Note: Descriptions are shown in the official language in which they were submitted.
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METHOD FOR CONVERTING A HYDROCARBONACEOUS MATERIAL TO
A FLUID HYDROCARBON PRODUCT COMPRISING
P-XYLENE
This application claims priority under 35 U.S.C. 119(e) to U.S. Provisional
Application Serial No. 61/655,605 filed June 5, 2012. The disclosure in this
provisional
application is incorporated herein by reference.
STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH OR
DEVELOPMENT.
The U.S. Government has a paid-up license in this invention and the right in
limited circumstances to require the patent owner to license others on
reasonable terms
as provided for by the terms of Grant No. CBET-0747996 awarded by the National
Science Foundation.
FIELD OF INVENTION
This invention relates to a method for converting a hydrocarbonaceous material
to a fluid hydrocarbon product comprising p-xylene, and more specifically, to
a method
for converting a hydrocarbonaceous material to a fluid hydrocarbon product
comprising
p-xylene via catalytic pyrolysis.
BACKGROUND
p-Xylene is used as a starting material for plasticizers and polyester fibers.
The
oxidation of p-xylene is used to commercially synthesize terephthalic acid.
Further
esterification of the acid with methanol forms dimethyl terephthalate. Both
monomers
may be used in the production of polyethylene terephthalate (PET) plastic
bottles and
polyester clothing.
SUMMARY
Catalytic pyrolysis, including 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
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method, the fluid hydrocarbon product comprises p-xylene, and may further
comprise
additional aromatics, olefins, and the like.
With this invention, 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, CO and CO2. The modified zeolite catalyst may comprise
pores
with pore mouth openings that have been reduced in size. The pyrolysis
products may
enter the pores in the modified zeolite catalyst where they may undergo
reaction. 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 include: 1) all the desired chemistry may
occur in a
single-step process, 2) the process may use a relatively inexpensive zeolite
catalyst, and
3) p-xylene may be produced at a relatively high level of production.
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. The problem therefore
is to
provide a method that allows for the production of p-xylene or xylenes with a
relatively
high selectivity to p-xylene. This invention provides a solution to this
problem. The
present invention provides for the use of a zeolite catalyst that has been
modified to
improve selectivity towards p-xylene. The zeolite catalyst may comprise
catalytic sites
on the external surface of the catalyst and pores with pore mouth openings.
The pores
may contain internal catalytic sites, some of which may be positioned near the
pore
mouth openings, and some of which may be internal catalytic sites positioned
away from
the pore mouth openings. The catalyst may be modified by treating it with an
effective
amount of a silicone compound to reduce the size of the pore mouth openings
and to
render at least some of the catalytic sites on the external surfaces of the
catalyst
inaccessible to the reactants. The treatment process may also be used to
render at least
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some of the catalytic sites in the pores near the pore mouth openings
inaccessible to the
reactants.
The zeolite catalyst may be treated by applying a silicone compound to the
surface of the catalyst. Application of the silicone compound may 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 as well as 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. Catalytic sites on the external
surface of the
catalyst as well as catalytic sites in the pores near the pore mouth openings
are believed
to be unselective in the production of xylenes, while the catalytic sites in
the pores away
from the pore mouth openings appear to provide for an increase in para
selectivity.
Inhibition or extinction of the activity of the catalytic sites on the
external surface of the
catalyst and in the pores of the catalyst near the pore mouth openings may
increase the
proportion of the catalytic reaction that occurs within the zeolite pores away
from the
pore mouth openings, hence an increase in selectivity to p-xylene.
The pores with pore mouth openings that have been reduced in size 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. This is illustrated in
Fig. 2. In
Fig. 2, a zeolite catalyst is shown that has a pore mouth opening that has
been reduced in
size to allow p-xylene, but not m- or o-xylene, to diffuse out of the zeolite
pore.
This invention relates to a method for producing a fluid hydrocarbon product
comprising p-xylene from a hydrocarbonaceous material, comprising: feeding the
hydrocarbonaceous material to a reactor; pyrolyzing within the reactor at
least a portion
of the hydrocarbonaceous material under reaction conditions sufficient to
produce a
pyrolysis product; and catalytically reacting at least a portion of the
pyrolysis product
under reaction conditions in the presence of a zeolite catalyst to produce the
fluid
hydrocarbon product; the zeolite catalyst comprising pores with pore mouth
openings
and catalytic sites on the external surface of the catalyst, and an effective
amount of a
treatment layer derived from a silicone compound to reduce the size of the
pore mouth
openings and to render at least some of the catalytic sites on the external
surface of the
catalyst inaccessible to the pyrolysis product. Catalytic sites may be
positioned in the
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pores near the pore mouth openings and the treatment layer may render at least
some of
these catalytic sites inaccessible to the pyrolysis product.
With the present invention, the selectivity to p-xylene in the xylenes may be
up to
100% when the only active catalytic sites are within the catalyst pores, and
the only
xylene to diffuse from the catalyst pores is p-xylene. In an embodiment, 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.
With the present invention, 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 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%.
The zeolite catalyst may be treated with the silicone 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 catalytic sites may also be
referred to as
acid sites. The covered or obscured catalytic sites may be referred to as
deactivated
catalytic sites. The silicone compound may have a molecular size that is
incapable of
entering the pores of the catalyst. During the catalyst treatment process, the
silicone
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
silicone compound may be at least about 15%, or at least about 25%, or at
least about
35%, or at least about 45%, or at least about 55%, or at least about 65%, or
at least about
75%, or at least about 85%, or at least about 90%, or at least about 95%, or
at least about
98%, or at least about 99%, of the available catalytic sites. The fraction of
catalytic sites
in the pores of the catalyst near the pore mouth openings that may be
deactivated by
treatment with the silicone compound may be at least about 20%, or at least
about 25%,
or at least about 30%, or at least about 33%, or at least about 35%, or at
least about 40%,
or at least about 45%, or at least about 49%, 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%,
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or at least about 80%, or at least about 85%, or at least about 90%, or at
least about 95%,
or at least about 99% of the available catalytic sites. The average diameter
of the pore
mouth openings subsequent to treatment with the silicone compound may be in
the range
from about 5 to about 10 angstroms, or from about 5.2 to about 7.4 angstroms.
The zeolite catalyst may comprise silica and alumina. The silica to alumina
molar ratio may be in the range from about 10:1 to about 50:1, or in the range
from about
10:1 to about 40:1, or in the range from about 10:1 to about 20:1, or about
15:1. The
zeolite catalyst may further comprise nickel, platinum, vanadium, palladium,
manganese,
cobalt, zinc, copper, chromium, gallium, an oxide of one or more thereof, or a
mixture of
two or more thereof
The silicone compound may comprise a compound containing at least one group
represented by the formula
I
¨0¨Si-
I
The silicone compound may be represented by the formula:
iRi 1
Si-0 I¨.
I
R2
wherein R1 and R2 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. R1 and/or R2 may
comprise
methyl, ethyl or phenyl. n may be a number in the range from about 3 to about
1000.
The silicone 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 silicone compound may comprise dimethylsilicone, diethylsilicone,
phenylmethylsilicone, methylhydrogensilicone, ethylhydrogen silicone,
phenylhydrogen
silicone, methylethyl silicone, phenylethyl silicone, diphenyl silicone,
methyltrifluoropropyl silicone, ethyltrifluoropropyl silicone, polydimethyl
silicone,
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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 silicone compound may comprise a tetraorthosilicate. The silicone
compound may comprise tetramethylorthosilicate, tetraethylorthosilicate, or a
mixture
thereof.
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. Fluidized
bed reactors
may be particularly advantageous.
The hydrocarbonaceous material may comprise a solid hydrocarbonaceous
material, a semi-solid hydrocarbonaceous material, a liquid hydrocarbonaceous
material,
or a mixture of two or more thereof 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, hemi-cellulose,
lignin, sugar cane
bagasse, glucose, wood, corn stover, or a mixture of two or more thereof. The
hydrocarbonaceous material may comprise furan and/or 2-methylfuran. 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 as the hydrocarbonaceous feed material.
The reactor may be operated at a temperature in the range from about 400 C to
about 600 C, or from about 425 C to about 500 C, or from about 440 C to
about
460 C.
The hydrocarbonaceous material may be fed to the reactor 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
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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 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 15 wt%, or
less than
about 10 wt%, of the pyrolysis product being coke.
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, methylethylbenzene,
trimethylbenzene, o-xylene, m-xylene, indanes naphthalene, methylnaphthelene,
dimethylnaphthalene, ethylnaphthalene, hydrindene, methylhydrindene,
dimethylhydrindene, or a mixture of two or more thereof.
The carbon yield of aromatics in the fluid hydrocarbon product may be at least
about 10%, or at least about 15%, or at least about 22%. The carbon yield of
olefins in
the fluid hydrocarbon product may be at least about 3%, or at least about 6%,
or at least
about 9%, or at least about 12 %. The mass yield of p-xylene may be at least
about 1.5
wt%, or at least about 2 wt%, or at least about 2.5 wt%, or at least about 3
wt%.
The catalytically reacting step may comprise a dehydration, decarbonylation,
decarboxylation, isomerization, oligomerization and/or dehydrogenation
reaction.
The pyrolyzing step and the catalytically reacting steps may be carried out in
a single vessel. Alternatively, the pyrolyzing step and the catalytically
reacting
steps may be carried out in separate vessels.
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
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where the present specification and a document incorporated by reference
include
conflicting and/or inconsistent disclosure, the present specification shall
control.
BRIEF DESCRIPTION OF THE DRAWINGS
Non-limiting embodiments of this invention will be described by way of example
with reference to the accompanying figures, which are schematic and are not
intended to
be drawn to scale. In the figures, each identical or nearly identical
component illustrated
is typically represented by a single numeral. For purposes of clarity, not
every
component is labeled in every figure, nor is every component of each
embodiment of the
invention shown where illustration is not necessary to allow those of ordinary
skill in the
art to understand the invention. In the figures:
FIGS. lA and 1B are schematic illustrations of a CFP process for converting a
solid hydrocarbonaceous material to a fluid hydrocarbon product.
FIG. 2 is a schematic illustration showing a pore with a pore mouth opening
that
has been reduced in size and the diffusion of p-xylene out of the pore.
DETAILED DESCRIPTION
All ranges and ratio limits disclosed in the specification and claims may be
combined in any manner. It is to be understood that unless specifically stated
otherwise,
references to "a," "an," and/or "the" may include one or more than one, and
that
reference to an item in the singular may also include the item in the plural.
The phrase "and/or" should be understood to mean "either or both" of the
elements so conjoined, i.e., elements that are conjunctively present in some
cases and
disjunctively present in other cases. Other elements may optionally be present
other than
the elements specifically identified by the "and/or" clause, whether related
or unrelated
to those elements specifically identified unless clearly indicated to the
contrary. Thus, as
a non-limiting example, a reference to "A and/or B," when used in conjunction
with
open-ended language such as "comprising" can refer, in one embodiment, to A
without B
(optionally including elements other than B); in another embodiment, to B
without A
(optionally including elements other than A); in yet another embodiment, to
both A and
B (optionally including other elements); etc.
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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. Only
terms clearly indicated to the contrary, such as "only one of" or "exactly one
of," or may
refer to the inclusion of exactly one element of a number or list of elements.
In general,
the term "or" as used herein shall only be interpreted as indicating exclusive
alternatives
(i.e. "one or the other but not both") when preceded by terms of exclusivity,
such as
"either," "one of," "only one of," or "exactly one of."
The phrase "at least one," in reference to a list of one or more elements,
should be
understood to mean at least one element selected from any one or more of the
elements
in the list of elements, but not necessarily including at least one of each
and every
element specifically listed within the list of elements and not excluding any
combinations of elements in the list of elements. This definition also allows
that
elements may optionally be present other than the elements specifically
identified within
the list of elements to which the phrase "at least one" refers, whether
related or unrelated
to those elements specifically identified. Thus, as a non-limiting example,
"at least one
of A and B" (or, equivalently, "at least one of A or B," or, equivalently "at
least one of A
and/or B") can refer, in one embodiment, to at least one, optionally including
more than
one, A, with no B present (and optionally including elements other than B); in
another
embodiment, to at least one, optionally including more than one, B, with no A
present
(and optionally including elements other than A); in yet another embodiment,
to at least
one, optionally including more than one, A, and at least one, optionally
including more
than one, B (and optionally including other elements); etc.
The transitional words or phrases, such as "comprising," "including,"
"carrying,"
"having," "containing," "involving," "holding," and the like, are to be
understood to be
open-ended, i.e., to mean including but not limited to.
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.
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The term "catalytic pyrolysis" refers to pyrolysis performed in the presence
of a
catalyst.
The terms "aromatics" or "aromatic compound" refer to a hydrocarbon
compound or compounds comprising one or more aromatic groups such as, for
example,
single aromatic ring systems (e.g., benzyl, phenyl, etc.) and/or 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, styrene,
cumene, 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 term "petrochemicals" is used herein to refer to chemicals, chemical
precursors, chemical intermediates, and the like, traditionally derived from
petroleum
sources. Petrochemicals include paraffins, olefins, aromatic compounds, and
the like.
For purposes of this application, when these materials are derived from
biomass, as well
as other non-petroleum sources (e.g., recycled plastics, municipal solid
waste, sugar cane
bagasse, wood, etc.), the term petrochemicals may be employed despite the fact
that the
chemicals, chemical precursors, chemical intermediates, and the like, may not
be derived
directly from petroleum.
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
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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 includes plant or animal matter used for production
of fibers,
chemicals, heat, and the like. Biomass may also include biodegradable waste or
byproducts that can be burnt as fuel or converted to chemicals. These may
include
municipal waste, green waste (the biodegradable waste comprised of garden or
park
waste such as grass or flower cuttings, hedge trimmings, and the like),
byproducts of
farming including animal manures, food processing wastes, sewage sludge, black
liquor
from wood pulp or algae, and the like. Biomass may be derived from plants,
including
miscanthus, spurge, sunflower, switchgrass, hemp, corn (maize), poplar,
willow,
sugarcane, and oil palm (palm oil), and the like. Biomass may be derived from
roots,
stems, leaves, seed husks, fruits, and the like. The particular plant or other
biomass
source used may not be important to the fluid hydrocarbon product produced in
accordance with the inventive method, although the processing of the biomass
may vary
according to the needs of the reactor and the form of the biomass.
The hydrocarbonaceous feed material for the inventive method may comprise a
solid hydrocarbonaceous material, a semi-solid hydrocarbonaceous material, a
liquid
hydrocarbonaceous material, 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 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,
hemi-cellulose,
lignin, sugar cane bagasse, glucose, wood, corn stover, or a mixture of two or
more
thereof The hydrocarbonaceous material may comprise furan and/or 2-
methylfuran. The
hydrocarbonaceous material may comprise pinewood. The hydrocarbonaceous
material
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may comprise pyrolysis oil derived from biomass, a carbohydrate derived from
biomass,
an alcohol derived from biomass, a biomass extract, a pretreated biomass, a
digested
biomass product, or a mixture of two or more thereof. Mixtures of two or more
of any of
the foregoing may be used.
The inventive method may comprise feeding the hydrocarbonaceous material to a
reactor. At least a portion of the hydrocarbonaceous material may be pyrolyzed
in the
reactor under reaction conditions sufficient to produce one or more pyrolysis
products. At
least a portion of the pyrolysis products may be catalytically reacted under
sufficient
conditions to produce the fluid hydrocarbon product. The reactor may comprise
a
continuously stirred tank reactor, a batch reactor, a semi-batch reactor, a
fixed bed
reactor, or a fluidized bed reactor. Advantageously, the reactor may comprise
a fluidized
bed reactor. The catalytic reaction step may be achieved by co-feeding the
catalyst with
the hydrocarbonaceous material. The catalyst may be fed separately. Part of
the catalyst
may be fed with the hydrocarbonaceous feed material and part of the catalyst
may be fed
separately.
The inventive method may be used for the production of fluid (e.g., a liquid,
a
supercritical fluid, and/or a gas) hydrocarbon products via a catalytic
pyrolysis process
(e.g., a CFP process). The fluid hydrocarbon product, or a portion thereof,
may
comprise a liquid at standard ambient temperature and pressure (SATP - i.e. 25
C and
100 kPa absolute pressure). The hycrocarbonaceous feed material may be
pyrolyzed at
intermediate temperatures (for example, in the range from about 400 C and
about
600 C), compared to temperatures typically used in the prior art. The
pyrolysis step
may be conducted for an effective amount of time to produce discrete,
identifiable fluid
hydrocarbon products. The inventive method may involve heating the
hydrocarbonaceous material (and optionally the catalyst) to the reaction
temperature at
relatively high heating rates (e.g., greater than about 50 C per second).
The inventive method may involve the use of specialized catalysts. These
catalysts are zeolite catalysts which contain silica and alumina. The catalyst
may be
characterized by pores with pore mouth openings wherein the pore mouth
openings
have been reduced in size, and catalytic sites on the external surface of the
catalyst that
have been covered or obscured. The catalyst may have catalytic sites in the
pores near
the pore mouth openings that have been covered or obscured. The catalyst may
be
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treated with a silicone compound to reduce the size of the pore mouth openings
and
cover or obscure catalytic sites on the external surface of the catalyst and
in the pores of
the catalyst near the pore mouth openings. The catalyst may comprise
relatively small
particles, which may be agglomerated. The composition fed to the reactor may
have a
relatively high catalyst to hydrocarbonaceous material mass ratio (e.g., from
about
0.33:1 to about 20:1, or from about 0.5:1 to about 20:1, or from about 2:1 to
about
10:1).
The expression that catalytic sites are positioned "in the pores near the pore
mouth openings" refers to catalytic sites within the pores that are rendered
inaccessible
to the pyrolysis product as a result of treatment with the silicone compound.
In an
embodiment, these catalytic sites may be positioned within the pores at a
depth of no
more than about 10 angstroms from the pore mouth openings, or no more than
about 7
angstroms, or no more than about 5 angstroms, or no more than about 2
angstroms,
from the pore mouth openings. These catalytic sites may be inhibited or
inoperative as
a result of treatment with the silicone compound. That is, these catalysts may
be
deactivated as a result of the treatment with the silicone compound.
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 bath reactor, a semi-batch reactor, a fixed bed reactor or a
fluidized bed
reactor. Multi-stage apparatuses may also be used for the production of fluid
hydrocarbon products in accordance with the invention.
The 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.
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The average particle size of the solid hydrocarbonaceous material may be less
than about
mm, less than about 2 mm, less than about 1 mm, less than about 500 microns,
less than
about 60 mesh (250 microns), less than about 100 mesh (149 microns), less than
about 140
mesh (105 microns), less than about 170 mesh (88 microns), less than about 200
mesh (74
5 microns), less than about 270 mesh (53 microns), or less than about 400
mesh (37
microns), or smaller.
It may be desirable to employ a feed material with an average particle size
above a
minimum amount in order to reduce the pressure required to pass the solid
hydrocarbonaceous feed material through the reactor. For example, it may be
desirable to
use a solid hydrocarbonaceous feed material with an average particle size of
at least about
400 mesh (37 microns), at least about 270 mesh (53 microns), at least about
200 mesh (74
microns), at least about 170 mesh (88 microns), at least about 140 mesh (105
microns), at
least about 100 mesh (149 microns), at least about 60 mesh (250 microns), at
least about
500 microns, a least about 1 mm, at least about 2 mm, at least about 5 mm, or
higher.
The hydrocarbonaceous material may comprise biomass. The
hydrocarbonaceous material may comprise plastic waste, recycled plastics,
agricultural
and/or municipal solid waste, food waste, animal waste, carbohydrates,
lignocellulosic
materials (e.g., wood chips or shavings), or a mixture of two or more thereof
The
hydrocarbonaceous material may comprise xylitol, glucose, cellobiose,
cellulose, hemi-
cellulose, lignin, or a mixture of two or more thereof. The hydrocarbonaceous
material
may comprise sugar cane bagasse, glucose, wood, corn stover, or a mixture of
two or
more thereof The hydrocarbonaceous material may comprise wood.
Biomass pyrolysis liquid or bio-oil may be formed during the pyrolyzing step
of
the inventive method. Biomass pyrolysis liquid may be dark brown and may
approximate to biomass in elemental composition. It may be composed of a very
complex mixture of oxygenated hydrocarbons with an appreciable proportion of
water
from both the original moisture and reaction product. Compositionally, biomass
pyrolysis oil may vary with the type of biomass, but is known to contain
oxygenated low
molecular weight alcohols (e.g., furfuryl alcohol), aldehydes (aromatic
aldehydes),
ketones (furanone), phenols (methoxy phenols) and water. Solid char may also
be
present, suspended in the oil. The liquid may be formed by rapidly quenching
the
intermediate products of flash pyrolysis of hemicellulose, cellulose and
lignin in the
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biomass. Chemically, the oil may contain several hundred different chemicals
in widely
varying proportions, ranging from formaldehyde and acetic acid to complex high
molecular weight phenols, anhydrosugars and other oligosaccharides. It may
have a
distinctive odor from low molecular weight aldehydes and acids, and may be
acidic with
a pH of about 1.5 to about 3.8, and can be an irritant.
The residence time of the catalyst in the reactor may be defined as the volume
of
the reactor filled with catalyst divided by the volumetric flow rate of the
catalyst through
the reactor. For example if a 3 liter reactor contains 2 liters of catalyst
and a flow of 0.4
liters per minute of catalyst is fed through the reactor, i.e., both fed and
removed, the
catalyst residence time will be 2/0.4 minutes, or 5 minutes.
The residence time of the catalyst in the reactor may be at least about 1
minute, at
least about 2 minutes, at least about 5 minutes, at least about 7 minutes, at
least about
10 minutes, at least about 15 minutes, at least about 20 minutes, at least
about 25
minutes, at least about 30 minutes, at least about 60 minutes, or at least
about
120 minutes. In some cases, the residence time of the catalyst in the reactor
may be less
than about 120 minutes, or from about 1 minute and about 120 minutes, or from
about 2
minutes to about 120 minutes, or from about 5 minutes to about 120 minutes, or
from
about 7 minutes to about 120 minutes, or from about 10 minutes to about 120
minutes, or
from about 12 minutes to about 120 minutes, or from about 15 minutes to about
120
minutes, or from about 20 minutes to about 120 minutes, or from about 30
minutes to
about 120 minutes, or from about 60 minutes to about 120 minutes. In some
cases, the
use of relatively long residence times may allow for additional chemical
reactions to
form desirable products. Long catalyst residence times may be achieved by, for
example, increasing the volume of the reactor and/or reducing the volumetric
flow rate of
the catalyst. The residence time of the catalyst may be relatively short,
e.g., less than
about 120 minutes, or less than about 60 minutes.
Contact time may be defined as the residence time of a material in a reactor
or
other device, when measured or calculated under standard conditions of
temperature and
pressure (i.e., 0 C and 100 kPa absolute pressure). For example, a 2 liter
reactor to
which is fed 3 standard liters per minute of gas has a contact time of 2/3
minute, or 40
seconds for that gas. 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
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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.
The term "conversion of a reactant" may refer to the reactant mole or mass
change between a material flowing into a reactor and a material flowing out of
the
reactor divided by the moles or mass of reactant in the material flowing into
the reactor.
For example, if 100 g of ethylene are fed to a reactor and 30 g of ethylene
are flowing
out of the reactor, the conversion is [ (100 ¨ 30) / 100] = 70% conversion of
ethylene.
The term "fluid" may refer to a gas, a liquid, a mixture of a gas and a
liquid, or a
gas or a liquid containing dispersed solids, liquid droplets and/or gaseous
bubbles. The
terms "gas" and "vapor" have the same meaning and are sometimes used
interchangeably. In some embodiments, it may be advantageous to control the
residence
time of the fluidization fluid in the reactor. The fluidization residence time
of the
fluidization fluid is defined as the volume of the reactor divided by the
volumetric flow
rate of the fluidization fluid under process conditions of temperature and
pressure.
The term "fluidized bed reactor" may be used to refer to reactors comprising a
vessel that contains a granular solid material (e.g., silica particles,
catalyst particles, etc.),
in which a fluid (e.g., a gas or a liquid) is passed through the granular
solid material at
velocities sufficiently high as to suspend the solid material and cause it to
behave as
though it were a fluid. The term "circulating fluidized bed reactor" may be
used to refer
to fluidized bed reactors in which the granular solid material is passed out
of the reactor,
circulated through a line in fluid communication with the reactor, and
recycled back into
the reactor.
Bubbling fluidized bed reactors and turbulent fluidized bed reactors may be
used.
In bubbling fluidized bed reactors, the fluid stream used to fluidize the
granular solid
material may be operated at a sufficiently low flow rate such that bubbles and
voids may
be observed within the volume of the fluidized bed during operation. In
turbulent
fluidized bed reactors, the flow rate of the fluidizing stream will be higher
than that
employed in a bubbling fluidized bed reactor. Examples of fluidized bed
reactors,
circulating fluidized bed reactors, bubbling and turbulent fluidized bed
reactors are
described in Kirk-Othmer Encyclopedia of Chemical Technology (online), Vol.
11,
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Hoboken, N.J.: Wiley Interscience, 2001, pages 791-825, these pages being
incorporated
herein by reference.
The terms "olefin" or "olefin compound" (a.k.a. "alkenes") may be used to
refer
to any unsaturated hydrocarbon containing one or more pairs of carbon atoms
linked by a
double bond. Olefins may include both cyclic and acyclic (aliphatic) olefins,
in which the
double bond is located between carbon atoms forming part of a cyclic (closed-
ring) or of
an open-chain grouping, respectively. In addition, olefins may include any
suitable
number of double bonds (e.g., monoolefins, diolefins, triolefins, etc.).
Examples of olefin
compounds may include ethene, propene, allene (propadiene), 1-butene, 2-
butene,
isobutene (2 methyl propene), butadiene, and isoprene, among others. Examples
of cyclic
olefins may include cyclopentene, cyclohexane, cycloheptene, among others.
Aromatic
compounds such as toluene are not considered olefins; however, olefins that
include
aromatic moieties are considered olefins, for example, benzyl acrylate or
styrene.
Pore size relates to the size of a molecule or atom that can penetrate into
the pores
of a material. As used herein, the term "pore size" for zeolites refers to the
Norman radii
adjusted pore size. 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, July 5-10, 1998;
(1999), 1, pp
409-414, which is incorporated herein by reference. As a specific exemplary
calculation,
the atomic radii for ZSM-5 pores are about 5.5-5.6 Angstroms, 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
Angstroms
larger than the atomic radii (about 6.2-6.3 Angstroms).
One of ordinary skill in the art will understand how to determine the pore
size
(e.g., minimum pore size, average of minimum pore sizes) in a catalyst. For
example, x-
ray diffraction (XRD) may be used to determine atomic coordinates. XRD
techniques for
the determination of pore size are described, for example, in Pecharsky, V.K.
et at,
"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) may include, for example, helium pycnometry or low
pressure
argon adsorption techniques. These and other techniques are described in
Magee, J.S. et
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at, "Fluid Catalytic Cracking: Science and Technology," Elsevier Publishing
Company,
July 1, 1993, pp. 185-195, which is incorporated herein by reference in its
entirety. 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 of which are incorporated herein by
reference.
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 may calculate, for example, the kinetic diameter of
a given
molecule. The type of catalyst may then be chosen such that the pores of the
catalyst
(e.g., Norman adjusted minimum radii) are sufficiently large to allow the
pyrolysis
product molecules to diffuse into and/or react with the catalyst. In some
embodiments,
the catalysts may be chosen such that their pore sizes are sufficiently small
to prevent
entry and/or reaction of pyrolysis products whose reaction would be
undesirable.
The catalyst may comprise any catalyst suitable for conducting the
catalytically
reacting step of the inventive method. The catalyst may be used to lower the
activation
energy (increase the rate) of the reaction conducted in the catalytically
reacting step
and/or improve the distribution of products or intermediates during the
reaction (for
example, a shape selective catalyst). Examples of reactions that can be
catalyzed include:
dehydration, dehydrogenation, isomerization, hydrogen transfer, aromatization,
decarbonylation, decarboxylation, aldol condensation, and combinations thereof
The
catalyst components may be acidic, neutral or basic.
The inventive method may comprise a CFP process. 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, 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,
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catalysts with average pore sizes of between about 5.5 Angstroms and about 6.5
Angstroms, or between about 5.9 Angstroms and about 6.3 Angstroms may be used.
In
some cases, catalysts with average pore sizes of between about 7 Angstroms and
about 8
Angstroms, or between about 7.2 Angstroms and about 7.8 Angstroms may be used.
The 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)A1P0-
31, SSZ-23, and the like. The catalyst may comprise silica and alumina, and
further
comprise one or more additional metals and/or a metal oxides. Suitable metals
and/or
oxides may include, for example, nickel, palladium, platinum, titanium,
vanadium,
chromium, manganese, iron, cobalt, zinc, copper, gallium, and/or any of their
oxides,
among others. In some cases promoter elements selected from the rare earth
elements,
i.e., elements 57-71, cerium, zirconium or their oxides, or combinations of
these may be
included to modify the activity, structure and/or stability of the catalyst.
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.
The catalyst may be treated or impregnated one or more times with 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 silicone compound, the
silicone compound
may be dispersed in a carrier, for example, an aqueous or organic liquid
carrier.
In each phase of the catalyst treatment process, the silicone compound may be
deposited on the external surface of the catalyst by any suitable method. For
example,
the silicone 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 silicone compound at a catalyst to silicone compound weight ratio in
the range
from about 1000:1 to about 1:10.
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The silicone compound may be provided in the form of a solution or an emulsion
under the conditions of contact with the catalyst. The deposited silicone
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. Patents 4,090,981; 5,243,117; 5,403,800, and
5,659,098, which are incorporated by reference herein.
The catalyst may be ex situ treated by single or multiple coatings with the
silicone compound, each coating followed by calcination. The coated catalyst
may
comprise silica and alumina. The silica to alumina molar ratio may be in the
range from
about 10:1 to about 50:1, or in the range from about 10:1 to about 40:1, or in
the range
from about 10:1 to about 20:1, or about 15:1. The coated catalyst may further
comprise
nickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper,
chromium,
gallium, an oxide of one or more thereof, or a mixture of two or more thereof
The term "silicone compound" is used herein to refer to any compound that
contains one or more Si-0 groups. The silicone compound may be a silicate
containing
one or more of Siatt 5i2076- or 5i601812- groups. These may include one or
more
tetraorthosilicates. The silicone compound may include one or more siloxanes
containing one or more silicon-oxygen backbones (-Si-O-Si-0-) 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 silicone compound may be
a
straight chain, branched chain or cyclical compound. The silicone compound may
be
monomeric, oligomeric or polymeric. The silicone compound may comprise a
compound containing at least one group represented by the formula
I
¨0--Si--
The silicone compound may be represented by the formula:
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Si-01¨
I
R2
:
wherein R1 and R2 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. R1 and/or R2 may
comprise
methyl, ethyl or phenyl. n may be a number in the range from about 3 to about
1000.
The silicone 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 silicone 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 silicone compound may comprise a tetraorthosilicate. The silicone
compound may comprise tetramethylorthosilicate, tetraethylorthosilicate, or a
mixture
thereof.
The kinetic diameter of the silicone compound may be larger than the pore
diameter of the catalyst in order to avoid entry of the silicone compound into
the pore
and any concomitant reduction in the internal activity of the catalyst.
The organic carrier for the silicone compound may comprise hydrocarbons such
as linear, branched, and cyclic alkanes having five or more carbons. The
carrier may
comprise a linear, branched or cyclic alkane having a boiling point greater
than about
70 C, and containing about 6 or more carbons. Optionally, mixtures of low
volatility
organic compounds, such as hydrocracker recycle oil, may also be employed as
carriers.
Low volatility hydrocarbon carriers for the silicone compound may comprise
decane,
dodecane, mixtures thereof, and the like.
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Following each deposition of the silicone compound, the catalyst may be
calcined
to decompose the molecular or polymeric species to a solid state species. The
catalyst
may be calcined at a rate of from about 0.2 C/minute to about 5 C/minute to
a
temperature greater than about 200 C, but below a temperature at which the
crystallinity
of the zeolite may be adversely affected. Generally, such temperature will be
below
about 600 C. The temperature of calcination may be in the range from about 350
C to
about 550 C. The catalyst may be maintained at the calcination temperature for
about 1
to about 24 hours, or about 2 to about 6 hours.
The catalyst may be treated with a tetraorthosilicate using a chemical liquid
deposition (CLD) process. The tetraorthosilicate may comprise
tetramethylorthosilicate,
tetraethylorthosilicate (TEOS), or a mixture thereof The CLD process may
comprise
dispersing the catalyst in a liquid medium (e.g., hexanes, alkanes, aromatics
or other
non-polar organic solvent) at a concentration in the range from about 0.1 to
about 20%
by weight, or from about 1 to about 10% by weight; adding the
tetraorthosilicate to
provide a mixture containing from about 0.01 to about 5% by weight, or from
about 0.1
to about 1% by weight of the tetraorthosilicate; refluxing the resulting
mixture at an
elevated temperature (e.g., in the range from about 50 to about 150 C, or
about 90 C)
with stirring; recovering the catalyst (e.g., via centrifuging) from the
liquid medium;
drying the catalyst; and then calcining the catalyst in air at a temperature
in the range
from about 100 to about 550 C, or from about 150 to about 325 C, for a time
period in
the range from about 1 to about 24 hours, or about 2 to about 12 hours. This
procedure
may be repeated any desired number of times (e.g., 1, 2, 3 additional times,
etc.), to
provide for the desired treatment layer derived from the tetraorthosilicate.
When using
TEOS as the tetraorthosilicate, this process may be referred to as a TEOS CLD
silylation
process.
While not wishing to be bound by theory, it is believed that the advantages of
treatment with silicone compounds are in part obtained by rendering active
catalytic sites
on the external surfaces of the catalyst substantially inaccessible to
reactants, while
increasing catalyst tortuosity by reducing the size of the pore mouth
openings. Active
catalytic sites existing on the external surface of the catalyst are believed
to isomerize the
para-isomer back to an equilibrium level with the other two isomers. Thus, by
reducing
the availability of these active catalytic sites, the relatively high
proportion of para-
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xylene may be maintained. It is believed that the silicone compounds of the
present
invention may block or otherwise render these external catalytic sites
unavailable to the
para-isomers by chemically modifying, covering, or obscuring the sites.
It may be beneficial to control the residence time of the reactants (e.g., the
solid
hydrocarbonaceous material and/or a non-solid reactant) and catalyst(s) in the
reactor
and/or under a defined set of reaction conditions (i.e. conditions under which
the
reactants may undergo pyrolysis or catalysis in a given reactor system).
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.
The term "reactant residence time" of a reactant in the reactor is defined as
the
amount of time the reactant spends in the reactor. Residence time may be based
on the
feed rate of reactant and is independent of rate of reaction. The reactant
residence time of
the reactants in a reactor may be calculated using different methods depending
upon the
type of reactor being used. For gaseous reactants, where flow rate into the
reactor is
known, this is typically a simple calculation. In the case of solid reactants
in which the
reactor comprises a packed bed reactor into which only reactants are
continuously fed
(i.e. no carrier or fluidizing flow is utilized), the reactant residence time
in the reactor
may be calculated by dividing the volume of the reactor by the volumetric flow
rate of
the hydrocarbonaceous material and fluid hydrocarbon product exiting the
reactor.
In cases where the reaction takes place in a reactor that is closed to the
flow of
mass during operation (e.g., a batch reactor), the batch residence time of the
reactants in
such may be reactor is defined as the amount of time elapsing between the time
at which
the temperature in the reactor containing the reactants reaches a level
sufficient to
commence a pyrolysis reaction (e.g., for CFP, typically about 300 C to about
1000 C
for many typical hydrocarbonaceous feedstock materials) and the time at which
the
reactor is quenched (e.g., cooled to a temperature below that sufficient to
support further
pyrolysis - e.g. typically about 300 C to about 1000 C for many
hydrocarbonaceous
feedstock materials).
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In some cases, e.g. for certain fluidized bed reactors, the reactor feed
stream(s)
may include feed stream(s) comprising auxiliary materials (i.e., matter other
than solid
hydrocarbonaceous materials and/or non-solid reactants). For example, in
certain cases
where fluidized beds are used as reactors, the feed stream may comprise
fluidization
fluid(s). In cases where circulating fluidized beds are used, catalyst and
fluidization fluid
may both be fed, recycled, or fed and recycled to the reactor. In such cases,
the reactant
residence time of the reactants in the reactor can be determined as the volume
of the
reactor divided by the volumetric flow rate of the reactants and reaction
product gases
exiting the reactor as with the packed bed situation described above; however,
since the
flow rate of the reactants and reaction product gases exiting the reactor may
not be
convenient to determine directly, the volumetric flow rate of the reactants
and reaction
product gases exiting the reactor may be estimated by subtracting the feed
volumetric
flow rate of the auxiliary materials (e.g., fluidization fluid, catalyst,
contaminants, etc.)
into the reactor from the total volumetric flow rate of the gas stream(s)
exiting the
reactor.
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.
For reactions involving conversion of a specific molecular reactant (ethene
for example),
selectivity is the percentage (on a mass basis unless specified otherwise) of
a selected
product divided by all the products produced. The selectivity for various
materials can
be determined using the following equations:
z-31 apYo cf7.,1ct-
(1) __________________________________ Overall selectivity=
eS,' c arL7t7n, n n tzx 100%
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n1,17 c,f carbcn: i.42n2n1.1-7-tiz- p-rod.t
(2) Aromatic selectivity ¨
moZE.s.or carbc,31 aromatk- prothz-i 17:Ex 100%
r
ZTS of crarhon bz. an- L-..t7f:.?.z,17 prod,.:t
(3) Olefin selectivity ¨
z-arbon 2:n af11.7e.f?ss rr.roduztox 100%
of ;.? ¨
(4) p-Xylene selectivity in xylenes = f x
100%
The term "yield" is used herein to refer to the amount of a product flowing
out of
a reactor divided by the amount of reactant flowing into the reactor, usually
expressed as
a percentage or fraction. Yields are often calculated on a mass basis, carbon
basis, or on
the basis of a particular feed component. Mass yield is the mass of a
particular product
divided by the weight of feed used to prepare that product. For example, if
500 grams of
biomass is fed to a reactor and 45 grams of p-xylene is produced, the mass
yield of p-
xylene would be 45/500 = 9% p-xylene . Carbon yield is the mass of carbon
found in a
particular product divided by the mass of carbon in the feed to the reactor.
For example,
if 500 grams of biomass that contains 40% carbon is reacted to produce 45 g of
p-xylene
that contains 90.6% carbon, the carbon yield is [(45 * 0.906)/(500 * 0.40)] =
20.4%.
Carbon yield from biomass is the mass of carbon found in a particular product
divided by
the mass of carbon fed to the reactor in a particular feed component. For
example, if 500
grams of biomass containing 40% carbon and 100 grams of CO2 are reacted to
produce
40 g of p-xylene (containing 90.6% carbon), the carbon yield on biomass is
[(40 *
0.906)/(500 * 0.40)] = 18.1%; note that the mass of CO2 does not enter into
the
calculation.
The embodiments described herein may also involve chemical process designs
used to perform catalytic pyrolysis. The processes may involve the use of one
or more
fluidized bed reactors (e.g., a circulating fluidized bed reactor, turbulent
fluidized bed
reactor, bubbling fluidized bed reactor, etc.). The process designs described
herein may
optionally involve specialized handling of the material fed to one or more
reactors. For
example, the feed material may be dried, cooled, and/or ground prior to
supplying the
material to a reactor. Other aspects of the invention may relate to product
compositions
produced using the process designs described herein.
Without being bound to a particular mode of action or order of steps of the
overall thermal/catalytic conversion process, catalytic pyrolysis is believed
to involve at
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least partial thermal pyrolysis of hydrocarbonaceous material (e.g., solid
biomass such as
cellulose) to produce one or more pyrolysis products (e.g., volatile organics,
gases, solid
coke, etc.) and catalytic reaction of at least a portion of the one or more
pyrolysis
products using a catalyst under reaction conditions sufficient to produce
fluid
hydrocarbon products. The catalytic reaction may involve volatile organics
entering into
a catalyst (e.g., a zeolite catalyst) where they are converted into, for
example, p-xylene as
well as other hydrocarbons such as aromatics and olefins, in addition to
carbon
monoxide, carbon dioxide, water, and coke. Inside or upon contact with the
catalyst, the
pyrolysis product may undergo a series of dehydration, decarbonylation,
decarboxylation, isomerization, oligomerization, and dehydrogenation reactions
that lead
to aromatics, olefins, CO, CO2 and water. 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%.
FIG. lA includes a schematic illustration of an exemplary chemical process
design used to perform catalytic pyrolysis, according to the inventive method.
The
process may comprise a CFP process. Referring to FIG. 1A, feed stream 10
includes a
solid hydrocarbonaceous material that can be fed to reactor 20. The solid
hydrocarbonaceous material may generally comprise at least carbon and
hydrogen. In
certain solid hydrocarbonaceous materials (e.g. wood), carbon may be the most
abundant
component by mass, while in others (e.g. glucose) oxygen may be more abundant
than
carbon. Certain solid hydrocarbonaceous materials may also comprise relatively
minor
proportions of other elements such as nitrogen and sulfur.
The feed streams to the reactor may be free of olefins, or may contain olefins
in an
insignificant amount (e.g., such that olefins make up less than about 1 wt%,
less than about
0.1 wt%, or less than about 0.01 wt% of the total weight of reactant fed to
the reactor). In
other embodiments, however, olefins may be present in one or more reactant
feed streams.
The solid hydrocarbonaceous material feed composition (e.g., in feed stream 10
of
FIG. 1A) 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
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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.
Moisture 12 may optionally be removed from the solid hydrocarbonaceous feed
composition prior to being fed to the reactor, e.g., by an optional dryer 14.
Removal of
moisture from the solid hydrocarbonaceous material feed stream may be
advantageous for
several reasons. For example, the moisture in the feed stream may require
additional
energy input in order to heat the solid hydrocarbonaceous material to a
temperature
sufficiently high to achieve pyrolysis. Variations in the moisture content of
the solid
hydrocarbonaceous feed may lead to difficulties in controlling the temperature
of the
reactor. In addition, removal of moisture from the solid hydrocarbonaceous
feed can
reduce or eliminate the need to process the water during later processing
steps.
The solid hydrocarbonaceous feed composition may be dried until the solid
hydrocarbonaceous feed composition comprises less than about 10%, less than
about 5%,
less than about 2%, or less than about 1% water by weight. Suitable equipment
capable of
removing water from the feed composition is known to those skilled in the art.
As an
example, the dryer may comprise an oven heated to a particular temperature
(e.g., at least
about 80 C, at least about 100 C, at least about 150 C, or higher) through
which the solid
hydrocarbonaceous feed composition may continuously, semi-continuously, or
periodically pass. The dryer may comprise a vacuum chamber into which the
solid
hydrocarbonaceous feed composition may be processed as a batch. The dryer may
combine elevated temperatures with vacuum operation. The dryer may be
integrally
connected to the reactor or may be provided as a separate unit from the
reactor.
The particle size of the solid hydrocarbonaceous feed composition may be
reduced
in an optional grinding system 16 prior to passing the solid hydrocarbonaceous
feed to the
reactor. The average diameter of the ground, solid hydrocarbonaceous feed
composition
exiting the grinding system may comprise no more than about 50%, no more than
about
25%, no more than about 10%, no more than about 5%, no more than about 2% of
the
average diameter of the feed composition fed to the grinding system. Large-
particle solid
hydrocarbonaceous feed material may be more easily transportable and less
messy than
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small-particle feed material. On the other hand, in some cases it may be
advantageous to
feed small particles of solid hydrocarbonaceous material to the reactor. The
use of a
grinding system allows for the transport of large-particle solid
hydrocarbonaceous feed
between the source and the process, while enabling the feed of small particles
to the
reactor.
Suitable equipment capable of grinding the solid hydrocarbonaceous feed
composition is known to those skilled in the art. For example, the grinding
system may
comprise an industrial mill (e.g., hammer mill, ball mill, etc.), a unit with
blades (e.g.,
chipper, shredder, etc.), or any other suitable type of grinding system. The
grinding system
may comprise a cooling system (e.g., an active cooling systems such as a
pumped fluid
heat exchanger, a passive cooling system such as one including fins, etc.),
which may be
used to maintain the solid hydrocarbonaceous feed composition at relatively
low
temperatures (e.g., ambient temperature) prior to introducing the solid
hydrocarbonaceous
feed composition to the reactor. The grinding system may be integrally
connected to the
reactor or may be provided as a separate unit from the reactor. While the
grinding step is
shown following the drying step in FIG. 1A, the order of these operations may
be reversed
in some embodiments. In still other embodiments, the drying and grinding steps
may be
achieved using an integrated unit.
Grinding and cooling of the solid hydrocarbonaceous material may be achieved
using separate units. Cooling of the solid hydrocarbonaceous material may be
desirable,
for example, to reduce or prevent unwanted decomposition of the solid
hydrocarbonaceous
feed material prior to passing it to the reactor. The solid hydrocarbonaceous
material may
be passed to a grinding system to produce a ground solid hydrocarbonaceous
material.
The ground solid hydrocarbonaceous material may then be passed from the
grinding
system to a cooling system and cooled. The solid hydrocarbonaceous material
may be
cooled to a temperature lower than about 300 C, lower than about 200 C,
lower than
about 100 C, lower than about 75 C, lower than about 50 C, lower than about
35 C, or
lower than about 20 C prior to introducing the solid hydrocarbonaceous
material into the
reactor. The cooling system may include an active cooling unit (e.g., a heat
exchanger)
capable of lowering the temperature of the solid hydrocarbonaceous material.
The two or
more of the drier, grinding system, and cooling system may be combined in a
single unit.
The cooling system may be directly integrated with one or more reactors.
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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. 1A, 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. In some cases, the reactor comprises a fluidized bed reactor, e.g., a
circulating
fluidized bed reactor. Fluidized bed reactors may, in some cases, provide
improved
mixing of the catalyst, solid hydrocarbonaceous material 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.
The reactor(s) may have any suitable size for performing the processes
described
herein. For example, the reactor may have a volume between about 0.1-1 L, 1-50
L,
50-100 L, 100-250 L, 250-500 L, 500-1000 L, 1000-5000 L, 5000-10,000 L, or
10,000-50,000 L. In some instances, the reactor may have a volume greater than
about 1
L, or in other instances, greater than about 10 L, 50 L, 100 L, 250 L, 500 L,
1,000 L, or
10,000 L. Reactor volumes greater than about 50,000 L may also be possible.
The
reactor may be cylindrical, spherical, or any other suitable shape.
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
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pore mouth opening size) may be controlled to achieve beneficial results, as
described
below.
The reactor(s) may be operated at any suitable temperature. In some instances,
it
may be desirable to operate the reactor(s) at intermediate temperatures,
compared to
temperatures typically used in many previous catalytic pyrolysis systems. For
example,
the reactor may be operated at temperatures of between about 400 C and about
600 C,
between about 425 C and about 500 C, or between about 440 C and about 460
C.
Operating the reactor(s) at these intermediate temperatures may allow one to
maximize
the amount of desirable products. The invention may not be limited to the use
of such
intermediate temperatures, however, and in other embodiments, lower and/or
higher
temperatures can be used.
The reactor(s) may also be operated at any suitable pressure. The reactor may
be
operated at a pressure of at least about 100 kPa, or at least about 200 kPa,
or at least
about 300 kPa, or at least about 400 kPa. The reactor may be operated at a
pressure
below about 600 kPa, or below about 400 kPa, or below about 200 kPa. The
reactor may
be operated at a pressure in the range from about 100 to about 600 kPa, or in
the range
from about 100 to about 400 kPa, or in the range from about 100 to about 200
kPa. The
invention may not be limited to the use of such pressures, however, and in
other
embodiments, lower and/or higher pressures may be employed.
It may be advantageous to heat the feed stream(s) at a relatively fast rate as
it
enters the reactor. High heating rates may be advantageous for a number of
reasons. For
instance, high heating rates may enhance the rate of mass transfer of the
reactants from
the bulk solid hydrocarbonaceous material to the catalytic reactant sites.
This may, for
example, facilitate introduction of volatile organic compounds formed during
the
pyrolysis of the solid hydrocarbonaceous material into the catalyst before
completely
thermally decomposing the solid hydrocarbonaceous material and/or the second
reactant
into generally undesired products (e.g., coke). In addition, high heating
rates may reduce
the amount of time the reactants are exposed to low temperatures (i.e.,
temperatures
between the temperature of the feed and the desired reaction temperature).
Prolonged
exposure of the reactants to low temperatures may lead to the formation of
undesirable
products via undesirable decomposition and/or reaction pathways. Examples of
suitable
heating rates for heating the feed stream(s) upon entering the reactor may
include, for
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example, greater than about 50 C/s, greater than about 100 C/s, greater than
about
200 C/s, greater than about 300 C/s, greater than about 400 C/s, greater
than about
500 C/s, greater than about 600 C/s, greater than about 700 C/s, greater
than about
800 C/s, greater than about 900 C/s, greater than about 1000 C/s, or
greater. In some
cases, the reactant(s) may be heated at a heating rate of between about 500
C/s and
about 1000 C/s. In some embodiments, the heating rate for heating the feed
stream(s)
upon entering the reactor may be between about 50 C/s and about 1000 C/s, or
between
about 50 C/s and about 400 C/s. The invention may not limited to the use of
such
heating rates, however, and in other embodiments, lower and/or higher heating
rates can
be used.
The mass-normalized space velocity of the hydrocarbonaceous material may be
selected to selectively produce a desired array of fluid hydrocarbon products.
As used
herein, the term "mass-normalized space velocity" of a component is defined as
the mass
flow rate of the component into the reactor (e.g., as measured in g/hr)
divided by the
mass of catalyst in the reactor (e.g., as measured in g) and has units of
inverse time. For
example, the mass-normalized space velocity of solid hydrocarbonaceous
material fed to
the reactor may be calculated as the mass flow rate of the solid
hydrocarbonaceous
material into the reactor divided by the mass of catalyst in the reactor. The
mass-
normalized space velocity of a component (e.g., the hydrocarbonaceous
material) in the
reactor may be calculated using different methods depending upon the type of
reactor
being used. For example, in systems employing batch or semi-batch reactors,
wherein
the solid hydrocarbonaceous material is not fed continuously to the reactor,
the solid
hydrocarbonaceous material does not have a mass-normalized space velocity. For
systems in which catalyst is fed to and/or extracted from the reactor during
reaction (e.g.,
circulating fluidized bed reactors), the mass-normalized space velocity may be
determined by calculating the average amount of catalyst within the volume of
the
reactor over a period of operation (e.g., steady-state operation).
The mass-normalized space velocity of the hydrocarbonaceous material fed to
the
reactor may be at a mass normalized space velocity of up to about 3 hour-1, or
up to
about 2 hour-1, or up to about 1.5 hour-1, or up to about 0.9 hour-1, or in
the range from
about 0.01 hour-1 to about 3 hour-1, or in the range from about 0.01 to about
2 hour-1, or
in the range from about 0.01 to about 1.5 hour-1, or in the range from about
0.01 to about
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0.9 hour-1, or in the range from about 0.01 hour-1 to about 0.5 hour-1, or in
the range
from about 0.1 hour-1 to about 0.9 hour-1, or in the range from about 0.1 hour-
1 to about
0.5 hour-1. The invention may not be limited to the use of such mass-
normalized space
velocities, however, and in other embodiments, lower and/or higher mass-
normalized
space velocities can be used.
The residence time of a reactant (e.g., the hydrocarbonaceous material) in the
reactor (i.e., the reactant residence time) may be at least about 1 second, at
least about
2 seconds, at least about 5 seconds, at least about 7 seconds, at least about
10 seconds, at
least about 15 seconds, at least about 20 seconds, at least about 25 seconds,
at least about
30 seconds, at least about 60 seconds, at least about 120 seconds, at least
about
240 seconds, or at least about 480 seconds. In some cases, the residence time
of a
reactant (e.g., the hydrocarbonaceous material) in the reactor may be less
than about 5
minutes, or from about 1 second and about 4 minutes, or from about 2 seconds
to about 4
minutes, or from about 5 seconds to about 4 minutes, or from about 7 seconds
to about 4
minutes, or from about 10 seconds to about 4 minutes, or from about 12 seconds
to about
4 minutes, or from about 15 seconds to about 4 minutes, or from about 20
seconds to
about 4 minutes, or from about 30 seconds to about 4 minutes, or from about 60
seconds
to about 4 minutes. Previous "fast pyrolysis" studies have, in many cases,
employed
systems with very short reactant residence times (e.g., less than 2 seconds).
In some
cases, however, the use of relatively longer residence times may allow for
additional
chemical reactions to form desirable products. Long residence times may be
achieved
by, for example, increasing the volume of the reactor and/or reducing the
volumetric
flow rate of the hydrocarbonaceous materials. It should be understood,
however, that in
some embodiments described herein, the residence time of the reactant (e.g.,
hydrocarbonaceous material) may be relatively shorter, e.g., less than about 2
seconds, or
less than about 1 second.
The contact time of the pyrolysis product (e.g., pyrolysis vapor) with the
catalyst
in the reactor may be at least about 1 second, at least about 2 seconds, at
least about
5 seconds, at least about 7 seconds, at least about 10 seconds, at least about
15 seconds,
at least about 20 seconds, at least about 25 seconds, at least about 30
seconds, at least
about 60 seconds, at least about 120 seconds, at least about 240 seconds, or
at least about
480 seconds. The contact time may be less than about 5 minutes, or from about
1 second
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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
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
5 from about 30 seconds to about 4 minutes, or from about 60 seconds to
about 4 minutes.
In certain cases where fluidized bed reactors are used, the feed material
(e.g., a
solid hydrocarbonaceous material) in the reactor may be fluidized by flowing a
fluid
stream through the reactor. In the exemplary embodiment of FIG. 1A, a fluid
stream 44
is used to fluidize the feed material in reactor 20. Fluid may be supplied to
the fluid
10 stream from a fluid source 24 and/or from the product streams of the
reactor via a
compressor 26. As used herein, the term "fluid" means a material generally in
a liquid,
supercritical, or gaseous state. Fluids, however, may also contain solids such
as, for
example, suspended or colloidal particles. In some embodiments, it may be
advantageous to control the residence time of the fluidization fluid in the
reactor. The
residence time of the fluidization fluid may be defined as the volume of the
reactor
divided by the volumetric flow rate of the fluidization fluid. The residence
time of the
fluidization fluid may be at least about 0.1 second, at least about 0.2
second, at least
about 0.5 second, at least about 1 second, at least about 2 seconds, at least
about
3 seconds, at least about 4 seconds, at least about 5 seconds, at least about
6 seconds, at
least about 8 seconds, at least about 10 seconds, at least about 12 seconds,
at least about
24 seconds, or at least about 48 seconds. The residence time of the
fluidization fluid
may be from about 0.1 second to about 48 seconds, from about 0.2 second to
about
48 seconds, from about 0.5 second to about 480 seconds, from about 1 second to
about
48 seconds, from about 3 seconds to about 48 seconds, from about 5 seconds to
about 48
seconds, from about 6 seconds to about 48 seconds, from about 8 seconds to
about 48
seconds, from about 10 seconds to about 48 seconds, from about 12 seconds to
about 48
seconds, or from about 24 seconds to about 48 seconds.
Suitable fluidization fluids that may be used in this invention include, for
example, inert gases (e.g., helium, argon, neon, etc.), hydrogen, nitrogen,
carbon
monoxide, and carbon dioxide, among others.
As shown in the illustrative embodiment of FIG. 1A, the products (e.g., fluid
hydrocarbon products) formed during the reaction of the reactants (e.g., the
solid
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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, and/or catalyst. In one set of embodiments,
the desired
reaction product(s) (e.g., liquid aromatic hydrocarbons, olefin hydrocarbons,
gaseous
products, etc.) may be recovered from an effluent stream of the reactor.
As shown in the illustrative embodiment of FIG. 1A, product stream 30 may be
fed to an optional solids separator 32. The solids separator may be used, in
some cases,
to separate the reaction products from catalyst (e.g., at least partially
deactivated catalyst)
present in the product stream. In addition, the solids separator may be used,
in some
instances, to remove coke and/or ash from the catalyst. In some embodiments,
the solids
separator may comprise optional purge stream 33, which may be used to purge
coke, ash,
and/or catalyst from the solids separator.
The equipment required to achieve the solids separation and/or decoking steps
can be readily designed by one of ordinary skill in the art. For example, the
solids
separator may comprise a vessel comprising a mesh material that defines a
retaining
portion and a permeate portion of the vessel. The mesh may serve to retain the
catalyst
within the retaining portion while allowing the reaction product to pass to
the permeate
portion. The catalyst may exit the solids separator through a port on the
retaining side of
the mesh while the reaction product may exit a port on the permeate side of
the mesh.
Other examples of solids separators and/or decokers are described in more
detail in Kirk-
Othmer Encyclopedia of Chemical Technology (Online), Vol. 11, Hoboken, N.J.:
Wiley-
Interscience, c2001-, pages 700-734; and C. D. Cooper and F. C. Alley. Air
Pollution
Control, A Design Approach, Second Ed. Prospect Heights, Illinois: Waveland
Press,
Inc. c1994, pages 127-149, which are incorporated herein by reference.
The solids separator may be operated at any suitable temperature. In some
embodiments, the solids separator may be operated at elevated temperatures.
For certain
reactions, the use of elevated temperatures in the solids separator can allow
for additional
reforming and/or reaction of the compounds from the reactor. This may allow
for the
increased formation of desirable products. While not wishing to be bound by
any theory,
it is believed that elevated temperatures in the solids separator may provide
enough
energy to drive endothermic reforming reactions. The solids separator may be
operated
at a temperature of, for example, between about 25 C and about 200 C,
between about
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200 C and about 500 C, between about 500 C and about 600 C, or between
about
600 C and about 800 C. In some cases, the solids separator may be operated
at
temperatures of at least about 500 C, at least about 600 C, at least 700 C,
at least
800 C, or higher.
It may be beneficial to control the residence time of the catalyst in the
solids
separator. The residence time of the catalyst in the solids separator may be
defined as
the volume of the solids separator divided by the volumetric flow rate of the
catalyst
through the solids separator. In some cases, relatively long residence times
of the
catalyst in the solids separator may be desired in order to facilitate the
removal of
sufficient amounts of ash, coke, and/or other undesirable products from the
catalyst. In
addition, by employing relatively long residence times of the catalyst in the
solids
separator, the pyrolysis products may be further reacted to produce desirable
products.
The residence time and temperature in the solids separator may together be
selected such
that a desired product stream is produced. The residence time of the catalyst
in the solids
separator may be at least about 1 second, at least about 5 seconds, at least
about 7
seconds, at least about 10 seconds, at least about 30 seconds, at least about
60 seconds, at
least about 120 seconds, at least about 240 seconds, at least about 300
seconds, at least
about 600 seconds, or at least about 1200 seconds. Methods for controlling the
residence
time of the catalyst in the solids separator are known by those skilled in the
art. For
example, in some cases, the interior wall of the solids separator may comprise
baffles
that serve to restrict the flow of catalyst through the solids separator
and/or increase the
path length of fluid flow in the solids separator. Additionally or
alternatively, the
residence time of the catalyst in the solids separator may be controlled by
controlling the
flow rate of the catalyst through the solids separator (e.g., by controlling
the flow rate of
the fluidizing fluid through the reactor).
The solids separator may have any suitable size. For example, the solids
separator may have a volume between about 0.1-1 L, 1-50 L, 50-100 L, 100-250
L,
250-500 L, 500-1000 L, 1000-5000 L, 5000-10,000 L, or 10,000-50,000 L. In some
instances, the solids separator may have a volume greater than about 1 L, or
in other
instances, greater than about 10 L, 50 L, 100 L, 250 L, 500 L, 1,000 L, or
10,000 L.
Solids separator volumes greater than 50,000 L are also possible. The solids
separator
may be cylindrical, spherical, or any other shape and may be circulating or
non-
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circulating. In some embodiments, the solids separator may comprise a vessel
or other
unit operation similar to that used for one or more of the reactor(s) used in
the process.
The flow path for the catalyst in the solids separator may comprise any
suitable
geometry. For example, the flow path may be substantially straight. In some
cases, the
solids separator may comprise a flow channel with a serpentine, meandering,
helical, or
any other suitable shape. The ratio of the length of the flow path of the
solids separator
(or, in certain embodiments, the path length of the catalyst through the
solids separator)
to the average diameter of the solids separator channel may comprise any
suitable ratio.
The ratio may be at least about 2:1, at least 5:1, at least 10:1, at least
50:1, at least 100:1,
or greater.
The solids separator may not be required in all embodiments. For example, for
situations in which catalytic fixed bed reactors are employed, the catalyst
may be
retained within the reactor, and the reaction products may exit the reactor
substantially
free of catalyst, thus negating the need for a separation step.
The separated catalyst may exit the solids separator via stream 34. A portion
of
the separated catalyst may be returned to the reactor via a return pipe, not
shown in FIG.
1A. The catalyst exiting the separator may be at least partially deactivated.
The
separated catalyst may be fed to a regenerator 36 in which any catalyst that
was at least
partially deactivated may be re-activated. The regenerator may comprise an
optional
purge stream 37, which may be used to purge coke, ash, and/or catalyst from
the
regenerator. Methods for activating catalyst are well-known to those skilled
in the art,
for example, as described in Kirk- Othmer Encyclopedia of Chemical Technology
(Online), Vol. 5, Hoboken, N.J. : Wiley-Interscience, c2001-, pages 255-322,
which are
incorporated herein by reference.
A portion of the catalyst may be removed from the reactor through a catalyst
exit
port (not shown in FIG. 1A.). 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. 1A). 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. 1A.).
An oxidizing agent may be fed to the regenerator via a stream 38, e.g., as
shown
in FIG. 1A. The oxidizing agent may originate from any source including, for
example,
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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. lA
comprises a vent
stream 40 which may include regeneration reaction products, residual oxidizing
agent,
etc.
The regenerator may be of any suitable size mentioned above in connection with
the reactor or the solids separator. In addition, the regenerator may be
operated at
elevated temperatures in some cases (e.g., at least about 300 C, 400 C, 500
C, 600 C,
700 C, 800 C, or higher). The residence time of the catalyst in the
regenerator may also
be controlled using methods known by those skilled in the art, including those
outlined
above. The mass flow rate of the catalyst through the regenerator may be
coupled to the
flow rate(s) in the reactor and/or solids separator in order to preserve the
mass balance in
the system.
The regenerated catalyst may exit the regenerator via stream 42. The
regenerated
catalyst may be recycled back to the reactor via recycle stream 47. In some
cases,
catalyst may be lost from the system or removed intentionally during
operation.
Additional "makeup" catalyst may be added to the system via a makeup stream
46. The
regenerated and makeup catalyst may be fed to the reactor with the
fluidization fluid via
recycle stream 47. Alternatively, the catalyst and fluidization fluid may be
fed to the
reactor via separate streams.
Referring back to solids separator 32 in FIG. 1A, 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. Examples of condensers that may be used are
described in more detail in Perry 's Chemical Engineers' Handbook, Section 11:
"Heat
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Transfer Equipment." 8th ed. New York : McGraw-Hill, c2008, which is
incorporated
herein by reference.
The condenser may also make use of pressure change to condense portions of the
product stream. In FIG. 1A, stream 54 may comprise the liquid fraction of the
reaction
products (e.g., water, aromatic compounds, olefin compounds, etc.), and stream
74 may
comprise the gaseous fraction of the reaction products (e.g., CO, CO2, F12,
etc.). In some
embodiments, the gaseous fraction may be fed to a vapor recovery system 70.
The vapor
recovery system may be used, for example, to recover any desirable vapors
within stream
74 and transport them via stream 72. In addition, stream 76 may be used to
transport
CO, CO2, and/or other non-recoverable gases from the vapor recovery system.
The
optional vapor recovery system may be placed in other locations. For example,
in some
embodiments, a vapor recovery system may be positioned downstream of purge
stream
54. One skilled in the art can select an appropriate placement for a vapor
recovery
system.
Other products (e.g., excess gas) may be transported to optional compressor 26
via stream 56, where they may be compressed and used as fluidization gas in
the reactor
(stream 22) and/or where they may assist in transporting the hydrocarbonaceous
material
to the reactor (streams 58) or may be used to transport catalyst to the
reactor (not shown),
or may be used to transport additional non-solid feeds to the reactor. In some
instances,
the liquid fraction may be further processed, for example, to separate the
water phase
from the organic phase, to separate individual compounds, etc.
It should be understood that, while the set of embodiments described by FIG.
lA
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
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production process (e.g., after passage through the reactor, after separation,
after
condensation, etc.).
The process may involve the use of more than one reactor. For instance,
multiple
reactors may be connected in fluid communication with each other, for example,
to
operate in series and/or in parallel, as shown in the exemplary embodiment of
FIG. 1B.
The process may comprise providing a solid hydrocarbonaceous material in a
first
reactor and pyrolyzing, within the first reactor, at least a portion of the
solid
hydrocarbonaceous material under reaction conditions sufficient to produce one
or more
pyrolysis products. A catalyst may be provided to the first reactor, and at
least a portion
of the one or more pyrolysis products in the first reactor may be
catalytically reacted
using the catalyst under reaction conditions sufficient to produce one or more
fluid
hydrocarbon products. The process may further comprise catalytically reacting
at least a
portion of the one or more pyrolysis products in a second reactor using a
catalyst under
reaction conditions sufficient to produce one or more fluid hydrocarbon
products. After
catalytically reacting at least a portion of the one or more pyrolysis
products in the
second reactor, the process may comprise the step of further reacting within
the second
reactor at least a portion of the one or more fluid hydrocarbon products from
the first
reactor to produce one or more other hydrocarbon products.
In FIG. 1B, the reaction product from reactor 20 may be transported to a
second
reactor 20'. Those skilled in the art are familiar with the use of multiple-
reactor systems
for the pyrolysis of organic material to produce organic products and such
systems are
known in the art. While FIG. 1B illustrates a set of embodiments in which the
reactors
are in fluid communication with each other, in some instances, the two
reactors may not
be in fluid communication. For example, a first reactor may be used to produce
a first
reaction product which may be transported to a separate facility for reaction
in a second
reactor. In some instances, a composition comprising a solid hydrocarbonaceous
material (with or without a catalyst) may be heated in a first reactor, and at
least a portion
of the solid hydrocarbonaceous material may be pyrolyzed to produce a
pyrolysis
product (and optionally at least partially deactivated catalyst). The first
pyrolysis
product may be in the form of a liquid and/or a gas. The composition
comprising the
first pyrolysis product may then be heated in a second reactor, which may or
may not be
in fluid communication with the first reactor. After the heating step in the
second
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reactor, a second pyrolysis product from the second reactor may be collected.
The
second pyrolysis product may be in the form of a liquid and/or a gas. In some
cases, the
composition comprising hydrocarbonaceous material that is fed into the first
reactor may
comprise, for example, a mixture of a solid hydrocarbonaceous material and a
solid
catalyst. The first pyrolysis product produced from the first reactor may be
different in
chemical composition, amount, state (e.g., a fluid vs. a gas) than the second
pyrolysis
product. For example, the first pyrolysis product may substantially include a
liquid,
while the second pyrolysis product may substantially include a gas. In another
example,
the first pyrolysis product may include a fluid product (e.g., a bio-oil,
sugar), and the
second pyrolysis product may comprise a relatively higher amount of aromatics
than the
first pyrolysis product. In some instances, the first pyrolysis product may
include a fluid
product (e.g., including aromatic compounds), and the second pyrolysis product
may
comprise a relatively higher amount of olefins than the first pyrolysis
product. In yet
another example, the first pyrolysis product may include a fluid product
(e.g., a bio-oil,
sugar), and the second pyrolysis product may comprise a relatively higher
amount of
oxygenated aromatic compounds than the first pyrolysis product.
One or more of the reactors in a multiple reactor configuration may comprise a
fluidized bed reactor (e.g., a circulating fluidized bed reactor, a turbulent
fluidized bed
reactor, etc.) or, in other instances, any other type of reactor (e.g., any of
the reactors
mentioned above). For example, the first reactor may comprise a circulating
fluidized
bed reactor or a turbulent fluidized bed reactor, and the second reactor
comprises a
circulating fluidized bed reactor or a turbulent fluidized bed reactor in
fluid
communication with the first reactor. In addition, the multiple reactor
configuration may
include any of the additional processing steps and/or equipment mentioned
herein (e.g., a
solids separator, a regenerator, a condenser, etc.). The reactors and/or
additional
processing equipment may be operated using any of the processing parameters
(e.g.,
temperatures, residence times, etc.) mentioned herein.
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
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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.
The catalyst particles described herein may comprise polycrystalline solids
(e.g.,
polycrystalline particles) in some cases. The catalyst particles may also
comprise single
crystals, in some embodiments. In certain cases, the particles may be distinct
and separate
physical objects that are stand-alone. In other cases, the particles may, at
least at certain
points in their preparation and/or use, comprise an agglomerate of a plurality
of individual
particles in intimate contact with each other.
A catalyst used in embodiments described herein (e.g., in the feed stream, in
the
reactor, etc.) may be of any suitable size. In some cases, it may be
advantageous to use
catalysts comprising relatively small catalyst particles, which may, as
mentioned
previously, in certain embodiments, be in the form of larger catalyst objects
that may be
comprised of a plurality of agglomerated catalyst particles. In some
embodiments, for
example, the use of small catalyst particles may increase the extent to which
the
hydrocarbonaceous material may contact the surface sites of the catalyst due
to, for
example, increased external catalytic surface area and decreased diffusion
distances
through the catalyst. In some cases, catalyst size and/or catalyst particle
size may be
chosen based at least in part on, for example, the type of fluid flow desired
and the catalyst
lifetime.
In some embodiments, the 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
500 microns, less than about 60 mesh (250 microns), less than about 100 mesh
(149
microns), less than about 140 mesh (105 microns), less than about 170 mesh (88
microns),
less than about 200 mesh (74 microns), less than about 270 mesh (53 microns),
or less than
about 400 mesh (37 microns), or smaller.
The catalyst may comprise particles having a maximum cross-sectional dimension
of less than about 5 microns, less than about 1 micron, less than about 500
nm, less than
about 100 nm, between about 100 nm and about 5 microns, between about 500 nm
and
about 5 microns, between about 100 nm and about 1 micron, or between about 500
nm and
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about 1 micron. Catalyst particles having the dimensions within the ranges
noted
immediately above may be agglomerated to form discrete catalyst objects having
dimensions within the ranges noted above. As used here, the "maximum cross-
sectional
dimension" of a particle refers to the largest dimension between two
boundaries of a
particle. One of ordinary skill in the art would be capable of measuring the
maximum
cross-sectional dimension of a particle by, for example, analyzing a scanning
electron
micrograph (SEM) of a catalyst preparation. In embodiments comprising
agglomerated
particles, the particles should be considered separately when determining the
maximum
cross-sectional dimensions. In such a case, the measurement may be performed
by
establishing imaginary boundaries between each of the agglomerated particles,
and
measuring the maximum cross-sectional dimension of the hypothetical,
individuated
particles that result from establishing such boundaries. In some embodiments,
a relatively
large number of the particles within a catalyst may have maximum cross-
sectional
dimensions that lie within a given range. For example, in some embodiments, at
least
about 50%, at least about 75%, at least about 90%, at least about 95%, or at
least about
99% of the particles within a catalyst have maximum cross-sectional dimensions
of less
than about 5 microns, less than about 1 micron, less than about 500 nm, less
than about
100 nm, between about 100 nm and about 5 microns, between about 500 nm and
about
5 microns, between about 100 nm and about 1 micron, or between about 500 nm
and about
1 micron.
A relatively large percentage of the volume of the catalyst can be occupied by
particles with maximum cross-sectional dimensions within a specific range, in
some cases.
For example, in some embodiments, at least about 50%, at least about 75%, at
least about
90%, at least about 95%, or at least about 99% of the sum of the volumes of
all the
catalyst used is occupied by particles having maximum cross-sectional
dimensions of
less than about 5 microns, less than about 1 micron, less than about 500 nm,
less than
about 100 nm, between about 100 nm and about 5 microns, between about 500 nm
and
about 5 microns, between about 100 nm and about 1 micron, or between about 500
nm and
about 1 micron.
In some embodiments, the particles within a catalyst may be substantially the
same size. For example, the catalyst may comprise particles with a
distribution of
dimensions such that the standard deviation of the maximum cross-sectional
dimensions
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of the particles is no more than about 50%, no more than about 25%, no more
than about
10%, no more than about 5%, no more than about 2%, or no more than about 1% of
the
average maximum cross-sectional dimensions of the particles. Standard
deviation
(lower-case sigma) may be calculated as:
E(D ¨ Davg) 2
a =li=1
n ¨ 1
wherein D, is the maximum cross-sectional dimension of particle i, D avg is
the average of
the maximum cross-sectional dimensions of all the particles, and n is the
number of
particles within the catalyst. The percentage comparisons between the standard
deviation
and the average maximum cross-sectional dimensions of the particles outlined
above can
be obtained by dividing the standard deviation by the average and multiplying
by 100%.
Using catalysts including particles within a chosen size distribution
indicated
above can lead to an increase in the yield and/or selectivity of aromatic
compounds
produced by the reaction of the hydrocarbonaceous material. For example, in
some
cases, using catalysts containing particles with a desired size range (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.).
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.
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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, July 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
Angstroms, 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 Angstroms larger than the atomic radii (about 6.2-6.3
Angstroms).
One of ordinary skill in the art will understand how to determine the pore
size
(e.g., minimum pore size, average 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, July
1, 1993, pp. 185-195, which is incorporated herein by reference in its
entirety. 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.
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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.
A screening method may be used to select catalysts with appropriate pore sizes
for
the conversion of specific pyrolysis product molecules. The screening method
may
comprise determining the size of pyrolysis product molecules desired to be
catalytically
reacted (e.g., the molecule kinetic diameters of the pyrolysis product
molecules). One of
ordinary skill in the art can calculate, for example, the kinetic diameter of
a given
molecule. The type of catalyst may then be chosen such that the pores of the
catalyst (e.g.,
Norman adjusted minimum radii) are sufficiently large to allow the pyrolysis
product
molecules to diffuse into and/or react with the catalyst. In some embodiments,
the
catalysts are chosen such that their pore sizes are sufficiently small to
prevent entry and/or
reaction of pyrolysis products whose reaction would be undesirable.
The catalyst may be selected from naturally-occurring zeolites, synthetic
zeolites
and combinations thereof. The catalyst may be a Mordenite Framework Inverted
(MFI)
type zeolite catalyst, such as a ZSM-5 zeolite catalyst. Catalysts comprising
ZSM-5 that
may be used with or without modification are available commercially. The
catalysts that
are provided for herein may comprise acid or catalytically active sites. While
not wishing
to be bound by theory, it is believed that various acid sites in ZSM-5 and
other zeolites are
catalytically active for reactions of the hydrocarbonaceous materials
including
dehydration, decarbonylation, decarboxylation, isomerization, oligomerization
and/or
dehydrogenation, hence the terms "acid sites" and "catalytically active sites"
may be used
interchangeably. Other types of useful zeolite catalysts may include
ferrierite, zeolite Y,
zeolite beta, modernite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)A1P0-
31, SSZ-23, mixtures of two or more thereof, and the like.
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, platinum, vanadium, palladium, manganese, cobalt, zinc,
copper,
chromium, gallium, and/or any of their oxides, 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
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comprise small amounts of structure stabilizing elements such as phosphorus,
lanthanum,
rare earths, and the like, typically at levels that are less than about 1 % by
weight of the
zeolite. The catalyst may be conditioned before operation in the process by a
wide range
of techniques known to those skilled in the art such as, but not limited to,
oxidation,
calcination, reduction, cyclic oxidation and reduction, steaming, hydrolysis,
and the like.
The metal and/or metal oxide may be incorporated into the lattice structure of
the catalyst.
For example, the metal and/or metal oxide may be included during the
preparation of the
catalyst, and the metal and/or metal oxide may occupy a lattice site of the
resulting catalyst
(e.g., a zeolite catalyst). As another example, the metal and/or metal oxide
may react or
otherwise interact with a zeolite such that the metal and/or metal oxide
displaces an atom
within the lattice structure of the zeolite.
In certain embodiments, a Mordenite Framework Inverted (MFI) zeolite catalyst
comprising gallium may be used. For example, a galloaluminosilicate MFI
(GaAlMFI)
zeolite catalyst may be used. One of ordinary skill in the art would be
familiar with
GaAlMFI zeolites, which may be thought of as aluminosilicate MFI zeolites in
which
some of the Al atoms have been replaced with Ga atoms. In some instances, the
zeolite
catalyst may be in the hydrogen form (e.g., H-GaAlMFI). The
galloaluminosilicate MFI
catalyst may be a ZSM-5 zeolite catalyst in which some of the aluminum atoms
have been
replaced with gallium atoms, in some embodiments.
In some instances, the ratio of moles of Si in the galloaluminosilicate
zeolite
catalyst to the sum of the moles of Ga and Al (i.e., the molar ratio expressed
as
Si:(Ga+A1)) in the galloaluminosilicate zeolite catalyst may be at least about
15:1, at least
about 20:1, at least about 25:1, at least about 35:1, at least about 50:1, at
least about 75:1,
or higher. In some embodiments, it may be advantageous to employ a catalyst
with a ratio
of moles of Si in the zeolite to the sum of the moles of Ga and Al of between
about 15:1
and about 100:1, from about 15:1 to about 75:1, between about 25:1 and about
80:1, or
between about 50:1 and about 75:1. In some instances, the ratio of moles of Si
in the
galloaluminosilicate zeolite catalyst to the moles of Ga in the
galloaluminosilicate zeolite
catalyst may be at least about 30:1, at least about 60:1, at least about
120:1, at least about
200:1, between about 30:1 and about 300:1, between about 30:1 and about 200:1,
between
about 30:1 and about 120:1, or between about 30:1 and about 75:1. The ratio of
the moles
of Si in the galloaluminosilicate zeolite catalyst to the moles of Al in the
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galloaluminosilicate zeolite catalyst may be at least about 10:1, at least
about 20:1, at least
about 30:1, at least about 40:1, at least about 50:1, at least about 75:1,
between about 10:1
and about 100:1, between about 10:1 and about 75:1, between about 10:1 and
about 50:1,
between about 10:1 and about 40:1, or between about 10:1 and about 30:1.
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 Angstrom pores and 7-8 Angstrom pores). In other cases,
a
mixture of two or more catalysts may be employed to establish the bimodal
distribution
(e.g., a mixture of two catalysts, each catalyst type including a distinct
range of average
pore sizes). In some embodiments, one of the one or more catalysts comprises a
zeolite
catalyst and another of the one or more catalysts comprises a non-zeolite
catalyst (e.g., a
mesoporous catalyst, a metal oxide catalyst, etc.).
For example, in some embodiments at least about 70%, at least about 80%, at
least about 90%, at least about 95%, at least about 98%, or at least about 99%
of the
pores of the one or more catalysts (e.g., a zeolite catalyst, a mesoporous
catalyst, etc.)
have smallest cross-sectional diameters that lie within a first size
distribution or a second
size distribution. In some cases, at least about 2%, at least about 5%, or at
least about
10% of the pores of the one or more catalysts have smallest cross-sectional
diameters
that lie within the first size distribution; and at least about 2%, at least
about 5%, or at
least about 10% of the pores of the one or more catalysts have smallest cross-
sectional
diameters that lie within the second size distribution. In some cases, the
first and second
size distributions are selected from the ranges provided above. In certain
embodiments,
the first and second size distributions are different from each other and do
not overlap.
An example of a non-overlapping range is 5.9-6.3 Angstroms and 6.9-8.0
Angstroms,
and an example of an overlapping range is 5.9-6.3 Angstroms and 6.1-6.5
Angstroms.
The first and second size distributions may be selected such that the range
are not
immediately adjacent one another, an example being pore sizes of 5.9-6.3
Angstroms and
6.9-8.0 Angstroms. An example of a range that is immediately adjacent one
another is
pore sizes of 5.9-6.3 Angstroms and 6.3-6.7 Angstroms.
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As a specific example, in some embodiments one or more catalysts is 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, 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 Angstroms and about
6.3 Angstroms or between about 7 Angstroms and about 8 Angstroms. 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 Angstroms
and about
6.3 Angstroms; 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 Angstroms and about 8 Angstroms.
In some embodiments, at least about 70%, at least about 80%, at least about
90%,
at least about 95%, at least about 98%, or at least about 99% of the pores of
the one or
more catalysts have smallest cross-sectional diameters between about 5.9
Angstroms and
about 6.3 Angstroms or between about 7 Angstroms and about 200 Angstroms. In
addition, at least about 2%, at least about 5%, or at least about 10% of the
pores of the
one or more catalysts have smallest cross-sectional diameters between about
5.9 Angstroms and about 6.3 Angstroms; and at least about 2%, at least about
5%, or at
least about 10% of the pores of the one or more catalysts have smallest cross-
sectional
diameters between about 7 Angstroms and about 200 Angstroms.
In some embodiments, at least about 70%, at least about 80%, at least about
90%,
at least about 95%, at least about 98%, or at least about 99% of the pores of
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
Angstroms
and about 6.3 Angstroms 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 Angstroms and about 200 Angstroms, between
about 7 Angstroms and about 100 Angstroms, between about 7 Angstroms and about
50
Angstroms, or between about 100 Angstroms and about 200 Angstroms. In some
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embodiments, the second catalyst may be mesoporous (e.g., have a pore size
distribution
of between about 2 nm and about 50 nm).
In some embodiments, the 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 described herein
(e.g., sugar
cane bagasse, glucose, wood, corn stover, cellulose, hemi-cellulose, lignin,
or any
others). 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 Angstroms and about 6.3
Angstroms. The second pore size distribution may be, for example, between
about 7
Angstroms and about 200 Angstroms, between about 7 Angstroms and about 100
Angstroms, between about 7 Angstroms and about 50 Angstroms, or between about
100
Angstroms and about 200 Angstroms. In some cases, the second catalyst may be
mesoporous or non-porous.
The first catalyst may be selective for catalytically reacting the first
component or
a derivative thereof to produce a fluid hydrocarbon product. In addition, the
second
catalyst may be selective for catalytically reacting the second component or a
derivative
thereof to produce a fluid hydrocarbon product. The method may further
comprise
pyrolyzing within the reactor at least a portion of the hydrocarbonaceous
material under
reaction conditions sufficient to produce one or more pyrolysis products and
catalytically
reacting at least a portion of the pyrolysis products with the first and
second catalysts to
produce the one or more hydrocarbon products. In some instances, at least
partially
deactivated catalyst may also be used.
In certain embodiments, a method used in combination with embodiments
described herein includes increasing the catalyst to hydrocarbonaceous
material mass ratio
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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, methylbenzene, methylethylbenzene,
trimethylbenzene,
xylenes, indanes, naphthalene, methylnaphthelene, dimethylnaphthalene,
ethylnaphthalene,
hydrindene, methylhydrindene, and dimethylhydrindene and combinations thereof.
In some embodiments, the reaction chemistry of a catalyst may be affected by
adding one or more additional compounds. For example, the addition of a metal
to a
catalyst may result in a shift in selective formation of specific compounds
(e.g., addition
of metal to alumina-silicate catalysts may result in the production of more
CO). In
addition, when the fluidization fluid comprises hydrogen, the amount of coke
formed on
the catalyst may be decreased.
The catalyst may comprise both silica and alumina. The silica (Si02) and
alumina
(A1203) 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 (Si02) to alumina
(A1203) molar ratio
of between about 10:1 and about 50:1, between about 10:1 and about 40:1, or
between
about 10:1 and about 20:1, or about 15:1.
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
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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.
In some embodiments, the articles and 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 18 wt%, at least about 20 wt%, at least about 25 wt%, at least
about 30 wt%,
at least about 35 wt%, at least about 39 wt%, between about 18 wt% and about
40 wt%,
between about 18 wt% and about 35 wt%, between about 20 wt% and about 40 wt%,
between about 20 wt% and about 35 wt%, between about 25 wt% and about 40 wt%,
between about 25 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.
As used herein, 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%.
As used herein, the term "aromatic compound" is used to refer to a hydrocarbon
compound comprising one or more aromatic groups such as, for example, single
aromatic ring systems (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-
ehtyl
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,
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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,
Cio-C14.
In some embodiments, aromatic compounds (especially p-xylene) may be
selectively produced when the mass-normalized space velocity of the solid
hydrocarbonaceous material fed to the reactor is up to about 3 hour-1, or up
to about 2
hour-1, or up to about 1.5 hour-1, or up to about 0.9 hour-1, or in the range
from about
0.01 hour-1 to about 3 hour-1, or in the range from about 0.01 to about 2 hour-
1, or in the
range from about 0.01 to about 1.5 hour-1, or in the range from about 0.01 to
about 0.9
hour-1, or in the range from about 0.01 hour-1 to about 0.5 hour-1, or in the
range from
about 0.1 hour-1 to about 0.9 hour-1, or in the range from about 0.1 hour-1 to
about
0.5 hour-1. In some instances, aromatic compounds (especially p-xylene) may be
selectively produced when the reactor is operated at a temperature of between
about
400 C and about 600 C (or between about 425 C and about 500 C, or between
about
440 C and about 460 C). In addition, certain heating rates (e.g., at least
about 50 C/s,
or at least about 400 C/s), high catalyst-to-feed mass ratios (e.g., at least
about 5:1),
and/or high silica to alumina molar ratios in the catalyst (e.g., at least
about 30:1) may be
used to facilitate selective production of aromatic compounds (especially p-
xylene).
Some such and other process conditions may be combined with a particular
reactor type,
such as a fluidized bed reactor (e.g., a circulating fluidized bed reactor),
to selectively
produce aromatic and/or olefin compounds.
Furthermore, in some embodiments, the catalyst may be chosen to facilitate
selective production of aromatic products (especially p-xylene). For example,
ZSM-5
may, in some cases, preferentially produce relatively higher amounts of
aromatic
compounds. In some cases, catalysts that include 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,
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in some embodiments, catalysts with average pore diameters between about
5.9 Angstroms and about 6.3 Angstroms may be particularly useful in producing
aromatic compounds. In addition, catalysts with average pore diameters between
about
7 Angstroms and about 8 Angstroms 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 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
2:1,
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.
In some embodiments, the catalyst to hydrocarbonaceous material mass ratio in
the feed is adjusted to produce desirable products and/or favorable yields. As
such, the
catalyst to hydrocarbonaceous material mass ratio may be, for example, at
least about
0.5:1, at least about 1:1, at least about 2:1, at least about 5:1, at least
about 10:1, at least
about 15:1, at least about 20:1, or higher in some embodiments; or, less than
about 0.5:1,
less than about 1:1, less than about 2:1, less than about 5:1, less than about
10:1, less than
about 15:1, or less than about 20:1 in other embodiments.
Furthermore, processes described herein may result in lower coke formation
than
certain existing methods. For example, in some embodiments, a pyrolysis
product can
be formed with less than about 30 wt%, less than about 25 wt%, less than about
20 wt%,
than about 15 wt%, or less than about 10 wt% of the pyrolysis product being
coke. The
amount of coke formed is measured as the weight of coke formed in the system
divided
by the weight of hydrocarbonaceous material used in forming the pyrolysis
product.
The following non-limiting examples are intended to illustrate various
aspects and features of the invention.
EXAMPLE
A series of zeolite catalysts are prepared and used in CFP processes for
converting furan, 2-methylfuran (2MF), and pinewood, to fluid hydrocarbon
products.
Catalysts
Four catalysts identified as ZSM, GaZSM, SD and GaSD are used. ZSM, which
is ZSM-5, (Si/Al= 15). GaZSM is made using ion exchange, where 1 g of ZSM is
refluxed in 100 mL of an aqueous solution of Ga(NO3)3 (0.01 M) at 70 C for 12
h. After
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ion exchange, the solution is dried at 110 C to form a dry powder. The dry
powder is
calcined under air at 550 C. SD is spray-dried HZSM-5. GaSD is made by
incipient
wetness impregnation of SD using a Ga(NO3)3 solution (0.43 M). The impregnated
GaSD is dried at 110 C and calcined under air at 550 C. The Ga content for
GaZSM and
GaSD is determined by inductively coupled plasma (ICP) analysis.
Chemical liquid deposition (CLD) employing tetraethylorthosilicate (TEOS) is
used to modify the catalysts and thereby reduce their pore-mouth opening
sizes. 1 g of
catalyst is dispersed in 25 mL of hexanes. Then 0.15 mL of TEOS is added. The
mixture is refluxed at 90 C for 1 h with stirring. The catalyst is recovered
by centrifuge.
The catalyst is dried at 100 C for 2 h and calcined at 500 C for 4 h in dry
air. The pore
mouth modification process is repeated two more times. The pore mouth modified
catalysts may be referred to as "silylated" catalysts and are identified below
with "*," for
example, ZSM*, GaZSM* and GaSD*. The modification process may be referred to
as
a TEOS CLD silyation process.
The catalyst samples are analyzed by temperature programmed desorption of
isopropyl amine (IPA) or 2,4,6-collidine (2,4,6-trimethylpyridine). Before
adsorption,
the sample is degassed for 2 h at 823 K. After cooling the sample to 393 K, it
is exposed
for 1 h to He that had been saturated with isopropylamine or 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 IPA or
2,4,6-
collidine.The sample is heated to 973 K at 10 K/min. The total amount of amine
desorbed is used to calculate the total number of acid sites, 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 IPA
desorption
detects acid sites within the zeolite pores as well as the 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 sylilation 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-
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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.
Table 1. Acid concentrations obtained by 2,4,6-collidine (kinetic diameter =
7.4
Angstroms) and isopropylamine (IPA, kinetic diameter = 5.2 Angstroms)
temperature
programmed desorption analysis.
Adsorbent Collidine Collidine Collidine Collidine IPA IPA IPA IPA
Catalyst ZSM-5 ZSM-5 Ga- Ga- ZSM- ZSM-5 GaZSM- GaZSM-
ZSM-5 ZSM-5 5 -5 -5
Acids Total Bronsted Total Bronsted Total Bronsted Total
Bronsted
Before 0.096 0.0548 0.051 0.0311 1.396 0.651 1.015
0.557
silylation
After 0.041 0.0344 0.032 0.0265 0.706 0.413 0.676
0.335
silylation
Reduction 57 37 37 15 49 37 33 40
(%)
Catalytic Conversion of Furan and 2MF
The catalytic reactions are carried out in a fixed-bed quartz reactor of 0.5
inch
(1.27 cm) O.D. The catalyst, which is in the form of a fixed-bed of
particulate solids, is
held in the reactor by a quartz frit. The catalyst bed is calcined at a
temperature of 600 C
with air flowing at a rate of 60 mL/min. After calcination the reactor is
purged by
helium at 408 mL/min for 5 min. Furan is pumped into the helium stream using a
syringe
pump. Prior to the test run, the furan bypasses the reactor for 30 min. The
helium
stream containing the furan is then switched to go through the reactor. An air
bath
condenser is used to trap the heavy products. Gas phase products are collected
by air
bags. All runs are conducted at atmospheric pressure. No pressure drop is
detected across
the catalyst bed. After the reaction process is completed, the reactor is
purged by helium
at a flow rate of 408 mL/min for 45 seconds at the reaction temperature. The
effluent is
collected using air bags. After each reaction is completed, spent catalyst is
regenerated at
a temperature of 600 C using air at a flow rate of 60 mL/min. CO formed during
regeneration is converted to CO2 by a copper converter (copper oxide) at a
temperature
of 240 C. CO2 is trapped by a CO2 trap. Coke yield is determined by measuring
the
weight change of the CO2 trap. Gas products are identified by GC-MS (Shimadzu-
2010)
and quantified by GC-FID/TCD (Shimadzu 2014 for gas samples, and HP-7890 for
liquid samples). All hydrocarbons in the gas phase products are quantified by
the GC-
FID. The CO and CO2 in the gas phase products are quantified by the GC-TCD.
The GC-
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FID is calibrated by C2 - C6 normal olefins standards (Scott Specialty Gas,
1000 ppm for
each olefin), furan, benzene, toluene, xylenes (gas phase standards are
prepared for these
aromatics that can vaporize at room temperature), ethylbenzene, styrene,
indene,
naphthalene, and benzofuran. The sensitivity of a hydrocarbon is assumed to be
proportional to the number of carbon molecules with similar structure (e.g.
styrene vs.
methylstyrenes; indene vs. methylindenes). The GC-TCD is calibrated by CO and
CO2
standards (Airgas, 6% CO2 and 14% CO, balanced by helium). Less than 0.05%
carbon
or the products are collected in the condenser. A majority of the products are
in either
the gas phase or coke deposited on the catalyst. Carbon balances close with >
90% for all
runs.
The reaction conditions for the furan conversion are a temperature of 550 C,
space velocity (WHSV) of 10.2 h-1, and a partial pressure of 6 torr. The furan
is pumped
with a pumping rate 0.58 mL/h, and the carrier gas is maintained at 408
mL/min. The
amount of catalyst that is loaded into the reactor is 53 mg.
The reaction process for 2MF is the same as the process for furan, except that
2%
propylene (1.986% propylene balanced by helium) rather than pure helium is
used as the
carrier gas. The reaction conditions for 2MF conversion are 600 C, WHSV of 5.7
h-1,
and partial pressure of 4.9 ton. 2MF is pumped with a pumping rate 0.57
mL/min, and
the flow of the carrier gas is maintained at 408 mL/min. The amount of
catalyst that is
loaded into the reactor is 92 mg.
The results are shown in Table 2.
Table 2. Summary of furan and 2MF conversions over ZSM, ZSM*, GaZSM, GaZSM*
and GaSD*
Feedstock 2MF 2MF 2MF 2MF Furan Furan Furan Furan
Furan
Carrier gas 2% 2% 2% 2% He He He He
He
propylene propylene propylene propylene
Catalyst ZSM ZSM* GaZSM GaZSM* ZSM ZSM* GaZSM GaZSM*
GaSD*
Temperature ( C) 600 600 600 600 550 550 550
550 550
Furan/2MF WHSV (h-1) 5.7 5.7 5.7 5.7 10.2 10.2 10.2
10.2 10.2
Pfuran/2MF (torr) 4.9 4.9 4.9 4.9 6.0 6.0 6.0
6.0 6.0
Olefins/furans molar 3.09 3.09 3.09 3.09
ratio
Furan/2MF conversion 99 89 98 74 33 32 41 24
24
(/o)
Propylene conversion (%) 31 22 31 21
Overall selectivity (%)
CO 5.9 6.6 6.0 9.7 13.1 9.3 13.9
9.5 10.6
CO2 0.1 0.5 0.7 0.5 1.6 3.4 2.0
1.6 2.7
Methane 0.0 0.7 1.0 0.4 0.0 0.0 0.0
0.0 0.0
Olefins 27.8 27.6 18.4 32.0 14.3 17.0 12.0
11.9 16.3
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Aromatics 59.6 53.3 68.5 32.4 37.4 39.0 42.7
47.7 49.3
Coke 6.2 9.4 4.7 13.6 28.5 23.3 23.9
22.8 16.3
Oxygenates 0.4 1.7 0.7 11.5 5.1 7.9 5.5
6.5 4.8
p-Xylene 5.1 14.9 6.6 4.6 1.0 1.9 0.9
1.2 3.7
Aromatic selectivity (%)
Benzene 24.4 20.8 34.4 32.7 18.0 17.7 26.7
24.9 19.2
Toluene 28.6 23.6 34.4 35.0 20.2 19.5 17.6
15.2 20.8
Xylene 26.9 30.3 16.5 14.7 5.1 5.6 3.5
2.7 8.5
Alkylbenzenesd 4.0 10.4 3.5 2.7 1.5 1.7 0.9
0.5 1.6
Styrenesb 2.6 4.8 5.0 4.3 7.8 7.7 8.6
5.9 7.3
Indenesc 10.2 7.9 3.6 3.5 23.7 19.7 23.5
10.1 14.6
Naphthalenesd 3.2 2.4 2.7 7.1 23.8 28.2 19.2
40.8 28.0
Olefin selectivity (%)
Ethylene 61.6 35.6 55.8 49.6 37.6 26.0 43.1
36.1 33.5
Propylene 37.8 44.5 43.6
34.0 39.7
C4 olefins 28.0 36.7 26.8 30.3 4.5 6.2 4.0
5.0 5.6
Allene 0.1 0.5 0.7 1.0 3.7 4.1 3.1
9.8 5.7
Cs olefins 7.3 16.3 14.2 15.0 10.8 11.2 4.7
11.0 9.4
C6 olefins 3.0 8.9 2.0 3.7 5.5 8.0 1.5
4.1 6.1
C7 olefins 0.0 2.0 0.5 0.3 0.0 0.0 0.0
0.0 0.0
Xylenes distribution (%)
p-Xylene 32 92 58 96 53 89 57 96
87
m-Xylene 49 6 34 3 38 9 35 3
10
o-Xylene 19 2 9 1 10 2 8 1
3
a: Ethylbenzene and trimethylbenzene
b: Styrene and methylstyrenes
c: Indene, methylindenes, and indane
d: Naphthalene, methylnaphthalene, and dihydronaphthalene
Figures A and B show overall p-xylene selectivity and xylenes distribution,
obtained from the conversions of 2MF and furan, respectively. In Figure A the
overall p-
xylene carbon selectivity obtained from ZSM is 5%. This value is increased to
15% by
using ZSM*. The silylation dramatically increases the para selectivity from
32% to 92%.
Similarly, the silylated GaZSM* also shows a significant increase ofpara
selectivity
from 58% (GaZSM) to 96%. However, the overall p-xylene selectivity for GaZSM*
is
lower than GaZSM. Table 2 shows that the conversion of 2MF for GaZSM* (74%) is
lower than that for GaZSM (98%). In addition, the 2MF conversion for ZSM*
(89%) is
also lower than that for ZSM (99%). The decrease of activity is believed to be
due to
some active sites in the external surface and in the surface near pore-
openings being
eliminated by silica deposition. Ga deposited on ZSM may have the ability to
increase
overall aromatics selectivity. This is also shown in Table 2 (2MF + propylene)
where
the aromatics selectivity is 60% for ZSM and 69% for GaZSM. However,
deactivation
caused by silylation on GaZSM* may be more severe than with ZSM*, according to
2MF conversion, suggesting that some active Ga species may be located at these
surfaces
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and may be killed by silylation. These Ga species imposes space confinement
that causes
an increase of para selectivity (Figure A, 58% for GaZSM and 32% for ZSM). The
silylation of GaZSM further imposes more space confinement and thus, gives a
better
para selectivity (96%).
For furan conversion (Figure B), increases of overall p-xylene selectivity and
amongst xylene species towards para are observed from silylated catalysts. The
para
selectivity is increased from 53% for ZSM to 89% for ZSM*, and from 57% for
GaZSM
to 96% for GaZSM*. Silylation also causes the activity to decrease as shown in
Table 1
where furan conversion is lower in silylated catalysts. The increase activity
in GaZSM
(41%) comparing with ZSM (33%) may be due to Ga species. However, the lowest
furan conversion observed on GaZSM* (24%) is again, due to the active sites
that are
killed by silica deposition. The significant decrease in overall xylene
selectivity for the
2MF + propylene reactions is believed to be due to the furan itself not being
a good
Diels-Alder reaction agent for xylene production. The results for furan
conversion using
GaSD* are shown in Table 2. This table shows that the para selectivity for
this catalyst
is 87% and suggests that silylation may be used for an FCC catalyst to
increase p-xylene
yield from biomass conversion in a fluidized-bed reactor. This is shown below
where
GaSD* is used for pinewood conversion in a bubbled fluidized-bed reactor.
Table 3
shows that an increase in para selectivity from 40% for GaSD to 72% for GaSD*.
Catalytic Conversion of Pinewood in a Fluidized-Bed Rreactor
CFP of pinewood is conducted in a fluidized bed reactor. The fluidized bed
reactor has a two-inch (5.08 cm) diameter, a height of ten inches (25.4 cm),
and is made
of 316 stainless steel. Inside the reactor, the catalyst bed is supported by a
distributor
plate made of stacked 316 stainless mesh (300 mesh). Solid pinewood is
introduced into
the reactor from a sealed feed-hopper. Test runs are conducted using GaSD and
GaSD*.
Prior to the test runs, the pinewood is ground and sieved to a particle sized
ranging
between 0.25 ¨ 1 mm. During the reaction, the catalyst is fluidized by helium
gas
flowing at 800 standard cubic centimeters (sccm) to enable the reactor to
operate in the
bubbling flow regime. The hopper and feed chamber are continuously purged with
helium at 200 sccm to maintain an inert environment. The total gas flow
through the
reactor is 1000 sccm helium. Both the reactor and the inlet gas stream are
heated to the
reaction temperature (550 C). The reactor is given two hours to reach this
temperature
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before the reaction is started. The effluent gas leaving the reactor flows
through a
cyclone to remove entrained particles. The effluent then flows into 7
condensers in series
to separate liquid and gas phase products. The first 3 condensers are placed
in an ice-
water bath with ethanol inside each condenser as a solvent, and the other 4
condensers
are surrounded by a dry ice and acetone bath (-55 C), without any solvent
inside the
condensers. Uncondensed gas phase products are collected in air bags at 5, 10,
20, and
30 minutes after the biomass first enters the reactor. The reaction time is 30
minutes.
After 30 minutes the reactor is purged using helium at a flow rate of 1000
sccm for 30
minutes to remove any CFP products other than coke on the catalyst. Liquid
products are
extracted from the condensers using ethanol. The catalyst is regenerated by
using air at
800 sccm for 3 hours in addition to the 200 sccm helium from the feed chamber
purge.
During regeneration, the effluent gases pass through a copper converter where
CO is
converted to CO2, and the CO2 is trapped by a CO2 trap. Gas phase products are
analyzed by a GC-FID/TCD (Shimadzu 2014). Liquid samples are analyzed by a GC-
FID (HP 7890). Coke yield is obtained by analyzing the weight change of the
CO2 trap.
The results are shown in Table 3.
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Table 3. Summary of pinewood conversions over GaSD and GaSD*.
Feed Stock SWP SWP SWP
Catalyst GaSD GaSD* GaSD*
TiC 550 550 550
WHSV (space velocity)/10 0.35 0.39 0.47
Overall Yield (carbon %)
Aromatics 23.2 14.8 13.3
Olefins 8.9 6.3 5.9
Methane 1.5 3.8 3.3
CO2 5.4 9.3 8.2
CO 17.2 22.9 19.6
Coke 33.3 33.4 30.0
Total 89.4 90.4 80.3
Aromatic Selectivity (%)
Benzene 22.0 27.6 25.9
Toluene 29.4 35.8 35.3
Xylenes 18.5 15.5 16.3
Naphthalene 14.6 8.9 4.4
Ethylbenzene 2.6 2.3 1.7
Styrene 2.2 1.1 1.6
Phenol 1.0 1.0 1.7
Benzofuran 1.3 1.8 2.2
Indene 1.2 2.7 3.7
Methylnaphthalene 4.6 1.7 3.1
Xylene distribution (%)
p-Xylene 40.0 70.7 72.4
m-Xylene 30.2 21.3 20.1
o-Xylene 29.8 7.97 7.50
Olefin Selectivity (%)
Ethylene 46.0 44.4 44.9
Propylene 47.1 46.3 44.3
Butylene 4.31 4.30 4.96
Butadiene 2.54 5.00 5.79
While the invention has been explained in relation to various embodiments, it
is
to be understood that various modifications thereof will become apparent to
those skilled
in the art upon reading the specification. Therefore, it is to be understood
that the
invention disclosed herein includes any such modifications that may fall
within the scope
of the appended claims.
