U.S. patent application number 14/569284 was filed with the patent office on 2015-06-18 for catalysts and processes for producing p-xylene from biomass.
The applicant listed for this patent is Anellotech, Inc.. Invention is credited to Yu-Ting Cheng, Terry Mazanec, Jian Shi, Ruozhi Song, Jr..
Application Number | 20150166899 14/569284 |
Document ID | / |
Family ID | 52424094 |
Filed Date | 2015-06-18 |
United States Patent
Application |
20150166899 |
Kind Code |
A1 |
Shi; Jian ; et al. |
June 18, 2015 |
Catalysts and Processes for Producing p-xylene from Biomass
Abstract
Biomass is converted to a fluid hydrocarbon product comprising
p-xylene by reaction over a zeolite catalyst. An iron-modified
zeolite catalyst having a siliceous coating and methods of making
the catalyst are also described.
Inventors: |
Shi; Jian; (Pearl River,
NY) ; Cheng; Yu-Ting; (Amherst, MA) ; Song,
Jr.; Ruozhi; (Wilmington, DE) ; Mazanec; Terry;
(Solon, OH) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
Anellotech, Inc. |
Pearl River |
NY |
US |
|
|
Family ID: |
52424094 |
Appl. No.: |
14/569284 |
Filed: |
December 12, 2014 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
61916180 |
Dec 14, 2013 |
|
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|
Current U.S.
Class: |
585/24 ; 422/129;
585/240; 585/242 |
Current CPC
Class: |
C10B 53/02 20130101;
B01J 29/072 20130101; B01J 2229/32 20130101; B01J 2229/123
20130101; C10B 49/22 20130101; B01J 37/0045 20130101; B01J 29/46
20130101; Y02P 30/20 20151101; Y02E 50/10 20130101; C10G 3/49
20130101; C10G 3/45 20130101; C10G 2300/1011 20130101; Y02P 20/145
20151101; C10B 57/06 20130101; C10G 2400/30 20130101; Y02E 50/14
20130101; C10G 1/02 20130101 |
International
Class: |
C10G 1/02 20060101
C10G001/02 |
Claims
1. A process for converting biomass to liquid hydrocarbons,
comprising: feeding biomass into a reactor; heating the biomass in
the presence of an aluminosilicate zeolite catalyst; and wherein
the aluminosilicate zeolite catalyst further comprises at least 0.2
wt % Fe wherein the Fe is not derived from biomass or reactor
walls, or wherein the aluminosilicate zeolite catalyst has been
manufactured to contain at least 0.2 wt % Fe or pretreated to
contain at least 0.2 wt % Fe; and converting the biomass to a
gaseous product stream comprising p-xylene.
2. The process of claim 1 wherein the aluminosilicate zeolite
catalyst further comprises at least 0.5, or at least 1.0, or at
least 1.5 wt % Fe, and in some embodiments up to 10 wt % Fe, up to
5 wt % Fe, or up to 3 wt % Fe.
3. The process of claim 1 wherein the gaseous product stream
comprises at least 10 wt % of aromatic compounds or at least 15 wt
% of aromatic compounds and in some embodiments up to 30 wt %
aromatics, in some embodiments up to 25 wt %; and/or wherein at
least 85% of the xylenes in the gaseous product stream is
p-xylene.
4. The process of claim 1 wherein the aluminosilicate zeolite
catalyst has been pretreated with a siliceous coating, or wherein
the aluminosilicate zeolite catalyst has a siliceous coating.
5. The process of any of claim 1 further comprising: removing the
catalyst from the reactor after it has been used to pyrolyze the
catalysis, heating the used catalyst in the presence of an oxygen
containing gas (preferably O2) to form a regenerated catalyst, and
returning the regenerated catalyst to the reactor, and again using
the catalyst to catalyze the conversion of the biomass to a gaseous
product stream comprising p-xylene.
6. A method of making a catalyst, comprising: providing a zeolite
catalyst; treating the catalyst to increase the iron content; and
applying a siliceous coating to the catalyst; resulting in an
Fe-modified, zeolite catalyst having a siliceous coating.
7. The method of claim 2, comprising: using the Fe-modified,
zeolite catalyst having a siliceous coating to catalyze the
pyrolysis of biomass; and subsequent to the pyrolysis of biomass,
regenerating the catalyst by heating in the presence of an oxygen
containing gas.
8. A hydrocarbon mixture, comprising: a biomass-derived (i.e.,
.sup.14C-containing) mixture of hydrocarbons comprising at least 10
mass % of xylenes; wherein the xylenes are made up of 85 to about
91% p-xylene.
9. The hydrocarbon mixture of claim 8 made by the process of any of
claims 1 to 5.
10. The hydrocarbon mixture of claim 8 comprising catalyst
particles of the type described herein.
11. A chemical system, comprising: a reactor, comprising an
Fe-modified zeolite catalyst; biomass; and a hydrocarbon product
stream comprising at least 10 mass % xylenes wherein at least 80%
of the xylenes are p-xylene.
12. The chemical system of claim 11 wherein the Fe-modified zeolite
catalyst comprises a siliceous coating.
13. The chemical system of claim 11 wherein the zeolite catalyst
comprises ZSM-5.
14. An aluminosilicate zeolite catalyst having a Si/Al molar ratio
of 100 or less, comprising: at least 0.2 wt % Fe wherein the Fe is
not derived from biomass or reactor walls, or wherein the
aluminosilicate zeolite catalyst has been manufactured to contain
at least 0.2 wt % Fe or pretreated to contain at least 0.2 wt % Fe;
and a siliceous coating.
15. The aluminosilicate zeolite catalyst of claim 14 wherein the
catalyst has Fe evenly distributed over the surface as measured by
SEM-EDS.
16. The aluminosilicate zeolite catalyst of claim 14 wherein the
catalyst comprises ZSM-5.
17. The aluminosilicate zeolite catalyst of claim 14 wherein the
catalyst has a surface ratio of Si/Fe in the ratio of 50:1 to 4:1;
preferably 30:1 to 5:1; in some embodiments 20:1 to 7:1.
18. The aluminosilicate zeolite catalyst of claim 14 wherein the Fe
is concentrated in clusters on the surface of the catalyst.
19. The aluminosilicate zeolite catalyst of claim 14 having a
Bronsted acidity of greater than 0.01, 0.05 or greater, preferably
a Bronsted acidity in the range of 0.01 to 0.2, preferably in the
range of 0.04 to 0.15, preferably in the range of 0.05 to 0.1
.mu.mol/mg.
Description
RELATED APPLICATIONS
[0001] This application claims the priority benefit of U.S.
Provisional Patent Application Ser. No. 61/916,180, filed Dec. 14,
2013.
FIELD OF INVENTION
[0002] This invention relates to a method for converting biomass to
a fluid hydrocarbon product comprising p-xylene by reaction over a
zeolite catalyst. An iron-modified zeolite catalyst having a
siliceous coating and methods of making the catalyst are also
described.
BACKGROUND
[0003] p-Xylene is used as a starting material for plasticizers and
polyester fibers. The oxidation of p-xylene is used to commercially
synthesize terephthalic acid. Further esterification of the acid
with methanol forms dimethyl terephthalate. Both monomers may be
used in the production of polyethylene terephthalate (PET) plastic
bottles and polyester clothing.
[0004] p-Xylene may be the most valuable of the xylenes (i.e., o-,
m- and p-xylenes). However, during the catalytic pyrolysis of
various hydrocarbonaceous materials, the xylenes may be formed with
the m-xylene selectivity and/or o-xylene selectivity being the same
as or higher than the p-xylene selectivity. The p-xylene that is
produced may also isomerize to m-xylene and/or o-xylene. As a
result, xylenes with undesirably high selectivities to m-xylene
and/or o-xylene may be formed. Thus, there is a well-known need for
the production of p-xylene or xylenes with a relatively high
selectivity to p-xylene.
SUMMARY
[0005] In a first aspect, the invention provides a process for
converting biomass to liquid hydrocarbons, comprising: feeding
biomass into a reactor; heating the biomass in the presence of an
aluminosilicate zeolite catalyst; and, wherein the aluminosilicate
zeolite catalyst further comprises at least 0.2 wt % Fe wherein the
Fe is not derived from biomass or reactor walls, or wherein the
aluminosilicate zeolite catalyst has been manufactured to contain
at least 0.2 wt % Fe or pretreated to contain at least 0.2 wt % Fe;
and converting the biomass to a gaseous product stream comprising
p-xylene.
[0006] In the present invention, an aluminosilicate zeolite means a
zeolite having a Si/Al molar ratio of 200:1 to 1:1, more preferably
150:1 to 1.5:1, in some embodiments 100:1 to 1:1, in some
embodiments 120:1 to 5:1. One preferred zeolite is ZSM-5. The
phrase "at least 0.2 wt % Fe" is determined by elemental analysis
of catalyst separated from biomass and separated from any ash, to
the extent practicable, where the elemental analysis is preferably
conducted by ICP. The term "pretreated" means treated prior to use
in a catalytic pyrolysis process. The phrase "wherein the Fe is not
derived from biomass or reactor walls" is to distinguish Fe made in
catalyst preparation or catalyst pretreatment from Fe that may be
deposited on the catalyst or may co-occur in ash as a result of the
pyrolysis process. It is believed that Fe that is not derived from
biomass or reactor walls will not be as effective as Fe added
during catalyst manufacture or pretreatment, and, in any case, such
adventitious Fe would be uncontrolled, variable, and unavailable
for initial processing. Such adventitious Fe will be
distinguishable from Fe added during catalyst preparation or
catalyst pretreatment by characterization techniques such as SEM
and Mossbauer spectroscopy.
[0007] In some preferred embodiments, the aluminosilicate zeolite
catalyst further comprises at least 0.5, or at least 1.0, or at
least 1.5 wt % Fe, and in some embodiments up to 10 wt % Fe, up to
5 wt % Fe, or up to 3 wt % Fe. The Fe is not derived, from biomass
or reactor walls; or the aluminosilicate zeolite catalyst has been
manufactured to contain or pretreated to contain the stated amount
or ranges of Fe.
[0008] Preferably, the gaseous product stream comprises at least 10
wt % of aromatic compounds or at least 15 wt % of aromatic
compounds and in some embodiments up to 30 wt % aromatics, in some
embodiments up to 25 wt %; and/or wherein at least 85% of the
xylenes in the gaseous product stream is p-xylene. Note that wt %
is identical to mass %.
[0009] In some preferred embodiments, the aluminosilicate zeolite
catalyst has a siliceous coating. A siliceous coating can be
applied to a zeolite surface by reaction with silicones or
siloxanes as described elsewhere herein and in the literature. A
siliceous coating can be identified can be identified by a higher
(at least 10% higher or at least 30% higher or at least 50% higher
or at least 100% higher Si/Al ratio) in the exterior 50 A (or
exterior 100 A) as compared to the Si/Al ratio at greater depths in
the catalyst. The preferred technique for analyzing the Si/Al ratio
is SEM/XPS before and after sputtering off 50 or 100 A.
[0010] Some embodiments further comprise removing the catalyst from
the reactor after it has been used to pyrolyze the catalysis,
heating the used catalyst in the presence of an oxygen containing
gas (preferably O2) to form a regenerated catalyst, and returning
the regenerated catalyst to the reactor, and again using the
catalyst to catalyze the conversion of the biomass to a gaseous
product stream comprising p-xylene. Surprisingly, the regenerated
catalyst was found to have superior selectivity to p-xylene as
compared to the freshly prepared catalyst.
[0011] In another aspect, the invention provides a method of making
a catalyst, comprising: providing a zeolite catalyst; treating the
catalyst to increase the iron content and applying a siliceous
coating to the catalyst. This results in an iron-modified, zeolite
catalyst having a siliceous coating. The step of applying a
siliceous coating can be conducted before, during or after step of
a treating the catalyst to increase the iron content; in some
embodiments the invention can be characterized by applying the
siliceous coating prior to treatment to increase Fe content.
Preferably the zeolite catalyst is an aluminosilicate catalyst
having a Si/Al ratio of 100 or less.
[0012] In any of the inventive catalysts, the Fe-modified, zeolite
catalyst having a siliceous coating is used to catalyze the
pyrolysis of biomass; and subsequent to the pyrolysis of biomass,
the invention can include a step of regenerating the catalyst by
heating in the presence of an oxygen containing gas.
[0013] In another aspect, the invention provides a hydrocarbon
mixture, comprising: a biomass-derived (i.e., .sup.14C-containing)
mixture of hydrocarbons comprising at least 10 mass % of xylenes;
wherein the xylenes are made up of 85 to about 91% p-xylene; or
comprising at least 1.5 mass % of xylenes; wherein the xylenes are
made up of at least 85% p-xylene. The invention includes a
hydrocarbon mixture made by any of the processes described herein.
The invention also includes a hydrocarbon mixture comprising
catalyst particles of the type described herein.
[0014] In a further aspect, the invention provides a chemical
system, comprising: a reactor, comprising: an iron-modified zeolite
catalyst, biomass, and a hydrocarbon product stream comprising at
least 10 mass % xylenes wherein at least 80% of the xylenes are
p-xylene; or comprising at least 1.5 mass % of xylenes; wherein the
xylenes are made up of at least 85% p-xylene. Preferably, the
Fe-modified zeolite catalyst comprises a siliceous coating.
Preferably, the Fe-modified zeolite catalyst comprises ZSM-5. The
hydrocarbon product stream may be a stream condensed from the
gaseous product stream (where the gaseous stream may include
suspended liquid droplets and solid particulates).
[0015] In another aspect, the invention provides an aluminosilicate
zeolite catalyst having a Si/Al molar ratio of 100 or less,
comprising: at least 0.2 wt % Fe wherein the Fe is not derived from
biomass or reactor walls, or wherein the aluminosilicate zeolite
catalyst has been manufactured to contain at least 0.2 wt % Fe or
pretreated to contain at least 0.2 wt % Fe; and a siliceous
coating. In some embodiments, the catalyst comprises one or more of
the following characteristics: Fe evenly distributed over the
surface as measured by SEM-EDS; the catalyst comprises ZSM-5;
wherein the catalyst has a surface ratio of Si/Fe in the ratio of
50:1 to 4:1; preferably 30:1 to 5:1; in some embodiments 20:1 to
7:1; wherein the Fe is concentrated in clusters on the surface of
the catalyst; wherein the catalyst has a Bronsted acidity of
greater than 0.01, 0.05 or greater, or a Bronsted acidity in the
range of 0.01 to 0.2, preferably in the range of 0.04 to 0.15,
preferably in the range of 0.05 to 0.1 .mu.mol/mg; wherein the
moles of Bronsted acid sites deactivated by the siliceous coating
is at least 0.015 or at least 0.03 per mole of Si added as measured
by desorption of isopropyl amine (IPA) in a temperature programmed
desorption experiment.
[0016] The invention includes any combination of the inventive
aspects. For example, the catalyst as defined in the descriptions
above can be present in any of the inventive processes or
systems.
[0017] Other advantages and novel features of the present invention
will become apparent from the following detailed description of
various non-limiting embodiments of the invention when considered
in conjunction with the accompanying figures. In cases where the
present specification and a document incorporated by reference
include conflicting and/or inconsistent disclosure, the present
specification shall control.
GLOSSARY
[0018] All ranges and ratio limits disclosed in the specification
and claims may be combined in any manner. It is to be understood
that unless specifically stated otherwise, references to "a," "an,"
and/or "the" may include one or more than one, and that reference
to an item in the singular may also include the item in the
plural.
[0019] The phrase "and/or" should be understood to mean "either or
both" of the elements so conjoined, i.e., elements that are
conjunctively present in some cases and disjunctively present in
other cases. Other elements may optionally be present other than
the elements specifically identified by the "and/or" clause,
whether related or unrelated to those elements specifically
identified unless clearly indicated to the contrary. Thus, as a
non-limiting example, a reference to "A and/or B," when used in
conjunction with open-ended language such as "comprising" can
refer, in one embodiment, to A without B (optionally including
elements other than B); in another embodiment, to B without A
(optionally including elements other than A); in yet another
embodiment, to both A and B (optionally including other elements);
etc.
[0020] The word "or" should be understood to have the same meaning
as "and/or" as defined above. For example, when separating items in
a list, "or" or "and/or" shall be interpreted as being inclusive,
i.e., the inclusion of at least one, but also including more than
one, of a number or list of elements, and, optionally, additional
unlisted items.
[0021] The term "aromatic compound" is used to refer to a
hydrocarbon compound comprising one or more aromatic groups such
as, for example, single aromatic ring systems (e.g., benzyl,
phenyl, etc.) and fused polycyclic aromatic ring systems (e.g.
naphthyl, 1,2,3,4-tetrahydronaphthyl, etc.). Examples of aromatic
compounds include, but are not limited to, benzene, toluene,
indane, indene, 2-ethyl toluene, 3-ethyl toluene, 4-ethyl toluene,
trimethyl benzene (e.g., 1,3,5-trimethyl benzene, 1,2,4-trimethyl
benzene, 1,2,3-trimethyl benzene, etc.), ethylbenzene,
methylbenzene, propylbenzene, xylenes (e.g., p-xylene, m-xylene,
o-xylene, etc.), naphthalene, methyl-naphthalene (e.g., 1-methyl
naphthalene, anthracene, 9.10-dimethylanthracene, pyrene,
phenanthrene, dimethyl-naphthalene (e.g., 1,5-dimethylnaphthalene,
1,6-dimethylnaphthalene, 2,5-dimethylnaphthalene, etc.),
ethyl-naphthalene, hydrindene, methyl-hydrindene, and
dymethyl-hydrindene. Single ring and/or higher ring aromatics may
be produced in some embodiments. The aromatic compounds may have
carbon numbers from, for example, C5-C14, C6-C8, C6-C12, C8-C12,
C10-C14.
[0022] The term "biomass" refers to living and recently dead
biological material. In accordance with the inventive method,
biomass may be converted, for example, to liquid fuel (e.g.,
biofuel or biodiesel) or to other fluid hydrocarbon products.
Biomass may include trees (e.g., wood) as well as other vegetation;
agricultural products and agricultural waste (e.g., corn stover,
bagasse, fruit, garbage, silage, etc.); energy crops (e.g.
switchgrass, miscanthus); algae and other marine plants; metabolic
wastes (e.g., manure, sewage); and cellulosic urban waste. Biomass
may be considered as comprising material that recently participated
in the carbon cycle so that the release of carbon in a combustion
process may result in no net increase averaged over a reasonably
short period of time. For this reason, peat, lignite, coal, shale
oil or petroleum may not be considered as being biomass as they
contain carbon that may not have participated in the carbon cycle
for a long time and, as such, their combustion may result in a net
increase in atmospheric carbon dioxide. The term biomass may refer
to plant matter grown for use as biofuel, but may also include
plant or animal matter used for production of fibers, chemicals,
heat, and the like. Biomass may also include biodegradable waste or
byproducts that can be burnt as fuel or converted to chemicals.
These may include municipal waste, green waste (the biodegradable
waste comprised of garden or park waste such as grass or flower
cuttings, hedge trimmings, and the like), byproducts of farming
including animal manures, food processing wastes, sewage sludge,
black liquor from wood pulp or algae, and the like. Biomass may be
derived from plants, including miscanthus, spurge, sunflower,
switchgrass, hemp, corn (maize), poplar, willow, sugarcane, and oil
palm (palm oil), and the like. Biomass may be derived from roots,
stems, leaves, seed husks, fruits, and the like. The particular
plant or other biomass source used may not be important to the
fluid hydrocarbon product produced in accordance with the inventive
method, although the processing of the biomass may vary according
to the needs of the reactor and the form of the biomass.
[0023] The term "catalytic pyrolysis" refers to pyrolysis performed
in the presence of a catalyst. Catalytic fast pyrolysis (CFP), is a
process that may be used to convert a hydrocarbonaceous material
(e.g., biomass) into a fluid hydrocarbon product using rapid
heating rates in the presence of a catalyst. With the inventive
method, the fluid hydrocarbon product comprises p-xylene, and may
further comprise additional aromatics, olefins, and the like.
[0024] Contact time is the residence time of a material in a
reactor or other device, when measured or calculated under standard
conditions of temperature and pressure (i.e., 0.degree. C. and 100
kPa absolute pressure). In some cases contact time can be expressed
for an individual component while in other cases contact time can
be expressed based on more than one component and in other cases
contact time can be expressed based on the entire reaction mixture
of all components. For example, a 2 liter reactor to which is fed 3
standard liters per minute of gas A has a contact time of 2/3
minute, or 40 seconds for gas A. For a chemical reaction, contact
time or residence time is based on the volume of the reactor where
substantial reaction is occurring; and would exclude volume where
substantially no reaction is occurring such as an inlet or an
exhaust conduit. For catalyzed reactions, the volume of a reaction
chamber is the volume where catalyst is present.
[0025] A hydrocarbonaceous feed material may comprise a solid, a
semi-solid, a liquid, or a mixture of two or more thereof. The
solids content of the hydrocarbonaceous feed may be up to about
100% by weight, or from about 30% to about 100% by weight, or from
about 50% to about 100%, or from about 70% to about 100%, or from
90% to about 100%, or from about 95% to about 100%, or from about
98% to about 100%, or from about 30% to about 95%, or from about
50% to about 95%, or from about 70% to about 95%, or from about 80%
to about 95%, or from about 85% to about 95%, or from about 90% to
about 95% by weight. The carbon content of the hydrocarbonaceous
feed may be up to about 90% by weight, or from 20% to 90% by
weight, or from 25% to 75%, or from 30% to 65%, or from 35% to 60%,
or from 40 to 50% by weight. The hydrocarbonaceous material may
comprise biomass. The hydrocarbonaceous material may comprise
plastic waste, recycled plastics, agricultural solid waste,
municipal solid waste, food waste, animal waste, carbohydrates,
lignocellulosic materials, xylitol, glucose, cellobiose, cellulose,
hemi-cellulose, lignin, sugar cane bagasse, glucose, wood, corn
stover, or a mixture of two or more thereof. The hydrocarbonaceous
material may comprise furan, 2-methylfuran, furfural, ethylene
glycol, glycerine, or any combination of these. The
hydrocarbonaceous material may comprise pinewood. The
hydrocarbonaceous material may comprise pyrolysis oil derived from
biomass, a carbohydrate derived from biomass, an alcohol derived
from biomass, a biomass extract, a pretreated biomass, a digested
biomass product, or a mixture of two or more thereof. Mixtures of
two or more of any of the foregoing may be used.
[0026] The term "conversion of a reactant" may refer to the
reactant mole or mass change between a material flowing into a
reactor and a material flowing out of the reactor divided by the
moles or mass of reactant in the material flowing into the reactor.
For example, if 100 g of ethylene are fed to a reactor and 30 g of
ethylene are flowing out of the reactor, the conversion is
[(100-30)/100]=70% conversion of ethylene.
[0027] The term "fluid" may refer to a gas, a liquid, a mixture of
a gas and a liquid, or a gas or a liquid containing dispersed
solids, liquid droplets and/or gaseous bubbles. The terms "gas" and
"vapor" have the same meaning and are sometimes used
interchangeably. In some embodiments, it may be advantageous to
control the residence time of the fluidization fluid in the
reactor. The fluidization residence time of the fluidization fluid
is defined as the volume of the reactor divided by the volumetric
flow rate of the fluidization fluid under process conditions of
temperature and pressure.
[0028] The term "fluidized bed reactor" may be used to refer to
reactors comprising a vessel that contains a granular solid
material (e.g., silica particles, catalyst particles, etc.), in
which a fluid (e.g., a gas or a liquid) is passed through the
granular solid material at velocities sufficiently high as to
suspend the solid material and cause it to behave as though it were
a fluid. The term "circulating fluidized bed reactor" may be used
to refer to fluidized bed reactors in which the granular solid
material is passed out of the reactor, circulated through a line in
fluid communication with the reactor, and recycled back into the
reactor. Bubbling fluidized bed reactors, circulating fluidized bed
reactors, or turbulent fluidized bed reactors may be used. Examples
of fluidized bed reactors, circulating fluidized bed reactors,
bubbling and turbulent fluidized bed reactors are described in
Fluidization Engineering, 2nd Edition, D. Kunii and O. Levenspiel,
Butterworth-Heinemann, 1991, Chapter 3, pages 61-94, these pages
being incorporated herein by reference.
[0029] The terms "olefin" or "olefin compound" (a.k.a. "alkenes")
may be used to refer to any unsaturated hydrocarbon containing one
or more pairs of carbon atoms linked by a double bond. Olefins may
include both cyclic and acyclic (aliphatic) olefins, in which the
double bond is located between carbon atoms forming part of a
cyclic (closed-ring) or of an open-chain grouping, respectively. In
addition, olefins may include any suitable number of double bonds
(e.g., monoolefins, diolefins, triolefins, etc.). Examples of
olefin compounds may include ethene, propene, allene (propadiene),
1-butene, 2-butene, isobutene (2 methyl propene), butadiene, and
isoprene, among others. Examples of cyclic olefins may include
cyclopentene, cyclohexane, cycloheptene, among others. Aromatic
compounds such as toluene are not considered olefins; however,
olefins that include aromatic moieties are considered olefins, for
example, benzyl acrylate or styrene.
[0030] The term "overall residence time" refers to the volume of a
reactor or device or specific portion of a reactor or device
divided by the exit flow of all gases out of the reactor or device
including fluidization gas, products, and impurities, measured or
calculated at the average temperature of the reactor or device and
the exit pressure of the reactor or device.
[0031] As used herein, the term "pore size" is used to refer to the
smallest cross-sectional diameter of a pore. The smallest
cross-sectional diameter of a pore may correspond to the smallest
cross-sectional dimension (e.g., a cross-sectional diameter) as
measured perpendicularly to the length of the pore. In some
embodiments, a catalyst with an "average pore size" or a "pore size
distribution" of X refers to a catalyst in which the average of the
smallest cross-sectional diameters of the pores within the catalyst
is about X. It should be understood that "pore size" or "smallest
cross sectional diameter" of a pore as used herein refers to the
Norman radii adjusted pore size well known to those skilled in the
art. Determination of Norman radii adjusted pore size is described,
for example, in Cook, M.; Conner, W. C., "How big are the pores of
zeolites?" Proceedings of the International Zeolite Conference,
12th, Baltimore, Jul. 5-10, 1998; (1999), 1, pp 409-414, which is
incorporated herein by reference in its entirety. As a specific
exemplary calculation, the atomic radii for ZSM-5 pores are about
5.5-5.6 .ANG., as measured by x-ray diffraction. In order to adjust
for the repulsive effects between the oxygen atoms in the catalyst,
Cook and Conner have shown that the Norman adjusted radii are 0.7
.ANG. larger than the atomic radii (about 6.2-6.3 .ANG.).
[0032] One of ordinary skill in the art will understand how to
determine the pore size (e.g., minimum pore size, average of
minimum pore sizes) in a catalyst. For example, x-ray diffraction
(XRD) can be used to determine atomic coordinates. XRD techniques
for the determination of pore size are described, for example, in
Pecharsky, V. K. et al, "Fundamentals of Powder Diffraction and
Structural Characterization of Materials," Springer
Science+Business Media, Inc., New York, 2005, incorporated herein
by reference in its entirety. Other techniques that may be useful
in determining pore sizes (e.g., zeolite pore sizes) include, for
example, helium pycnometry or low pressure argon adsorption
techniques. These and other techniques are described in Magee, J.
S. et al, "Fluid Catalytic Cracking: Science and Technology,"
Elsevier Publishing Company, Jul. 1, 1993, pp. 185-195, which is
incorporated herein by reference. Pore sizes of mesoporous
catalysts may be determined using, for example, nitrogen adsorption
techniques, as described in Gregg, S. J. at al, "Adsorption,
Surface Area and Porosity," 2nd Ed., Academic Press Inc., New York,
1982 and Rouquerol, F. et al, "Adsorption by powders and porous
materials. Principles, Methodology and Applications," Academic
Press Inc., New York, 1998, both incorporated herein by reference
in their entirety. Unless otherwise indicated, pore sizes referred
to herein are those determined by x-ray diffraction corrected as
described above to reflect their Norman radii adjusted pore
sizes.
[0033] The terms "pyrolysis" and "pyrolyzing" refer to the
transformation of a material (e.g., a solid hydrocarbonaceous
material) into one or more other materials (e.g., volatile organic
compounds, gases, coke, etc.) by heat, without oxygen or other
oxidants or without significant amounts of oxygen or other
oxidants, and with or without the use of a catalyst.
[0034] The term "reactant residence time" of a reactant in the
reactor is defined as the amount of time the reactant spends in the
reactor. Residence time may be based on the feed rate of reactant
and is independent of rate of reaction. The reactant residence time
of the reactants in a reactor may be calculated using different
methods depending upon the type of reactor being used. For gaseous
reactants, where flow rate into the reactor is known, this is
typically a simple calculation. In the case of solid reactants in
which the reactor comprises a packed bed reactor into which only
reactants are continuously fed (i.e. no carrier or fluidizing flow
is utilized), the reactant residence time in the reactor may be
calculated by dividing the volume of the reactor by the volumetric
flow rate of the hydrocarbonaceous material and fluid hydrocarbon
product exiting the reactor. In cases where the reaction takes
place in a reactor that is closed to the flow of mass during
operation (e.g., a batch reactor), the batch residence time of the
reactants in such may be reactor is defined as the amount of time
elapsing between the time at which the temperature in the reactor
containing the reactants reaches a level sufficient to commence a
pyrolysis reaction (e.g., for CFP, typically about 300.degree. C.
to about 1000.degree. C. for many typical hydrocarbonaceous
feedstock materials) and the time at which the reactor is quenched
(e.g., cooled to a temperature below that sufficient to support
further pyrolysis--e.g. typically about 300.degree. C. to about
1000.degree. C. for many hydrocarbonaceous feedstock
materials).
[0035] The residence time of the catalyst in a fluidized bed
reactor may be defined as the volume of the reactor filled with
catalyst divided by the volumetric flow rate of the catalyst
through the reactor. For example if a 3 liter reactor contains 2
liters of catalyst and a flow of 0.4 liters per minute of catalyst
is fed through the reactor, i.e., both fed and removed, the
catalyst residence time will be 2/0.4 minutes, or 5 minutes.
[0036] The term "silicon-containing compound" is used herein to
refer to any compound that contains one or more Si--O groups. The
silicon-containing compound may be a silicate containing one or
more of SiO44-, Si2O76- or Si6O1812-groups. These may include one
or more tetraorthosilicates. The silicon-containing compound may
include one or more siloxanes containing one or more silicon-oxygen
backbones (--Si--O--Si--O--) with organic (e.g., hydrocarbon) side
groups attached to the silicon atoms. These may include one or more
siloxane polymers (e.g., polydimethyl siloxane). The
silicon-containing compound may be a straight chain, branched chain
or cyclical compound. The silicon-containing compound may be
monomeric, oligomeric or polymeric. The silicon-containing compound
may comprise a compound containing at least one group represented
by the formula
##STR00001##
The silicon-containing compound may be represented by the
formula:
##STR00002##
wherein R.sub.1 and R.sub.2 independently comprise hydrogen,
halogen, hydroxyl, alkyl, alkoxyl, halogenated alkyl, aryl,
halogenated aryl, aralkyl, halogenated aralkyl, alkaryl or
halogenated alkaryl; and n is a number that is at least 2. R.sub.1
and/or R.sub.2 may comprise methyl, ethyl or phenyl. n may be a
number in the range from about 3 to about 1000.
[0037] The term "selectivity" refers to the amount of production of
a particular product in comparison to a selection of products.
Selectivity to a product may be calculated by dividing the amount
of a particular product by the amount of a number of products
produced. For example, if 75 grams of aromatics are produced in a
reaction and 20 grams of benzene are found in these aromatics, on a
mass basis the selectivity to benzene amongst aromatic products is
20/75=26.7%. Selectivity may be calculated on a mass basis, as in
the aforementioned example, or it may be calculated on a carbon
basis where the selectivity is calculated by dividing the amount of
carbon that is found in a particular product by the amount of
carbon that is found in a selection of products. Unless specified
otherwise, for reactions involving biomass as a reactant,
selectivity is on a mass basis. The carbon selectivities for
various materials can be determined using the following
equations:
Overall selectivity = moles of carbon in a product moles of carbon
in all products .times. 100 % ##EQU00001## Aromatic selectivity =
moles of carbon in an aromatic product moles of carbon in all
aromatic products .times. 100 % ##EQU00001.2## Olefin selectivity =
moles of carbon in an olefinic product moles of carbon in all
olefins products .times. 100 % ##EQU00001.3## p - Xylene
selectivity in xylenes = moles of p - xylene isomer moles of all
xylene isomers .times. 100 % ##EQU00001.4##
[0038] The term "yield" is used herein to refer to the amount of a
product flowing out of a reactor divided by the amount of reactant
flowing into the reactor, usually expressed as a percentage or
fraction. Yields are often calculated on a mass basis, carbon
basis, or on the basis of a particular feed component. Mass yield
is the mass of a particular product divided by the weight of feed
used to prepare that product. For example, if 500 grams of biomass
is fed to a reactor and 45 grams of p-xylene is produced, the mass
yield of p-xylene would be 45/500=9% p-xylene. Carbon yield is the
mass of carbon found in a particular product divided by the mass of
carbon in the feed to the reactor. For example, if 500 grams of
biomass that contains 40% carbon is reacted to produce 45 g of
p-xylene that contains 90.6% carbon, the carbon yield is
[(45*0.906)/(500*0.40)]=20.4%. Carbon yield from biomass is the
mass of carbon found in a particular product divided by the mass of
carbon fed to the reactor in a particular feed component. For
example, if 500 grams of biomass containing 40% carbon and 100
grams of CO2 are reacted to produce 40 g of p-xylene (containing
90.6% carbon), the carbon yield on biomass is
[(40*0.906)/(500*0.40)]=18.1%; note that the mass of CO2 does not
enter into the calculation.
BRIEF DESCRIPTION OF THE DRAWINGS
[0039] FIG. 1. is a schematic illustration of a CFP process for
converting a solid hydrocarbonaceous material to a fluid
hydrocarbon product
[0040] FIG. 2 is a graph of the acidity of silicone treated zeolite
vs the amount of SiO.sub.2 added to the zeolite for two different
zeolites and two different coating compounds.
DETAILED DESCRIPTION
[0041] Referring to FIG. 1, feed stream 10 includes a solid
hydrocarbonaceous material (typically biomass) that can be fed to
reactor 20. Certain solid hydrocarbonaceous materials may also
comprise relatively minor proportions of other elements such as
nitrogen and sulfur. The feed streams to the reactor may be free of
olefins, or may contain olefins in an insignificant amount (e.g.,
such that olefins make up less than about 1 wt %, less than about
0.1 wt %, or less than about 0.01 wt % of the total weight of
reactant fed to the reactor). In other embodiments, however,
olefins may be present in one or more reactant feed streams.
[0042] The solid hydrocarbonaceous material feed composition (e.g.,
in feed stream 10 of FIG. 1) may comprise a mixture of solid
hydrocarbonaceous material and a catalyst. The mixture may
comprise, for example, a solid catalyst and a solid
hydrocarbonaceous material. In other embodiments, a catalyst may be
provided separately from the solid hydrocarbonaceous material
(e.g., by co-feeding the catalyst via an independent catalyst
inlet). A variety of catalysts may be used. For example, in some
instances, zeolite catalysts with varying molar ratios of silica to
alumina, and/or varying pore sizes and/or pore opening sizes,
and/or varying catalytically active metals and/or metal oxides, may
be used.
[0043] A hydrocarbonaceous material may be fed to a reactor (e.g.,
a fluidized-bed reactor) where the hydrocarbonaceous material may
first thermally decompose to form one or more pyrolysis products.
The pyrolysis products may comprise one or more pyrolysis vapors.
These pyrolysis products may react in the presence of a modified
zeolite catalyst to form one or more aromatic compounds as well as
olefin compounds, water, CO, and CO.sub.2. The modified zeolite
catalyst may comprise acid sites on its external surface and in its
pores. The modified zeolite may be modified to change the number of
acid sites on its external surface and may be modified to reduce
the size or number of the pore mouth openings. The modified zeolite
may be modified by the incorporation of one or more promoter
elements to improve catalyst stability, activity, or selectivity.
The pyrolysis products may enter the pores in the modified zeolite
catalyst where they may undergo reaction or react on the surface of
the modified zeolite. The products formed in the catalyst pores may
then diffuse out of the pores. The aromatic compounds may comprise
p-xylene or xylenes with a relatively high selectivity towards
p-xylene. Advantages of this process may include one or more of the
following: 1) all the desired chemistry may occur in a single-step
process, 2) the process may use a relatively inexpensive zeolite
catalyst, 3) p-xylene may be produced at a relatively high level of
production and selectivity amongst the xylenes, and 4) the
activity, selectivity, or activity and selectivity of the catalyst
may be maintained or improved with repeated cycling.
[0044] The hydrocarbonaceous material may comprise solids of any
suitable size. In some cases, it may be advantageous to use
hydrocarbonaceous solids with relatively small particle sizes.
Small-particle solids may, in some instances, react more quickly
than larger solids due to their relatively higher surface area to
volume ratios compared to larger solids. In addition, small
particle sizes may allow for more efficient heat transfer within
each particle and/or within the reactor volume. This may prevent or
reduce the formation of undesired reaction products. Moreover,
small particle sizes may provide for increased solid-gas and
solid-solid contact, leading to improved heat and mass transfer. In
some embodiments, the mass average particle size of the solid
hydrocarbonaceous material is less than about 5 mm, less than about
2 mm, less than about 1 mm, less than about 500 microns, less than
about 250 microns (60 mesh), less than about 149 microns (100
mesh), less than about 105 microns (140 mesh), less than about 88
microns (170 mesh), less than about 74 microns (200 mesh), less
than about 53 microns (270 mesh), or less than about 37 microns
(400 mesh), or smaller.
[0045] It may be desirable to employ a feed material with an
average particle size above a minimum amount in order to reduce the
pressure required to pass the solid hydrocarbonaceous feed material
through the reactor. For example, it may be desirable to use a
solid hydrocarbonaceous feed material with an average particle size
of at least about 37 microns (400 mesh), at least about 53 microns
(270 mesh), at least about 74 microns (200 mesh), at least about 88
microns (170 mesh), at least about 105 microns (140 mesh), at least
about 149 microns (100 mesh), at least about 250 microns (60 mesh),
at least about 0.5 mm, a least about 1 mm, at least about 2 mm, at
least about 5 mm, or higher.
[0046] In some embodiments, catalyst and hydrocarbonaceous material
may be present in any suitable ratio. For example, the catalyst and
hydrocarbonaceous material may be present in any suitable mass
ratio in cases where the feed composition (e.g., through one or
more feed streams comprising catalyst and hydrocarbonaceous
material or through separate catalyst and hydrocarbonaceous
material feed streams), comprises catalyst and hydrocarbonaceous
material (e.g., circulating fluidized bed reactors). As another
example, in cases where the reactor is initially loaded with a
mixture of catalyst and hydrocarbonaceous material (e.g., a batch
reactor), the catalyst and hydrocarbonaceous material may be
present in any suitable mass ratio. In some embodiments involving
circulating fluidized bed reactors, the mass ratio of the catalyst
to hydrocarbonaceous material in the feed stream--i.e., in a
composition comprising a catalyst and a hydrocarbonaceous material
provided to a reactor--may be at least about 0.5:1, at least about
1:1, at least about 2:1, at least about 5:1, at least about 10:1,
at least about 15:1, at least about 20:1, or higher. In some
embodiments involving circulating fluidized bed reactors, the mass
ratio of the catalyst to hydrocarbonaceous material in the feed
stream may be less than about 0.5:1, less than about 1:1, less than
about 2:1, less than about 5:1, less than about 10:1, less than
about 15:1, or less than about 20:1; or from about 0.5:1 to about
20:1, from about 1:1 to about 20:1, or from about 5:1 to about
20:1. Employing a relatively high catalyst to hydrocarbonaceous
material mass ratio may facilitate introduction of the volatile
organic compounds, formed from the pyrolysis of the feed material,
into the catalyst before they thermally decompose to coke. Not
wishing to be bound by any theory, this effect may be at least
partially due to the presence of a stoichiometric excess of
catalyst sites within the reactor.
[0047] The reactor may comprise a continuously stirred tank
reactor, a batch reactor, a semi-batch reactor, a fixed bed
reactor, or a fluidized bed reactor. Advantageously, the reactor
may comprise a fluidized bed reactor. The catalytic reaction step
may be achieved by co-feeding the catalyst with the
hydrocarbonaceous material. The catalyst may be fed separately.
Part of the catalyst may be fed with the hydrocarbonaceous feed
material and part of the catalyst may be fed separately.
[0048] The inventive method preferably comprises a catalytic fast
pyrolysis (CFP) process. Aspects of a CFP process have been
described in U.S. Pat. No. 8,277,643, U.S. Pat. No. 8,864,984, US
Patent Application 2012/0203042 A1, US Patent Application
2013/0060070 A1, US Patent Application 2014/0027265 A1 and US
Patent Application 2014/0303414 A1 all incorporated herein in full
by reference.
[0049] For CFP processes, particularly advantageous catalysts may
include those containing internal porosity selected according to
pore size (e.g., mesoporous and pore sizes typically associated
with zeolites), e.g., average pore sizes of less than about 100
Angstroms (.ANG.), less than about 50 .ANG., less than about 20
.ANG., less than about 10 .ANG., less than about 5 .ANG., or
smaller. In some embodiments, catalysts with average pore sizes of
from about 5 .ANG. to about 100 .ANG. may be used. In some
embodiments, catalysts with average pore sizes of between about 5.5
.ANG. and about 6.5 .ANG., or between about 5.9 .ANG. and about 6.3
.ANG. may be used. In some cases, catalysts with average pore sizes
of between about 7 .ANG. and about 8 .ANG., or between about 7.2
.ANG. and about 7.8 .ANG. may be used.
[0050] The reactor may be operated at a temperature in the range
from about 400.degree. C. to about 650.degree. C., or from about
500.degree. C. to about 600.degree. C., or from about 525.degree.
C. to about 575.degree. C.
[0051] The hydrocarbonaceous material may be fed to the reactor at
a mass normalized space velocity of up to about 12 hour.sup.-1, or
up to about 6 hour.sup.-1, or up to about 3 hour.sup.-1% or up to
about 1.5 hour.sup.-1, or up to about 0.9 hour.sup.-1, or in the
range from about 0.01 hour.sup.- to about 12 hour.sup.-, or in the
range from about 0.01 to about 2 hour.sup.-1, or in the range from
about 0.01 to about 1.5 hour.sup.-1, or in the range from about
0.01 to about 0.9 hour.sup.-1, or in the range from about 0.01
hour-1 to about 0.5 hour.sup.-1, or in the range from about 0.1
hour.sup.- to about 0.9 hour.sup.-1, or in the range from about 0.1
hour.sup.- to about 0.5 hour.sup.-.
[0052] The reactor may be operated at a pressure of at least about
100 kPa, or at least about 200 kPa, or at least about 300 kPa, or
at least about 400 kPa. The reactor may be operated at a pressure
below about 600 kPa, or below about 400 kPa, or below about 200
kPa. The reactor may be operated at a pressure in the range from
about 100 to about 600 kPa, or in the range from about 100 to about
400 kPa, or in the range from about 100 to about 200 kPa. The
method may be conducted under reaction conditions that minimize
coke production. The pyrolysis product may be formed with less than
about 30 wt %, or less than about 25 wt %, or less than about 20 wt
%, or less than about 10 wt %, of the pyrolysis product being
coke.
[0053] The catalyst may comprise any catalyst suitable for
conducting the catalytically reacting step of the inventive method.
The catalyst may be used to lower the activation energy (increase
the rate) of the reaction conducted in the catalytically reacting
step and/or improve the distribution of products or intermediates
during the reaction (for example, a shape selective catalyst).
Examples of reactions that can be catalyzed include: dehydration,
dehydrogenation, isomerization, hydrogen transfer, aromatization,
decarbonylation, decarboxylation, aldol condensation, and
combinations thereof. The catalyst components may be acidic,
neutral or basic.
[0054] The catalyst may be selected from naturally occurring
zeolites, synthetic zeolites and combinations thereof. The catalyst
may comprise a ZSM-5 zeolite catalyst. The catalyst may comprise
acid sites. These acid sites may also be referred to as
catalytically active sites. Other zeolite catalysts that may be
used may include ferrierite, zeolite Y, zeolite beta, mordenite,
MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)AlPO-31, SSZ-23,
and the like. The catalyst may comprise silica and alumina, and
further comprise one or more additional promoter elements such as
metals and/or metal oxides. Suitable metals and/or oxides may
include, for example, nickel, palladium, silver, platinum,
palladium, titanium, vanadium; chromium, manganese, iron, cobalt,
zinc, copper, gallium, sodium, potassium, magnesium, calcium, the
rare earth elements, i.e., elements 57-71, i.e., La, Ce, Pr, Nd,
Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm, Yb, Lu, or zirconium, hafnium,
tantalum, phosphorus, and/or any of their oxides, among others. In
addition, in some cases, properties of the catalysts (e.g., pore
structure, type and/or number of catalytic sites, etc.) may be
chosen to selectively produce a desired product. Preferable
promoter elements are Ga, La and Fe. A particularly preferable
promoter element is Fe.
[0055] The promoter element(s) may be added in an amount sufficient
to provide an effect on the yield, selectivity, or both yield and
selectivity of the biomass conversion process. The amount of
promoter elements can be from 0.1% by weight to 30% by weight, or
0.2% to 20%, or 0.5% to 10%, or 1.0% to 5%, or at least 0.2%, or at
least 0.5%, or at least 1.0%. The amount of promoter element(s)
added to a zeolite catalyst can be adjusted with respect to the
alumina content of the zeolite or number of acid sites of the
zeolite. The ratio of the added promoter element(s) to the number
of acid sites on a mole to mole basis (moles of promoter element(s)
to moles of acid sites on the unpromoted catalyst) can be from 0.05
to 20.0, or from 0.1 to 10, or from 0.5 to 6, or at least 0.1, or
at least 0.5, or at least 2, or at least 6.0.
[0056] In some embodiments, the mass average diameter (as measured
by conventional sieve analysis) of catalyst objects, which may in
certain instances each comprise a single catalyst particle or in
other instances comprise an agglomerate of a plurality of
particles, may be less than about 5 mm, less than about 2 mm, less
than about 1 mm, less than about 0.5 mm, less than about 250
microns (60 mesh), less than about 149 microns (100 mesh), less
than about 105 microns (140 mesh), less than about 88 microns (170
mesh), less than about 74 microns (200 mesh), less than about 53
microns (270 mesh), or less than about 37 microns (400 mesh), or
smaller.
[0057] The catalyst may comprise particles having a maximum
cross-sectional dimension of less than about 5 microns, less than
about 1 micron, less than about 500 nm, less than about 100 nm,
between about 100 nm and about 5 microns, between about 500 nm and
about 5 microns, between about 100 nm and about 1 micron, or
between about 500 nm and about 1 micron. Catalyst particles having
the dimensions within the ranges noted immediately above may be
agglomerated to form discrete catalyst objects having dimensions
within the ranges noted above. As used here, the "maximum
cross-sectional dimension" of a particle refers to the largest
dimension between two boundaries of a particle. One of ordinary
skill in the art would be capable of measuring the maximum
cross-sectional dimension of a particle by, for example, analyzing
a scanning electron micrograph (SEM) of a catalyst preparation. In
embodiments comprising agglomerated particles, the particles should
be considered separately when determining the maximum
cross-sectional dimensions. In such a case, the measurement may be
performed by establishing imaginary boundaries between each of the
agglomerated particles, and measuring the maximum cross-sectional
dimension of the hypothetical, individuated particles that result
from establishing such boundaries. In some embodiments, a
relatively large number of the particles within a catalyst may have
maximum cross-sectional dimensions that lie within a given range.
For example, in some embodiments, at least about 50%, at least
about 75%, at least about 90%, at least about 95%, or at least
about 99% of the particles within a catalyst have maximum
cross-sectional dimensions of less than about 5 microns, less than
about 1 micron, less than about 500 nm, less than about 100 nm,
between about 100 nm and about 5 microns, between about 500 nm and
about 5 microns, between about 100 nm and about 1 micron, or
between about 500 nm and about 1 micron.
[0058] A relatively large percentage of the volume of the catalyst
can be occupied by particles with maximum cross-sectional
dimensions within a specific range, in some cases. For example, in
some embodiments, at least about 50%, at least about 75%, at least
about 90%, at least about 95%, or at least about 99% of the sum of
the volumes of all the catalyst used is occupied by particles
having maximum cross-sectional dimensions of less than about 5
microns, less than about 1 micron, less than about 500 nm, less
than about 100 nm, between about 100 nm and about 5 microns,
between about 500 nm and about 5 microns, between about 100 nm and
about 1 micron, or between about 500 nm and about 1 micron.
[0059] Using catalysts including particles within a chosen size
distribution indicated above can lead to an increase in the yield
and/or selectivity of aromatic compounds produced by the reaction
of the hydrocarbonaceous material. For example, in some cases,
using catalysts containing particles with a desired size range
(e.g., any of the size distributions outlined above) can result in
an increase in the amount of aromatic compounds in the reaction
product of at least about 5%, at least about 10%, or at least about
20%, relative to an amount of aromatic compounds that would be
produced using catalysts containing particles with a size
distribution outside the desired range (e.g., with a large
percentage of particles larger than 1 micron, larger than 5
microns. etc. as above).
[0060] Alternatively, catalysts may be selected according to pore
size (e.g., mesoporous and pore sizes typically associated with
zeolites), e.g., average pore sizes of less than about 100
Angstroms, less than about 50 Angstroms, less than about 20
Angstroms, less than about 10 Angstroms, less than about 5
Angstroms, or smaller. In some embodiments, catalysts with average
pore sizes of from about 5 Angstroms to about 100 Angstroms may be
used. In some embodiments, catalysts with average pore sizes of
between about 5.5 Angstroms and about 6.5 Angstroms, or between
about 5.9 Angstroms and about 6.3 Angstroms may be used. In some
cases, catalysts with average pore sizes of between about 7
Angstroms and about 8 Angstroms, or between about 7.2 Angstroms and
about 7.8 Angstroms may be used.
[0061] The zeolite catalyst typically comprises silicon and
aluminum. The silicon to aluminum molar ratio may be at least 10:1
or at least 15:1, or at least 25:1, in the range from about 10:1 to
about 240:1, or in the range from about 10:1 to about 40:1, or in
the range from about 20:1 to about 50:1, or about 30:1, or at least
about 30:1. The zeolite catalyst may further comprise nickel,
palladium, silver, platinum, palladium, titanium, vanadium,
chromium, manganese, iron, cobalt, zinc, copper, gallium, sodium,
potassium, magnesium, calcium, zirconium, lanthanum, cerium,
phosphorus, an oxide of one or more thereof, or a mixture of two or
more thereof.
[0062] A screening method may be used to select catalysts with
appropriate pore sizes for the conversion of specific pyrolysis
product molecules. The screening method may comprise determining
the size of pyrolysis product molecules desired to be catalytically
reacted (e.g., the molecular kinetic diameters of the pyrolysis
product molecules). One of ordinary skill in the art may calculate,
for example, the kinetic diameter of a given molecule. The type of
catalyst may then be chosen such that the pores of the catalyst
(e.g., Norman adjusted minimum radii) are sufficiently large to
allow the pyrolysis product molecules to diffuse into and/or react
with the catalyst. In some embodiments, the catalysts may be chosen
such that their pore sizes are sufficiently small to prevent entry
and/or reaction of pyrolysis products whose reaction would be
undesirable.
[0063] The catalyst may be treated or impregnated one or more times
with a coating compound such as a silicone compound to reduce the
size of the pore mouth openings in the catalyst as well as cover or
obscure catalytic sites on the external surface of the catalyst and
inside the pores of the catalyst near the pore mouth openings. The
covering of the catalytic sites with the treatment layer may
inhibit and/or extinguish their catalytic activity. In order to
facilitate a more controlled application of the coating compound,
the coating compound may be dispersed in a carrier, for example, an
aqueous or organic liquid carrier.
[0064] In each phase of the catalyst treatment process, the coating
compound may be deposited on the external surface of the catalyst
by any suitable method. For example, the coating compound may be
dissolved in an organic carrier, mixed with the catalyst, and then
dried by evaporation or vacuum distillation. The catalyst may be
contacted with the coating compound at a catalyst to coating
compound weight ratio in the range from about 1000:1 to about 1:10.
The content of the coating compound in the coated catalyst may be
from 0.1% to 30% by weight, or 0.5% to 20% by weight, or 1% to 12%
by weight, or 4% to 8% by weight, or at least 2%, or at least 4% or
at least 6%, or at least 8% by weight of the final catalyst
weight.
[0065] The zeolite catalyst may be treated with the
silicon-containing compound to reduce the size of the pore
openings, and cover or obscure catalytic sites on the external
surface of the catalyst. This treatment process may also be used to
cover or obscure catalytic sites in the pores near the pore mouths
openings. The covered or obscured catalytic sites may be referred
to as deactivated catalytic sites. A silicone coating compound
preferably has a molecular size that is incapable of entering the
pores of the catalyst. During the catalyst treatment process, the
silicon-containing compound may be applied to the catalyst and
subsequently calcined. This process may be repeated until the
desired level of treatment is provided. The fraction of catalytic
sites on the external surface of the catalyst that may be
deactivated by treatment with the coating compound may be at least
about 5%, or at least about 10%, or at least about 15%, or at least
about 25%, or at least about 35%, or at least about 45%, or at
least about 55%, or at least about 65%, or at least about 75%, or
at least about 85%, or at least about 90%, or at least about 95%,
or at least about 98%, or at least about 99%, of the available
catalytic sites on the external surface of the catalyst. The
fraction of catalyst sites on the external surface of the catalyst
that remain as acid sites after coating with the coating compound
may be at least about 95%, at least about 90%, at least about 80%,
at least about 70%, at least about 60%, at least about 50%, at
least about 40%, at least about 30%, at least about 20%, at least
about 10%, at least about 5%, at least about 1%, or between 1% and
95%, or between 5% and 90%, or between 50% and 90%, or between 60%
and 90% of the active sites on the external surface of the
non-coated material. The fraction of catalytic sites on the
external surface of the catalyst may be measured by a temperature
programmed desorption experiment using 2,4,6-collidine
(2,4,6-trimethylpyridine). A sample of the material to be measured
is degassed for 2 h at 823 K. After cooling the sample to 393 K, it
is exposed for 1 h to He (helium) that had been saturated with
2,4,6-collidine at room temperature by flowing pure He through a
bubbler containing the amine. Then the sample is held at 393 K with
He flow for 2 h to remove physisorbed 2,4,6-collidine. The sample
is heated to 973 K at 10 K./min. The total amount of
2,4,6-collidine desorbed is used to calculate the total number of
acid sites on the external surface of the catalyst, and the amount
of amine that desorbs between about 580 K and 650 K is used to
calculate the number of Bronsted acid sites for each catalyst. Due
to the size of 2,4,6 collidine, it does not enter ZSM-5 pores.
Therefore, desorption of 2,4,6-collidine only detects acid sites on
the external surface of the catalyst or in the pores near the pore
mouth openings. The decrease in the 2,4,6-collidine adsorption that
occurs with the catalyst treatment shows the decrease in the number
of acid sites on the external surface and in or near the pore mouth
openings. The decrease in these external sites is believed to be a
factor in the production of m- and o-xylene and in reducing the
re-equilibration of p-xylene formed in the pores to m- or o-xylene,
and thus improving the selectivity to p-xylene.
[0066] The catalyst may be treated with a tetraorthosilicate or
other coating compound using a chemical liquid deposition (CLD)
process or any of a number of other coating processes known to
those skilled in the art.
[0067] The coating compound may be provided in the form of a
solution or an emulsion under the conditions of contact with the
catalyst. The deposited coating compound may cover, and reside
substantially exclusively on, the external surface of the catalyst,
blocking external sites and partially blocking pore mouths and
sites in or near the pore mouths openings. Examples of methods of
depositing silicone compounds on the surface of zeolites may be
found in U.S. Pat. Nos. 4,090,981; 5,243,117; 5,403,800, and
5,659,098, which are incorporated by reference herein.
[0068] The coated catalyst may comprise silica and alumina. The
silicon to aluminum molar ratio may be in the range from about 10:1
to about 240:1, or in the range from about 10:1 to about 40:1, or
in the range from about 20:1 to about 50:1, or about 30:1, or at
least 30:1.
[0069] The catalyst may be treated in-situ with a silicon-based
coating by flowing a hydrocarbon solution containing the silicone
compound over the catalyst prior to introduction of biomass feed.
If the hydrocarbon solution contains toluene, then the p-xylene
selectivity increase during the coating process can be monitored
and adjusted to the desired selectivity level.
[0070] During the addition of a metal or metal oxide compound, the
metal compound may be applied to the catalyst and subsequently
calcined. This process may be repeated until the desired level of
metal is provided. The fraction of acid sites on the catalyst that
may be deactivated by treatment with the metal compound may be at
least about 5%, or at least about 10%, or at least about 15%, or at
least about 25%, or at least about 35%, or at least about 45%, or
at least about 55%, or at least about 65%, or at least about 75%,
or at least about 85%, or at least about 90%, or at least about
95%, or at least about 98%, or at least about 99%, of the available
catalytic sites. The fraction of acid sites on the catalyst that
remain as acid sites after treatment with the metal may be at least
about 95%, at least about 90%, at least about 80%, at least about
70%, at least about 60%, at least about 50%, at least about 40%, at
least about 30%, at least about 20%, at least about 10%, at least
about 5%, at least about 1%, or between 1% and 95%, or between 5%
and 90%, or between 50% and 90%, or between 60% and 90% of the
active sites of the material before metal addition.
[0071] The coating compound may have a number average molecular
weight in the range from about 80 to about 20,000, or from about
150 to 10,000. The coating compound may be a Dynasylan from Evonik
Industries AG, such as Hydrosil 2909, Hydrosil 2627, or any of the
other Dynasylan products, or a silicone emulsifier product such as
Dow Corning.RTM. 5329 Silicone Emulsifier or any similar product
obtainable from Dow Corning, or a modified polysiloxane such as KF
6015 or any of those obtainable from Shin Etsu, or the like.
[0072] The coating compound may comprise dimethylsilicone,
diethylsilicone, phenylmethylsilicone, methylhydrogensilicone,
ethylhydrogen silicone, phenylhydrogen silicone, methylethyl
silicone, phenylethyl silicone, diphenyl silicone,
methyltrifluoropropyl silicone, ethyltrifluoropropyl silicone,
polydimethyl silicone, tetrachloro-phenylmethyl silicone,
tetrachlorophenylethyl silicone, tetrachlorophenylhydrogen
silicone, tetrachlorophenylphenyl silicone, methylvinyl silicone,
hexamethyl cyclotrisiloxane, octamethyl cyclotetrasiloxane,
hexaphenyl cyclotrisiloxane, octaphenyl cyclotetrasiloxane, or a
mixture of two or more thereof. The coating compound may comprise a
tetraorthosilicate. The coating compound may comprise
tetramethylorthosilicate, tetraethylorthosilicate, or a mixture
thereof.
[0073] The kinetic diameter of the coating compound may be larger
than the pore diameter of the catalyst in order to avoid entry of
the silicone compound into the pore and any concomitant reduction
in the internal activity of the catalyst.
[0074] The organic carrier for the coating compound may comprise
hydrocarbons such as linear, branched, and cyclic alkanes having
five or more carbons. The carrier may comprise a linear, branched
or cyclic alkane having a boiling point greater than about
70.degree. C., and containing about 6 or more carbons. Optionally,
mixtures of low volatility organic compounds, such as hydrocracker
recycle oil, may also be employed as carriers. Low volatility
hydrocarbon carriers for the coating compound may comprise decane,
dodecane, mixtures thereof, and the like. A preferred carrier is
water.
[0075] Following each deposition of the coating compound, the
catalyst may be calcined to decompose the molecular or polymeric
species to a solid state species. The catalyst may be calcined at a
rate of from about 0.2.degree. C./minute to about 5.degree.
C./minute to a temperature greater than about 200.degree. C., but
below a temperature at which the crystallinity of the zeolite may
be adversely affected. Generally, such temperature will be below
about 600.degree. C. The temperature of calcination may be in the
range from about 350.degree. C. to about 550.degree. C. The
catalyst may be maintained at the calcination temperature for about
1 to about 24 hours, or about 2 to about 6 hours, or longer.
[0076] The catalyst may be treated with a water soluble silicone
compound or polysiloxane in a similar fashion. The soluble silicone
compound, silane, or polysiloxane may be added to the catalyst in
any manner such as incipient wetness, adsorption, spray coating, or
wet coating, or any other process known to those skilled in the
art. In a preferred procedure, the soluble silicone compound or
polysiloxane is dispersed in a solution of water at from 1 to 95%
by weight concentration of coating compound, or from 10 to 80% by
weight, or from 20 to 65% by weight, or from 40 to 60% by weight,
or at least 20% b/w, or at least 35% b/w, or at least 50% b/w, or
employed neat (i.e., with no water). The mixture is impregnated
into a dried catalyst by incipient wetness, i.e., just enough is
added to nearly fill the pores, but not enough to cause observable
wetness as is known to those skilled in the art. The resulting
impregnated material is dried in air, preferably for at least 30
minutes at greater than 100 C, heated to calcination temperature at
a heating rate of at least 1 C/minute, or at least 2 C/min, or at
least 5 C/min, or at least 10 C/min, or more rapidly, and calcined
in air at a temperature sufficient to decompose the organic
functionality of the polysiloxane, preferable at least 500 C, or at
least 550 C, or at least 600 C, for 1 hour, or 2 hours, or 3 hours,
or 5 hours, or longer.
[0077] The silicone or other coating compound can be largely
deposited on the surface of the catalyst and will be found to be
enriched on the surface of the catalyst relative to the whole
catalyst. Detection of the coating material can be performed by
Low-Energy Ion Scattering spectroscopy (LEIS), sometimes referred
to simply as ion scattering spectroscopy (ISS), which is a
surface-sensitive analytical technique used to characterize the
chemical and structural makeup of materials. In some cases the
surface coating may be detected by X-ray Photoelectron Spectroscopy
(XPS, sometimes called Electron Spectroscopy for Chemical Analysis,
ESCA) or by Scanning Electron Microscopy/Energy Dispersive X-ray
spectroscopy (SEM/EDX), or Transmission Electron Microscopy (TEM).
Another useful technique is Secondary Ion Mass Spectroscopy (SIMS)
in which a sample is bombarded with an ionic species (eg Ar+) and
the resulting ejected ions are detected. With SIMS a depth profile
of the various components can be obtained. For a silicone coating
LEIS, SEM/EDX, TEM, or SIMS can be used to show that the silica to
alumina ratio (SAR) of the surface atomic layers is higher than the
SAR of the bulk zeolite particles. For example, if the SAR
(SiO2/Al2O3) ratio of the bulk zeolite is 30:1, the SAR of the
surface will be at least 40:1 or at least 50:1 or at least
100:1.
[0078] For catalysts that are both coated with a coating compound
and promoted by addition of metal or metal oxide promoters, either
the catalyst can be coated first and then promoted or the catalyst
can be promoted first and then coated with the coating compound.
Catalysts that are coated first with the coating compound and
promoted with a metal or metal oxide promoter in a later step are
preferred.
[0079] The zeolite catalyst may be treated with the coating
compound to reduce the size of the pore openings, block pore
openings, and cover or obscure catalytic sites on the external
surface of the catalyst. This treatment process may also be used to
cover or obscure catalytic sites in the pores near the pore mouths
openings. The covered or obscured catalytic sites may be referred
to as deactivated catalytic sites. The coating compound may have a
molecular size that is incapable of entering the pores of the
catalyst or only partially entering the pores. During the catalyst
treatment process, the coating compound may be applied to the
catalyst and subsequently calcined. This process may be repeated
until the desired level of treatment is provided. The fraction of
catalytic sites on the external surface of the catalyst that may be
deactivated by treatment with the coating compound may be at least
about 15%, or at least about 25%, or at least about 35%, or at
least about 45%, or at least about 55%, or at least about 65%, or
at least about 75%, or at least about 85%, or at least about 90%,
or at least about 95%, or at least about 98%, or at least about
99%, of the available catalytic sites.
[0080] The coating compound may be a silicon-containing compound
having a number average molecular weight in the range from about 80
to about 20,000, or from about 150 to 10,000.
[0081] The coating compound may comprise oxides of Ti, V, Zr, Cr,
Mo, or Mn or mixtures of these. The coating may comprise carbon or
coke from prior exposure of the catalyst to CFP or other coke
forming conditions.
[0082] The pores with pore mouth openings that have been reduced in
size or number may allow for an increase in para selectivity for
the xylenes. This may be due to the fact that the reduced pore
mouth openings may allow p-xylene to diffuse out of the pores while
the diffusion of m-xylene and o-xylene may be restricted.
[0083] Reduction of the number of pore mouth openings may have a
similar effect. Due to the more rapid diffusion of p-xylene within
the zeolite when compared to isomers m- and o-xylene, closing pore
mouths increases the diffusion length of molecules entering and
leaving the zeolite; ie a linear pore that is closed on one end has
twice the diffusion path length compared to a pore that is open on
both ends. Increased pore length permits the faster diffusing
p-xylene to diffuse out of the pore more quickly than the o- and
m-xylenes.
[0084] Promoters can be added to the zeolite that impact the pore
and channel sizes as well. Promoter elements situated in or near
the pore mouth opening are capable of reducing the effective
diameter of the pore. Pores with reduced effective diameters have
slower rates of diffusion. For the xylene isomers, the rates of
diffusion of the o- and m-xylene isomers are more reduced than for
the p-xylene isomer. Again the faster relative rate of diffusion of
the p-xylene isomer permits it to diffuse out of the pores more
rapidly than the m- and o-xylene isomers, thus increasing the
selectivity of p-xylene among the xylenes in the product. Promoters
on the zeolite external surface can likewise inhibit the activity
of the external sites that are non-shape-selective, hence
increasing p-xylene selectivity.
[0085] The selectivity to p-xylene in the xylenes may be up to
100%, or, p-xylene may be produced in preference to o-xylene and/or
m-xylene, but some o-xylene and/or m-xylene may nevertheless be
produced. The fluid hydrocarbon product produced using the
foregoing method may comprise xylenes and may be characterized by a
p-xylene selectivity in the xylenes of at least about 50%, or at
least about 55%, or at least about 60%, or at least about 65%, or
at least about 70%, or at least about 75%, or at least about 80%,
or at least about 85%, or at least about 90%, or at least about
95%.
[0086] The method may further comprise the step of recovering the
fluid hydrocarbon product. The fluid hydrocarbon product may
further comprise, in addition to p-xylene, other aromatic compounds
and/or olefin compounds. The fluid hydrocarbon product may further
comprise benzene, toluene, ethylbenzene, styrene,
methylethylbenzene, trimethylbenzene, o-xylene, m-xylene, indanes,
naphthalene, methylnaphthalene, dimethylnaphthalene,
ethylnaphthalene, hydrindene, methylhydrindene, dimethylhydrindene,
or a mixture of two or more thereof.
[0087] The carbon yield of aromatics in the fluid hydrocarbon
product may be at least about 13%, or at least about 17%, or at
least about 20%. The carbon yield of olefins in the fluid
hydrocarbon product may be at least about 7%, or at least about 9%,
or at least about 11%. The mass yield of p-xylene may be at least
about 1.5 wt %, or at least about 2 wt %, or at least about 2.5 wt
%, or at least about 3 wt %.
[0088] The inventive method may comprise a single-stage method for
the pyrolysis of the hydrocarbonaceous material. This method may
comprise providing or using a single-stage pyrolysis apparatus. A
single-stage pyrolysis apparatus may be one in which pyrolysis and
subsequent catalytic reactions are carried out in a single vessel.
The single-stage pyrolysis apparatus may comprise a continuously
stirred tank reactor, a batch reactor, a semi-batch reactor, a
fixed bed reactor or a fluidized bed reactor. Multi-stage
apparatuses may also be used for the production of fluid
hydrocarbon products in accordance with the invention.
[0089] The catalysts provided for herein may be particularly suited
for producing xylenes with a relatively high selectivity to
p-xylene in the xylenes of at least about 40%, or at least about
45%, or at least about 50%, or at least about 55%, or at least
about 60%, or at least about 65%, or at least about 70%, or at
least about 75%, or at least about 80%, or at least about 85%, or
at least about 90%.
[0090] Referring to FIG. 1, feed stream 10 includes a solid
hydrocarbonaceous material that can be fed to reactor 20. Moisture
12 may optionally be removed from the solid hydrocarbonaceous feed
composition prior to being fed to the reactor, e.g., by an optional
dryer 14. Removal of moisture from the solid hydrocarbonaceous
material feed stream may be advantageous for several reasons. For
example, the moisture in the feed stream may require additional
energy input in order to heat the solid hydrocarbonaceous material
to a temperature sufficiently high to achieve pyrolysis. Variations
in the moisture content of the solid hydrocarbonaceous feed may
lead to difficulties in controlling the temperature of the reactor.
In addition, removal of moisture from the solid hydrocarbonaceous
feed can reduce or eliminate the need to process the water during
later processing steps.
[0091] The particle size of the solid hydrocarbonaceous feed
composition may be reduced in an optional grinding system 16 prior
to passing the solid hydrocarbonaceous feed to the reactor. The
hydrocarbonaceous material may be transferred to reactor 20. The
reactor may be used, in some instances, to perform catalytic
pyrolysis of at least a portion of the first reactant comprising
the hydrocarbonaceous material under reaction conditions sufficient
to produce one or more pyrolysis products. In the illustrative
embodiment of FIG. 1, the reactor comprises any suitable reactor
known to those skilled in the art. For example, in some instances,
the reactor may comprise a continuously stirred tank reactor
(CSTR), a batch reactor, a semi-batch reactor, or a fixed bed
catalytic reactor, among others. Preferably, the reactor comprises
a fluidized bed reactor, e.g., a circulating fluidized bed reactor,
bubbling bed reactor, or riser reactor. Fluidized bed reactors may,
in some cases, provide improved mixing of the catalyst, solid
biomass during pyrolysis and/or subsequent reactions, which may
lead to enhanced control over the reaction products formed. The use
of fluidized bed reactors may also lead to improved heat transfer
within the reactor. In addition, improved mixing in a fluidized bed
reactor may lead to a reduction of the amount of coke adhered to
the catalyst, resulting in reduced deactivation of the catalyst in
some cases.
[0092] Higher yields of desired product formation, lower yields of
coke formation, and/or more controlled product formation (e.g.,
higher production of p-xylene relative to other products) may be
achieved when particular combinations of reaction conditions and
system components are implemented in methods and systems described
herein. For example, conditions such as the mass normalized space
velocity(ies) (e.g., of the solid hydrocarbonaceous material and/or
the fluidization fluid), the temperature of the reactor and/or
solids separator, the reactor pressure, the heating rate of the
feed stream(s), the catalyst to solid hydrocarbonaceous material
mass ratio, the residence time of the hydrocarbonaceous material in
the reactor, the residence time of the reaction products in the
solids separator, and/or the catalyst type (as well as silica to
alumina molar ratio and pore mouth opening size) may be controlled
to achieve beneficial results.
[0093] The reactor(s) may be operated at any suitable temperature.
In some instances, it may be desirable to operate the reactor(s) at
intermediate temperatures, compared to temperatures typically used
in many previous catalytic pyrolysis systems. For example, the
reactor may be operated at temperatures of between about
400.degree. C. and about 650.degree. C., between about 425.degree.
C. and about 600.degree. C., or between about 525.degree. C. and
about 575.degree. C. Operating the reactor(s) at these intermediate
temperatures may allow one to maximize the amount of desirable
products. The invention may not be limited to the use of such
intermediate temperatures, however, and in other embodiments, lower
and/or higher temperatures can be used.
[0094] The reactor(s) may also be operated at any suitable
pressure. The reactor may be operated at a pressure of at least
about 100 kPa, or at least about 200 kPa, or at least about 300
kPa, or at least about 400 kPa. The reactor may be operated at a
pressure below about 600 kPa, or below about 400 kPa, or below
about 200 kPa. The reactor may be operated at a pressure in the
range from about 100 to about 600 kPa, or in the range from about
100 to about 400 kPa, or in the range from about 100 to about 200
kPa. The invention may not be limited to the use of such pressures,
however, and in other embodiments, lower and/or higher pressures
may be employed.
[0095] The mass-normalized space velocity of the hydrocarbonaceous
material may be selected to selectively produce a desired array of
fluid hydrocarbon products. As used herein, the term
"mass-normalized space velocity" of a component is defined as the
mass flow rate of the component into the reactor (e.g., as measured
in g/hr) divided by the mass of catalyst in the reactor (e.g., as
measured in g) and has units of inverse time. For example, the
mass-normalized space velocity of solid hydrocarbonaceous material
fed to the reactor may be calculated as the mass flow rate of the
solid hydrocarbonaceous material into the reactor divided by the
mass of catalyst in the reactor. The mass-normalized space velocity
of a component (e.g., the hydrocarbonaceous material) in the
reactor may be calculated using different methods depending upon
the type of reactor being used. For example, in systems employing
batch or semi-batch reactors, wherein the solid hydrocarbonaceous
material is not fed continuously to the reactor, the solid
hydrocarbonaceous material does not have a mass-normalized space
velocity. For systems in which catalyst is fed to and/or extracted
from the reactor during reaction (e.g., circulating fluidized bed
reactors), the mass-normalized space velocity may be determined by
calculating the average amount of catalyst within the volume of the
reactor over a period of operation (e.g., steady-state
operation).
[0096] The mass-normalized space velocity of the hydrocarbonaceous
material fed to the reactor may be at a mass normalized space
velocity of up to about 3 hour-1, or up to about 2 hour-1, or up to
about 1.5 hour-1, or up to about 0.9 hour-1, or in the range from
about 0.01 hour-1 to about 3 hour-1, or in the range from about
0.01 to about 2 hour-1, or in the range from about 0.01 to about
1.5 hour-1, or in the range from about 0.01 to about 0.9 hour-1, or
in the range from about 0.01 hour-1 to about 0.5 hour-1, or in the
range from about 0.1 hour-1 to about 0.9 hour-1, or in the range
from about 0.1 hour-1 to about 0.5 hour-1. The invention may not be
limited to the use of such mass-normalized space velocities,
however, and in other embodiments, lower and/or higher
mass-normalized space velocities can be used.
[0097] The residence time of a reactant (e.g., the
hydrocarbonaceous material) in the reactor (i.e., the reactant
residence time) may be at least about 1 second, at least about 2
seconds, at least about 5 seconds, at least about 7 seconds, at
least about 10 seconds, at least about 15 seconds, at least about
20 seconds, at least about 25 seconds, at least about 30 seconds,
at least about 60 seconds, at least about 120 seconds, at least
about 240 seconds, or at least about 480 seconds. In some cases,
the residence time of a reactant (e.g., the hydrocarbonaceous
material) in the reactor may be less than about 5 minutes, or from
about 1 second and about 4 minutes, or from about 2 seconds to
about 4 minutes, or from about 5 seconds to about 4 minutes, or
from about 7 seconds to about 4 minutes, or from about 10 seconds
to about 4 minutes, or from about 12 seconds to about 4 minutes, or
from about 15 seconds to about 4 minutes, or from about 20 seconds
to about 4 minutes, or from about 30 seconds to about 4 minutes, or
from about 60 seconds to about 4 minutes. Previous "fast pyrolysis"
studies have, in many cases, employed systems with very short
reactant residence times (e.g., less than 2 seconds). In some
cases, however, the use of relatively longer residence times may
allow for additional chemical reactions to form desirable products.
Long residence times may be achieved by, for example, increasing
the volume of the reactor and/or reducing the volumetric flow rate
of the hydrocarbonaceous materials. It should be understood,
however, that in some embodiments described herein, the residence
time of the reactant (e.g., hydrocarbonaceous material) may be
relatively shorter, e.g., less than about 2 seconds, or less than
about 1 second.
[0098] The contact time of the pyrolysis product (e.g., pyrolysis
vapor) with the catalyst in the reactor may be at least about 1
second, at least about 2 seconds, at least about 5 seconds, at
least about 7 seconds, at least about 10 seconds, at least about 15
seconds, at least about 20 seconds, at least about 25 seconds, at
least about 30 seconds, at least about 60 seconds, at least about
120 seconds, at least about 240 seconds, or at least about 480
seconds. The contact time may be less than about 5 minutes, or from
about 1 second and about 4 minutes, or from about 2 seconds to
about 4 minutes, or from about 5 seconds to about 4 minutes, or
from about 7 seconds to about 4 minutes, or from about 10 seconds
to about 4 minutes, or from about 12 seconds to about 4 minutes, or
from about 15 seconds to about 4 minutes, or from about 20 seconds
to about 4 minutes, or from about 30 seconds to about 4 minutes, or
from about 60 seconds to about 4 minutes.
[0099] In certain cases where fluidized bed reactors are used, the
feed material (e.g., a solid hydrocarbonaceous material) in the
reactor may be fluidized by flowing a fluid stream through the
reactor. In the exemplary embodiment of FIG. 1, a fluid stream 44
is used to fluidize the feed material in reactor 20. Fluid may be
supplied to the fluid stream from a fluid source 24 and/or from the
product streams of the reactor via a compressor 26. As used herein,
the term "fluid" means a material generally in a liquid,
supercritical, or gaseous state. Fluids, however, may also contain
solids such as, for example, suspended or colloidal particles. In
some embodiments, it may be advantageous to control the residence
time of the fluidization fluid in the reactor. The residence time
of the fluidization fluid may be defined as the volume of the
reactor divided by the volumetric flow rate of the fluidization
fluid. The residence time of the fluidization fluid may be at least
about 0.1 second, at least about 0.2 second, at least about 0.5
second, at least about 1 second, at least about 2 seconds, at least
about 3 seconds, at least about 4 seconds, at least about 5
seconds, at least about 6 seconds, at least about 8 seconds, at
least about 10 seconds, at least about 12 seconds, at least about
24 seconds, or at least about 48 seconds. The residence time of the
fluidization fluid may be from about 0.1 second to about 48
seconds, from about 0.2 second to about 48 seconds, from about 0.5
second to about 480 seconds, from about 1 second to about 48
seconds, from about 3 seconds to about 48 seconds, from about 5
seconds to about 48 seconds, from about 6 seconds to about 48
seconds, from about 8 seconds to about 48 seconds, from about 10
seconds to about 48 seconds, from about 12 seconds to about 48
seconds, or from about 24 seconds to about 48 seconds.
[0100] Suitable fluidization fluids that may be used in this
invention include, for example, inert gases (e.g., helium, argon,
neon, etc.), hydrogen, nitrogen, steam, carbon monoxide, and carbon
dioxide, among others.
[0101] As shown in the illustrative embodiment of FIG. 1, the
products (e.g., fluid hydrocarbon products) formed during the
reaction of the reactants (e.g., the solid hydrocarbonaceous
material) exit the reactor via a product stream 30. In addition to
the reaction products, the product stream may, in some cases,
comprise unreacted reactant(s), fluidization fluid, char, ash,
and/or catalyst. In one set of embodiments, the desired reaction
product(s) (e.g., liquid aromatic hydrocarbons, olefin
hydrocarbons, gaseous products, etc.) may be recovered from an
effluent stream of the reactor.
[0102] As shown in the illustrative embodiment of FIG. 1, product
stream 30 may be fed to an optional solids separator 32. The solids
separator may be used, in some cases, to separate the reaction
products from catalyst (e.g., at least partially deactivated
catalyst) present in the product stream. In addition, the solids
separator may be used, in some instances, to remove coke and/or ash
from the catalyst. In some embodiments, the solids separator may
comprise optional purge stream 33, which may be used to purge coke,
ash, and/or catalyst from the solids separator.
[0103] The solids separator may not be required in all embodiments.
For example, for situations in which catalytic fixed bed reactors
are employed, the catalyst may be retained within the reactor, and
the reaction products may exit the reactor substantially free of
catalyst, thus negating the need for a separation step.
[0104] The separated catalyst may exit the solids separator via
stream 34. A portion of the separated catalyst may be returned to
the reactor via a return pipe, not shown in FIG. 1. The catalyst
exiting the separator may be at least partially deactivated. The
separated catalyst may be fed to a regenerator 36 in which any
catalyst that was at least partially deactivated may be
re-activated. The regenerator may comprise an optional purge stream
37, which may be used to purge coke, ash, and/or catalyst from the
regenerator. Methods for activating catalyst are well-known to
those skilled in the art, for example, as described in Kirk-Othmer
Encyclopedia of Chemical Technology (Online), Vol. 5, Hoboken,
N.J.: Wiley-Interscience, c2001-, pages 255-322, which are
incorporated herein by reference.
[0105] A portion of the catalyst may be removed from the reactor
through a catalyst exit port (not shown in FIG. 1.). The catalyst
removed from the reactor may be partially deactivated and passed
via a conduit into regenerator 36, or into a separate regenerator
(not shown in FIG. 1). Removed catalyst that has been regenerated
may be returned to the reactor via stream 47, or may be returned to
the reactor separately from the fluidization gas via a separate
stream (not shown in FIG. 1.).
[0106] An oxidizing agent may be fed to the regenerator via a
stream 38, e.g., as shown in FIG. 1. The oxidizing agent may
originate from any source including, for example, a tank of oxygen,
atmospheric air, steam, among others. In the regenerator, the
catalyst may be re-activated by reacting the catalyst with the
oxidizing agent. The deactivated catalyst may comprise residual
carbon and/or coke, which may be removed via reaction with the
oxidizing agent in the regenerator. The regenerator in FIG. 1
comprises a vent stream 40 which may include regeneration reaction
products, residual oxidizing agent, etc.
[0107] The regenerator may be of any suitable size mentioned above
in connection with the reactor or the solids separator. In
addition, the regenerator may be operated at elevated temperatures
in some cases (e.g., at least about 300.degree. C., 400.degree. C.,
500.degree. C., 600.degree. C., 700.degree. C., 800.degree. C., or
higher). The residence time of the catalyst in the regenerator may
also be controlled using methods known by those skilled in the art,
including those outlined above. The mass flow rate of the catalyst
through the regenerator may be coupled to the flow rate(s) in the
reactor and/or solids separator in order to preserve the mass
balance in the system.
[0108] The regenerated catalyst may exit the regenerator via stream
42. The regenerated catalyst may be recycled back to the reactor
via recycle stream 47. In some cases, catalyst may be lost from the
system or removed intentionally during operation. Additional
"makeup" catalyst may be added to the system via a makeup stream
46. The regenerated and makeup catalyst may be fed to the reactor
with the fluidization fluid via recycle stream 47. Alternatively,
the catalyst and fluidization fluid may be fed to the reactor via
separate streams.
[0109] Referring to solids separator 32 in FIG. 1, the reaction
products (e.g., fluid hydrocarbon products) may exit the solids
separator via stream 48. In some cases, a fraction of stream 48 may
be purged via purge stream 60. The contents of the purge stream may
be fed to a combustor or a water-gas shift reactor, for example, to
recuperate energy that would otherwise be lost from the system. In
some cases, the reaction products in stream 48 may be fed to an
optional condenser 50. The condenser may comprise a heat exchanger
which condenses at least a portion of the reaction product from a
gaseous to a liquid state. The condenser may be used to separate
the reaction products into gaseous, liquid, and solid fractions.
The operation of condensers is well known to those skilled in the
art. The condenser may also make use of pressure change to condense
portions of the product stream. In FIG. 1, stream 54 may comprise
the liquid fraction of the reaction products (e.g., water, aromatic
compounds, olefin compounds, etc.), and stream 74 may comprise the
gaseous fraction of the reaction products (e.g., CO, CO.sub.2,
H.sub.2, etc.). In some embodiments, the gaseous fraction may be
fed to a vapor recovery system 70. The vapor recovery system may be
used, for example, to recover any desirable vapors within stream 74
and transport them via stream 72. In addition, stream 76 may be
used to transport CO, CO.sub.2, and/or other gases from the vapor
recovery system. The optional vapor recovery system may be placed
in other locations. For example, in some embodiments, a vapor
recovery system may be positioned downstream of purge stream 54.
One skilled in the art can select an appropriate placement for a
vapor recovery system.
[0110] Other products (e.g., excess gas) may be transported to
optional compressor 26 via stream 56, where they may be compressed
and used as fluidization gas in the reactor (stream 22) and/or
where they may assist in transporting the hydrocarbonaceous
material to the reactor (streams 58) or may be used to transport
catalyst to the reactor (not shown), or may be used to transport
additional non-solid feeds to the reactor. In some instances, the
liquid fraction may be further processed, for example, to separate
the water phase from the organic phase, to separate individual
compounds, etc.
[0111] It should be understood that, while the set of embodiments
described by FIG. 1 includes a reactor, solids separator,
regenerator, condenser, etc., not all embodiments will involve the
use of these elements. For example, in some embodiments, the feed
stream(s) may be fed to a catalytic fixed bed reactor, reacted, and
the reaction products may be collected directly from the reactor
and cooled without the use of a dedicated condenser. In some
instances, while a dryer, grinding system, solids separator,
regenerator, condenser, and/or compressor may be used as part of
the process, one or more of these elements may comprise separate
units not fluidically and/or integrally connected to the reactor.
In other embodiments one or more of the dryer, grinding system,
solids separator, regenerator, condenser, and/or compressor may be
absent. In some embodiments, the desired reaction product(s) may be
recovered at any point in the production process (e.g., after
passage through the reactor, after separation, after condensation,
etc.).
[0112] Catalyst components useful in the context of this invention
can be selected from any catalyst known in the art, or as would be
understood by those skilled in the art made aware of this
invention. Functionally, catalysts may be limited only by the
capability of any such material to promote and/or effect
dehydration, dehydrogenation, isomerization, hydrogen transfer,
aromatization, decarbonylation, decarboxylation, aldol condensation
and/or any other reaction or process associated with or related to
the pyrolysis of a hydrocarbonaceous material. Catalyst components
can be considered acidic, neutral or basic, as would be understood
by those skilled in the art.
[0113] A screening method may be used to select catalysts with
appropriate pore sizes for the conversion of specific pyrolysis
product molecules. The screening method may comprise determining
the size of pyrolysis product molecules desired to be catalytically
reacted (e.g., the molecule kinetic diameters of the pyrolysis
product molecules). One of ordinary skill in the art can calculate,
for example, the kinetic diameter of a given molecule. The type of
catalyst may then be chosen such that the pores of the catalyst
(e.g., Norman adjusted minimum radii) are sufficiently large to
allow the pyrolysis product molecules to diffuse into and/or react
with the catalyst. In some embodiments, the catalysts are chosen
such that their pore sizes are sufficiently small to prevent entry
and/or reaction of pyrolysis products whose reaction would be
undesirable.
[0114] The catalyst may be selected from naturally-occurring
zeolites, synthetic zeolites and combinations thereof. The catalyst
may be a Mordenite Framework Inverted (MFI) type zeolite catalyst,
such as a ZSM-5 zeolite catalyst. Catalysts comprising ZSM-5 that
may be used with or without modification are available
commercially. The catalysts that are provided for herein may
comprise acid or catalytically active sites. While not wishing to
be bound by theory, it is believed that various acid sites in ZSM-5
and other zeolites are catalytically active for reactions of the
hydrocarbonaceous materials including dehydration, decarbonylation,
decarboxylation, isomerization, oligomerization and/or
dehydrogenation, hence the terms "acid sites" and "catalytically
active sites" may be used interchangeably. Other types of useful
zeolite catalysts may include ferrierite, zeolite Y, zeolite beta,
modernite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)AlPO-31,
ZSM-11, SSZ-23, mixtures of two or more thereof, and the like.
[0115] The zeolite catalyst may contain binders and fillers in
addition to the zeolite. Typical binders or fillers include
amorphous materials such as alumina, silica, silica-alumina,
titania, aluminum phosphate, kaolin, attapulgite clay and various
types of clays. Preferred zeolite catalysts comprise at least 25%,
or at least 30%, or at least 35%, or at least 40%, or at least 45%,
or at least 50%, or at least 60% by weight zeolite in the catalyst
particle.
[0116] The catalyst may comprise, in addition to alumina and
silica, one or more additional metals and/or a metal oxides.
Suitable metals and/or oxides may include, for example, nickel,
palladium, silver, platinum, palladium, titanium, vanadium,
chromium, manganese, iron, cobalt, zinc, copper, gallium, sodium,
potassium, magnesium, calcium, zirconium, lanthanum, cerium,
phosphorus, an oxide of one or more thereof, or a mixture of two or
more thereof, among others. The metal and/or metal oxide can be
impregnated into the catalyst (e.g., in the interstices of the
lattice structure of the catalyst), in some embodiments. The metal
or metal oxide can be added to the zeolite by any of a number of
techniques known to those skilled in the art, such as, but not
limited to, impregnation, ion exchange, vapor deposition, and the
like. The zeolite may comprise small amounts of structure
stabilizing elements such as phosphorus, lanthanum, rare earths,
and the like, typically at levels that are less than about 4% by
weight of the zeolite. The catalyst may be conditioned before
operation in the process by a wide range of techniques known to
those skilled in the art such as, but not limited to, oxidation,
calcination, reduction, cyclic oxidation and reduction; steaming,
hydrolysis, and the like. The metal and/or metal oxide may be
incorporated into the lattice structure of the catalyst. For
example, the metal and/or metal oxide may be included during the
preparation of the catalyst, and the metal and/or metal oxide may
occupy a lattice site of the resulting catalyst (e.g., a zeolite
catalyst). As another example, the metal and/or metal oxide may
react or otherwise interact with a zeolite such that the metal
and/or metal oxide displaces an atom within the lattice structure
of the zeolite.
[0117] In addition, in some cases, properties of the catalysts
(e.g., pore structure, type and/or number of acid sites, etc.) may
be chosen to selectively produce a desired product. It may be
desirable, in some embodiments, to employ one or more catalysts to
establish a bimodal distribution of pore sizes. In some cases, a
single catalyst with a bimodal distribution of pore sizes may be
used (e.g., a single catalyst that contains predominantly 5.9-6.3
.ANG. pores and 7-8 .ANG. pores). In other cases, a mixture of two
or more catalysts may be employed to establish the bimodal
distribution (e.g., a mixture of two catalysts, each catalyst type
including a distinct range of average pore sizes). In some
embodiments, one of the one or more catalysts comprises a zeolite
catalyst and another of the one or more catalysts comprises a
non-zeolite catalyst (e.g., a mesoporous catalyst, a metal oxide
catalyst, etc.).
[0118] For example, in some embodiments at least about 70%, at
least about 80%, at least about 90%, at least about 95%, at least
about 98%, or at least about 99% of the pores of the one or more
catalysts (e.g., a zeolite catalyst, a mesoporous catalyst, etc.)
have smallest cross-sectional diameters that lie within a first
size distribution or a second size distribution. In some cases, at
least about 2%, at least about 5%, or at least about 10% of the
pores of the one or more catalysts have smallest cross-sectional
diameters that lie within the first size distribution; and at least
about 2%, at least about 5%, or at least about 10% of the pores of
the one or more catalysts have smallest cross-sectional diameters
that lie within the second size distribution. In some cases, the
first and second size distributions are selected from the ranges
provided above. In certain embodiments, the first and second size
distributions are different from each other and do not overlap. An
example of a non-overlapping range is 5.9-6.3 .ANG. and 6.9-8.0
.ANG., and an example of an overlapping range is 5.9-6.3 .ANG. and
6.1-6.5 .ANG.. The first and second size distributions may be
selected such that the ranges are not immediately adjacent one
another, an example being pore sizes of 5.9-6.3 .ANG. and 6.9-8.0
.ANG.. An example of a range that is immediately adjacent one
another is pore sizes of 5.9-6.3 .ANG. and 6.3-6.7 .ANG.. As a
specific example, in some embodiments one or more catalysts are
used to provide a bimodal pore size distribution for the
simultaneous production of aromatic and olefin compounds. That is,
one pore size distribution may advantageously produce a relatively
high amount of aromatic compounds, particularly p-xylene, and the
other pore size distribution may advantageously produce a
relatively high amount of olefin compounds. In some embodiments, at
least about 70%, at least about 80%, at least about 90%, at least
about 95%, at least about 98%, or at least about 99% of the pores
of the one or more catalysts have smallest cross-sectional
diameters between about 5.9 .ANG. and about 6.3 .ANG. or between
about 7 .ANG. and about 8 .ANG.. In addition, at least about 2%, at
least about 5%, or at least about 10% of the pores of the one or
more catalysts have smallest cross-sectional diameters between
about 5.9 .ANG. and about 6.3 .ANG.; and at least about 2%, at
least about 5%, or at least about 10% of the pores of the one or
more catalysts have smallest cross-sectional diameters between
about 7 .ANG. and about 8 .ANG..
[0119] In some embodiments, it least about 70%, at least about 80%,
at least about 90%, at least about 95%, at least about 98%, or at
least about 99% of the pores of the one or more catalysts have
smallest cross-sectional diameters between about 5.9 .ANG. and
about 6.3 .ANG. or between about 7 .ANG. and about 200 .ANG.. In
addition, at least about 2%, at least about 5%, or at least about
10% of the pores of the one or more catalysts have smallest
cross-sectional diameters between about 5.9 .ANG. and about 6.3
.ANG.; and at least about 2%, at least about 5%, or at least about
10% of the pores of the one or more catalysts have smallest
cross-sectional diameters between about 7 .ANG. and about 200
.ANG.. In some embodiments, at least about 70%, at least about 80%,
at least about 90%, at least about 95%, at least about 98%, or at
least about 99% of the pores of the one or more catalysts have
smallest cross-sectional diameters that lie within a first
distribution and a second distribution, wherein the first
distribution is between about 5.9 .ANG. and about 6.3 .ANG. and the
second distribution is different from and does not overlap with the
first distribution. In some embodiments, the second pore size
distribution may be between about 7 .ANG. and about 200 .ANG.,
between about 7 .ANG. and about 100 .ANG., between about 7 .ANG.
and about 50 .ANG., or between about 100 .ANG. and about 200 .ANG..
In some embodiments, the second catalyst may be mesoporous (e.g.,
have a pore size distribution of between about 2 nm and about 50
nm).
[0120] In some embodiments, a bimodal distribution of pore sizes
may be beneficial in reacting two or more hydrocarbonaceous feed
material components. For example, some embodiments comprise
providing a solid hydrocarbonaceous material comprising a first
component and a second component in a reactor, wherein the first
and second components are different. Examples of compounds that may
be used as first or second components include any of the
hydrocarbonaceous materials. For example, the first component may
comprise one of cellulose, hemi-cellulose and lignin, and the
second component comprises one of cellulose, hemicellulose and
lignin. The method may further comprise providing first and second
catalysts in the reactor. In some embodiments, the first catalyst
may have a first pore size distribution and the second catalyst may
have a second pore size distribution, wherein the first and second
pore size distributions are different and do not overlap. The first
pore size distribution may be, for example, between about 5.9 .ANG.
and about 6.3 .ANG.. The second pore size distribution may be, for
example, between about 7 .ANG. and about 200 .ANG., between about 7
.ANG. and about 100 .ANG., between about 7 .ANG. and about 50
.ANG., or between about 100 .ANG. and about 200 .ANG.. In some
cases, the second catalyst may be mesoporous or non-porous. The
first catalyst may be selective for catalytically reacting the
first component or a derivative thereof to produce a fluid
hydrocarbon product. In addition, the second catalyst may be
selective for catalytically reacting the second component or a
derivative thereof to produce a fluid hydrocarbon product. The
method may further comprise pyrolyzing within the reactor at least
a portion of the hydrocarbonaceous material under reaction
conditions sufficient to produce one or more pyrolysis products and
catalytically reacting at least a portion of the pyrolysis products
with the first and second catalysts to produce the one or more
hydrocarbon products. In some instances, at least partially
deactivated catalyst may also be used.
[0121] In certain embodiments, a method used in combination with
embodiments described herein includes increasing the catalyst to
hydrocarbonaceous material mass ratio of a composition to increase
production of identifiable aromatic compounds. As illustrated
herein, representing but one distinction over certain prior
catalytic pyrolysis methods, articles and methods described herein
can be used to produce discrete, identifiable aromatic, biofuel
compounds selected from but not limited to benzene, toluene,
propylbenzene, ethylbenzene, styrene, methylbenzene,
methylethylbenzene, trimethylbenzene, xylenes, indanes,
naphthalene, methylnaphthelene, dimethylnaphthalene,
ethylnaphthalene, hydrindene, methylhydrindene, and
dimethylhydrindene and combinations thereof.
[0122] In some embodiments, the reaction chemistry of a catalyst
may be affected by adding one or more additional compounds. For
example, the addition of a metal to a catalyst may result in a
shift in selective formation of specific compounds (e.g., addition
of metal to alumina-silicate catalysts may result in the production
of more CO or CO.sub.2). In addition, when the fluidization fluid
comprises hydrogen, the amount of coke formed on the catalyst may
be decreased.
[0123] The catalyst may comprise both silica and alumina. The
silica (SiO.sub.2) and alumina (Al.sub.2O.sub.3) in the catalyst
may be present in any suitable molar ratio. For example, in some
cases, the catalyst in the feed may comprise a silica (SiO.sub.2)
to alumina (Al.sub.2O.sub.3) molar ratio (SAR) of between about
10:1 to about 240:1, or in the range from about 10:1 to about 40:1,
or in the range from about 20:1 to about 50:1, or about 30:1, or
greater than 30:1.
[0124] In some embodiments, the methods described herein may be
configured to selectively produce aromatic compounds (e.g.,
p-xylene) in a single-stage, or alternatively, a multi-stage
pyrolysis apparatus. For example, in some embodiments, the mass
yield of the aromatic compounds in the fluid hydrocarbon product
may be at least about 13 wt %, at least about 17 wt %, at least
about 20 wt %, at least about 30 wt %, between about 13 wt % and
about 40 wt %, between about 13 wt % and about 35 wt %, between
about 17 wt % and about 40 wt %, between about 17 wt % and about 35
wt %, between about 20 wt % and about 40 wt %, between about 20 wt
% and about 35 wt %, between about 30 wt % and about 40 wt %, or
between about 30 wt % and about 35 wt %. The mass yield of p-xylene
may be at least about 1.5% by weight, or at least about 2% by
weight, or at least about 2.5% by weight, or at least about 3% by
weight. The "mass yield" of aromatic compounds or p-xylene in a
given product is calculated as the total weight of the aromatic
compounds or p-xylene present in the fluid hydrocarbon product
divided by the weight of the solid hydrocarbonaceous material used
in forming the reaction product, multiplied by 100%.
[0125] In some embodiments, aromatic compounds (especially
p-xylene) may be selectively produced when the mass-normalized
space velocity of the solid hydrocarbonaceous material fed to the
reactor is up to about 3 hour-1, or up to about 2 hour-1, or up to
about 1.5 hour-1, or up to about 0.9 hour-1, or in the range from
about 0.01 hour-1 to about 3 hour-1, or in the range from about
0.01 to about 2 hour-1, or in the range from about 0.01 to about
1.5 hour-1, or in the range from about 0.01 to about 0.9 hour-1, or
in the range from about 0.01 hour-1 to about 0.5 hour-1, or in the
range from about 0.1 hour-1 to about 0.9 hour-1, or in the range
from about 0.1 hour-1 to about 0.5 hour-1. In some instances,
aromatic compounds (especially p-xylene) may be selectively
produced when the reactor is operated at a temperature of between
about 400.degree. C. and about 650.degree. C. (or between about
425.degree. C. and about 600.degree. C., or between about
500.degree. C. and about 575.degree. C.). In addition, certain
heating rates (e.g., at least about 50.degree. C./s, or at least
about 400.degree. C./s), high catalyst-to-feed mass ratios (e.g.,
at least about 5:1), and/or high silica to alumina molar ratios in
the catalyst (e.g., at least about 30:1) may be used to facilitate
selective production of aromatic compounds (especially p-xylene).
Some such and other process conditions may be combined with a
particular reactor type, such as a fluidized bed reactor (e.g., a
bubbling fluidized bed, a turbulent fluidized bed, a fast fluidized
bed, circulating fluidized bed reactor), to selectively produce
aromatic and/or olefin compounds.
[0126] Furthermore, in some embodiments, the catalyst may be chosen
to facilitate selective production of aromatic products (especially
p-xylene). For example, ZSM-5 may, in some cases, preferentially
produce relatively higher amounts of aromatic compounds. In some
cases, catalysts that include Bronsted acid sites may facilitate
selective production of aromatic compounds. In addition, catalysts
with well-ordered pore structures may facilitate selective
production of aromatic compounds. For example, in some embodiments,
catalysts with average pore diameters between about 5.9 .ANG. and
about 6.3 .ANG. may be particularly useful in producing aromatic
compounds. In addition, catalysts with average pore diameters
between about 7 .ANG. and about 8 .ANG. may be useful in producing
olefins. In some embodiments, a combination of one or more of the
above process parameters may be employed to facilitate selective
production of aromatic and/or olefin compounds. The ratio of
aromatics to olefins produced on a carbon basis may be, for
example, between about 0.1:1 and about 10:1, between about 0.2:1
and about 5:1, between about 0.5:1 and about 4:1, between about
0.1:1 and about 0.5:1, between about 0.5:1 and about 1:1, between
about 1:1 and about 5:1, or between about 5:1 and about 10:1.
[0127] Furthermore, processes described herein may result in lower
coke formation than certain existing methods. For example, in some
embodiments, a pyrolysis product can be formed with less than about
30 wt %, less than about 25 wt %, less than about 20 wt %, than
about 15 wt %, or less than about 10 wt % of the pyrolysis product
being coke. The amount of coke formed is measured as the weight of
coke formed in the system divided by the weight of
hydrocarbonaceous material used in forming the pyrolysis
product.
[0128] The catalyst of the present invention may comprise at least
0.2 wt % Fe, or at least 0.3% Fe, or at least 0.5% Fe, or at least
1.0% Fe, or at least 1.5% Fe, or at least 1.7% Fe, or at least 2.0%
Fe, or from 0.2% to 10% Fe, or from 0.3% to 2.0% Fe, or from 0.5%
to 1.7% Fe. The phrase "at least 0.2 wt % Fe" is determined by
elemental analysis of catalyst separated from biomass and separated
from any ash, to the extent practicable, where the elemental
analysis is preferably conducted by ICP. The term "pretreated"
means treated prior to use in a catalytic pyrolysis process.
[0129] The catalyst of the present invention may comprise an
aluminosilicate zeolite wherein the catalyst has been treated with
sufficient iron such that the Fe:acid site ratio, defined as the
number of moles of Fe in the catalyst divided by the number of
moles of acid sites in the catalyst that has not been treated with
Fe, is at least 0.1, or at least 0.5, or at least 1, or at least 5,
or at least 6, or less than 20, or less than 15, or less than 10,
or less than 7, or from 0.1 to 20, or from 0.5 to 10 wherein the Fe
is determined by elemental analysis and the acid site concentration
is determined by desorption of isopropylamine, where the elemental
analysis for Fe is preferably conducted by ICP.
[0130] A siliceous coating can be applied to a zeolite surface by
reaction with silicones or siloxanes as described elsewhere herein
and in the literature. A siliceous coating can be identified can be
identified by a higher (at least 10% higher or at least 30% higher
or at least 50% higher or at least 100% higher) in the exterior 50
A (or exterior 100 A) as compared to the Si/Al ratio at greater
depths in the catalyst. The preferred technique for analyzing the
Si/Al ratio is SEM/XPS before and after sputtering off 50 or 100 A,
or by cross-sectional analysis by SEM/XPS.
[0131] The following non-limiting examples are intended to
illustrate various aspects and features of the invention.
EXAMPLES
Catalyst A
[0132] A sample of a spray-dried catalyst containing 50%
NH.sub.4-ZSM-5 with a silica binder was obtained commercially
(Zeolite A) and calcined at 550 C for 2 hours in air to form
Catalyst A.
Catalyst B
[0133] A 200 g portion of the calcined Catalyst A was impregnated
with 105 g of aqueous solution of 14.3 wt % Fe(NO.sub.3).sub.3 in
deionized water to achieve a loading of 1.7% Fe by weight. The
impregnated sample was dried at 120 C for two hours in air and then
oven calcined five hours at 600 C in air with a temperature ramp
rate of 10 C/min. The catalyst thus obtained was tested in a
fluidized bed reactor for the conversion of biomass to aromatics
and olefins.
Catalyst C
[0134] A second batch of Fe-impregnated catalyst was prepared by
incipient wetness using the same procedure as Catalyst B. The
catalyst thus obtained was tested in a fluidized bed reactor for
the conversion of biomass to aromatics and olefins.
Catalyst D
[0135] A 200 g portion of calcined Catalyst A was mixed with 500 ml
of dried hexane in a 1 liter round-bottom flask. The mixture was
heated to reflux under N.sub.2 with stirring. While refluxing, 30
ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12
ml/h through a syringe pump. Upon completion of adding TEOS, the
mixture was refluxed for additional 30 minutes, and then was cooled
to ambient temperature. The solid material was collected by vacuum
filtration, dried one hour at 120 C and then calcined five hours at
600 C in air. The calcined material has a nominal composition of
4.0% SiO.sub.2/ZSM-5 and was tested in a fluidized bed reactor for
biomass conversion to aromatics and olefins. Another 200 g batch
was prepared following the same procedure and was used for the
preparation of Catalyst E.
Catalyst E
[0136] A 200 g portion of Catalyst D was impregnated with 105 g of
aqueous solution of 14.3 wt % Fe(NO.sub.3).sub.3 in deionized water
to achieve an iron loading of 1.7% Fe by weight. The impregnated
sample was dried at 120 C for two hours and then oven calcined five
hours at 600 C in air with a temperature ramp rate of 10 C/min. The
catalyst thus obtained has a nominal composition of 1.7% Fe, 4.0%
SiO.sub.2/ZSM-5 and was tested in a fluidized bed reactor for the
conversion of biomass to aromatics and olefins.
Catalyst F
[0137] A 200 g portion of calcined Catalyst A was mixed with 500 ml
of dried hexane in a 1 liter round-bottom flask. The mixture was
heated to reflux under N2 with stirring. While refluxing, 45 ml of
tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h
through a syringe pump. Upon completion of adding TEOS, the mixture
was refluxed for additional 30 minutes, and then was cooled to
ambient temperature. The solid material was collected by vacuum
filtration, dried one hour at 120 C and then calcined five hours at
600 C in air. The calcined material has a nominal composition of
5.7% SiO.sub.2/ZSM-5 and was tested in fluidized bed reactor for
biomass conversion to aromatics and olefins. Another 200 g batch
was prepared following the same procedure and was used for the
preparation of Catalyst G.
Catalyst G
[0138] A 200 g portion of Catalyst F was impregnated with 105 g of
aqueous solution of 14.3 wt % Fe(NO.sub.3).sub.3 in deionized water
to achieve an iron loading of 1.7% Fe by weight. The impregnated
sample was dried at 120 C for two hours and then oven calcined five
hours at 600 C in air with a temperature ramp rate of 10 C/min. The
catalyst thus obtained has a nominal composition of 1.7% Fe, 5.7%
SiO.sub.2/ZSM-5 and was tested in a fluidized bed reactor for the
conversion of biomass to aromatics and olefins.
Catalyst H
[0139] An aminofunctional oligomeric siloxane (Hydrosil 2627) was
diluted with deionized water at 1:1 weight ratio. A 400 g portion
of spray-dried NH.sub.4ZSM-5 (50% ZSM-5) obtained commercially was
impregnated with 194 g of the silane polymer/water solution. The
impregnated sample was dried two hours at 120 C and was then
calcined at 600 C for five hours in air with a temperature ramp
rate of 10 C/min. The catalyst thus obtained has a nominal
composition of 4% SiO.sub.2/ZSM-5. A portion of the calcined
material was tested in fluidized bed reactor for the conversion of
biomass to aromatics and olefins and the other portion was used for
preparing Catalyst I.
Catalyst I
[0140] A 200 g portion of Catalyst H was impregnated with 105 g of
aqueous solution of 14.3 wt % Fe(NO3)3 in deionized water to
achieve a loading of 1.7% Fe by weight. The impregnated sample was
dried at 120 C for two hours and then oven calcined five hours at
600 C with a temperature ramp rate of 10 C/min. The catalyst thus
obtained has a nominal composition of 1.7% Fe, 4% SiO.sub.2/ZSM-5
and it was tested in a fluidized bed reactor for the conversion of
biomass to aromatics and olefins.
Catalyst J
[0141] A sample of a second commercially obtained spray-dried
zeolite containing H-ZSM-5 with clay, an alumina binder, and some
residual carbon (Zeolite B) was calcined at 550 C for 2 hours in
air to form Catalyst J.
Catalyst K
[0142] A 200 g portion of calcined Catalyst J was mixed with 500 ml
of dried hexane in a 1 liter round-bottom flask. The mixture was
heated to reflux under N.sub.2 with stirring. While refluxing, 30
ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12
ml/h through a syringe pump. Upon completion of adding TEOS, the
mixture was refluxed for an additional 60 minutes, and then was
cooled to ambient temperature. The solid material was collected by
vacuum filtration, dried one hour at 120 C and then calcined five
hours at 600 C in air. The procedures were repeated to double the
amount of SiO2 loadings. The resulting material was impregnated
with 105 g of aqueous solution of 14.3 wt % Fe(NO3)3 in deionized
water to achieve a loading of 1.7% Fe by weight. The impregnated
sample was dried at 120 C for two hours and then oven calcined five
hours at 600 C with a temperature ramp rate of 10 C/min. The
catalyst thus obtained has a nominal composition of 1.7% Fe, 7.48%
SiO.sub.2/ZSM-5 and it was tested in a fluidized bed reactor for
the conversion of biomass to aromatics and olefins.
Catalyst L
[0143] A 200 g portion of calcined Catalyst J was mixed with 500 ml
of dried hexane in a 1 liter round-bottom flask. The mixture was
heated to reflux under N2 with stirring. While refluxing, 45 ml of
tetraethoxysilane (TEOS) was introduced to the mixture at 12 ml/h
through a syringe pump. Upon completion of adding TEOS, the mixture
was refluxed for an additional 30 minutes, and then was cooled to
ambient temperature. The solid material was collected by vacuum
filtration, dried one hour at 120 C and then calcined five hours at
600 C in air. The catalyst thus obtained has a nominal composition
of 5.7% SiO.sub.2/ZSM-5 and it was tested in a fluidized bed
reactor for the conversion of biomass to aromatics and olefins.
Catalyst M
[0144] A 200 g portion of calcined Catalyst J was mixed with 500 ml
of dried hexane in a 1 liter round-bottom flask. The mixture was
heated to reflux under N.sub.2 with stirring. While refluxing, 30
ml of tetraethoxysilane (TEOS) was introduced to the mixture at 12
ml/h through a syringe pump. Upon completion of adding TEOS, the
mixture was refluxed for an additional 60 minutes, and then was
cooled to ambient temperature. The solid material was collected by
vacuum filtration, dried one hour at 120 C and then calcined five
hours at 600 C in air. The procedures were repeated to double the
amount of SiO.sub.2 loadings. The catalyst thus obtained has a
nominal composition of 7.48% SiO.sub.2/ZSM-5 and it was tested in a
fluidized bed reactor for the conversion of biomass to aromatics
and olefins.
Catalysts N, O, P, Q, and R
[0145] Catalysts N, O, P, Q, and R were prepared using the
procedure of Catalyst D, except on a smaller scale with only 10 g
of Catalyst A and the appropriate amount of TEOS.
Catalysts S and T
[0146] Catalysts S and T were prepared using the procedure of
Catalyst H except on a smaller scale with only 10 g of Catalyst A
and the appropriate amount of Hydrosil 2627.
Catalyst U
[0147] Catalyst U was prepared by calcining a sample of Catalyst G
retrieved from Experiment 22 and then ion-exchanging it with a
solution of NH4NO3 in water for 2 hours. The catalyst was dried at
120 C to give Catalyst U.
Catalyst V
[0148] A sample of a third commercially obtained spray-dried
zeolite containing H-ZSM-5 with clay, an alumina binder, and some
residual carbon (Zeolite C) was calcined at 550 C for 2 hours in
air to form Catalyst V.
Catalysts W, X, Y, and Z
[0149] For each of these catalysts, a sample of Catalyst V was
impregnated with a measured amount of TEOS in accord with the
procedure used to make Catalyst L, except at a smaller scale. The
catalysts were calcined 5 hours at 600 C in air to give catalysts
W, X, Y, and Z.
Catalyst AA
[0150] A sample of a fourth commercially obtained spray-dried
zeolite containing --H-ZSM-5 was calcined at 600 C for 2 hours in
air to form Catalyst AA.
Catalyst AB--4% SiO2/ZSM5
[0151] A 10 g sample of the catalyst AA was suspended in a 200 mL
hexane in a round-bottom flask. The solution was heated under
stirring until reflux. A de-humidified N2 (60 mL/min) was used to
purge the reflux flask and the condenser throughout the entire
process. The mixture was allowed to reflux for 1 hour at
90-110.degree. C. and the system was filled with dry N2. A 1.5 mL
portion of TEOS, corresponding to 1.85 wt % Si (4.0 wt %
SiO.sub.2), was added into the mixture. Another 60 min was given
under stirring and refluxing after the addition of TEOS to complete
the silylation. The mixture was filtered to recover the catalyst.
The catalyst was dried at 120.degree. C. for 1 h and calcined at
600.degree. C. for 5 h.
Catalyst AC--4% La/4% SiO2/ZSM-5
[0152] Lanthanum was loaded onto the catalyst by the incipient
wetness method. A weighed portion of La(NO.sub.3).sub.3 (hydrated
form) was dissolved in de-ionized water to obtain the desired
concentration. A sample of Catalyst AB was impregnated by the
lanthanum nitrate solution until the catalyst pores were filled
with the solution (45 mL solution per 100 g of catalyst). The
La-loaded catalyst was dried at 120.degree. C. for 2 h and calcined
at 600.degree. C. for 5 h. In this example, 4 wt % La/was deposited
on the catalyst.
Catalyst AD--4% La/ZSM-5
[0153] Lanthanum was loaded onto the catalyst by the incipient
wetness method. A weighed portion of La(NO3)3 (hydrated form) was
dissolved in de-ionized water to obtain the desired concentration.
A sample of Catalyst AA was impregnated by the lanthanum nitrate
solution until the catalyst pores were filled with the solution (45
mL solution per 100 g of catalyst). The La-loaded catalyst was
dried at 120.degree. C. for 2 h and calcined at 600.degree. C. for
5 h. In this example, 4 wt % La/was deposited on the catalyst.
Catalyst AE--4% SiO2/ZSM-5
[0154] A 15 g sample of the catalyst AA was suspended in a 300 mL
hexane in a round-bottom flask. The solution was heated under
stirring until reflux. A de-humidified N2 (60 mL/min) was used to
purge the reflux flask and the condenser throughout the entire
process. The mixture was allowed to reflux for 1 hour at
90-110.degree. C. and the system was filled with dry N2. A 1.58 mL
portion of tetramethoxysilane (MEOS), corresponding to 1.85 wt % Si
(4.0 wt % SiO.sub.2), was added into the mixture. Another 60 min
was given under stirring and refluxing after the addition of MEOS
to complete the silylation. The mixture was filtered to recover the
catalyst. The catalyst was dried at 120.degree. C. for 1 h and
calcined at 600.degree. C. for 5 h.
Catalyst AF--4% SiO2/ZSM-5
[0155] The procedure of Catalyst AE was repeated.
Catalyst AG--4% SiO2/4% La/ZSM-5
[0156] A 15 g sample of the catalyst AH was suspended in a 300 mL
hexane in a round-bottom flask. The solution was heated under
stirring until reflux. A de-humidified N.sub.2 (60 mL/min) was used
to purge the reflux flask and the condenser throughout the entire
process. The mixture was allowed to reflux for 1 hour at
90-110.degree. C. and the system was filled with dry N.sub.2. A
1.58 mL portion of tetramethoxysilane (MEOS), corresponding to 1.85
wt % Si (4.0 wt % SiO.sub.2), was added into the mixture. Another
60 min was given under stirring and refluxing after the addition of
MEOS to complete the silylation. The mixture was filtered to
recover the catalyst. The catalyst was dried at 120.degree. C. for
1 h and calcined at 600.degree. C. for 5 h.
Catalyst AH--4% La/4% SiO2/ZSM-5
[0157] Lanthanum was loaded onto the catalyst by the incipient
wetness method. A weighed portion of La(NO3)3 (hydrated form) was
dissolved in de-ionized water to obtain the desired concentration.
A sample of Catalyst AE was impregnated by the lanthanum nitrate
solution until the catalyst pores were filled with the solution (45
mL solution per 100 g of catalyst). The La-loaded catalyst was
dried at 120.degree. C. for 2 h and calcined at 600.degree. C. for
5 h. In this example, 4 wt % La/was deposited on the catalyst.
Catalyst AI 4.6% SiO2/ZSM-5
[0158] A sample of Catalyst A was silylated by Chemical Liquid
Deposition (CLD) using tetraethoxysilane (TEOS). A 314 g sample of
the parent catalyst was suspended in a 700 mL dried hexane in a
round-bottom flask. The solution was heated under stirring until
reflux. A de-humidified N2 (60 mL/min) was used to purge the reflux
flask and the condenser throughout the entire process. After
reflux, another 1 hour was given to ensure the temperature was
stabilized at 90-110.degree. C. and the system was filled with dry
N.sub.2. A 56.25 mL portion of TEOS, corresponding to 2.2 wt % Si
(4.6 wt % SiO.sub.2), was introduced into the mixture at 12 mL/h
using a syringe pump (SyringePump, NE-300). The mixture was stirred
under reflux an additional 60 minutes to complete the silylation.
The mixture was filtered to recover solid catalyst. The catalyst
was dried at 120.degree. C. for 1 h and calcined at 600.degree. C.
for 5 h.
Catalyst AJ. 0.5% Fe/4.6% SiO2/ZSM-5
[0159] A sample of Catalyst AI was impregnated with Fe by incipient
wetness. A portion of Fe(NO3)3 (hydrated form) was dissolved in
de-ionized water to obtain a solution of 4.2 wt % Fe(NO3)3.
Catalyst F was then impregnated by the iron nitrate solution until
the catalyst pores were filled with the solution (45 mL solution
per 100 g of catalyst). The catalyst was dried at 120.degree. C.
for 2 h and calcined at 600.degree. C. for 5 h.
Catalyst AK. 5% Fe/4.6% SiO2/ZSM-5
[0160] A sample of Catalyst AI was impregnated with Fe by incipient
wetness. A portion of Fe(NO.sub.3).sub.3 (hydrated form) was
dissolved in de-ionized water to obtain a solution of 42 wt %
Fe(NO.sub.3).sub.3. Catalyst F was then impregnated by the iron
nitrate solution until the catalyst pores were filled with the
solution (45 mL solution per 100 g of catalyst). The catalyst was
dried at 120.degree. C. for 2 h and calcined at 600.degree. C. for
5 h.
Example 1
[0161] Catalytic fast pyrolysis (CFP) experiments were conducted in
a fluidized bed reactor. The fluidized bed reactor was 1.94 inches
in diameter (ID) and 24 inches in height and was made of 316
stainless steel. Inside the reactor, the catalyst bed was supported
by a distributor plate made of 316 stainless steel plate with 1/16
inch circular openings.
[0162] Biomass was charged to the biomass hopper and its flow rate
was controlled by an augur inside the hopper that delivers the
biomass to the feed tube. A curved 1/4-inch OD 316-stainless steel
tube extended from the feed hopper to the biomass inlet port. A
series of sieve trays made of perforated 316 stainless steel with
1/8 inch openings and 42% open area were installed inside the
reactor. There were six sieve trays attached to a central, threaded
rod with a 1-inch spacing between the sieve trays. The reactor was
loaded with 172.97 g of catalyst A prior to the experiment and the
catalyst was calcined in-situ in air at the flow rate of 1.5 SLPM
and 3 SLPM of N.sub.2 for 2 hours at 600.degree. C. The biomass
feed was sieved to 20-40 mesh particle size. A 383.9 gram portion
of hardwood pellets (46.67% C) was weighed and loaded into the
hopper system. The reactor was purged with a flow of N2 at 3.0 SLPM
for 30 minutes prior to starting the biomass conversion.
[0163] The reactor was heated to 575.degree. C. and the
fluidization gas feeding tube was heated to approximately
500.degree. C. The solid biomass was introduced into the reactor
from a side feeding tube with N2 flow. Gas flow rate through the
biomass screw auger feed tube was 3.0 SLPM. The biomass flow rate
was adjusted to approximately 1.4 g/min and 41.49 g of biomass was
fed during the 30 minute experiment. During reaction, 1.5 SLPM of
N.sub.2 was passed into the reactor through the distributor plate
to fluidize the catalyst in addition to the feeding tube N.sub.2
flow.
[0164] The reactor effluent exited the reactor from the top through
a heated cyclone (350.degree. C.) to remove solid particles,
including small catalyst and char particles. The effluent exiting
the cyclone flowed into a product collection system that included
two bubblers and three condensers. The bubblers were placed in an
ice water bath and charged with 150 ml of isopropanol inside as
solvent; the three condensers contained no solvent and were placed
inside a Dry Ice/isopropanol bath. The uncondensed gas phase
products that exited the last condenser were collected in gas bags.
The reaction time was typically 30 min and two gas bag samples were
taken at 15 and 30 minutes time on stream after initiating the feed
of biomass. After each bag was taken, the total gas flow rate was
measured with a bubble flow meter; at least 4 measurements were
made and the average was used for performance calculations. The gas
bags samples were analyzed by injection into a Shimadzu GC 2014
equipped with TCD and FID detectors that had been calibrated with
analytical standard gas mixtures.
[0165] The contents of the two bubblers were combined into a first
sample for analysis. The three condensers were rinsed with
isopropanol to collect the contents, and the emptied bubblers were
rinsed with isopropanol; the condenser rinse and bubbler rinse were
combined into a second sample. The volumes of the two liquid
samples were measured and weights determined. Liquid samples were
analyzed by injection into a Shimadzu GC 2010-plus equipped with an
FID detector, and 60 meter Rtx-1 capillary column from Restek
[0166] After the experiment the reactor was flushed an additional
15 minutes with N2 to ensure that the condensable products were
swept into the product collection train, and then allowed to cool.
The solid catalyst and char were removed from the reactor and
sieved through a 60 mesh sieve. The 169.26 g portion passing
through the sieve contained most of the catalyst. A small portion
of this material was analyzed for carbon content on a Total Organic
Carbon analyzer (TOC). The 4.19 g of larger particles were mostly
char derived from biomass. This fraction was also analyzed for
carbon content on a TOC. The carbon yield of aromatics was
determined to be 26.47% of the carbon fed.
Example 2
[0167] The experiment in Example 1 was repeated with 173.46 g of
Catalyst B and 46.59 g of hardwood.
Example 3
[0168] The experiment in Example 1 was repeated with 173 g of
Catalyst C and 46.4 g of hardwood and a temperature of 525 C.
[0169] A comparison of Examples 2 and 3 with Example 1 shows the
effect of iron promotion on the selectivity of the xylenes. A
measurable increase in the selectivity to p-xylene was observed,
from 49.0% selectivity to p-xylene without iron to 52.3 or 54.2%
selectivity to p-xylene for two separate catalyst preparations
containing iron in Examples 2 and 3, respectively.
[0170] Example 3 shows that the increase in the selectivity to
p-xylene with the iron promoted catalyst extends to lower
temperature (525 C) as well as at 575 C.
Example 4
[0171] The experiment in Example 1 was repeated with 178.3 g of
Catalyst D and 50.2 g of newsprint (41.70% Carbon) and a
temperature of 525 C.
Example 5
[0172] The catalyst in Example 4 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 4 was repeated with regenerated Catalyst D
and 63.72 g of newsprint.
[0173] Examples 4 and 5 show the impact of coating the catalyst
with SiO.sub.2 on the selectivity to p-xylene. With 4% SiO.sub.2
coating the selectivity to p-xylene increased from 49.0% from
Example 1 to 71.5 and 75.5% selectivity to p-xylene in Examples 4
and 5, respectively.
Example 6
Cycle 1
[0174] The experiment in Example 1 was repeated with 173.0 g of
Catalyst E and 52.29 g of hardwood and a temperature of 550 C.
[0175] Example 6 shows the surprising effect of the addition of
iron to a catalyst that is coated with 4% SiO2 on the selectivity
to p-xylene. Without the added iron the selectivity to p-xylene was
71.5 and 75.5 (Examples 4 and 5, respectively), whereas with the
addition of 1.7% Fe the selectivity to p-xylene increased to 84.1%
among the xylenes.
Example 7
Cycle 2
[0176] The catalyst in Example 6 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 6 was repeated with regenerated Catalyst E
and 48.1 g of hardwood at a temperature of 550 C.
[0177] Example 7 shows the surprising effect on the selectivity to
p-xylene of cycling of the catalyst. As the catalyst was cycled
through biomass upgrading and regeneration cycles, the selectivity
to p-xylene increased from the initial 84.1% selectivity to
p-xylene in the first cycle (Example 6) to 84.9% in the second
cycle (Example 7).
Example 8
Cycle 3
[0178] The catalyst in Example 7 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 6 was repeated with regenerated Catalyst E
and 50.24 g of hardwood and a temperature of 525 C.
[0179] The results of Example 8 show that even at lower temperature
(525 C) the selectivity to p-xylene of a catalyst that has been
cycled (85.2%) is greater than the selectivity to p-xylene of the
fresh catalyst (84.1%).
Example 9
Cycle 4
[0180] The catalyst in Example 8 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 8 was repeated with regenerated Catalyst E
and 52.76 g of hardwood at a temperature of 525 C.
Example 10
Cycle 5
[0181] The catalyst in Example 9 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 6 was repeated with regenerated Catalyst E
and 50.19 g of newsprint and a temperature of 600 C.
[0182] Experiment 10 further shows that the improved selectivity to
p-xylene is maintained with different feedstocks, in this case
newsprint, as well as with hardwood.
Example 11
Cycle 6
[0183] The catalyst in Example 10 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 6 was repeated with regenerated Catalyst E
and 45.1 g of hardwood and a temperature of 568 C.
[0184] Examples 7 through 11 show the surprising effect on the
selectivity to p-xylene of repeated cycling of the catalyst. As the
catalyst was cycled through biomass upgrading and regeneration
cycles, the selectivity to p-xylene increased from the initial
84.1% selectivity to p-xylene in the first cycle (Example 6) to
84.9%, 85.2%, 86.7%, 85.9%, and 86.0% p-xylene in successive cycles
(Examples 7 through 11). These surprising results show that cycling
the catalyst improves the p-xylene selectivity compared to the
84.1% selectivity to p-xylene in the fresh catalyst, and that the
improved selectivity to p-xylene is maintained for multiple cycles
of biomass conversion and catalyst regeneration.
[0185] Example 11 shows that the yield of aromatics of a catalyst
that has been cycled 6 times and operated at 568 C (18.43%
aromatics, Example 11) is higher than the fresh catalyst when
operated at 550 C (17.93% aromatics, Example 6). Moreover, the
yield of most desirable products, ie aromatics plus olefins, is
greater with the catalyst that has been cycled 6 times (31.06%
carbon yield of aromatics plus olefins, Example 11) than the fresh
catalyst operated at 550 C (25.75% carbon yield of aromatics plus
olefins, Example 6).
Example 12
[0186] The Experiment in Example 1 was repeated with 173.0 g of
Catalyst F and 49.45 g of newsprint and a temperature of 525 C.
Example 13
[0187] The catalyst in Example 12 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 12 was repeated with regenerated Catalyst F
and 48.2 g of newsprint and a temperature of 525 C.
[0188] Examples 12 and 13 show that a catalyst coated with 5.7% by
weight SiO.sub.2 has a greater selectivity to p-xylene (79.4% and
83.8% selectivity to p-xylene in Examples 12 and 13, respectively)
than a catalyst coated with 4% by weight SiO.sub.2 (71.5% and 75.5%
selectivity to p-xylene in Examples 4 and 5, respectively).
Moreover, for both the first cycle of biomass conversion with the
catalyst (79.4% selectivity to p-xylene in Example 12 vs 71.5%
selectivity to p-xylene in Example 4) and for a catalyst in the
second cycle (83.8% selectivity to p-xylene in Example 13 vs 75.5%
selectivity to p-xylene in Example 5), the selectivity to p-xylene
is greater for the catalyst with 5.7% by weight SiO.sub.2 than for
the catalyst with 4% by weight SiO.sub.2.
Example 14
Cycle 1
[0189] The Experiment in Example 1 was repeated with 173.0 g of
Catalyst G and 49.97 g of hardwood and a temperature of 525 C. The
results show that for the catalyst with 5.7% by weight SiO.sub.2
and 1.7% by weight Fe, the selectivity to p-xylene (86.5%) was
greater than for a catalyst with 1.7% Fe (52.3% and 54.2%
selectivity to p-xylene in Examples 2 and 3, respectively). The
results show that for the catalyst with 5.7% by weight SiO2 and
1.7% by weight Fe, the selectivity to p-xylene (86.5%) was greater
than for a catalyst with 5.7% by weight SiO.sub.2 (79.4% and 83.8%
selectivity to p-xylene in Examples 12 and 13, respectively). This
surprising result shows the synergistic effect of a catalyst that
is promoted with Fe in addition to being coated with SiO.sub.2.
Example 15
Cycle 2
[0190] The catalyst in Example 14 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 14 was repeated with regenerated Catalyst G
and 51.76 g of newsprint and a temperature of 525 C.
Example 16
Cycle 3
[0191] The catalyst in Example 15 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 15 was repeated with regenerated Catalyst G
and 59.25 g of hardwood.
[0192] Example 16 shows the surprising effect on the selectivity to
p-xylene of cycling of the catalyst. As the catalyst was cycled
through biomass upgrading and regeneration cycles, the selectivity
to p-xylene increased from the initial 86.5% selectivity to
p-xylene in the first cycle (Example 14) to 87.9% in the third
cycle (Example 16).
Example 17
Cycle 4
[0193] The catalyst in Example 16 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 16 was repeated with regenerated Catalyst G
and 48.25 g of hardwood.
[0194] Example 17 shows the surprising effect on the selectivity to
p-xylene of cycling of the catalyst. As the catalyst was cycled
through biomass upgrading and regeneration cycles, the selectivity
to p-xylene increased from the 87.9% selectivity to p-xylene in the
third cycle (Example 16) to 89.3% in the fourth cycle (Example
17).
Example 18
Cycle 5
[0195] The catalyst in Example 17 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 17 was repeated with regenerated Catalyst G
and 49.05 g of hardwood.
Example 19
Cycles 6 through 12
[0196] The catalyst in Example 18 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 18 was repeated with regenerated Catalyst G
and 45 to 60 g of hardwood 6 additional times, regenerating the
catalyst with air and N2 for 2 hours at 600 C after each exposure
to hardwood. The Experiment in Example 18 was repeated a seventh
time with regenerated Catalyst G and 46.46 g of hardwood. At this
juncture the catalyst has participated in 12 cycles of biomass
conversion.
Example 20
Cycle 13
[0197] The catalyst in Example 19 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 18 was repeated with regenerated Catalyst G
and 48.88 g of hardwood.
Example 21
Cycle 14
[0198] The catalyst in Example 20 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 20 was repeated with regenerated Catalyst G
and 47.62 g of hardwood.
Example 22
Cycle 15
[0199] The catalyst in Example 21 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 21 was repeated with regenerated Catalyst G
and 46.87 g of hardwood.
[0200] Examples 16 through 22 show the surprising effect on the
selectivity to p-xylene of repeated cycling of the catalyst. As the
catalyst was cycled through biomass upgrading and regeneration
cycles, the selectivity to p-xylene increased from the initial
84.1% selectivity to p-xylene in the first cycle (Example 14) to
87.9%, 89.3%, 89.3%, 89.1%, 89.3%, 88.8%, 89.6% and 89.3%
selectivity to p-xylene in 3, 4, 5, 12, 13, 14, and 15 cycles
(Examples 16 through 22). These surprising results show that
cycling the catalyst improves the p-xylene selectivity compared to
the 86.5% selectivity to p-xylene in the first cycle with the
catalyst, and that the improved selectivity to p-xylene is
maintained for multiple cycles of biomass conversion and catalyst
regeneration.
[0201] Examples 21 and 22 show that the yields of aromatics of a
catalyst that has been cycled 14 and 15 times and operated at 525 C
(20.79% aromatics, Example 21, 19.17% aromatics, Example 22) are
higher than the fresh catalyst when operated at 525 C (18.82%
aromatics, Example 14). Moreover, the yield of most desirable
products, ie aromatics plus olefins, is greater with the catalyst
that has been cycled 14 or 15 times (31.77% and 30.42% carbon yield
of aromatics plus olefins, Example 21 and Example 22, respectively)
than the first cycle with the catalyst operated at 525 C (28.91%
carbon yield of aromatics plus olefins, Example 14).
Example 23
[0202] The Experiment in Example 1 was repeated with 135.8 g of
Catalyst H and 48 g of hardwood and a temperature of 525 C.
[0203] Example 23 shows that the catalyst can be coated with any of
a variety of sources of silica and still provide increased
selectivity of p-xylene. Example 23 shows that a silicone polymer
can be used as the SiO.sub.2 source and provide high selectivity to
p-xylene (84.0% in Example 23) compared to the non-coated material
(49.0% in Example 1).
Example 24
[0204] The catalyst in Example 23 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 23 was repeated with regenerated Catalyst H
and 48 g of hardwood.
[0205] Example 24 shows that a catalyst coated with polymer-derived
SiO.sub.2 that has been cycled through a biomass conversion and
catalyst regeneration cycle retains a high selectivity to p-xylene
(84.4%), and that the selectivity to p-xylene increases with
cycling (84.4% vs 84.0% for the first cycle, Example 23).
Example 25
[0206] The Experiment in Example 23 was repeated with 173.02 g of
Catalyst I and 24 g of hardwood and a temperature of 525 C.
[0207] Example 25 shows that addition of iron to a catalyst coated
with SiO.sub.2, wherein the SiO.sub.2 was derived from a silicone
polymer, provides an increased selectivity to p-xylene (88.1%
selectivity to p-xylene in Example 25) compared to the coated
catalyst without Fe addition (84.0% selectivity to p-xylene in
Example 23).
Example 26
[0208] The catalyst in Example 25 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 25 was repeated with regenerated Catalyst I
and 26 g of hardwood.
[0209] Example 26 shows that a catalyst coated with polymer-derived
SiO.sub.2 that has been cycled through a biomass conversion and
catalyst regeneration cycle retains a higher selectivity to
p-xylene (87.8%) compared to the catalyst without iron addition
that has been cycled through a biomass conversion and catalyst
regeneration cycle (84.4% in Example 24).
[0210] Moreover Examples 6-11, 14-22, and 25-26 show that iron
addition to a SiO.sub.2 coated catalyst increases the selectivity
to p-xylene for different SiO2 loadings and sources of the
SiO.sub.2.
Example 27
[0211] The Experiment in Example 23 was repeated with 170.87 g of
Catalyst J and 51 g of hardwood and a temperature of 571 C. The
results show that with Catalyst J the selectivity to p-xylene is
46.6%.
Example 28
[0212] The catalyst in Example 27 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 27 was repeated with regenerated Catalyst J
and 27 g of hardwood and a temperature of 572 C.
[0213] The results of Example 28 show that the selectivity to
p-xylene with Catalyst J that has no coating or iron addition
remains low and declines relative to the first cycle (45.0%
selectivity to p-xylene in the second cycle, Example 28 vs 46.6%
selectivity to p-xylene in the first cycle).
Example 29
[0214] The Experiment in Example 27 was repeated with 173.02 g of
catalyst K and 57 g of hardwood and a temperature of 523 C.
[0215] Results of Example 29 show that a catalyst that contains
clay and is coated with 7.48% SiO.sub.2 and has added iron provides
a higher selectivity to p-xylene (90.7% in Example 29) than the
uncoated, unpromoted catalyst (Examples 27-28 where the
selectivities to p-xylene were 46.4 and 45.0, respectively) or than
the coated, unpromoted catalyst (Examples 34-35, where the
selectivities to p-xylene were 87.5 and 88.3, respectively).
Example 30
[0216] The catalyst in Example 29 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 29 was repeated with regenerated Catalyst K
and 60 g of hardwood and a temperature of 522 C.
[0217] Example 30 shows that a catalyst coated with 7.48% SiO2 and
promoted with Fe that has been cycled through a biomass conversion
and catalyst regeneration cycle retains a high selectivity to
p-xylene (91.3%), and that the selectivity to p-xylene unexpectedly
increases with cycling (91.3% vs 90.7%% for the first cycle,
Example 29). Example 30 shows that with an appropriate combination
of catalyst components, catalyst binder, SiO.sub.2 coating content,
and promoter elements, the selectivity to p-xylene is increased
with cycling in biomass conversion and catalyst regeneration
cycles.
Example 31
[0218] The Experiment in Example 23 was repeated with 170.87 g of
Catalyst L and 51 g of hardwood and a temperature of 526 C.
[0219] The results of Example 31 show that coating a
clay-containing catalyst with 5.7% by weight SiO2 can increase the
selectivity to p-xylene (85.3% in Example 31) compared to the
uncoated catalyst (46.6% in Example 27). The results of Example 31
show that the increase in selectivity to p-xylene produced by
SiO.sub.2 coating is general for ZSM-5 containing catalysts with
different binders and from different sources.
Example 32
[0220] The catalyst in Example 31 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 31 was repeated with regenerated Catalyst L
and 27 g of hardwood and a temperature of 527 C.
[0221] Example 32 shows that a catalyst coated with SiO.sub.2 that
has been cycled through a biomass conversion and catalyst
regeneration cycle retains a high selectivity to p-xylene (76.4%).
Example 32 shows that the selectivity to p-xylene with a catalyst
containing ZSM-5 and comprising clay and coated with 5.7% SiO.sub.2
does not increase with cycling (76.4%% vs 85.3%% for the first
cycle, Example 31).
Example 33
[0222] The catalyst in Example 32 was regenerated in the reactor by
reaction with air and N2 over the course of 2 hours at 600 C. The
Experiment in Example 31 was repeated with regenerated Catalyst L
and 24 g of hardwood and a temperature of 553 C.
[0223] Example 33 shows that a catalyst coated with SiO.sub.2 that
has been cycled through a biomass conversion and catalyst
regeneration cycle retains a high selectivity to p-xylene (74.2%).
Example 33 shows that the selectivity to p-xylene with a ZSM-5
containing catalyst comprising clay and coated with 5.7% SiO.sub.2
does not increase with cycling (74.2% for the third cycle vs 85.3%%
for the first cycle, Example 31, and 76.4% for the second cycle,
Example 32).
[0224] Examples 32 and 33 show that the increase in selectivity to
p-xylene is limited to particular combinations of catalyst
components, catalyst binder, SiO.sub.2 coating content, and
presence of promoter elements.
Example 34
[0225] The Experiment in Example 23 was repeated with 162.9 g of
Catalyst M and 45 g of hardwood and a temperature of 522 C.
[0226] The results of Example 34 show that coating a
clay-containing catalyst with 7.48% by weight SiO2 can increase the
selectivity to p-xylene (87.5% in Example 34) compared to the
uncoated catalyst (46.6% in Example 27). The results of Example 34
show that the increase in selectivity to p-xylene produced by
SiO.sub.2 coating is general for ZSM-5 containing catalysts with
different binders and from different sources. The results of
Example 34 show that with some catalysts a greater amount of
SiO.sub.2 coating increases the selectivity to p-xylene compared to
a lesser amount of SiO.sub.2; Example 34 with a catalyst having
7.48% SiO.sub.2 coating provided 87.5% selectivity to p-xylene
compared to Example 31 with a catalyst having 5.7% SiO.sub.2
coating that provided 85.3% selectivity to p-xylene.
Example 35
[0227] The catalyst in Example 34 was regenerated in the reactor by
reaction with air and N2 over the course of 3 hours at 600 C. The
Experiment in Example 34 was repeated with regenerated Catalyst M
and 45 g of hardwood and a temperature of 525 C.
[0228] Example 35 shows the surprising effect on the selectivity to
p-xylene of cycling of the catalyst. As the catalyst was cycled
through biomass upgrading and regeneration cycles, the selectivity
to p-xylene unexpectedly increased from the initial 87.5% in the
first cycle (Example 34) to 88.3% in the second cycle (Example 35).
Example 35 shows that with an appropriate combination of catalyst
components, catalyst binder, SiO.sub.2 coating content, and
promoter elements, the selectivity to p-xylene is increased with
cycling in biomass conversion and catalyst regeneration cycles.
Examples 36-51
[0229] In order to determine the acid site concentration on various
materials the desorption of isopropylamine from the materials is
measured by means of Isopropyl-amine Temperature-Programmed
Desorption (IPA-TPD). For the IPA-TPD experiments, a TGA instrument
(Shimadzu TGA-50) is adjusted to read zero with an empty platinum
sample cell. The sample cell is then filled with a sample of
catalyst powder (30-50 mg). The catalyst is pre-treated at 500 C
under 50 mL/min N2. It is then cooled to 120 C under a 50 mL/min
flow of N2. Isopropylamine (IPA) is fed into the TGA chamber at
this temperature by flowing a 2nd portion of N2 gas (<10 mL/min)
through a bubbler filled with liquid IPA while monitoring the
weight of the sample. The feed of IPA is stopped when the catalyst
is saturated as indicated by no more weight increase. The flows of
N2 are maintained through the chamber, but bypassing the IPA
bubbler, for an additional 120 min to remove weakly adsorbed'IPA.
The TGA chamber is then heated up to 700.degree. C. at a ramping
rate of 10.degree. C./min to obtain desorption curves, and the
weight is monitored as a function of temperature.
[0230] In the desorption curve, the sharp desorption at 270-380 C
is assigned to IPA decomposition into propylene and NH3 occurring
on the Bronsted acid sites. The peak area under the desorption
curve measured from 270 to 380 C is used for quantifying the number
of Bronsted acid sites for a particular sample.
[0231] Examples 36-41 and FIG. 2 show that the addition of a
coating, e.g., silicone coating derived from TEOS, reduces the
concentration of acid sites on the catalyst and that more coating
reduces the concentration further.
[0232] Example 41 and FIG. 2 show that at some level of SiO.sub.2
coating (ie 11.46% SiO.sub.2 in Example 41 compared to 7.80 wt % in
Examples 39 and 40 and lower wt % in Examples 37-38), the
concentration of acid sites is no longer significantly reduced by
addition of more SiO.sub.2.
[0233] Examples 42 and 43 show that a polysiloxane coating reduces
the concentration of acid sites compared to the concentration of
acid sites in the parent catalyst in Example 36. Example 43 shows
that increasing the amount of coating with a polysiloxane from
3.96% to a larger amount does not further reduce the concentration
of acid sites.
[0234] Example 42 shows that the use of a polysiloxane reduces the
concentration of acid sites more than an equivalent amount of
SiO.sub.2 derived from TEOS (compare to Example 37). Example 43 and
FIG. 2 shows that addition of polysiloxane beyond 3.96% by weight
does not significantly reduce the concentration of acid sites in
the catalyst.
[0235] Example 44 shows that the addition of Fe to a silicone
coated catalyst further reduces the concentration of acid sites
compared to the catalyst without Fe in Example 39.
[0236] Example 45 shows that cycling the coated and Fe-promoted
catalyst multiple times through biomass upgrading and catalyst
regeneration reduces the concentration of acid sites and with
Example 22 shows that the carbon yields of aromatics (19.17%) and
olefins (11.25%) remain high and the selectivity to p-xylene
remains high (89.3%), and is higher than in the first cycle.
Example 45 shows that a catalyst with reduced concentration of acid
sites (5.9% of the parent zeolite) can be effectively used to
convert biomass to useful products.
[0237] Example 46 shows that the concentration of acid sites in a
silicone coated and Fe-promoted catalyst that has been cycled
multiple times can be increased by an ion exchange with acid
solution (NH.sub.4NO.sub.3).
[0238] Example 47 through 50 and FIG. 2 shows that a commercially
obtained spray-dried zeolite containing H-ZSM-5 with clay, an
alumina binder, and some residual carbon (Zeolite C) shows a
reduction in acid concentration with added coating of SiO.sub.2.
Examples 48 through 50 show that addition of SiO.sub.2 beyond 5.89%
by weight SiO.sub.2 does not significantly further reduce the
concentration of acid sites
[0239] Examples 36-43 and examples 47-51 show that the moles of
Bronsted acid sites deactivated by the siliceous coating is at
least 0.015 or at least 0.03, or at least 0.04, or at least 0.06
per mole of Si added as measured by desorption of isopropyl amine
(IPA) in a temperature programmed desorption experiment.
Example 52
[0240] The 2-methylfuran conversion was carried out in 1/2'' O.D.
stainless steel tubular reactor. A 1-gram sample of Catalyst AA was
held by a piece of quartz wool inside the reactor. N2 was used as
the carrier gas. 2-methylfuran (b.p.=65.degree. C.) was dripped
into the top of the reactor by an HPLC pump and it was carried with
a down-flow of N.sub.2 into the catalyst bed. The product stream
was collected in gas bags at 4, 7, 10, 15, and 20 minutes time on
stream. Coke deposited on the catalyst was quantified by using a
TOC instrument (Shimadzu SSM-5000A). The results in Table 4 are
averages of the samples collected after 7 and 10 minutes have
elapsed since the introduction of 2-methylfuran.
Example 53
[0241] The experiment from Example 52 was repeated with Catalyst AB
in place of Catalyst AA. The experiment was repeated after
regenerating the catalyst by exposure to air at 650 C. The cycle
was repeated for a total of 3 experiments with Catalyst AB. The
results in Table 4 are the average of the three experiments.
[0242] Example 53 shows that coating a zeolite with 4 wt %
SiO.sub.2 using TEOS as the source of the coating results in high
selectivity to p-xylene (92.3%) compared to the uncoated catalyst
(75.3%).
Example 54
[0243] The experiment in Example 52 was repeated with Catalyst AC
in place of Catalyst AA. Example 54 shows that addition of 4% by
weight of La to a catalyst coated with 4% by weight SiO.sub.2
increases the aromatic yield (16.6% vs 12.3%) compared to the
catalyst that has no La added and that the high selectivity to
p-xylene is maintained (92.1% vs 93.3%).
Example 55
[0244] The experiment in Example 52 was repeated with Catalyst AD
in place of Catalyst AA. Example 55 shows that the addition of 4%
by weight of La to an uncoated catalyst improves the aromatics
yield (16.1% vs 11.7%) and the p-xylene selectivity (80.2% vs
75.3%).
Example 56
[0245] The experiment in Example 52 was repeated with Catalyst AE
in place of Catalyst AA. Example 56 shows that a catalyst coated
with SiO.sub.2 using MEOS as the source of SiO.sub.2 shows increase
p-xylene selectivity compared to the uncoated catalyst (95.5% vs
75.3%).
Example 57
[0246] The experiment in Example 52 was repeated with Catalyst AF
in place of Catalyst AA. Example 57 shows that a catalyst prepared
at larger scale provides similar improvement in p-xylene
selectivity as the catalyst prepared at the smaller scale.
Example 58
[0247] The experiment in Example 52 was repeated with Catalyst AG
in place of Catalyst AA. Example 58 shows that when the catalyst
that is promoted first with 4% by weight La and then silicone
coated the yield of aromatics is increased compared to the coated
catalysts (14.7% vs 9.6%) but the p-xylene selectivity is lower
than the catalyst that has only been coated (80.2% vs 75.3%).
Example 58 shows that a La-promoted and then silicone coated
catalyst has higher yield of aromatics (14/7% vs 11.7%) and
p-xylene selectivity (86.2% vs 75.3%) compared to an unpromoted,
uncoated catalyst.
Example 59
[0248] The experiment in Example 52 was repeated with Catalyst AH
in place of Catalyst AA. Example 59 shows that when the catalyst
that is first coated with 4% by weight SiO.sub.2 and then promoted
with 4% La the yield of aromatics (14.8 vs 14.7%) and the p-xylene
selectivity (95.3% vs 86.2%) are higher than when the catalyst is
first impregnated with La and then coated with SiO.sub.2; Example
59 shows that the sequence of catalyst promotion and coating is
critical to obtaining an improved catalyst and that the sequence in
which the catalyst is first coated with silicone and then promoted
with metal addition provides high yield of aromatics and higher
selectivity to p-xylene compared to the catalyst prepared by
promotion followed by coating.
Example 60
[0249] The furfural conversion was carried out in 1/2'' O.D.
stainless steel tubular reactor. A 1-gram sample of Catalyst F was
held by a piece of quartz wool inside the reactor. N.sub.2 was used
as the carrier gas. Furfural (b.p.=162.degree. C.) was dripped into
the top of the reactor by an HPLC pump and it was carried with a
down-flow of N.sub.2 into the catalyst bed. The product stream was
collected in gas bags at 4, 8, 12, and 16 minutes time on stream.
Coke deposited on the catalyst was quantified by using a TOC
instrument (Shimadzu SSM-5000A).
[0250] The results in Table 4 are averages of the samples collected
after 8 and 12 minutes have elapsed since the introduction of
furfural.
Example 61
[0251] The experiment of Example 60 was repeated with Catalyst AI
in place of Catalyst F.
[0252] The results in Table 5 show that addition of 0.5% Fe to the
coated catalyst increased the selectivity of p-xylene (83.2% vs
74.9%) and yield of p-xylene (1.70% vs 1.36%) by the addition of
Fe.
Example 62
[0253] The experiment in Example 60 was repeated with Catalyst AJ
in place of Catalyst F.
[0254] The results in Table 5 show that addition of 5.0% Fe to the
coated catalyst increased the selectivity of p-xylene (84.6% vs
74.9%) and yield of p-xylene (2.04% vs 1.36%) by the addition of 5%
Fe compared to the Fe-free catalyst, and that the addition of 5.0%
Fe increased the selectivity of p-xylene (84.6% vs 83.2%) and yield
of p-xylene (2.04% vs 1.70%) by the addition of 5.0% Fe compared to
the 0.5% Fe catalyst. The results in Table 5 show that the addition
of Fe to the catalyst decreased the ratio of CO/CO.sub.2 with
increasing Fe content indicating a more efficient rejection of
oxygen from the product mixture with added Fe.
TABLE-US-00001 TABLE 1 Catalyst Preparations. The Wt % SiO2 is the
fraction of SiO2 added in the treatment step. Run Metal Catalyst
Numbers Zeolite Additive Wt % SiO2 SiO2 Source A 1 A none none none
B 2 A 1.7% Fe none none C 3 A 1.7% Fe None none D 4-5 A none 4 TEOS
E 6-11 A 1.7% Fe 4 TEOS F 12-13 A none 5.7 TEOS G 14-22 A 1.7% Fe
5.7 TEOS H 23-24 A none 4 H2O soluble polymer I 25-26 A 1.7% Fe 4
H2O soluble polymer J 27-28 B none none none K 29-30 B 1.7% Fe 7.48
TEOS L 31-33 B none 5.7 TEOS M 34-35 B none 7.48 TEOS
TABLE-US-00002 TABLE 6 Bronsted acid concentration obtained from
IPA-TPD experiments Bronsted acid % of Acid Bronsted acid Fe
concentration Sites in Fe added Fe: sites deactivated content
(mmol/mg) .times. Parent (mmol/mg) .times. Acid site per Fe added
Catalyst (wt %) 1000 Zeolite 1000 ratio (1) (mol/mol) (2) A 0
0.2000 100 -- -- AI 0 0.1445 72 -- -- AJ 0.5 0.0901 45 0.089 0.62
0.61 AK 5.0 0.0414 21 0.89 6.2 0.12 (1) Moles of Fe per mole of
acid sites in the Fe-free catalyst AI. (2) Moles of Bronsted acid
sites deactivated per moles of Fe added as determined by
IPA-TPD.
TABLE-US-00003 TABLE 1 Experimental testing results of catalysts in
fluidized bed reactor tests. Xylene Run Biomass Yield, % Carbon
Selectivity, % No Catalyst Feed T, C. Cycle Aromatics Olefins CO
CH.sub.4 CO.sub.2 pX mX oX 1 A Hardwood 575 1 26.47 7.5 18.19 2.50
7.39 49.0 40.8 10.2 2 B Hardwood 575 1 22.32 9.2 12.91 2.67 7.24
52.3 36.7 11.1 3 C Hardwood 525 1 20.49 8.1 15.85 2.36 7.88 54.2
35.8 10 4 D Newsprint 525 1 21.98 6.4 11.41 1.57 4.25 71.5 22.6 5.9
5 D Newsprint 525 2 17.90 7.5 13.24 1.82 4.97 75.5 19.5 5.0 6 E
Hardwood 550 1 17.93 7.82 11.01 2.43 5.64 84.1 13.0 2.9 7 E
Hardwood 550 2 16.83 7.73 10.88 2.41 5.57 84.9 12.1 3.0 8 E
Hardwood 525 3 16.07 9.39 13.22 2.61 6.69 85.2 12.1 2.7 9 E
Hardwood 525 4 15.19 8.12 11.59 2.18 5.69 86.7 10.8 2.5 10 E
Newsprint 600 5 17.98 11.33 18.23 4.71 8.21 85.9 11.3 2.8 11 E
Hardwood 568 6 18.43 12.63 18.74 4.72 8.38 86.0 10.8 3.2 12 F
Newsprint 525 1 24.69 8.33 14.27 2.07 5.57 79.4 16.7 3.9 13 F
Newsprint 525 2 24.37 8.41 14.48 2.11 5.65 83.8 13.1 3.2 14 G
Hardwood 525 1 18.82 10.09 13.22 2.70 7.03 86.5 11.2 2.4 15 G
Newsprint 525 2 19.79 9.60 13.31 2.12 6.10 86.2 11.2 2.6 16 G
Hardwood 525 3 16.67 10.10 13.79 2.77 6.92 87.9 9.8 2.4 17 G
Hardwood 525 4 19.20 11.21 14.46 2.88 7.23 89.3 8.6 2.1 18 G
Hardwood 525 5 19.25 11.35 15.29 3.00 7.58 89.1 8.8 2.1 19 G
Hardwood 525 12 17.45 9.68 12.17 2.49 6.15 89.3 -- -- 20 G Hardwood
525 13 17.01 11.61 14.90 2.63 7.92 88.8 -- -- 21 G Hardwood 525 14
20.79 10.98 13.91 2.81 7.06 89.6 -- -- 22 G Hardwood 525 15 19.17
11.25 14.26 2.87 7.24 89.3 -- -- 23 H Hardwood 525 1 22.94 10.28
14.05 2.17 6.62 84.0 11.0 5.1 24 H Hardwood 525 2 24.43 10.43 14.27
2.21 6.84 84.4 10.6 5.0 25 I Hardwood 525 1 22.02 10.41 13.37 2.40
7.36 88.1 9.3 2.5 26 I Hardwood 525 2 22.07 10.72 13.49 2.35 7.03
87.8 9.7 2.4 27 J Hardwood 571 1 16.08 8.96 18.93 3.68 4.90 46.6
38.8 14.5 28 J Hardwood 572 2 16.16 7.91 16.12 3.37 4.18 45.0 40.2
14.8 29 K Hardwood 523 1 14.15 9.91 13.47 2.43 5.50 90.7 7.0 2.6 30
K Hardwood 522 2 13.48 19.43 14.29 5.90 5.69 91.3 7.0 2.2 31 L
Hardwood 526 1 19.11 10.48 15.97 2.69 5.30 85.3 11.6 3.1 32 L
Hardwood 527 2 21.95 11.41 17.22 2.42 5.85 76.4 18.7 4.9 33 L
Hardwood 553 3 21.03 12.17 19.61 2.94 6.20 74.2 20.4 5.4 34 M
Hardwood 522 1 14.03 9.21 16.72 2.97 5.27 87.5 10.0 2.6 35 M
Hardwood 525 2 13.23 9.24 16.79 2.98 5.29 88.3 9.2 2.2
TABLE-US-00004 TABLE 2 IPA-TPD experimental results for catalysts.
Bronsted Acidity - Wt % Moles Si loading acidity % of Bronsted acid
sites Coating SiO.sub.2 [mmol/mg] [mmol/mg] Parent deactivated per
Si Example Catalyst Zeolite Promoter Compound Added (.times.1000)
(.times.1000) Zeolite added [mol/mol] 36 A A none na -- -- 0.2000
100 -- 37 N A Si TEOS 3.96 0.661 0.1749 87 0.038 38 O A Si TEOS
5.89 0.982 0.1623 81 0.038 39 P A Si TEOS 7.80 1.300 0.1448 72
0.042 40 Q A Si TEOS 7.80 1.300 0.1424 71 0.044 41 R A Si TEOS
11.46 1.911 0.1419 71 0.030 42 S A Si Polysiloxane 3.96 0.661
0.1382 69 0.094 43 T A Si Polysiloxane 5.89 0.982 0.1364 68 0.065
44 G A Si, Fe TEOS 5.7 1.000 0.1177 58.8 na 45 G* (1) A Si, Fe TEOS
5.7 1.000 0.0118 5.9 na 46 U A Si, Fe TEOS 5.7 1.000 0.0349 17.5 na
47 V C none na -- -- 0.0854 100 -- 48 W C Si TEOS 3.96 0.661 0.0628
73 0.034 49 X C Si TEOS 5.89 0.982 0.0523 61 0.034 50 Y C Si TEOS
7.80 1.300 0.0504 59 0.027 51 Z C Si TEOS 11.46 1.911 0.0530 62
0.017 (1). Sample G* is a sample from Experiment 22 that has been
subjected to 15 cycles of biomass conversion and regeneration in
air.
TABLE-US-00005 TABLE 3 Results of 2-methyl-furan (2MF) conversion
experiments. 2MF Yield, Carbon % Coating SiO.sub.2 conversion
Aromatics + Selectivity % Example Catalyst Sequence Source (%)
Aromatics Olefins Olefins pXylene pX mX oX 52 AA None na 86.2 11.7
12.6 24.34 0.97 75.3 21.7 3.1 53 AB Si only TEOS 90.4 12.3 13.4
17.63 1.27 93.3 5.4 1.6 54 AC La on Si TEOS 98.8 16.6 18.1 34.71
1.26 92.1 6.6 1.3 55 AD La only TEOS 96.1 16.1 16.1 32.18 1.23 80.2
17.2 2.6 56 AE Si only MEOS 82.0 9.4 10.8 20.15 0.98 95.5 3.5 0.9
57 AF Si only MEOS 80.8 9.6 11.6 21.25 1.02 96.0 3.0 1.1 58 AG Si
on La MEOS 98.1 14.7 15.0 29.74 1.16 86.2 12.2 1.6 59 AH La on Si
MEOS 96.4 14.8 15.8 30.61 1.37 95.3 4.2 0.5
TABLE-US-00006 TABLE 4 Results of Furfural Conversion Experiments.
Furfural Yield, % Carbon Wt % conversion CO/CO.sub.2 Aromatics +
Selectivity, % Example Catalyst Fe (%) ratio Aromatics Olefins
Olefins pXylene pX mX oX 60 AI 0.0 100 9.77 24.85 7.58 32.43 1.36
74.9 21.2 3.9 61 AJ 0.5 100 7.00 23.00 8.95 31.95 1.70 83.2 14.1
2.7 62 AK 5.0 100 5.67 24.44 8.78 33.22 2.04 84.6 13.1 2.3
* * * * *