Note: Descriptions are shown in the official language in which they were submitted.
CA 02859935 2014-08-20
HIGH CONVERSION AND SELECTIVITY ODH PROCESS
FIELD OF THE INVENTION
The present invention relates to oxidative dehydrogenation of lower paraffins
in a
high conversion and high selectivity process. To date the art has emphasized
that
oxidative dehydrogenation reactions must be carried out in a reaction mixture
below the
lower oxidative combustion limits. As the reaction is "oxygen starved" the
conversion
per pass tends to be low. However at the other end of the spectrum one might
operate
above the upper oxidative combustion limits. Such a process has very short
dwell
times in the reactor, for both the feed stream and the catalyst, and provides
a once
through high conversion and high selectivity process. Preferably the reaction
is
conducted in an apparatus of the design for a fluidized bed catalyst cracker.
BACKGROUND OF THE INVENTION
The concept of oxidative dehydrogenation of paraffins to olefins (ODH) has
been
around since at least the late 1960's. Steam cracking of paraffins was a well
established technology and commercially practiced well prior to the 1960's.
The
perceived benefit of ODH is lower operating temperatures which in turn reduce
greenhouse gas emissions. The downside to ODH processes is the potential for a
decomposition (decomp). Industrial scale facilities are expensive and
corporations shy
away from processes which may result in a decomp. As a result ODH technology
has
had a difficult time gaining traction.
There are a number of United States patents assigned to Petro-Tex Chemical
Corporation issued in the late 1960's that disclose the use of various
ferrites in a steam
cracker to produce olefins from paraffins. The patents include United States
patents
3,420,911 and 3,420,912 in the names of Woskow et al. The patents teach
introducing
ferrites such as zinc, cadmium, and manganese ferrites (i.e. mixed oxides with
iron
oxide). The ferrites are introduced into a dehydrogenation zone at a
temperature from
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=
about 250 C up to about 750 C at pressures less than 100 psi (689.476 kPa) for
a time
less than 2 seconds, typically from 0.005 to 0.9 seconds. The reaction appears
to take
place in the presence of steam that may tend to shift the equilibrium in the
"wrong"
direction. Additionally the reaction does not take place in the presence of a
catalyst.
GB 1,213,181, which seems to correspond in part to the above Petro-Tex
patents, discloses that nickel ferrite may be used in the oxidative
dehydrogenation
process. The reaction conditions are comparable to those of above noted Petro-
Tex
patents.
United States patent 6,891,075 issued May 10, 2005 to Liu, assigned to Symyx
Technologies, Inc. teaches a catalyst for the oxidative dehydrogenation of a
paraffin
(alkane) such as ethane. The gaseous feedstock comprises at least the alkane
and
oxygen, but may also include diluents (such as argon, nitrogen, etc.) or other
components (such as water or carbon dioxide). The dehydrogenation catalyst
comprises at least about 2 weight % of NiO and a broad range of other elements
preferably Nb, Ta, and Co. While NiO is present in the catalyst it does not
appear to be
the source of the oxygen for the oxidative dehydrogenation of the alkane
(ethane).
United States patent 6,521,808 issued Feb. 18,2003 to Ozkan, et al, assigned
to the
Ohio State University teaches sol-gel supported catalysts for the oxidative
dehydrogenation of ethane to ethylene. The catalyst appears to be a mixed
metal
system such as Ni-Co-Mo, V-Nb-Mo possibly doped with small amounts of Li, Na,
K,
Rb, and Cs on a mixed silica oxide/titanium oxide support. Again the catalyst
does not
provide the oxygen for the oxidative dehydrogenation, rather gaseous oxygen is
included in the feed.
United States Patent no 4,450,313, issued May 22, 1984 to Eastman et al.,
assigned to Phillips Petroleum Company discloses a catalyst of the composition
Li20-
Ti02, which is characterized by a low ethane conversion not exceeding 10%, in
spite of
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a rather high selectivity to ethylene (92%). The major drawback of this
catalyst is the
high temperature of the process of oxidative dehydrogenation, which is close
to or
higher than 650 C.
The preparation of a supported catalyst usable for low temperature oxy-
dehydrogenation of ethane to ethylene is disclosed in the US Patent 4,596,787
A, 24
June, 1986 assigned to UNION CARBIDE CORP. A supported catalyst for the low
temperature gas phase oxydehydrogenation of ethane to ethylene is prepared by
(a)
preparing a precursor solution having soluble and insoluble portions of metal
compounds; (b) separating the soluble portion; (c) impregnating a catalyst
support with
the soluble portion and (d) activating the impregnated support to obtain the
catalyst.
The calcined catalyst has the composition MoaVbNbcSbaXe. X is nothing or Li,
Sc, Na,
Be, Mg, Ca, Sr, Ba, Ti, Zr, Hf, Y, Ta, Cr, Fe, Co, Ni, Ce, La, Zn, Cd, Hg, Al,
TI, Pb, As,
Bi, Te , U, Mn and/or W; a is 0.5-0.9, b is 0.1-0.4, c is 0.001-0.2, d is
0.001-0.1, e is
0.001-0.1 when X is an element. The patent fails to teach or suggest a co-
comminution
of the catalyst and the support.
[0016] Another example of the low temperature oxy-dehydrogenation of ethane to
ethylene using a calcined oxide catalyst containing molybdenum, vanadium,
niobium
and antimony is described in the US Patent 4,524,236 A, 18 June, 1985 and
4,250,346
A, 10 February, 1981, both assigned to UNION CARBIDE CORP. The calcined
catalyst contains MoaVbNbcSbdXe in the form of oxides. The catalyst is
prepared from a
solution of soluble compounds and/or complexes and/or compounds of each of the
metals. The dried catalyst is calcined by heating at 220-550 C in air or
oxygen. The
catalyst precursor solutions may be supported onto a support, e.g. silica,
aluminum
oxide, silicon carbide, zirconia, titania or mixtures of these. The
selectivity to ethylene
may be greater than 65% for a 50% conversion of ethane.
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The US patents numbers 6,624,116, issued Sept. 23, 2003 to Bharadwaj, et al.
and 6,566,573 issued May 20, 2003 to Bharadwaj, et al. both assigned to Dow
Global
Technologies Inc. disclose Pt-Sn-Sb-Cu-Ag monolith systems that have been
tested in
an autothermal regime at T>750 C, the starting gas mixture contained hydrogen
(H2:
02 = 2: 1, GHSV = 180 000 h-1). The catalyst composition is different from
that of the
present invention and the present invention does not contemplate the use of
molecular
hydrogen in the feed.
US Patents 4,524,236 issued June 18, 1985 to McCain assigned to Union
Carbide Corporation and 4,899,003, issued February 6, 1990 to Manyik et at,
assigned
to Union Carbide Chemicals and Plastics Company Inc. disclose mixed metal
oxide
catalysts of V-Mo-Nb-Sb. At 375-400 C the ethane conversion reached 70% with
the
selectivity close to 71-73%. However, these parameters were achieved only at
very low
gas hourly space velocities less than 900 h-1 (i.e. 720 h-1).
United States Patent 7,319,179 issued January 15,2008 to Lopez-Nieto et al.
assigned to Consejo Superior de Investigaciones Cientificas and Universidad
Politecnica de Valencia, discloses Mo-V-Te-Nb-0 oxide catalysts that provided
an
ethane conversion of 50-70% and selectivity to ethylene up to 95% (at 38%
conversion)
at 360-400 C. The catalysts have the empirical formula MoTehViNbjAk0,, where
A is a
fifth modifying element. The catalyst is a calcined mixed oxide (at least of
Mo, Te, V
and Nb), optionally supported on: (i) silica, alumina and/or titania,
preferably silica at
20-70 wt% of the total supported catalyst or (ii) silicon carbide. The
supported catalyst
is prepared by conventional methods of precipitation from solutions, drying
the
precipitate then calcining.
The preparation of a Mo-Te-V-Nb composition is described in WO 2005058498
Al, published 30 June, 2005 (corresponding to U.S. published application
2007149390A1). Preparation of the catalyst involves preparing a slurry by
combining
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an inert ceramic carrier with at least one solution comprising ionic species
of Mo, V, Te,
and Nb, drying the slurry to obtain a particulate product, precalcining the
dried product
at 150-350 C in an oxygen containing atmosphere and calcining the dried
product at
350-750 C under inert atmosphere. The catalyst prepared exhibits the activity
and
selectivity in the oxidation reaction comparable to the non-supported
catalyst.
A process for preparation of ethylene from gaseous feed comprising ethane and
oxygen involving contacting the feed with a mixed oxide catalyst containing
vanadium,
molybdenum, tantalum and tellurium in a reactor to form effluent of ethylene
is
disclosed in WO 2006130288 Al, 7 December, 2006, assigned to Celanese Int.
Corp.
The catalyst has a selectivity for ethylene of 50-80% thereby allowing
oxidation of
ethane to produce ethylene and acetic acid with high selectivity. The catalyst
has the
formula MoiV0.3Tao1Te0.30,. The catalyst is optionally supported on a support
selected
from porous silicon dioxide, ignited silicon dioxide, kieselguhr, silica gel,
porous and
nonporous aluminum oxide, titanium dioxide, zirconium dioxide, thorium
dioxide,
lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide,
cerium
dioxide, zinc oxide, boron oxide, boron nitride, boron carbide, boron
phosphate,
zirconium phosphate, aluminum silicate, silicon nitride, silicon carbide, and
glass,
carbon, carbon-fiber, activated carbon, metal-oxide or metal networks and
corresponding monoliths; or is encapsulated in a material (preferably silicon
dioxide
(Si02), phosphorus pentoxide (P205), magnesium oxide (MgO), chromium trioxide
(Cr203), titanium oxide (Ti02), zirconium oxide (Zr02) or alumina (A1203).
However, the
methods of preparation of the supported compositions involve the procedures of
wet
chemistry (solutions are impregnated into the solid support and then the
materials are
dried and calcined).
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, .
U.S. patent 5,202,517 issued April 13, 1993 to Minet et at., assigned to
Medalert
Incorporated teaches a ceramic tube for use in the conventional
dehydrogenation of
ethane to ethylene. The "tube" is a ceramic membrane, the ethane flows inside
the
tube and hydrogen diffuses out of the tube to improve the reaction kinetics.
The
reactive ceramic is 5 microns thick on a 1.5 to 2 mm thick support.
U.S. 6,818,189 issued Nov. 16, 2004 to Adris et al., assigned to SABIC teaches
in the passage bridging columns 9 and 10 a process in which ceramic pellets
are
packed around a tubular reactor and different reactants flow around the
outside and
inside of the tube. The patent is directed to the oxidative dehydrogenation of
ethane to
ethylene.
There is a significant amount of art on the separation of ethylene and ethane
using silver or copper ions in their + 1 oxidation state. See U.S. 6,518,476
at Col. 5,
lines 10-15 and Col. 16 line 12¨ Col. 17 line 57. NOVA Chemicals has also
disclosed
separation of olefins from non-olefins using ionic liquids (dithiolene in CA
2415064 (now
abandoned)). Also see United States patent 6,120,692 to Exxon; United States
patent
6,518,476 issued Feb 11, 2003 to Union Carbide at columns 16 and 17 the
abstract of
JP 59172428 published Sept 29, 1984 and the abstract of JP 59172427 published
Sept. 29, 1984.
United States patent 8,107,825 issued Sept 13, 2011 to Kuznicki et al.
assigned
to the University of Alberta contains a good outline of prior art for
separation of ethane
from ethylene and an adsorption method using ETS -10.
United States patent 7,411,107 issued Aug. 12, 2008 to Lucy et al., assigned
to
BP Chemicals Limited discloses a process for the separation of acetic acid
from an
oxidative dehydrogenation process to convert ethane to ethylene and acetic
acid. The
process uses a reversible complex of a metal salt (e.g. Cu or Ag) to separate
ethylene
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(Col. 8). The patent then discloses the acetic acid may be separated from the
liquids
by a distillation (Col. 13 lines 35 to 40).
United States Patent application 20110245571 in the name of NOVA Chemicals
(International) S.A. teaches oxidative dehydrogenation of ethane in a
fluidized bed in
contact with a bed of regenerative oxides to provide oxygen to the reactor. In
this
process "free" oxygen is not directly mixed with the feedstock reducing the
likelihood of
"decompositions".
United States patent 3,904,703 issued Sept. 9, 1975 to Lo et al., assigned to
El
Paso Products Company teaches a zoned or layered oxidative reactor in which
following a zone for oxidative dehydrogenation there is an "oxidation zone"
following a
dehydrogenation zone to oxidize hydrogen to water. Following the oxidation
zone there
is an adsorption bed to remove water from the reactants before they enter a
subsequent dehydrogenation zone. This is to reduce the impact of water on
downstream dehydrogenation catalysts.
U.S. patent application 2010/0256432 published Oct 7, 2010 in the name of
Arnold et al., assigned to Lummus discloses at paragraphs 86-94 methods to
remove
residual oxygen from the product stream. A combustible such as hydrogen or a
hydrocarbon may be added to the product stream to eliminate residual oxygen.
The
patent refers to a catalyst but does not disclose its composition. As noted
above it may
then be necessary to treat the product stream to eliminate water.
United States Patent Application 2004/0010174 (now abandoned) published Jan.
15, 2004 in the name of Wang et al., assigned to ConocoPhillips Company
discloses
using a circulating fluidized bed (CFB) reactor (similar in design to an FCC
reactor) to
conduct an oxidative dehydrogenation. The disclosure teaches at paragraph 40
the
catalyst acts to carry oxygen into the reactor as lattice oxygen or as
adsorbed oxygen.
The disclosure teaches away from adding air or oxygen to the feed stream.
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. ,
United States patent 8,519,210 issued Aug. 27, 2013 to Arnold et al., assigned
to Lummus Technology Inc. teaches that the concentration of oxygen in the feed
may
be limited to ,with a margin below, the minimum oxygen for combustion,
typically by
including steam or inert gases to dilute the feed to below flammability
limits.
The present invention seeks to provide a one pass process to oxidatively
dehydrogenate lower paraffins (alkanes, preferably n-alkanes) to produce alpha
olefins.
SUMMARY OF THE INVENTION
In one embodiment the present invention provides a process for the oxidative
dehydrogenation of one or more alkanes selected from the group consisting of
ethane
and propane and mixtures thereof in the presence of a supported catalyst
selected from
the group consisting of:
i) catalysts of the formula
VxMoyNID,TemMenOp
wherein Me is a metal selected from the group consisting of Ta, Ti, W, Hf, Zr,
Sb and
mixtures thereof; and
x is from 0.1 to 3 provided that when Me is absent x is greater than 0.5;
y is from 0.5 to 1.5;
z is from 0.001 to 3;
m is from 0.001 to 5;
n is from 0 to 2
and p is a number to satisfy the valence state of the mixed oxide catalyst
ii) catalysts of the formula
MoaVbNbcTee0d
wherein:
a is from 0.75 to 1.25, preferably from 0.90 to 1.10;
b is from 0.1 to 0.5, preferably from 0.25 to 0.4;
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c is from 0.1 to 0.5, preferably from 0.1 to 0.35;
e is from 0.1 to 0.35 preferably from 0.1 to 0.3, and
d is a number to satisfy the valence state of the mixed oxide catalyst on a
metal oxide
support; comprising:
a) passing through an oxidative dehydrogenation reactor containing a
fluidized bed
of said catalyst said one or more alkanes and oxygen at a temperature from 250
C to
450 C, a pressure from 3.447 to 689.47 kPa (0.5 to 100 psi) preferably, from
103.4 to
344.73 kPa (15 to 50 psi) and a residence time of said one or more alkanes in
said
reactor from 0.002 to 10 seconds, and reducing said catalyst, said catalyst
having an
average residence time in the dehydrogenation reactor of less than 30 seconds;
b) feeding said reduced catalyst to a regeneration reactor and passing
a stream of
air optionally with additional nitrogen at a temperature from 250 C to 400 C
and
pressures from 3.447 to 689.47 kPa (0.5 to 100 psi) [preferably, from 103.4 to
344.73
kPa (15 to 50 psi) through said bed to oxidize said catalyst; and
c) passing said oxidized catalyst back to said oxidative dehydrogenation
reactor
wherein the amount of oxygen in the feed to said reactor is above the upper
flammability limit for said feed. The conversion of alkane to alkene is not
less than 50%
per pass and the selectivity for the conversion of alkane to alkene is not
less than 0.9.
In the further embodiment the process comprises passing the product stream
through one or more oxygen scavenging reactors. Preferably, reactors area
operated in
parallel and one is being oxidized and another is being reduced to lower
oxidation state
of the metals in the catalyst.
In some embodiments oxygen scavenging reactors use the same catalyst used
in oxidative dehydrogenation reactors.
In a further embodiment the oxidative dehydrogenation reactor comprises a
riser
and the regeneration reactor is a separate fluidized bed reactor, said
regeneration
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reactor being connected with said riser to flow oxidized catalyst back to said
riser (e.g.
CFB [circulating fluidized bed] type reactor).
In a further embodiment the top of said riser comprises a distributor system
to
improve temperature control in the reactor [to minimize combustion of the
alkane feed
and] to maintain the overall selectivity of the reactor above 90%.
A further embodiment comprises passing low temperature steam into said
catalyst flow into said riser to cool the catalyst to control the heat balance
of the
oxidative dehydrogenation reactor.
In a further embodiment there is a downcomer between said oxidative
dehydrogenation reactor and said regeneration reactor to flow reduced catalyst
from
said oxidative dehydrogenation reactor to said regeneration reactor.
A further embodiment comprises passing low temperature steam [counter
current to the flow of catalyst through said downcomer] to strip entrained
alkane feed
and product.
A further embodiment comprises passing air or a mixture of air and nitrogen
through the regeneration reactor in an amount to substantially extract the
oxygen from
the air or a mixture of air and nitrogen and generating a gas product stream
comprising
not less than 85 vol.-`)/0 of nitrogen.
A further embodiment comprises recycling a portion of the oxygen reduced
effluent stream from the regenerator reactor and optionally cooling it and
recycling it to
the regenerator reactor.
A further embodiment comprises separating said alkene product from the
oxidative dehydrogenation reactor from water in the product stream from the
oxidative
dehydrogenation unit.
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= ,
A further embodiment comprises passing unused nitrogen from the effluent
stream from the catalyst regeneration reactor to a site integrated unit
operation using
nitrogen as a part of the feedstock.
In a further embodiment two or more fixed bed reactors are used as scavengers
having piping and valves so that the feed to the fluidized bed oxidative
dehydrogenation
reactor passes through one or more of the fixed bed reactors having an
oxidative
dehydrogenation which is oxidized to depleted the catalyst of oxygen, and
passing the
product stream through one or more of the fixed bed reactors having an
oxidative
dehydrogenation catalyst depleted of oxygen, to remove residual oxygen from
the
product by reaction and switching the flow of product stream to reactors to
oxygen
depleted reactors and the flow of feed stream to oxygen rich reactors.
In a further embodiment the site integrated unit operation is selected from an
ammonia plant and an acrylonitrile plant, urea plant and, an ammonium nitrate
plant.
In a further embodiment the residence time of the catalyst in the oxidative
dehydrogenation reactor is less than 30 seconds (preferably less than 10 more
desirable less than 5 seconds).
In a further embodiment the residence time of the catalyst in the regeneration
reactor is less than 3 minutes.
In a further embodiment the ratio of residence time of the catalyst in the
regenerator to the residence time of the catalyst in the oxidative
dehydrogenation
catalyst is not less than 3.
In a further embodiment the product stream from the oxidative dehydrogenation
reactor and at least a portion of the effluent stream from the regenerator
reactor are
passed through separate steam generators to recover heat.
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In a further embodiment the product stream from the oxidative dehydrogenation
reactor is cooled and passed through a column to separate combustion products
from
alkene.
In a further embodiment the product stream from the oxidative dehydrogenation
reactor is cooled and passed through an amine unit to remove CO2.
In a further embodiment the support is selected from the group consisting of
silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide,
zirconium
dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide,
barium
oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, yttrium oxide.
In a further embodiment the alkane is ethane.
In a further embodiment the catalyst is of the formula
MoaVbNbcTee0d
wherein:
a is from 0.90 to 1.10;
b is from 0.25 to 0.4;
c is from 0.1 to 0.3;
e is from 0.1 to 0.3 , and
d is a number to satisfy the valence state of the mixed oxide catalyst on a
metal oxide
support.
In a further embodiment in the oxidative dehydrogenation reactor the
conversion
to ethylene is greater than 60%.
In a further embodiment in the oxidative dehydrogenation reactor the
selectivity
to ethylene is greater than 75%.
BRIEF DESCRIPTION OF THE DRAWINGS
Figure 1 is a schematic drawing of a CFB reactor useful in accordance with the
present invention.
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Figures 2 a and b shows the conversion and selectivity of a feed stream
comprising ethylene and 25% mole.-% of oxygen at a temperature of 355 C in
the
presence of a catalyst in accordance with the present invention over a time up
to 60
seconds.
Figure 3 a and b shows the conversion and selectivity of a feed stream
comprising ethylene and 25% mole. -%of oxygen at a temperature of 355 C in
the
presence of a catalyst in accordance with the present invention over a time up
to 60
seconds.
Figure 4 is a plot showing time dependence of the ethane and 02 conversion (a)
and selectivity of ethylene formation (b) after the gas flow switch [air to
gas mixture] on
the Mo-V-Te-Nb-Ox catalyst at 398 C. [2400 h-1] Dotted lines correspond to the
equilibrium values.
Figure 5 is a plot showing time dependence of the ethane and 02 conversion (a)
and selectivity of ethylene formation (b) after the gas flow switch [air to
gas mixture] on
the Mo-V-Te-Nb- Ox catalyst at 398 C. [600 1-11] Dotted lines correspond to
the
equilibrium values.
Figure 6 is a plot showing time dependence of the selectivity of ethylene
formation after the gas flow switch [air to gas mixture] on the Mo-V-Te-Nb-Ox
catalyst
at 398 C at different flow rates.
Figure 7 is a plot showing dependence of the ethane conversion on the amount
of ethane supplied into the reactor at different rates after the gas flow
switch [air to gas
mixture] on the Mo-V-Te-Nb-Ox catalyst at 398 C.
Figure 8 is a plot showing time dependence of ethane conversion and 02
residual content (a) and selectivity of ethylene formation (b) after the gas
flow switch
[02 to C2H6] on the Mo-V-Te-Nb-Ox catalyst at 355 C.
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Figure 9 is a plot showing time dependence of ethane conversion and 02
residual content (a) and selectivity of ethylene formation (b) after the gas
flow switch
[02 to C2H6] on the Mo-V-Te-Nb-Ox catalyst at 397 C.
Figures 10, 11, and 12 illustrate how a series of three fixed bed catalysts
may be
used to scavenge oxygen from the product stream in an oxidative
dehydrogenation
reactor.
DETAILED DESCRIPTION
Numbers ranges
Other than in the operating examples or where otherwise indicated, all numbers
or expressions referring to quantities of ingredients, reaction conditions,
etc. used in the
specification and claims are to be understood as modified in all instances by
the term
"about." Accordingly, unless indicated to the contrary, the numerical
parameters set
forth in the following specification and attached claims are approximations
that can vary
depending upon the properties that the present invention desires to obtain. At
the very
least, and not as an attempt to limit the application of the doctrine of
equivalents to the
scope of the claims, each numerical parameter should at least be construed in
light of
the number of reported significant digits and by applying ordinary rounding
techniques.
Notwithstanding that the numerical ranges and parameters setting forth the
broad scope of the invention are approximations, the numerical values set
forth in the
specific examples are reported as precisely as possible. Any numerical values,
however, inherently contain certain errors necessarily resulting from the
standard
deviation found in their respective testing measurements.
Also, it should be understood that any numerical range recited herein is
intended
to include all sub-ranges subsumed therein. For example, a range of "1 to 10"
is
intended to include all sub-ranges between and including the recited minimum
value of
1 and the recited maximum value of 10; that is, having a minimum value equal
to or
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greater than 1 and a maximum value of equal to or less than 10. Because the
disclosed numerical ranges are continuous, they include every value between
the
minimum and maximum values. Unless expressly indicated otherwise, the various
numerical ranges specified in this application are approximations.
All compositional ranges expressed herein are limited in total to and do not
exceed 100 percent (volume percent or weight percent) in practice. Where
multiple
components can be present in a composition, the sum of the maximum amounts of
each component can exceed 100 percent, with the understanding that, and as
those
skilled in the art readily understand, the amounts of the components actually
used will
conform to the maximum of 100 percent.
Catalysts:
There are a number of catalysts which may be used in accordance with the
present invention. The following catalyst systems may be used individually or
in
combination. One of ordinary skill in the art would understand that
combinations should
be tested at a laboratory scale to determine if there are any antagonistic
effects when
catalyst combinations are used.
The oxidative dehydrogenation catalyst of the present invention may be
selected
from the group consisting of:
i) catalysts of the formula:
NixAaBbDdOe
wherein
x is a number from 0.1 to 0.9 preferably from 0.3 to 0.9, most preferably from
0.5 to
0.85, most preferably 0.6 to 0.8;
a is a number from 0.04 to 0.9;
b is a number from 0 to 0.5;
d is a number from 0 to 0Ø5;
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=
e is a number to satisfy the valence state of the catalyst;
A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Y, Zn, Zr, Si
and Al or
mixtures thereof;
B is selected from the group consisting of La, Ce, Pr, Nd, Sm, Sb, Sn, Bi, Pb,
TI, In, Te,
Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir, Au, Hg and mixtures
thereof;
D is selected from the group consisting of Ca, K, Mg, Li, Na, Sr, Ba, Cs, and
Rb and
mixtures thereof; and
0 is oxygen; and
ii) catalysts of the formula:
MofXgYh
wherein
X is selected from the group consisting of Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V,
W and
mixtures thereof;
Y is selected from the group consisting of Bi, Ce, Co, Cu, Fe, K, Mg V, Ni, P,
Pb, Sb, Si,
Sn, Ti, U and mixtures thereof;
f= 1;
g is 0 to 2;
h is 0 to 2, with the proviso that the total value of h for Co, Ni, Fe and
mixtures thereof
is less than 0.5;
and catalysts of formula iii) below,
and mixtures thereof.
In one embodiment the catalyst is the catalyst of formula i) wherein x is from
0.5
to 0.85, a is from 0.15 to 0.5, b is from 0 to 0.1 and d is from 0 to 0.1. In
catalyst i)
typically A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Zr,
Si, Al and
mixtures thereof, B is selected from the group consisting of La, Ce, Nd, Sb,
Sn, Bi, Pb,
Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir and mixtures thereof
and D is
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selected from the group consisting of Ca, K, Mg, Li, Na, Ba, Cs, Rb and
mixtures
thereof.
In an alternative embodiment the catalyst is catalyst ii). In some embodiments
of
this aspect of the invention typically X is selected from the group consisting
of Ba, Ca,
Cr, Mn, Nb, Ti, Te, V, W and mixtures thereof, Y is selected from the group
consisting
of Bi, Ce, Co, Cu, Fe, K, Mg V, Ni, P, Pb, Sb, Sn, Ti and mixtures thereof.
= One additional particularly useful family of catalysts iii) comprise one
or more
catalysts selected from the group consisting of a mixed oxide catalyst of the
formula
VxMoyN13,TemMen0p,
wherein Me is a metal selected from the group consisting of Ti, Ta, Sb, Hf, W,
Y, Zn, Zr,
La, Ce, Pr, Nd, Sm, Sn, Bi, Pb Cr, Mn, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os,
Ir, Au,
and mixtures thereof; and
x is from 0.1 to 3, preferably from 0.5 to 2.0 most preferably from 0.75 to
1.5 and
provided that when Me is absent x is greater than 0.5;
y is from 0.5 to 1.5, preferably from 0.75 to 1.0;
z is from 0.001 to 3, preferably from 0.1 to 2, most preferably from 0.5 to
1.5.
m is from 0.001 to 5, preferably from 1 to 4.
n is from 0 to 2, preferably n is 0, however when Me is present n is
preferably from 0.5;
to 1.5 and
p is a number to satisfy the valence state of the mixed oxide catalyst.
In one embodiment the catalyst has the formula:
MoaVbNbcTee0d
wherein:
a is from 0.90 to 1.10, preferably 0.95 to 1.1;
b is from 0.25 to 0.4, preferably 0.3 to 0.35;
c is from 0.1 to 0.3, preferably 0.1 to 0.15;
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e is from 0.1 to 0.3, preferably 0.1 to 0.25; and
d is a number to satisfy the valence state of the mixed oxide catalyst on a
metal oxide
support.
In a further embodiment in the catalyst the ratio of x:m is from 0.3 to 10,
most
preferably from 0.5 to 8, desirably from 0.5 to 6.
The methods of preparing the catalysts are known to those skilled in the art.
For example, the catalyst may be prepared by mixing aqueous solutions of
soluble metal compounds such as hydroxides, sulphates, nitrates, halides,
salts of
lower (C1_5) mono or di carboxylic acids and ammonium salts or the metal acid
per se.
For instance, the catalyst could be prepared by blending solutions such as
ammonium
metavanadate, niobium oxalate, ammonium molybdate, telluric acid etc. The
resulting
solution is then dried typically in air at 100-150 C and calcined in a flow of
inert gas
such as those selected from the group consisting of N2, He, Ar, Ne and
mixtures thereof
at 200-600 C, preferably at 300-500 C. The calcining step may take from 1 to
20,
typically from 5 to 15 usually about 10 hours. The resulting oxide is a
friable solid
typically insoluble in water.
The Support / Binder
There are several ways the oxidative dehydrogenation catalyst may be
supported or bound.
Preferred components for forming ceramic supports and for binders include
oxides of titanium, zirconium, aluminum, magnesium, silicon, phosphates, boron
phosphate, zirconium phosphate and mixtures thereof, for both fluidized and
fixed bed
reactors. In the fluidized bed typically catalyst is generally spray dried
with the binder,
typically forming spherical particles ranging in size (effective diameter)
from 40-100 um.
However, one needs to be careful to insure that particles area sufficiently
robust to
minimize the attrition in the fluidized bed.
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The support for the catalyst for the fixed bed may further be ceramic
precursor
formed from oxides, dioxides, nitrides, carbides selected from the group
consisting of
silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide,
zirconium
dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide,
barium
oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, boron nitride,
boron carbide,
yttrium oxide, aluminum silicate, silicon nitride, silicon carbide and
mixtures thereof.
In one embodiment the support for the fixed bed may have a low surface area
less than 20 m2/g, alternatively, less than 15 m2/g, alternatively, less than
3.0 m2/g for
the oxidative dehydrogenation catalyst. Such support may be prepared by
compression molding. At higher pressures the interstices within the ceramic
precursor
being compressed collapse. Depending on the pressure exerted on the support
precursor the surface area of the support may be from about 20 to 10 m2/g.
The low surface area support could be of any conventional shape such as
spheres, rings, saddles etc.
In the present invention the oxidized catalyst in the fluidized bed contains
one or
more of lattice oxygen and adsorbed oxygen. The supported catalyst together
with
added air or preferably oxygen passes together through a oxidative
dehydrogenation
reactor, and the catalyst is reduced when ethane is converted to ethylene.
Then the
supported catalyst is passed by the downcomer to the regeneration reactor
where it is
oxidized.
It is important that the support be dried prior to use (i.e. before adding
catalyst).
Generally, the support may be heated at a temperature of at least 200 C for
up to 24
hours, typically at a temperature from 500 C to 800 C for about 2 to 20
hours,
preferably 4 to 10 hours. The resulting support will be free of adsorbed water
and
should have a surface hydroxyl content from about 0.1 to 5 mmol/g of support,
preferably from 0.5 to 3 mmol/g.
HATrevor\TTSpec\2014007Canada.docx
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' .
The amount of the hydroxyl groups on silica may be determined according to the
method disclosed by J. B. Pen i and A. L. Hensley, Jr., in J. Phys. Chem., 72
(8), 2926,
1968, the entire contents of which are incorporated herein by reference.
The dried support for the fixed bed catalyst may then be compressed into the
required shape by compression molding. Depending on the particle size of the
support,
it may be combined with an inert binder to hold the shape of the compressed
part. .
Loadings
Typically the catalyst loading on the support for the fixed bed catalyst
provides
from 1 to 30 weight % typically from 5 to 20 weight %, preferably from 8 to 15
weight %
of said catalyst and from 99 to 70 weight %, typically from 80 to 95 weight %,
preferably
from 85 to 92 weight /0, respectively, of said support.
The catalyst may be added to the support in any number of ways. For example
the catalyst could be deposited from an aqueous slurry onto one of the
surfaces of the
low surface area support by impregnation, wash-coating, brushing or spraying.
The
catalyst could also be co-precipitated from a slurry with the ceramic
precursor (e.g.
alumina) to form the low surface area supported catalyst.
The catalyst loading for the fluidized bed may be chosen based on a number of
factors including the volume of bed, the flow rate of alkane through the bed,
energy
balance in the bed, binder type, etc. For the fluidized bed catalyst loading
may cover a
wide range of values ranging from 10 wt.-% up to 90 wt.-%, typically above 20
wt.-%
desirably above 35 wt.-%.
In the present invention the feed to the oxidative dehydrogenation reactor
includes oxygen in an amount above the upper explosive/ignition limit. For
example for
ethane oxidative dehydrogenation, typically the oxygen will be present in an
amount of
not less than about 5 mole-% preferably about 18 mole-%, for example from
about 22
to 27mole-%,or 23 to 26 mole-% . It is desirable not to have too great an
excess of
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oxygen as this may reduce selectivity arising from combustion of feed or final
products.
Additionally too high an excess of oxygen in the feed stream may require
additional
separation steps at the downstream end of the reaction.
The process of the present invention will be described in conjunction with
figure
1 which schematically illustrates a circulating fluidized bed reactor.
In one embodiment the reactor system 1 comprises a fluidized bed oxidative
dehydrogenation reactor 3 and a regenerator reactor 4. The fluidized bed
oxidative
dehydrogenation reactor riser 2 and the regeneration reactor 4 are joined by a
downcomer 10 which conducts clean oxidized supported catalyst from the
regenerator
reactor 4 to the oxidative dehydrogenation reactor riser 2. Each of the
fluidized bed
oxidative dehydrogenation reactor riser 2, the fluidized bed 6 in the
dehydrogenation
reactor 3 and the regeneration reactor 4 contain fluidized bed of catalyst
particles 5, 6
and 7 respectively. In the oxidative dehydrogenation reactor riser 2 and the
regeneration reactor 4 above fluidized catalyst beds 5 and 7 are disengagement
zones
8 and 9, respectively.
The inlet 11 to downcommer 10 is attached to the regenerator reactor 4
generally at a point between about 1/3 to 2/3 the height of the fluidized bed
7. The
downcommer 10 enters the bottom of the oxidative dehydrogenation reactor riser
2.
The reactor riser 2 extends up into the dehydrogenation reactor 3 above the
fluid level
of the fluidized bed 6 (typically 1/3 to 2/3 of height). The reactor riser 2,
flares to form an
inverted cone disperser 12 to provide a disengagement zone for catalyst from
the
product. Optionally, a disperser plate 13 may be used above the cone 12. The
disperser may have a shape other than an inverted cone; however, care must be
taken
to ensure a substantially uniform gas flow around the disperser.
The oxidative dehydrogenation reactor operates at temperatures below 450 C
typically from 350 to 450 , pressures from 3.447 to 689.47 kPa (0.5 to 100
psi)
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preferably, from 103.4 to 344.73 kPa (15 to 50 psi) and a residence time of
the one or
more alkanes in the oxidative dehydrogenation reactor riser 2 from 0.002 to 20
seconds.
Flared section 12 of the riser should be sufficiently broad to cause the
catalyst
particles to drop in the fluidized bed zone 6, the disperser plate 13 should
be high
enough to minimize catalyst attrition.
Port 14 in the riser 10 permits the introduction of low temperature steam, at
a
temperature at least about 25 C desirably 50 C lower than the temperature of
the
oxidative dehydrogenation reactor. In some embodiments steam has a temperature
from about 200 C to about 400 C, in further embodiments the temperature may
be
from about 300 C to 350 C. The steam cools the catalyst coming from the
regenerator reactor 4 and also removes any entrained or absorbed impurities
(e.g.
ethylene or air).
Port 15 at or towards the base of the oxidative dehydrogenation reactor riser
2 is
an inlet for the hydrocarbon feed typically high purity ethane mixed with
oxygen or an
oxygen containing gas. The hydrocarbon feed and oxygen could be combined
proximate and upstream of the oxidative dehydrogenation reactor. As this is a
fluidized
bed reactor it is necessary that the upward flow of hydrocarbon feed and
oxygen
containing gas be sufficiently well distributed to fluidize the bed of
catalyst particles to
minimize hot spots.
The process of the present invention may be used to generate ethylene from
relatively pure feedstock.
The ethane individually should comprise about 95 wt.% of ethane, preferably
98 wt.% of ethane and not more than about 5 wt.% of associated hydrocarbons
such as
methane. Preferably the feed is oxygen having a relatively high purity in some
embodiments above 90% pure, in further embodiments greater than 95 % pure.
While
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air may be used as a source for oxygen it may give rise to downstream
separation
issues.
In the further embodiment of the invention the reactor of the present
invention
may be used to replace and ethane / ethylene splitter or off-gas from refinery
or other
hydrocarbon processing process in which case the feedstock can comprise from
10 -
80 vol.-% ethylene and balance ethane.
To maintain a viable fluidized bed, the mass gas flow rate through the bed
must
be above the minimum flow required for fluidization, and preferably from about
1.5 to
about 10 times Umf and more preferably from about 2 to about 6 times Umf. Umf
is used
in the accepted form as the abbreviation for the minimum mass gas flow
required to
achieve fluidization, C. Y. Wen and Y. H. Yu, "Mechanics of Fluidization",
Chemical
Engineering Progress Symposium Series, Vol. 62, p. 100-111 (1966). Typically
the
superficial gas velocity required ranges from 0.3 to 5 m/s.
At the upper end of the oxidative dehydrogenation reactor, below disengagement
zone 9 is port 16 which permits the spent catalyst stream to settle and leave
the
reactor. At the top of the reactor 3, there are cyclones 17 to remove any
catalyst fines,
which were not settled in disengagement zone 9.
The average residence time of the supported catalyst in the oxidative
dehydrogenation reactor riser 2 is less than about 30 seconds in some cases
less than
15 seconds in some cases from 1 to 6 seconds. The port 16 connects downcomer
18
with the oxidative dehydrogenation reactor 3 and the regeneration reactor 4.
Port 19 in
the downcomer 18 is positioned proximate the regeneration reactor 4. Port 19
allows
the introduction of steam at a temperature from about 300 C to 500 C, in
some
embodiments from 350 C to 450 C to flow counter current to the stream of
spent
catalyst to remove entrained feedstock and product. In some cases the steam
may
also burn of surface coke on the catalyst particles. The flow rate of the
steam in the
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downcomer should be sufficiently low to prevent the supported catalyst from
being
pushed back into disengagement zone 9.
The regeneration reactor is also a fluidized bed reactor. Port 20 at the
bottom of
the regeneration reactor permits air and in some cases recycled cooled
nitrogen back
into the reactor. The regeneration reactor is typically operated at
temperatures from
250 C to 400 C and pressures from 3.447 to 689.47 kPa (0.5 to 100 psi),
preferably,
from 103.4 to 344.73 kPa (15 to 50 psi). The residence time of the supported
catalyst
in the regeneration reactor is less than 3 minutes. Typically the ratio of the
residence
time of the catalyst in the regenerator reactor to the residence time in the
oxidative
dehydrogenation reactor is not less than 3.
Port 21 on the upper portion of the regenerator reactor 4, above the fluidized
bed
of supported catalyst particles permits the off gas to leave the reactor.
There may be
cyclones as described for the oxidative dehydrogenation reactor 3 in the upper
section
of the regenerator reactor 4 to remove any catalyst fines from nitrogen
product stream.
Air and optionally nitrogen which may be cooled are passed through the
regeneration
reactor. The oxygen is substantially taken from the air. The off gas will
comprise from
85 to 100% of nitrogen.
The above description of the circulating fluidized bed reactor has been
largely
schematic. There may be various valves, filters, etc.at the ports. The
selection of
appropriate valves would be well known to those skilled in the art. Similarly
there may
be suitable fans and compression means used to force gases through the system.
The
selection of appropriate fans or compressors or expanders for cooling would be
known
to one of ordinary skill in the art.
It is desirable to recover as much energy as possible from the oxidative
dehydrogenation reaction and the regeneration reaction. The ethylene feed and
the
co-products (e.g. CO2 and CD) from the oxidative reactor are fed to separate
steam
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CA 02859935 2014-08-20
generators to generate steam. Part of the steam may be recycled back to the
process.
The steam could be injected in the riser to cool the catalyst particles. The
steam could
also be injected into the downcomer to burn off any coke and to entrain any
absorbed
or adsorbed feed or products.
The oxygen containing stream passing through the regenerator is substantially
depleted of oxygen on exit from the reactor (e.g. the exit stream comprises
not less
than about 90% of nitrogen). If nitrogen is also used as a component of the
feed
stream a part of the product stream may be recycled to the inlet for the
regeneration
reactor. The portion of the product stream from the regeneration reactor may
be
subject to one or more cooling or refrigeration steps to maintain the heat
balance in the
regenerator.
The process of the present invention should be operated to have a conversion
of
not less than 80% (to ethylene) and a selectivity of not less than 90%,
preferably
greater than 95% to ethylene.
Separation of Product Streams.
The stream 22 exiting the dehydrogenation reaction comprises ethylene, water
(vapour ¨ steam) and a small amounts of ethane, unconsumed oxygen and off
gases
typically CO and CO2 The issue of separation needs to be considered in the
context of
the intended use for the ethylene.
There are a number of processes which may use dilute ethylene such as
polymerization process. However, this approach needs to be balanced with the
effect of
polar molecules such as CO and CO2 and oxygen on the catalyst used for the
polymerization. It may be preferable to separate the polar molecules prior to
separation
of ethylene and ethane. The polar molecules may be separated by an adsorption
bed
such as a zeolite bed. In the simplest embodiment, depending on the ratios of
the
components the bed could be regenerated and all the components fed to a burner
to
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, .
burn the CO. However, at a chemical complex there are other unit operations
which
could use CO as a feed (various carboxylic acid and anhydride processes
(acetic acid,
meth acrylic acid and maleic anhydride). If there is a significant amount of
CO and
CO2 the components could be separated. There are a number of well-known
methods
in the art to separate CO2 and CO. The stream would be cooled and washed and
then
passed through an adsorber such as activated carbon (to remove impurities from
the
002) or a liquid amine separator or a liquid carbonate separator to absorb the
CO2. CO
could be separated by a number of techniques. Depending on the volume a vacuum
separation method using activated carbon as an adsorbent may be suitable, a
membrane separation may be suitable and adsorption on copper ions (on a
suitable
support) may be suitable.
Oxygen removal ¨ fixed bed
In one embodiment there may be two or more fixed bed reactors having an
oxidative dehydrogenation catalyst which releases or takes up oxygen are used
as
scavengers to accommodate the product flow out of the circulating fluidized
bed
oxidative dehydrogenation reactor. The fixed bed reactors have piping and
valves so
that the feed to the fluidized bed oxidative dehydrogenation reactor passes
through one
or more of the fixed bed reactors having a catalyst containing oxygen which is
consumed or given up. This is not so much of an issue with the pre-reactor
operating
in oxidative dehydrogenation mode since any excess alkane not dehydrogenated
in the
pre-reactor will be converted in the fluidized bed oxidative dehydrogenation
reactor.
The key issue is depleting the catalyst in the fixed bed reactor of oxygen.
The piping
and valves flow the product stream through one or more of the fixed bed
reactors
having oxidative dehydrogenation catalysts which are depleted of oxygen. The
depleted fixed bed catalyst scavenges oxygen from the product stream. As noted
the
valves and piping of the streams can be operated so that feed streams flow
through the
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oxygenated fixed bed catalyst reactor and the product stream flows through one
or
more of the oxygen depleted fixed beds catalyst reactors.
The oxidative dehydrogenation catalyst containing oxygen may have oxygen as
lattice oxygen, adsorbed oxygen or adsorbed oxygen on the catalyst, the
support or
both. The oxidative dehydrogenation catalyst depleted of oxygen has a reduced,
preferably, about 60 % less oxygen in the catalyst and support as lattice
oxygen,
adsorbed oxygen or adsorbed oxygen on the catalyst, the support or both.
Preferably, at the exit of the fluidized bed oxidative dehydrogenation reactor
is
an oxygen sensor. Additionally, there should be an oxygen sensor at the exit
for the
dehydrogenated product from each fixed bed reactor to determine the oxygen
level
leaving the product leaving that fixed bed reactor. When the oxygen level
rises at the
dehydrogenated product outlet of the fixed bed reactor operating in scavenger
mode it
indicates the catalyst have substantially taken up reactive oxygen (and may be
returned
to use as a pre-reactor). The amount of reactive oxygen uptake by the oxygen
depleted catalyst in the pre-reactor operation in oxygen scavenging or
chemisorption
mode should be not less than about 1.5%, typically about 2% of the total
oxygen in the
catalyst (this will also correspond to the amount of reactive oxygen available
for release
from the catalyst in the pre-reactors in oxidative dehydrogenation mode).
One mode for operation using three pre-reactors is illustrated schematically
in
Figures 10, 11, and 12 (in which like parts have like numbers) and the table
below. In
Figures 10, 11, and 12 the valves are not shown. The main reactor
configuration is the
same however the switching of the valves causes the pre-reactor, scavenger
reactor
and the guard reactor to appear to "switch" places. One pre-reactor operates
as such
and converts part of the feed stream to ethylene. One oxygen depleted pre-
reactor
acts as a primary oxygen scavenger or chemisorption reactor and a second pre-
reactor
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(also oxygen depleted acts as a guard or secondary oxygen scavenger or
chemisorption reactor).
Operation
Process Step Process streams flow sequence
Step 1 (Figure 10): Ethane (50) is routed to a fixed bed reactor (51)
(preferably
oxygen saturated). Some of the ethane is converted to ethylene
and the product together with oxygen (52) is routed to the fluidized
bed oxidative dehydrogenation reactor (53), where most or all
ethane is converted to ethylene. The product is cooled in a
condenser (54) to a temperature from 50 C to 270 C, and
optionally water is knocked out of the product stream in knock out
drum (55) (adsorbed by one or more guard beds). The cooled
product stream is routed to a primary scavenging oxygen depleted
fixed bed reactor which acts as a lead oxygen scavenger reactor
(56). Oxygen scavenging / chemisorption is exothermic, the
product stream from the primary oxygen scavenging reactor may
be cooled in a condenser (57) and routed through water knock out
drum (58) to the secondary or guard oxygen scavenger fixed bed
reactor (59) (oxygen depleted pre-reactor) - (cooling down may
not be required, since the only reason for cooling is to reduce any
oxidation reaction of the final product (60) (e.g. production of CO
and CO2 or both), in the secondary or guard oxygen depleted pre-
reactor initially there is a very low level of reactive oxygen
(typically less 50, preferably less than 25, most preferably less
than 10 ppm of reactive oxygen in the feed stream) will be
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=
present; A slightly elevated temperature (2 C to 5 C higher) will
help to remove it to very low level without converting the product
to CO and CO2. Oxygen sensors, not shown, are active on inlets
to the lead (primary) and fixed bed oxygen scavenging guard
reactor (secondary oxygen scavengers and the outlet of the guard
scavenger. The operation is to go to step 2 when the oxygen
content in the product stream exiting the guard reactor exceeds
specified value.
Step 2: (Figure 11) Changes from Step 1 (Figure 10): The former fixed bed
oxidative
dehydrogenation reactor (pre-reactor) (51) now becomes guard
scavenger (59); former guard scavenger (59) now becomes lead
scavenger (56), former lead scavenger (56) becomes pre-reactor
(51). Operation is the same as described for the Step 1.
Step 3: (Figure 12) Changes from Step 2 (figure 11): Fixed bed oxidative
dehydrogenation reactor ( Pre-reactor (51) becomes guard
scavenger reactor (59); former guard scavenger becomes lead
scavenger reactor (56,) and former lead scavenger becomes the
pre-reactor (51). Operation is the same as described for the Step
1.
Step 4: (Figure 10) Return to Step 1.
In an alternate embodiment the oxygen may be separated from the product
stream using cryogenic methods. However, this adds both capital and operating
costs
to the process.
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The above scavenging process is more fully described in Canadian patent
application 2,833,822, filed Nov. 21, 2013 the text of which is herein
incorporated by
reference.
Residual gases from the downcomer would also be subject to the same
separation techniques to recover them.
As noted above it may not be necessary to separate the ethane from the
ethylene at this stage however, if desired there are a number of techniques
that may be
used.
The most common techniques would be to use a cryogenic C2 splitter.
Other separation techniques in the following.
One method of separation of the product stream is by absorption. The gaseous
product stream comprising primarily ethane and ethylene may be contacted in a
counter current flow with a heavier paraffinic oil such as mineral seal oil or
medicinal
white oil at a pressure up to 800 psi ( about 5.5X103 kPa) and at temperatures
from
about 25 F to 125 F ( about -4 C to about 52 C). The ethylene and lower
boiling
components are not absorbed into the oil. The ethane and higher boiling
components
are absorbed into the oil. The ethylene and lower boiling components may then
be
passed to the C2 splitter if required. The absorption oil may be selectively
extracted
with a solvent such as furfural, dimethyl formamide, sulfur dioxide, aniline,
nitrobenzene, and other known solvents to extract any heavier paraffins. This
process
is more fully described in U.S. patent 2,395,362 issued May 15, 1945 to
Welling
assigned to Phillips Petroleum Company, the contents of which are herein
incorporated
by reference.
Another separation method is an adsorption method. The adsorbent
preferentially adsorbs one of the components in the product stream. The
adsorption
method typically comprises a train of two or more adsorption units so that
when a unit
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has reached capacity the feed is directed to an alternate unit while the fully
loaded unit
is regenerated typically by one or more of a change in temperature or pressure
or both.
There is a significant amount of art on the separation of ethylene and ethane
using silver or copper ions in their + 1 oxidation state. The olefins are
preferentially
absorbed into a complexing solution that contains the complexing agent
selected from
silver (I) or copper (I) salts. dissolved in a solvent.. Some silver
absorbents include
silver nitrate, silver fluoroborate, silver fluorosilicate, silver
hydroxyfluoroborate, and
silver trifluoroacetate. Some copper absorbents include cuprous nitrate;
cuprous
halides such as cuprous chloride; cuprous sulfate; cuprous sulfonate; cuprous
carboxylates; cuprous salts of fluorocarboxylic acids, such as cuprous
trifluoroacetate
and cuprous perfluoroacetate; cuprous fluorinated acetylacetonate; cuprous
hexafluoroacetylacetonate; cuprous dodecylbenzenesulfonate; copper-aluminum
halides, such as cuprous aluminum tetrachloride; CuAICH3C13 ; CuAlC2 H5C13 ;
and
cuprous aluminum cyanotrichloride. If the product stream has been dried prior
to
contact with the liquid adsorbent, the absorbent should be stable to
hydrolysis. The
complexing agent preferably is stable and has high solubility in the solvent.
After one
adsorbent solution is substantially loaded the feed of product stream is
switched to a
further solution. The solution of adsorbent which is fully loaded is then
regenerated
through heat or pressure changes or both. This releases the ethylene.
These types of processes are described in U.S. patents 6,581,476 issued Feb.
11, 2003 to Culp et al. assigned to Union Carbide Chemicals & Plastics
Corporation
and U.S. 5,859,304 issued Jan. 12, 1999 to Barchas et al., assigned to Stone
and
Webster Engineering the contents of which are herein incorporated by
reference.
In an alternative to the solution process supports such as zeolite 4A, zeolite
X,
zeolite Y, alumina and silica, may be treated with a copper salt, to
selectively remove
carbon monoxide and/or olefins from a gaseous mixture containing saturated
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hydrocarbons (i.e. paraffins) such as ethane and propane. U.S. Patent
4,917,711
issued April 17, 1990 to Xie et al., assigned to Peking University describes
the use of
such supported adsorbents, the contents of which are incorporated herein by
reference.
Similarly, U.S. Patents 6,867,166 issued March 15, 2005 and 6,423,881 and July
23, 2002 ty Yang et al., assigned to the Regents of the University of
Michigan, which
are herein incorporated by reference, describe the use of copper salts and
silver
compounds supported alternatively on silica, alumina, MCM-41 zeolite, 4A
zeolite,
carbon molecular sieves, polymers such as Amberlyst-35 resin, and alumina to
selectively adsorb olefins from gaseous mixtures containing olefins and
paraffins. Both
kinetic and thermodynamic separation behavior was observed and modeled. The
adsorption of the olefin takes place at pressures from 1 to 35 atmospheres,
preferably
less than 10 atmospheres, most preferably less than 2 atmospheres at
temperatures
from 0 to 50 C, preferably from 25 to 50 C and the desorption occurs at
pressures
from 0.01 to 5 atmospheres, preferably 0.1 to 0.5 at temperatures from 70 C
to 200 C,
preferably from 100 C to 120 C.
In a further embodiment the adsorbent may be a physical adsorbent selected
from the group consisting of natural and synthetic zeolites without a silver
or copper
salt.
In general, the adsorbent may be alumina, silica, zeolites, carbon molecular
sieves, etc. Typical adsorbents include alumina, silica gel, carbon molecular
sieves,
zeolites, such as type A and type X zeolite, type Y zeolite, etc. The
preferred
adsorbents are type A zeolites, and the most preferred adsorbent is type 4A
zeolite.
Type 4A zeolite, i.e. the sodium form of type A zeolite, has an apparent pore
size
of about 3.6 to 4 Angstrom units. This adsorbent provides enhanced selectivity
and
capacity in adsorbing ethylene from ethylene-ethane mixtures and propylene
from
propylene-propane mixtures at elevated temperatures. This adsorbent is most
effective
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for use in the invention when it is substantially unmodified, i.e. when it has
only sodium
ions as its exchangeable cations. However, certain properties of the
adsorbent, such as
thermal and light stability, may be improved by partly exchanging some of the
sodium
ions with other cations (other than silver or copper). Accordingly, it is
within the scope of
the preferred embodiment of the invention to use a type 4A zeolite in which
some of the
sodium ions attached to the adsorbent are replaced with other metal ions,
provided that
the percentage of ions exchanged is not so great that the adsorbent loses its
type 4A
character. Among the properties that define type 4A character are the ability
of the
adsorbent to selectively adsorb ethylene from ethylene-ethane mixtures and
propylene
from propylene-propane gas mixtures at elevated temperatures, and to
accomplish this
result without causing significant oligomerization or polymerization of the
alkenes
present in the mixtures. In general, it has been determined that up to about
25 percent
(on an equivalent basis) of the sodium ions in 4A zeolite can be replaced by
ion
exchange with other cations without divesting the adsorbent of its type 4A
character.
Cations that may be ion exchanged with the 4A zeolite used in the alkene-
alkane
separation include, among others, potassium, calcium, magnesium, strontium,
zinc,
cobalt, manganese, cadmium, aluminum, cerium, etc. When exchanging other
cations
for sodium ions it is preferred that less than about 10 percent of the sodium
ions (on an
equivalent basis) be replaced with such other cations. The replacement of
sodium ions
may modify the properties of the adsorbent. For example, substituting some of
the
sodium ions with other cations may improve the stability of the adsorbent. As
disclosed
in U.S. patent 5,744,687 issued April 28, 1998 to Ramachandran et al.,
assigned to the
BOO Group, Inc. the contents of which are herein incorporated by reference.
A particularly preferred zeolite is ZSM -5.
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In addition to zeolites there are a number of titanosilicate homologues
referred to
as ETS compounds.
US Pat. No. 5,011,591 discloses the synthesis of a large pore diameter
titanosilicate designated "ETS-10". In contrast to ETS-4 and CTS-1 (referenced
below),
the large pore titanosilicate material, ETS-10, which has pore diameters of
about 8 A,
cannot kinetically distinguish light olefins from paraffins of the same carbon
number.
Nevertheless, high degrees of selectivity have been reported for the
separation of
ethylene from ethane using as prepared ETS-10 zeolites; see: Al-Baghli and
Loughlin in
J. Chem. Eng. Data 2006, v51, p248. The authors demonstrate that Na-ETS-10 is
capable of selectively adsorbing ethylene from a mixture of ethylene and
ethane under
thermodynamic conditions, even at ambient temperature. Although, the reported
selectivity for ethylene adsorption using Na-ETS-10 was high at ambient
temperature,
the adsorption isotherms for ethylene and ethane had highly rectangular shapes
consistent with a low pressure swing capacity. Consequently, Na-ETS-10 is not
readily
applicable to pressure swing absorption processes (PSA), at least at lower or
ambient
temperatures.
However, cationic modification of as prepared Na-ETS-10 provides an adsorbent
for the PSA separation of olefins and paraffins having the same number of
carbon
atoms, at ambient temperatures. the mono-, di- and tri-valent cations are
selected from
the group 2-4 metals, a proton, ammonium compounds and mixtures thereof. Some
specific non-limiting examples of mono-, di, or tri-valent cations that can be
used in the
current invention include, Li, K+, Cs, Mg2+, Ca2+, Sr, Ba2+, Sc3+, 113+, La3+,
Cu, Zn2+,
Cd2+, Ag+, Au, H+, NH4, and NR4+ where R is an alkyl, aryl, alkylaryl, or
arylalkyl
group. The cationic modifiers are generally added to unmodified Na-ETS-10 in
the form
of a salt or an acid. The anionic counterion associated with the cationic
modifier is not
specifically defined, provided that is does not adversely affect the
modification (i.e.
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cation exchange) reactions. Suitable 1nions include but are not limited to
acetate,
carboxylate, benzoate, bromate, chlorate, perch lorate, chorite, citrate,
nitrate, nitrite,
sulfates, and halide (F, Cl, Br, I) and mixtures thereof. Suitable acids
include inorganic
and organic acids, with inorganic acids being preferred. United States patent
8.017,825
issued September 13, 2011 to Kuznicki et al, assigned to the Governors of the
University of Alberta discloses the technology, the text of which is herein
incorporated
by reference.
As described in US Pat. No. 6,517,611, heat treatment of ETS-4 gave a
controlled pore volume zeolite material, dubbed "CTS-1" which is a highly
selective
absorbent for olefin/paraffin separations. The CTS-1 zeolite, which has pore
diameters
of from about 3-4 A, selectively adsorbed ethylene from a mixture of ethylene
and
ethane through a size exclusion process. The pore diameter of CTS-1, allowed
diffusion of ethylene, while blocking diffusion of ethane which was too large
to enter the
pores of the CTS-1 zeolite, thereby providing a kinetic separation. The CTS-1
adsorbent was successfully applied to a PSA process in which ethylene or
propylene
could be separated from ethane or propane respectively.
The above adsorbents may be used in pressure swing adsorption units.
Typically, the range of absolute pressures used during the adsorption step can
be from
about 10 kPa to about 2,000 kPa, (about 1.5 to about 290 pounds per square
inch (psi))
preferably from about 50 kPa to about 1000 kPa (from about 7.2 to about 145
psi) . The
range of pressures used during the release of adsorbate (i.e. during the
regeneration
step) can be from about 0.01 kPa to about 150 kPa (about 0.0015 to about 22
psi),
preferably from about 0.1 kPa to about 50 kPa (about 0.015 to about 7.3 psi).
In
general, the adsorption step can be carried out at from ambient temperatures
to above
about 200 C, preferably less than 150 C, most preferably less than 100 C,
provided
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that the temperatures do not exceed temperatures at which chemical reaction of
the
olefin, such as a oligomerization or polymerization takes place.
Another class of adsorbents is ionic liquids. Olefins and paraffins can be
separated using ionic liquids of the formula a metal dithiolene selected from
the group
of complexes of the formulae:
(i) M[S2 C2 (R1 R2)]2;
/S
R2
____________________________________________ 2
and
(ii) M[S2 06 (R3 R4 R6 R7)]2.
R3
R4
100 R5
R6
-
wherein M is selected from the group consisting of Fe, Co, Ni, Cu, Pd and Pt;
and R1,
R2, R3, R4, R5, and R6 are independently selected from the group consisting of
a
hydrogen atom, electron-withdrawing groups including those that are or contain
heterocyclic, cyano, carboxylate, carboxylic ester, keto, nitro, and sulfonyl
groups,
hydrocarbyl radicals selected from the group consisting of C1_6, alkyl groups,
C6_8, alkyl
groups, C2-8, alkenyl groups and C6_8 aryl groups which hydrocarbyl radicals
are
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unsubstituted or fully or partly substituted, preferably those substituted by
halogen
atoms. The ionic liquid may be used with a non-reactive solvent or co solvent.
The
solvent may be selected from the group conventional aromatic solvents,
typically
toluene. Adsorption pressures may range from 200 psig to 300 psig (1.3X103 to
2X103
kPag), preferably below 250 psig (1.7X103kPag) and adsorption temperatures may
range from ambient to 200 C, preferably below 150 C, and the olefin may be
released
from the ionic liquid by one or more of lowering the pressure by at least 50
psi (3.4X102
kPa) and increasing the temperature by not less than 15 C.
The nitrogen from the regeneration rector, not recycled to the regeneration
reactor could be used in a number of downstream unit operations. Potential
downstream unit operations include an ammonia plant, an acrylonitrile plant, a
urea
plant and an ammonium nitrate plant.
The following non-limiting examples demonstrate the present invention.
The catalyst used in the experiments was of the formula
MoaVbNbcTee0d
wherein:
a is from 0.90 to 1.10;
b is from 0.25 to 0.4;
c is from 0.1 to 0.3;
e is from 0.1 to 0.3 , and
d is a number to satisfy the valence state of the mixed oxide catalyst.
The reactor used in the experiments consisted of a quartz tube reactor. The
sample size was typically about 0.5 cm3, 0.17g. The particle size for the
catalyst was
0.2 -0.7 mm.
The reactor was initially operated in a regeneration (oxidation of the
catalyst)
mode. The reactor was heated to a temperature from 355 C to 397 C. in air
for 30
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minutes. Then the gas flow was switched to a mixture of 75 vol.% ethane and 25
vol.
% oxygen. The flow rate of the mixture of ethane and oxygen was varied over
300/600/1200 cm3 (Stp) per hour. The reaction took place during the first
minute of the
passage of the reactants over the oxidized catalyst bed. The catalyst bed was
then
reoxidized and then a mixture of ethane and oxygen were passed over the
oxidized
catalyst. The gas leaving the reactor was analyzed to measure the residual
oxygen
and the amount of ethane, ethylene and by-products in the product gas.
Experiment #1
(air <=> gas mixture [75%C2H64-25%02]) 355 C
Figures 2 and 3 demonstrate a time dependence of the ethane and 02 conversion
as
well as the selectivity of ethylene formation upon the gradual reduction of
the pre-
oxidized catalyst by the reaction mixture supplied at 600 cc/hrat two
different
temperatures.
As one can see (Figs. 2, 3), all transient processes take place during the
first
operation minute in our testing conditions. The effect is not pronounced at -
355 C, only
a slight increase of the conversion without any selectivity loss can be noted
(Fig. 2).
The same effect of the conversion rise becomes much stronger at 400 C, but in
this
case it is accompanied by a substantial loss of the selectivity due to
additional
formation of CO2 (Fig. 3). It is necessary to mention that the more actively
occurring
process at 400 C is accompanied by measurable self-heating of the catalyst
layer (-5-
6 C measured on the wall of the reactor). Some contribution of undesirable
complete
oxidation in the gas phase cannot be excluded.
Experiment 2
Experiment 1 was repeated at 398 C.
Comparing experiment 1 and 2, the maximum conversion was increased to
above 70%.
Experiment 3
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To clarify the invention, an additional test was carried out with varied gas
flow
rates (300 and 1200 cc/min) using same condition as example 1 and 2, except
that the
flow rate was 1200 cm3/h. Results obtained are presented in Figures 4 and 5.
The data obtained (Figs. 3-5), shows that the selectivity is related to the
feed
flow rate (space velocity). Reduction of the gas flow rate down to 600 h-1
causes a
temporary drop of the selectivity down to 75% (Fig. 5b). Again, the process is
accompanied by considerable self-heating of the catalytic layer after the gas
switch to
the reaction mixture (-6-7 C measured on the wall of the reactor). The
selectivity
curves are summarized and compared in Fig. 6. It is interesting to note at
short
residence time, despite the high conversion, very little gas phase oxygen is
consumed.
So, the contribution of undesirable complete oxidation with temporary heating
of the
catalyst bed becomes more and more pronounced upon the rise of the contact
time
(Fig. 6). At the same time, an increase of the gas flow rate up to 2400 I-11
permits us to
avoid a considerable contribution of total oxidation (Fig. 6).
For quantitative comparison of the conversion data, all the results obtained
at
flow rates differing by a factor of 2 are presented in Fig. 7 using an
absolute scale (i.e.,
as a function of the amount of the ethane fed through the reactor). All three
curves look
quite similar (Fig. 7). It is evident that the increased starting conversion
of ethane (70-
80%) is caused by the presence of extra-oxygen stored in the pre-oxidized
catalyst, and
the transient process shown in Figure 7 is related with the gradual loss of
this additional
oxygen. The results obtained permit us to calculate the amount of the
"reactive" lattice
oxygen involved in the reaction during the transient process. Depletion of
oxygen from
the catalyst is the same for all three tests and can be evaluated as -1% from
the total
lattice oxygen of our Mo-V-Te-Nb-Ox catalyst.
Experiment 4
Periodical redox cycle (pure 02<:=> pure C2H6): reference testing
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To clarify the invention, this experiment was done in the absence of oxygen in
the ethane stream under the same other conditions (i.e. flow rate and
temperature). In
this test, the catalyst charge placed into a quartz reactor was heated to at
given
temperature (354 or 397 C) in pure 02 flow, kept for 30 min, then the gas
flow (600
cm3/h) was switched to pure C2H6, and the sample of the outgoing gas was
analyzed
after a given time. After reoxidation of the catalyst for 30 min measurements
were
repeated several times with varied time interval, and resulting response
curves of
products were received (up to 3 min). Figures 8 and 9 demonstrate a time
dependence
of ethane conversion and residual 02 content as well as selectivity of
ethylene
formation upon the catalyst gradual reduction by the ethane at two different
temperatures.
Transient processes take place during 1-2 minutes in our testing conditions
(Figs. 8,9). Reaction is accompanied by a measurable selectivity loss. Effect
is quite
pronounced even at -350 C (Fig. 8,b) and becomes stronger at 400 C (Fig. 9,b).
It is
important to note that reaction is accompanied by a measurable self-heating of
the
catalyst layer (4-8 C measured on the outer wall of the reactor). Back switch
to 02 flow
for the catalyst reoxidation is also accompanied by some catalyst heating (3-4
C). In
addition, this heating seems to be non-uniform but moving throughout the layer
during
reaction. Taking into account that this over-heating of the catalyst is
considerably
stronger inside the catalyst bed the role of non-isothermal conditions
provided by switch
between two pure gases could be important.
The above examples also illustrates that the conversions and selectivity using
a
pulse mode of ODH are not as effective as a present invention.
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