 36.94 27.02Aromatic 11.84 18.15 30.30 39, 11 Olefins 6, 38 6, 12 16.32 13.19
Selectivity in relation to the various types of aromatic compounds varied. As shown in Table 13, selectivity for benzene increased from 29% when ZSM-5 was used as the catalyst to 40% when GaAlMFl 5 was used as the catalyst. Its selectivity among aromatics increased from 29% to 41%.
Table 13. Selectivity ofcarbon (%) of aromatics Products ZSM-5 GaAlMFl Benzene 28.95 40.67 Toluene 26, 42 17.76
118/119
Ethylbenzene 1.32 0.40 Xylene 4.86 2.80 Styrene 9, 12 8.22 Benzofuran 6, 63 5, 10 Indeno 14.91 13.57 Naphthalene 7.79 11.48 Total 100.0 100.00
Although various embodiments of the present invention have been described and illustrated here, those of ordinary skill in the art will readily provide for a variety of other means and / or structures for carrying out the functions and / or obtaining the results and / or one or more of the advantages here. described, and each such variation and / or modification is considered to be within the scope of the present invention. More generally, those skilled in the art will easily appreciate that all parameters,
dimensions, materials and configurations described here are intended be exemplary, and that the parameters, dimensions, materials, and / or actual configurations will depend
specific application or applications for which the teachings of the present invention are / are used. Those skilled in the art will recognize, or be able to determine using, no more than routine experimentation, many equivalents to the specific embodiments of the invention described herein. It is, therefore, to be understood that the modalities described above are presented by way of example only and that, within the scope of the appended and equivalent claims thereof, the invention may be practiced in a manner other than those specifically described and claimed. This invention is directed to each individual characteristic, system, article, material, kit, and / or the method described here. In addition, any combination of two or more of these characteristics,
119/119 systems, articles, materials, kits, and / or methods, if such characteristics, systems, articles, materials, kits, and / or methods are not mutually inconsistent, is included within the scope of the present invention.
The indefinite article one, as used here in the specification and in the claims, unless expressly stated to the contrary, should be understood as at least one.

权利要求:
Claims (15)
[1]
1. Method for producing one or more fluid hydrocarbon products from a solid hydrocarbon material characterized by:
feeding a solid hydrocarbonaceous material to a reactor, where the reactor is a fluidized bed reactor and where the solid hydrocarbonaceous material comprises a biomass material;
pyrolize within the reactor at least a portion of the solid hydrocarbon material under sufficient reaction conditions to produce one or more pyrolysis products;
reacting catalytically within the reactor at least a portion of one or more pyrolysis products under sufficient reaction conditions to produce one or more fluid hydrocarbon products comprising olefins and aromatic compounds, wherein the catalytic reaction is catalyzed through a plurality of catalytic particles , and at least 50% of the sum of the total catalyst volume is occupied by particles having maximum cross-sectional dimensions less than 1 micron;
separating at least a portion of the olefins into one or more fluid hydrocarbon products to produce a recycling stream comprising at least the separate olefins and a product stream; and feed at least a portion of the recycle stream to the reactor, where the fluid hydrocarbon product comprises an amount of aromatic compounds of at least 15% of the total amount of the hydrocarbonaceous material used to form the pyrolysis product and which is calculated as the weight of the aromatic compounds present in the fluid hydrocarbon product divided by the weight of the hydrocarbonaceous material used to form the pyrolysis products.
Petition 870180142735, of 10/19/2018, p. 10/14
[2]
2/3
2. Method according to claim 1, characterized in that the solid hydrocarbon material is fed to the reactor at a standardized mass space velocity between 0.01 h ^-1 and 10 h ^-1
[3]
Method according to claim 1, characterized by the ratio between the mass of carbon in the hydrocarbonaceous material and the mass of carbon in the olefins of the recycle stream being between
2: 1 and
20: 1.
[4]
Method according to claim 2, characterized in that it further comprises varying the normalized spatial velocity in mass of the hydrocarbon material to selectively produce products other than fluid hydrocarbons.
[5]
Method according to any one of claims 1 to 4, characterized in that the fluid hydrocarbon products contain more aromatic compounds than olefinic compounds.
[6]
Method according to any one of claims 1 to 5, characterized in that the hydrocarbonaceous material comprises sugarcane bagasse, glucose, wood, corn straw or combinations thereof.
[7]
Method according to any one of claims 1 to 6, characterized in that the priolysis step occurs at a temperature between 500 ° C and 1000 ° C.
[8]
Method according to any one of claims 1 to 7, characterized in that the catalytic reaction step takes place at a temperature between 500 ° C and 1000 ° C.
[9]
9. Method, according to claim 1, characterized by the pyrolysis and catalytic reaction steps taking place in the same reactor.
Petition 870180142735, of 10/19/2018, p. 11/14
3/3
[10]
10. Method according to claim 1, characterized by the pyrolysis and catalytic reaction steps taking place in different reactors.
[11]
Method according to claim 1, characterized in that the hydrocarbon material and the catalyst are fed into one or more feed streams to the reactor, so that the mass ratio between the catalyst and the hydrocarbon material in the one or more feed streams between 0.5: 1 and 20: 1.
[12]
12. Method according to claim 1, characterized in that the residence time of the hydrocarbon material inside the reactor is between 2 and 480 seconds.
[13]
Method according to claim 1, characterized in that the hydrocarbonaceous material comprises a first component and a second component, wherein the first and second components are different; and further understand:
supplying a first and a second catalyst to the reactor, where the first catalyst is selective to catalytically react the first component or a derivative thereof to obtain a fluid hydrocarbon product, and where the second catalyst is selective to catalytically react the second component or a derivative thereof to obtain a fluid hydrocarbon product.
[14]
14. Method according to claim 1, characterized by the contents of the reactor in which the pyrolysis step occurs are heated to a heating rate greater than 50 ° C / s.
[15]
15. Method, according to claim 1, characterized by the contents of the reactor in which the catalytic reaction step occurs are heated to a heating rate greater than 50 ° C / s.

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法律状态:
2018-07-24| B07A| Application suspended after technical examination (opinion) [chapter 7.1 patent gazette]|

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2019-01-29| B09A| Decision: intention to grant [chapter 9.1 patent gazette]|

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2019-03-06| B16A| Patent or certificate of addition of invention granted [chapter 16.1 patent gazette]|Free format text: PRAZO DE VALIDADE: 20 (VINTE) ANOS CONTADOS A PARTIR DE 09/09/2010, OBSERVADAS AS CONDICOES LEGAIS. |

优先权:

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申请号 | 申请日 | 专利标题

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US24101809P| true| 2009-09-09|2009-09-09|

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US61/241,018|2009-09-09|

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PCT/US2010/002472|WO2011031320A2|2009-09-09|2010-09-09|Systems and processes for catalytic pyrolysis of biomass and hydrocarbonaceous materials for production of aromatics with optional olefin recycle, and catalysts having selected particle size for catalytic pyrolysis|

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