U.S. patent application number 14/634319 was filed with the patent office on 2015-08-27 for systems and methods for partial or complete oxidation of fuels.
The applicant listed for this patent is OHIO STATE INNOVATION FOUNDATION. Invention is credited to Elena Chung, Liang-Shih Fan, Mandar Kathe, Andrew Tong, William Wang.
Application Number | 20150238915 14/634319 |
Document ID | / |
Family ID | 52630538 |
Filed Date | 2015-08-27 |
United States Patent
Application |
20150238915 |
Kind Code |
A1 |
Fan; Liang-Shih ; et
al. |
August 27, 2015 |
SYSTEMS AND METHODS FOR PARTIAL OR COMPLETE OXIDATION OF FUELS
Abstract
A system used for converting multiple fuel feedstocks may
include three reactors. The reactor system combination can be so
chosen that one of the reactors completely or partially converts
the fuel while the other generates the gaseous product required by
utilizing the gaseous product from the second reactor. The
metal-oxide composition and the reactor flow-patterns can be
manipulated to provide the desired product. A method for optimizing
the system efficiency where a pressurized gaseous fuel or a
pressurized utility is used for applications downstream can be used
to any system processing fuels and metal-oxide.
Inventors: |
Fan; Liang-Shih; (Columbus,
OH) ; Kathe; Mandar; (Columbus, OH) ; Wang;
William; (Hilliard, OH) ; Chung; Elena;
(Columbus, OH) ; Tong; Andrew; (Massillon,
OH) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
OHIO STATE INNOVATION FOUNDATION |
Columbus |
OH |
US |
|
|
Family ID: |
52630538 |
Appl. No.: |
14/634319 |
Filed: |
February 27, 2015 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
61945257 |
Feb 27, 2014 |
|
|
|
62041703 |
Aug 26, 2014 |
|
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Current U.S.
Class: |
422/630 |
Current CPC
Class: |
C10J 2300/0916 20130101;
C01B 2203/0216 20130101; C10J 2300/1659 20130101; C10J 2300/0969
20130101; F23C 2900/99008 20130101; C10J 2300/0986 20130101; C10J
3/06 20130101; C10J 2300/1665 20130101; Y02E 50/10 20130101; C10J
3/725 20130101; C01B 3/344 20130101; Y02E 50/18 20130101; C01B
2203/0222 20130101; B01J 8/0278 20130101; C10J 2300/1807 20130101;
B01J 2208/023 20130101; Y02E 20/34 20130101; Y02E 20/346 20130101;
C10G 2/32 20130101 |
International
Class: |
B01J 8/02 20060101
B01J008/02 |
Claims
1. A system for the production of syngas, comprising: a first
reactor comprising a plurality of oxygen carrying particles
comprising a first metal oxide, wherein the first reactor is
configured to provide a counter-current contact mode between the
first metal oxide and a first fuel to reduce the first metal oxide
to a second metal oxide; a second reactor in communication with the
first reactor, the second reactor configured to oxidize the second
metal oxide to a third metal oxide, and further configured to
reduce the third metal oxide to a fourth metal oxide with a second
fuel to provide a partially or fully oxidized gaseous fuel
comprising one or more of CO, CO.sub.2, H.sub.2, and H.sub.2O,
wherein the second metal oxide is oxidized to the third metal oxide
using an enhancing gas of CO.sub.2 and H.sub.2O, the partially or
fully oxidized gaseous fuel, or a combination thereof, to generate
syngas; and a third reactor in communication with the second
reactor, the third reactor configured to regenerate the first metal
oxide by oxidizing the fourth metal oxide with an oxygen
source.
2. The system of claim 1, wherein the counter-current contact mode
between the first metal oxide and the first fuel is such that the
first metal oxide moves downward and the first fuel moves
upward.
3. The system of claim 1, wherein the first metal oxide is
introduced to the top of the first reactor, and the first fuel is
introduced to the bottom of the first reactor.
4. The system of claim 1, wherein the second reactor is configured
to provide a counter-current contact mode between the second metal
oxide and the enhancing gas, and a counter-current contact mode
between the third metal oxide and the second fuel.
5. The system of claim 4, wherein the second metal oxide is
introduced to the top of the second reactor, the enhancing gas is
introduced to the middle of the second reactor, and the second fuel
is introduced to the bottom of the second reactor.
6. The system of claim 1, wherein the second reactor is configured
to provide a co-current contact mode between the second metal oxide
and the enhancing gas, and a co-current contact mode between the
third metal oxide and the second fuel.
7. The system of claim 6, wherein the second metal oxide is
introduced to the top of the second reactor, the enhancing gas is
introduced to the middle of the second reactor, and the second fuel
is introduced to the top of the second reactor.
8. The system of claim 6, wherein the second metal oxide is
introduced to the top of the second reactor, the enhancing gas is
introduced to the top of the second reactor, and the second fuel is
introduced to the top or the middle of the second reactor.
9. The system of claim 1, wherein at least a portion of the
enhancing gas is derived from the first reactor resulting from the
reduction of the first metal oxide with the first fuel.
10. The system of claim 1, wherein at least a portion of the
enhancing gas is derived from oxidation of a carbon-containing or
hydrogen-containing source in the third reactor, a fourth reactor,
or a combination thereof.
11. The system of claim 1, wherein the third reactor is
communication with the first reactor, wherein at least a portion of
the second metal oxide is circulated directly to the third reactor,
and oxidation of a carbon-containing or hydrogen-containing source
with an oxygen source in the third reactor generates a stream of
enhancing gas, wherein at least a portion of the enhancing gas
generated in the third reactor is used in the second reactor as an
enhancing gas.
12. The system of claim 1, comprising a fourth reactor in
communication with the first reactor, configured to generate a
stream of enhancing gas comprising CO.sub.2 and H.sub.2O.
13. The system of claim 12, wherein at least a portion of the
enhancing gas generated in the fourth reactor is used in the second
reactor as an enhancing gas.
14. The system of claim 1, wherein the first fuel is a solid fuel
selected from biomass, coal, pet-coke, solid hydrocarbon-based
waste products, or a combination thereof.
15. The system of claim 1, wherein the first fuel is a gaseous fuel
selected from natural gas, gasified coal, a light hydrocarbon
off-gas stream, or a combination thereof.
16. The system of claim 1, wherein the second fuel is a solid fuel
selected from biomass, coal, pet-coke, solid hydrocarbon-based
waste products, or a combination thereof.
17. The system of claim 1, wherein the second fuel is a gaseous
fuel selected from natural gas, gasified coal, a light hydrocarbon
off-gas stream, or a combination thereof.
18. The system of claim 1, wherein the second reactor is in
communication with a Fischer-Tropsch or methanol synthesis system
that produces a light hydrocarbon tail-gas, wherein the second
reactor provides syngas to the Fischer-Tropsch or methanol
synthesis system and the light hydrocarbon tail-gas is optionally
recycled to first reactor, the second reactor, or a combination
thereof.
19. The system of claim 1, wherein the first metal oxide has
formula FeO.sub.aTi.sub.x or FeO.sub.aAl.sub.x, the second metal
oxide has formula FeO.sub.bTix or FeO.sub.bAl.sub.x, the third
metal oxide has formula Fe.sub.cTi.sub.x or FeO.sub.cTi.sub.x, and
the fourth metal oxide has formula FeO.sub.dTi.sub.x or
FeO.sub.dAlx, wherein 1.5>a>b, b<c>d, 1.5>c, and x
is 0.01 to 5.
20. The system of claim 1, wherein the first metal oxide has
formula FeO.sub.aTiO.sub.2 or FeO.sub.aAl.sub.2O.sub.3, the second
metal oxide has formula FeO.sub.bTiO.sub.2 or
FeO.sub.bAl.sub.2O.sub.3, the third metal oxide has formula
FeO.sub.cTiO.sub.2 or FeO.sub.cAl.sub.2O.sub.3, and the fourth
metal oxide has formula FeO.sub.dTiO.sub.2 or
FeO.sub.dAl.sub.2O.sub.3, wherein 1.5>a>b, b<c>d, and
1.5>c.
Description
CROSS REFERENCE TO RELATED APPLICATION
[0001] This application claims priority to U.S. Provisional
Application 61/945,257, filed Feb. 27, 2014; and U.S. Provisional
Application 62/041,703, filed Aug. 26, 2014, each of which is
incorporated herein by reference in its entirety.
TECHNICAL FIELD
[0002] The present disclosure relates to systems and methods for
converting carbon-based fuels such as methane-rich sources like
natural gas or shale gas, syngas, biomass, and coal to value-added
products with oxidation-reduction metal-oxides. The exemplary
embodiments detail the oxidation-state of the metal oxide in
multiple reactor configurations and ways to utilize multiple
carbonaceous sources.
BACKGROUND
[0003] The rise in human population is related to the rise in
global energy demand for value-added products such as gasoline,
jet-fuel, diesel, and synthetic intermediates for polyester,
polyethylene, Teflon, etc. Conversion of fossil fuels (e.g. natural
gas, coal) to value-added products can be used to meet the growing
energy demand. Given the abundance of natural fossil resources
worldwide and the potential benefits of economic liquid fuel
production, the fuel projects have seen some of the highest capital
investments for a single fuel processing project worldwide. The
cost-intensive nature of the conventional technology has led to
considerable research in developing alternatives for
fuel-to-liquids conversion.
[0004] Conventional technologies for liquid fuel production from
coal and natural gas utilize a two-step process. The initial step
involves converting the fuel to a synthesis gas (syngas) composing
an appropriate H.sub.2/CO ratio that can vary from 1.0 for
co-firing to 2.0 or greater for the Fischer-Tropsch reaction or
methanol synthesis. This initial step, also known as the
syngas-generating step, is capital and energy intensive in terms of
the overall plant capital cost and the syngas generation
efficiency, respectively.
[0005] As energy demands rise due to global population increase,
developing systems and system components that can convert fuels
efficiently are a necessity. This need also opens the opportunity
to develop processes that can flexibly operate using multiple types
of hydrocarbon feeds and/or can reduce the demands of conventional
cost-intensive process unit operations.
SUMMARY
[0006] In one aspect, disclosed is a system for conversion of fuels
to produce syngas, the system comprising a series of reactors. In
one embodiment, a reactor system is used to produce syngas by
reaction with an oxygen carrier material while the fuel source
consists of two feedstocks. The reactor system is designed to
partially or fully oxidize a given fuel and the resultant gaseous
product stream can act as an enhancing agent for gasifying the
second fuel. The gas-solid contact mode can be designed to react
the metal-oxide with the first and second fuel to convert them to a
high purity, syngas with a flexible H.sub.2/CO ratio. A third
reactor is used to re-oxidize the metal oxide oxygen carrier and
complete the loop auto-thermally. The present disclosure also
includes a system and method for increasing the overall system
thermal efficiency and decreasing the cost by applying a unique
system configuration of compressors/expanders.
[0007] In another aspect, disclosed is a system for conversion of
fuel (e.g., coal) to produce high quality syngas. The specified
operating condition obtained by combination of oxygen carrier and
reactor contact pattern selectively oxidizes the fuel feed-stock to
a syngas suitable for liquids and/or chemicals production. The
combination disclosed increases the carbon utilization efficiency
and reduces the costs associated with additional reforming while
producing syngas. The operating strategy includes a combination of
a suitable oxygen carrier and its temperature swing for sustainable
heat management, a co-current downward reaction mode obtained by a
specific flow ratio of the oxygen carrier to the solid fuel, and a
specific steam flow. The operating condition is characterized by a
critical point of operation, wherein the oxygen transfer from the
oxygen carrier to the solid fuel and from the steam to the oxygen
carrier is maximized while minimizing the oxygen transfer from
steam to the fuel. At the specified operating condition, the oxygen
carrier undergoes an overall loss of oxygen in the reducer reactor,
while gaining it in the oxidizer reactor. A part of the swing is
utilized to produce energy for satisfying the parasitic energy
requirement of the system. The current disclosure also provides a
method used in-conjunction with the specific operating condition to
enhance the savings on the net operating energy.
[0008] Additional features include the advantages and process
configurations with detailed description for each system disclosed.
The method disclosure has a general description followed by a
detailed example of analysis demonstrating how the method can be
applied.
BRIEF DESCRIPTION OF THE DRAWINGS
[0009] The following detailed description of specific embodiments
of the present disclosure can be best understood when read in
conjunction with the following drawings, where like structure is
indicated with like reference numerals.
[0010] FIG. 1 is a configuration where an oxidation state swing is
used in the second reactor, according to one or more embodiments
described herein.
[0011] FIG. 2 illustrates the oxidation state swing of metal-oxide
using light hydrocarbon recycle or any waste hydrocarbon stream
from the liquid fuel production facility or otherwise, according to
one or more embodiments described herein.
[0012] FIG. 3 illustrates the use of the first reactor to produce
syngas and the second reactor to generate enhancing gas while
over-reducing the metal-oxide, according to one or more embodiments
described herein.
[0013] FIG. 4 illustrates the oxidation state swing of metal-oxide
using light hydrocarbon recycle or any waste hydrocarbon stream
from the liquid fuel production facility or otherwise, according to
one or more embodiments described herein.
[0014] FIG. 5 is a configuration with the enhancing gas obtained
from the second reactor according to one or more embodiments
described herein.
[0015] FIG. 6 is a similar configuration as FIG. 5 with light
hydrocarbon recycle used an enhancing gas for either fuel,
according to one or more embodiments described herein.
[0016] FIG. 7 is the oxidation state swing in the second reactor
with co-current downward gas-solid flow, according to one or more
embodiments described herein.
[0017] FIG. 8 is a variation of FIG. 7 to use a light hydrocarbon
recycle from the liquid fuel production facility, flare gas recycle
or otherwise, according to one or more embodiments described
herein.
[0018] FIG. 9 illustrates syngas production in a co-current
downward flow reactor producing syngas with metal-oxide oxidation
state swing, according to one or more embodiments described
herein.
[0019] FIG. 10 shows a variation of FIG. 9 using light
hydrocarbons, according to one or more embodiments described
herein.
[0020] FIG. 11 illustrates a combination of partial/full oxidation
to generate enhancing gas in combination with metal re-oxidation or
any combination thereof, according to one or more embodiments
described herein.
[0021] FIG. 12 shows oxidation of a carbon-containing and/or
hydrogen-containing source in a fourth reactor system, a
combination of partial/full oxidation to generate enhancing gas in
combination with metal re-oxidation or any combination thereof with
light hydrocarbon recycle from liquid fuel production facility or
otherwise, according to one or more embodiments described
herein.
[0022] FIG. 13 illustrates a two-reactor system for high
temperature separation of char, oxygen carrier particles, and
abraded particle and ash fines, according to one or more
embodiments described herein.
[0023] FIG. 14 illustrates a two reactor system for syngas
generation with a gas compressor and expansion used on the gas
stream for the second reactor, according to one or more embodiments
described herein.
[0024] FIG. 15 shows a schematic of a chemical looping reactor
scheme operated at 1 atm, heat integration Scheme 1
[0025] FIG. 16 shows a schematic of a chemical looping reactor
scheme operated at 1 atm, heat integration Scheme 2.
[0026] FIG. 17 shows an overall schematic representative of Case 3
to Case 6 (5 to 25 atm), heat integration Scheme 2.
[0027] FIG. 18 shows an overall flow diagram and the heat
integration optimization schematic for Case 7, heat integration
Scheme 2.
[0028] FIG. 19 shows the overall % operating cost for specific case
over the ASU operating cost for same methane input baseline.
[0029] FIG. 20 shows the basic reactor system set-up analyzed and
used in the disclosure for reducer reactor for various
configurations described.
[0030] FIG. 21a illustrates the variation in net-duty (kW/mol C) as
a function of Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input
of methane and PRB coal at an iso-thermal temperature of 900 C for
the reducer reactor.
[0031] FIG. 21b illustrates the variation in the H.sub.2/CO ratio
and M value as a function of Fe.sub.2O.sub.3/C ratio variation for
a 50% HHV input of methane and PRB coal and an iso-thermal
temperature of 900 C for the reducer reactor.
[0032] FIG. 22a illustrates the variation in the H.sub.2/CO ratio
and M value as a function of Fe.sub.2O.sub.3/C ratio variation for
a 50% HHV input of methane and PRB coal and varying temperatures
for the reducer reactor.
[0033] FIG. 22b illustrates the variation in the net-duty (kW/mol
C) as a function of Fe.sub.2O.sub.3/C ratio variation for a 50% HHV
input of methane and PRB coal and varying temperatures for the
reducer reactor.
[0034] FIG. 23a illustrates the variation in the H.sub.2/CO ratio
as a function of Fe.sub.2O.sub.3/C ratio variation for a 50% HHV
input of methane and PRB coal and varying temperatures for the
reducer reactor for varying amount of H.sub.2O injection.
[0035] FIG. 23b illustrates the variation in the M value
((H.sub.2--CO.sub.2)/(CO+CO.sub.2)) as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0036] FIG. 24a illustrates the variation in the net reducer duty
(kW/mole C) as a function of Fe.sub.2O.sub.3/C ratio variation for
a 50% HHV input of methane and PRB coal for the reducer reactor for
varying amount of H.sub.2O injection.
[0037] FIG. 24b illustrates the variation in the % Syngas
(%(CO+H.sub.2) in gas-stream) as a function of Fe.sub.2O.sub.3/C
ratio variation for a 50% HHV input of methane and PRB coal for the
reducer reactor for varying amount of H.sub.2O injection.
[0038] FIG. 25 illustrates the variation in the CO/CO.sub.2 ratio
of syngas produced as a function of Fe.sub.2O.sub.3/C ratio
variation for a 50% HHV input of methane and PRB coal for the
reducer reactor for varying amount of H.sub.2O injection.
[0039] FIG. 26 illustrates the variation in the H.sub.2/CO ratio of
syngas produced as a function of Fe.sub.2O.sub.3/C ratio variation
for a 50% HHV input of methane and PRB coal for the reducer reactor
for varying amount of H.sub.2O injection.
[0040] FIG. 27 illustrates the variation in the temperature of the
solids coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0041] FIG. 28 illustrates the variation in the CO content in the
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0042] FIG. 29 illustrates the variation in the CO.sub.2 content in
the syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0043] FIG. 30 illustrates the variation in the H.sub.2 content in
the syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0044] FIG. 31 illustrates the variation in the Carbon deposition
in the reduced solids coming out of reducer reactor produced as a
function of Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input
of methane and PRB coal for the reducer reactor for varying amount
of H.sub.2O injection.
[0045] FIG. 32 illustrates the variation in the M value of the
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of H.sub.2O
injection.
[0046] FIG. 33 shows the basic reactor system set-up analysed with
CO.sub.2 recycle from an Acid Gas Removal (AGR) type system.
[0047] FIG. 34a illustrates the variation in the temperature of the
solids coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of CO.sub.2
recycle with steam injection.
[0048] FIG. 34b illustrates the Carbon deposition as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor; for varying amount of
CO.sub.2 recycle with steam injection.
[0049] FIG. 35a shows the shifted M value
((H.sub.2--CO.sub.2(NET))/(CO+CO.sub.2(NET))) in the syngas coming
out of reducer reactor produced as a function of Fe.sub.2O.sub.3/C
ratio variation for a 50% HHV input of methane and PRB coal for the
reducer reactor for varying amount of CO.sub.2 recycle with steam
injection.
[0050] FIG. 35b shows the variation in the H.sub.2/CO ratio in the
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of with
CO.sub.2 recycle with steam injection.
[0051] FIG. 36a shows the variation in the Carbon deposition in the
solids coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0052] FIG. 36b shows the variation in the Temperature of the
solids coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0053] FIG. 37a shows the variation in the CO content of syngas
coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0054] FIG. 37b shows the variation in the H.sub.2 content of
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0055] FIG. 37c shows the variation in the CO.sub.2 content of
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0056] FIG. 38a shows the variation in the M value
((H.sub.2--CO.sub.2)/(CO+CO.sub.2)) of syngas coming out of reducer
reactor produced as a function of Fe.sub.2O.sub.3/C ratio variation
for a 50% HHV input of methane and PRB coal for the reducer reactor
for varying amount of steam injection.
[0057] FIG. 38b shows the variation in the CO/CO.sub.2 value of
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0058] FIG. 38c shows the variation in the H.sub.2/CO value of
syngas coming out of reducer reactor produced as a function of
Fe.sub.2O.sub.3/C ratio variation for a 50% HHV input of methane
and PRB coal for the reducer reactor for varying amount of steam
injection.
[0059] FIG. 39 shows the M value out of the reactor configuration
proposed as a function of varying the % HHV of coal types and
methane.
[0060] FIG. 40 shows the % Syngas (% (CO+H.sub.2)) out of the
reactor configuration proposed as a function of varying the % HHV
of coal types and methane.
[0061] FIG. 41 shows the H.sub.2/CO ratio in the syngas coming out
of the reactor configuration proposed as a function of varying the
% HHV of coal types and methane.
[0062] FIG. 42 illustrates a reactor configuration which includes
fuel injection in the oxidizer for satisfying the heat balance
while getting a higher quality of syngas.
[0063] FIG. 43 shows the percentage carbon utilization for the
cases shown in Table 19 and Table 20.
[0064] FIG. 44 shows an illustration of the methodology used in
conjunction with the configurations disclosed for reducing the net
operating energy of the system.
[0065] FIG. 45 shows an illustration of the application of the
methodology in FIG. 44 for a specific case.
DETAILED DESCRIPTION
[0066] In certain embodiments, disclosed is the use of a specific
metal-oxide composition to perform the selective partial oxidation
of a fuel feedstock. The chemical looping processes typically use a
metal-oxide to perform reduction-oxidation cycles. The overall fuel
reforming reaction is exothermic. The first step is the reaction of
fuel feed with the metal oxide to partially oxidize it to synthesis
gas stream consisting predominantly of H.sub.2 and CO and is
generally endothermic or slightly exothermic. The second step is
the exothermic reaction of the metal oxide with an oxygen
containing reactant. The metal-oxide can be developed and directed
towards the particular application it is used for.
[0067] In certain embodiments, disclosed is a specific operating
strategy obtained by a unique combination of a suitable oxygen
carrier, a co-current downward reaction mode obtained by a specific
flow ratio of the oxygen carrier to the solid fuel, the chosen flow
enhancer gas injection, and the temperature swing of the
near-adiabatic operation. The configurations described use a
metal-oxide oxygen carrier in conjunction with a unique reactor
configuration to convert fuels like coal, biomass etc. to a
H.sub.2-rich syngas. The configurations have two basic reactors in
which the conversions take place. The first reactor, also known as
reducer converts the fuel mixture to a syngas stream with high
carbon-utilization. The first reactor in the process of partially
oxidizing the fuel reduces the metal-oxide oxygen carrier. The
metal-oxide oxygen carrier is re-oxidized back in the oxidizer
using an oxygen source like air, steam etc.
[0068] Syngas production is often used in determining the overall
economics of a liquids/chemicals production facility. The syngas
production systems can make a significant contribution to the
capital cost and the operating cost of the plant. The disclosed
configurations and methodologies provide improvements in the
overall economics of the process, and provide greater flexibility
for providing efficient and economic syngas production systems and
methods.
1. DEFINITIONS
[0069] Unless otherwise defined, all technical and scientific terms
used herein have the same meaning as commonly understood by one of
ordinary skill in the art. In case of conflict, the present
document, including definitions, will control. Preferred methods
and materials are described below, although methods and materials
similar or equivalent to those described herein can be used in
practice or testing of the present invention. All publications,
patent applications, patents and other references mentioned herein
are incorporated by reference in their entirety. The materials,
methods, and examples disclosed herein are illustrative only and
not intended to be limiting.
[0070] The terms "comprise(s)," "include(s)," "having," "has,"
"can," "contain(s)," and variants thereof, as used herein, are
intended to be open-ended transitional phrases, terms, or words
that do not preclude the possibility of additional acts or
structures. The singular forms "a," "an" and "the" include plural
references unless the context clearly dictates otherwise. The
present disclosure also contemplates other embodiments
"comprising," "consisting of" and "consisting essentially of," the
embodiments or elements presented herein, whether explicitly set
forth or not.
[0071] The conjunctive term "or" includes any and all combinations
of one or more listed elements associated by the conjunctive term.
For example, the phrase "an apparatus comprising A or B" may refer
to an apparatus including A where B is not present, an apparatus
including B where A is not present, or an apparatus where both A
and B are present. The phrases "at least one of A, B, . . . and N"
or "at least one of A, B, . . . N, or combinations thereof" are
defined in the broadest sense to mean one or more elements selected
from the group comprising A, B, . . . and N, that is to say, any
combination of one or more of the elements A, B, . . . or N
including any one element alone or in combination with one or more
of the other elements which may also include, in combination,
additional elements not listed.
[0072] For the recitation of numeric ranges herein, each
intervening number there between with the same degree of precision
is explicitly contemplated. For example, for the range of 6-9, the
numbers 7 and 8 are contemplated in addition to 6 and 9, and for
the range 6.0-7.0, the number 6.0, 6.1, 6.2, 6.3, 6.4, 6.5, 6.6,
6.7, 6.8, 6.9, and 7.0 are explicitly contemplated.
2. SYNGAS PRODUCTION SYSTEMS
[0073] Syngas production systems play a role in the overall cost of
a Fischer Tropsch plant configuration for liquid fuel production
from natural gas. Typical syngas production systems include
steam-methane reforming, auto-thermal oxidation and partial direct
oxidation of the gaseous fuel. The present disclosure provides
systems and methods to produce high-quality products like syngas,
with lower fuel consumption per unit heating value of product
produced.
[0074] In one embodiment illustrated in FIG. 1, a three reactor
system is used for flexibly producing syngas from a combination of
fuels. This reactor system consists of 3 reactors to perform
multiple reduction-oxidation cycles. During the reduction cycle,
those skilled in the art can design the gas-solid contact mode to
partially or fully oxidize the fuels in combination to produce
high-quality syngas with a variable H.sub.2:CO ratio.
[0075] As illustrated in FIG. 1, the first reactor uses solid fuels
such as lignite, bituminous, sub-bituminous anthracite, pet-coke,
and/or biomass. The second reactor uses a gaseous fuel such as
natural gas, shale-gas or syngas, etc to react and provide the
counter-current contact mode. The first reactor can be a packed
moving bed, rotary kiln and/or downer to simulate the
counter-current gas-solid contact mode. The solid fuel is
introduced into the first reactor system to simulate a
counter-current contact mode with the oxygen carrier. The
metal-oxide oxygen carrier flows from the top to the bottom in the
first reactor. The fuel residence time and the gas-solid contact
pattern is designed to entirely convert the solid fuel to a gaseous
form when reaching the top of the reactor. The preferred products
are completely oxidized to include CO.sub.2 and H.sub.2O. The metal
oxide donates oxygen to the solid fuel to reform and/or fully
oxidize it. In certain embodiments, the oxygen carrier is written
as FeO.sub.aTiO.sub.2 the product outlet will be FeO.sub.bTiO.sub.2
governed by the relationship a>b. For the specific case shown in
FIG. 1, the overall defining relationship is
1.5>a>b>0.1.
[0076] The second reactor receives the metal-oxide from the first
reactor. The second reactor can be a packed moving bed, rotary
kiln, a downer or a combination thereof. In certain embodiments, a
series of fluidized beds simulating multiple stages of equilibrium
can be used. The fully or partially oxidized gaseous products from
the first reactor are fed into the middle of this reactor. The
gaseous fuel is fed from the bottom. In embodiments when shale and
natural gas are used, the predominant gaseous fuel is methane. The
gaseous product injection from the first reactor should be
manipulated such that it is sent in a location where all the
methane has been converted to a mixture of CO, CO.sub.2, H.sub.2,
and H.sub.2O. The gaseous fuel and the gaseous products flow
counter-currently to the flow of the solid oxygen carrier particles
from the bottom to the top in FIG. 1. The final product is syngas.
The counter-current equilibrium differential can produce a syngas
stream with a flexible composition, which can be used downstream
for producing liquid fuels. The syngas will have a H.sub.2/CO ratio
between 1 and 4.
[0077] The metal-oxide oxidation state swing is a unique
characteristic of this reactor operation. In the top section of the
second reactor, the metal oxide gains oxygen as it reforms the
fully or partially oxidized product to syngas. This involves
gaining oxygen to convert CO.sub.2 to CO and H.sub.2O to H.sub.2.
In certain embodiments, the oxygen carrier introduced into the
second reactor has a formula FeO.sub.bTiO.sub.2. At the location of
where the gaseous fuel is introduced to the second reactor, the
oxygen carrier has a formula FeO.sub.cTiO.sub.2. In the top section
of the second reactor, the oxygen carrier gains oxygen due to its
reaction the product gases from the first reactor. Therefore, the
governing equation is c>b. The same oxygen carrier will donate
oxygen to partially or fully oxidize the gaseous fuels injected
from the bottom of the reactor. In some embodiments, the formula
FeO.sub.dTiO.sub.2 represents the oxygen carrier at this stage and
the governing equation is d<c. Thus, the overall governing
equation is 1.5>a>b and b<=c=>d and 1.5>c. In case
when c is fully oxidized, the value of c can range from b to 1.33.
The governing equations are important in simulating the oxygen
swing of the oxygen carrier.
[0078] The third reactor uses air and/or an oxygen-containing
source such as oxygen from an air-separation to re-oxidize the
metal-oxide to its full oxidation state. From FIG. 1, the
metal-oxide enters the third reactor in the form of
FeO.sub.dTiO.sub.2 and exits in the form FeO.sub.aTiO.sub.2, where
1.5>=a>d>0.1. This value `a` corresponds to the value of
the oxidation state introduced into the first reactor and thereby
completing the redox cycle.
[0079] The overall reactor operation is auto-thermal or near
auto-thermal with the inert percentage of the metal-oxide carrier
varying between 5 to 100% depending on the capacity and the
flow-rates used. The average operating pressure of the system can
vary between 1 and 80 atm. The operating temperature of the reactor
system can be 600 C to 1,300 C for isothermal operation. For an
overall adiabatic operation the operating temperature is
manipulated so no additional heat reactor system is required. The
gaseous fuel is pre-heated to maintain the auto-thermal temperature
profile. The solid fuel can preferably be processed to have a
suitable pellet size for ease of injection and reaction. In certain
exemplary embodiments, the metal-oxide oxygen carrier is
FeO.sub.aTiO.sub.2, with the inert TiO.sub.2 composing of 20 wt %
and with an operating pressure of 20 atm. The temperature in the
first, second and third reactor will range from 1200 C to 900 C,
900 C to 700 C, and 700 C to 1200 C, respectively. The
corresponding syngas composition can be designed to have a
H.sub.2/CO ratio of between 4:1 to 2:1 and the CO.sub.2 and
H.sub.2O % will be less than 10 v/v % of the total product
flow.
[0080] In another embodiment as illustrated in FIG. 2, the syngas
from the system is sent to the Fischer-Tropsch and/or methanol
synthesis system. The light hydrocarbon tail-gas which comes from
the various unit-operations in the liquid fuel production is
recycled to combine with the gaseous fuel and pre-heated to
increase the overall carbon-efficiency process to greater than 99%.
The overall system can produce excess energy from the high-quality
heat extraction from the third reactor and the air-product stream.
This will offset the burning requirement for the light hydrocarbon
from the down-stream liquid fuel production complex and thereby
some of the gas can be recycled to the chemical looping reactor
schematic.
[0081] It should be noted that although it is stated that the first
reactor can use solid fuels, liquid fuels such as naphtha,
gasoline, or residual oil could also be used either in a co-current
or a counter-current manner. The reactor system design will change
to account for the pressure change in vaporization and the
respective residence time requirement to complete the oxidation
prior to introducing the product gas stream into the second
reactor. From FIG. 2, an enhancing gas is introduced into the first
reactor for promoting char gasification in solid fuel. The
enhancing gas can be supplied by providing a slip-stream from the
outlet of the first reactor. The volume percentage of the
slip-stream as compared to the product outlet stream will vary
between 0 to 30%. In the case where the light hydrocarbons from the
liquid fuel production facility are recycled, a portion of the
recycle stream can be used in combination with the recycled
slip-stream out of the gaseous product of the first reactor. The
C.sub.1-C.sub.4 hydrocarbons can be 0 to 100% of the enhancing gas
volumetric flow rate with the complimentary being the gaseous
product of the first reactor. The dotted arrows in FIG. 2 indicate
the possible locations/ways in which enhancing gas is injected.
[0082] In yet another embodiment illustrated in FIG. 3, the first
reactor is used to partially oxidize the solid fuel source while
the second reactor fully oxidizes the gaseous fuel to CO.sub.2 and
H.sub.2O. The first and the second reactors utilize a fuel source
while the third reactor reoxidizes the reduced metal oxide to its
full-oxidation state. In the first two reactors the metal oxide
moves downwards. In the first reactor, the gas and solid move
co-currently while in the second reactor, the gas and solid move
counter-currently. In this embodiment, the solid fuel is partially
converted to syngas while the gaseous fuel is partially or fully
oxidized to increase the gasification efficiency.
[0083] From FIG. 3, a packed moving bed reactor design is used for
both the first and second reactor to provide multiple stages for
the oxygen carrier and thereby separate the two fuel partial
oxidation reaction front for easier large-scale integration. The
first reactor can utilize solid fuels like coal, pet-coke, biomass,
etc. The metal-oxide oxygen carrier is introduced at the top of the
reactor while it exits the reactor at the bottom. In certain
embodiments, the reactor is a packed downward moving bed, a rotary
kiln or a down-comer. In other embodiments, a fluidized bed can be
configured to provide multiple equilibrium stages and perform the
same function as the moving bed reactor. The solid fuel is
introduced at the top of the reactor along a level below the
metal-oxide injection. The gasifying product from the second
reactor is introduced at the top to gasify the solid fuel. It is
injected at a level above the solid fuel. In the embodiment
illustrated in FIG. 3, the highest oxidation state of oxygen
carrier is in contact with the fresh solid feed introduced. This
leads to a higher driving force for oxygen donation and
correspondingly faster kinetics and a smaller reactor size. The
metal oxide will donate oxygen to the solid char and the gaseous
products and drive the conversion to produce syngas. In certain
embodiments, the oxygen carrier is iron based and has the formula
of FeO.sub.aTiO.sub.2 and the outlet metal oxide has a formula of
FeO.sub.bTiO.sub.2, the governing equation relating the value of
`a` and `b` will be 1.5=>a>b=>0.75. The solid fuel flows
co-currently with the gas phase products and reactants.
[0084] The second reactor takes in a gaseous fuel such as natural
gas or shale gas and fully-oxidizes it to a mixture of CO.sub.2 and
H.sub.2O. The gaseous fuel is pre-heated and injected at the bottom
of the second reactor. The gaseous fuel travels upwards
counter-currently while extracting oxygen from the oxygen carrier
reducing it to a lower oxidation state. The oxygen carrier flows
downwards counter-currently with the gas phase products and
reactants. The reduction of the oxygen carrier has a catalyzing
effect initially in converting methane. The greater the reduction
of the oxygen carrier can also help reduce the overall solids
circulation rate by creating a larger exothermic reaction in the
third reactor to fully oxidize the oxygen carrier. The second
reactor can be configured to produce a fully or partially oxidized
gaseous stream. The final product is a mixture which serves as a
gasifying agent to convert the solid fuel to a flexible mixture of
syngas with the H.sub.2/CO ratio varying from 1:1 to 4:1. In
certain embodiments, the oxygen carrier entering the second reactor
from the first reactor has a formula of FeO.sub.bTiO.sub.2 and the
outlet will have a formula of FeO.sub.cTiO.sub.2. The governing
equation for this reactor in terms of the oxygen carrier oxidation
states is 0.1<c<b.
[0085] In FIG. 3, the third reactor uses oxygen from air or from
equipment such as an air separation unit and/or vacuum distillation
unit to re-oxidize the metal oxide to the original oxidation state.
In certain embodiments, the inlet metal-oxide has a formula of
FeO.sub.cTiO.sub.2 and the outlet has a formula of
FeO.sub.aTiO.sub.2. The overall governing equation will be a>c.
The overall system governing equation will be
1.5>=a>b>c>0.1 and c<a. The overall reactor system
can be operated in an iso-thermal or adiabatic mode of operation.
The iso-thermal operation can include a range of operating
temperatures from 700 C to 1,300 C. The adiabatic operation can be
auto-thermal with respect to the oxygen carrier if the endothermic
or near endothermic reactions in the first and second reactor are
balanced or less than the heat extracted from the third reactor
step. In certain embodiments, the pre-heating of the gaseous fuel
stream and the air-stream assists in the auto-thermal operation of
the reactor system. The solid fuel can have a suitable pellet size
relative to the oxygen carrier for ease of injection and good
mixing. In certain embodiments, the oxygen carrier is
FeO.sub.aTiO.sub.2 with 80 wt % TiO.sub.2. The operating pressure
of the system is set at 20 atm and the air and gaseous fuel are
pre-heated to 600 C. The first reactor has a temperature range from
1200 C to 1000 C. In the second reactor the temperature ranges from
1000 C to 700 C. The third reactor has a temperature range from 700
C to 1300 C.
[0086] In another embodiment as illustrated in FIG. 4, the light
hydrocarbons (C.sub.1-C.sub.4) from the liquid fuel production
units can be utilized along with the gaseous fuels. This embodiment
increases the carbon-efficiency to a value greater than 90%. The
overall energy requirement can be offset and the carbon emission
removed by extracting heat and using it in a
gas-turbine/steam-turbine system to generate electricity to
compensate for parasitic energy consumption. In another variation,
the gaseous fuels and/or solid fuels can be substituted with a
liquid fuel such as waste gasoline or petroleum in a co-current
downward flow or counter-current upward flow. The possible
configurations for the hydrocarbon recycle are shown as dotted
lines in FIG. 4.
[0087] In another embodiment illustrated in FIG. 5, a system to
generate syngas in a three reactor system is disclosed. Similar to
FIG. 3, the first reactor converts the solid fuel to syngas in a
co-current gas-solid flow. The oxygen carrier donates oxygen to the
fuel to partially oxidize it to a high quality syngas. The gaseous
products from the second reactor enhance the gasification of the
solid fuels in the first reactor. The reduced metal oxide is sent
to the second reactor. The second reactor operates in a co-current
gas-solid contact pattern. The gaseous fuel is injected from the
top and exits the bottom. Both the reactors can be moving bed,
downer and/or a rotary kiln. The possible benefit of this operation
mode is that the fuel receives oxygen in the first step from a
relatively oxygen rich oxygen carrier. This leads to a higher
driving force and correspondingly a lower residence time
requirement and a smaller size of the reactor. In certain
embodiments, FeO.sub.bTiO.sub.2 enters the second reactor at the
top and exits as FeO.sub.cTiO.sub.2 at the bottom; the governing
equation for the system will be c<b. It should be noted that
though this embodiment is similar to FIG. 4 except for the contact
mode in the second reactor--it is co-current downwards as opposed
to counter-current upwards. Correspondingly the value of `c` at the
outlet of the second reactor in FIG. 3 is lower than in FIG. 5.
[0088] In the embodiment illustrated in FIG. 5, the operating
pressure of the system can vary from 1 atm to 100 atm. The
auto-thermal operation can have a temperature range from 1300 C to
700 C in the first and second reactor. The third reactor schematic
is similar to that in FIG. 3. The syngas quality will have a
H.sub.2/CO ratio between 1:1 to 1:4 and a combined CO.sub.2,
H.sub.2O content of less than or equal to 15%.
[0089] In another embodiment illustrated in FIG. 6, the
C.sub.1-C.sub.4 off-gas stream from various unit operations in a
liquid chemicals production refinery can be used to off-set the
power generation unit in the refinery complex. High quality heat
can be extracted and the carbon efficiency can achieve greater than
90% for this configuration. The inert percentage can vary between
20 to 80% for near auto-thermal operation. The oxygen can be
provided by air or pure oxygen from an air-separation unit, vacuum
distillation unit or oxygen tanks.
[0090] In another embodiment illustrated in FIG. 7, the first and
second reactors operate in a counter and co-current gas-solid flow,
respectively. The first and second reactor can be configured to
have a downward flow of the oxygen carrier. Similar to the case
disclosed in FIG. 2, the enhancing gas for the first reactor can
include up to 30 v/v % slip-stream from the outlet of the first
reactor. In another embodiment illustrated in FIG. 8, the light
hydrocarbons are recycled to the first reactor for greater
end-product production efficiency. The light hydrocarbons can range
between 0 to 100 v/v % of the enhancing gas stream.
[0091] From FIG. 7 in comparison to FIG. 1, the metal oxide enters
from the top and exits out of the second reactor, undergoing a
partial oxidation in the top-section and partial reduction in the
bottom section. The gaseous fuel will enter from the top of the
second reactor in this schematic. The gaseous product from the
first reactor will enter the middle of the second reactor at a
place such that all the methane is converted to a gas mixture
composing predominantly CO, CO.sub.2, H.sub.2, and H.sub.2O. In
certain embodiments, the metal oxide is represented as
FeO.sub.bTiO.sub.2 exiting the first reactor and enters the second
where TiO.sub.2 is the inert (85 wt %), and FeO.sub.cTiO.sub.2 is
the metal oxide state where the gaseous products are introduced in
the second reactor. This is the phase where the metal-oxide donates
oxygen and hence the governing equation is 0<c<b. The oxygen
carrier exits the second reactor as FeO.sub.dTiO.sub.2 and the
governing equation will be d>c as the metal-oxide will take in
oxygen from the gaseous product CO.sub.2 to convert it to CO and
H.sub.2O to convert it to H.sub.2. The difference in FIG. 7 and
FIG. 8 as compared to FIG. 1 and FIG. 2 is the mode of operation is
co-current instead of counter-current. This helps to contact the
injected methane to react with a higher oxidation state of iron
than a lower one. The catalytic effect of the iron-phase along with
the oxygen donation capacity will help in lowering the residence
time requirement. The syngas produced is high quality with <25%
of H.sub.2O and CO.sub.2 present and the H.sub.2/CO ratio can vary
between 1:1 and 4:1. The CO/CO.sub.2 ratio can be as high as 9, but
in general is greater than 3. The operating temperature range is
similar to the iso-thermal and the auto-thermal conditions
presented in FIG. 1 and FIG. 2. The operating pressures can vary
from 1 atm to 100 atm. The secondary metal oxide in the oxygen
carrier required for satisfying the heat balance can vary from 5%
to 95%.
[0092] In yet another embodiment, a variation of the first and
second reactors is used for producing syngas as illustrated in FIG.
9. The first reactor is used for providing enhancing gas by partial
or complete oxidation of a carbonaceous source including gaseous
fuels like methane, shale gas, liquid fuels like gasoline, diesel
or solid fuels like coal, biomass, solid waste etc. The product gas
from the first reactor is introduced to the second reactor as an
enhancer gas for solid fuel gasification. In certain embodiments, a
gaseous fuel such as natural gas or shale gas is also introduced to
the second reactor as illustrated in FIG. 10. The product gas from
first reactor is used to reform the gaseous fuel and/or partially
oxidize the oxygen carrier.
[0093] From FIG. 9, the carbonaceous fuel reacts in a
counter-current gas-solid contact mode with the oxygen carrier in
the first reactor. The objective of the first reactor is to enable
the metal-oxide to donate oxygen in a counter-current manner with
the fuel. The carbonaceous fuel is converted to gaseous product
which is sent to the top of the second reactor. In certain
embodiments, the oxygen carrier is of the form FeO.sub.aTiO.sub.2
where the TiO.sub.2 percentage is 85 w t % and the outlet
metal-oxide is of the form FeO.sub.bTiO.sub.2 and the governing
equation for the system is that 0.1<=b<a<=1.5.
[0094] As illustrated in FIG. 10, the gases produced in the second
reactor flow co-currently with the solid fuel and oxygen carrier
gasifying the solid fuel which is introduced in the middle of the
reactor. In the top half of the reactor, the metal-oxide oxygen
carrier gains oxygen by converting the CO.sub.2 and H.sub.2O to CO
and H.sub.2 respectively. In certain embodiments, the oxygen
carrier entering the second reactor has the formula
FeO.sub.bTiO.sub.2 and the oxygen carrier oxidation state where the
solid fuel is injected is FeO.sub.cTiO.sub.2, the governing
equation for this section of the reactor in terms of metal oxide is
1.33>=c>=b. The solid fuel is injected in the middle and the
syngas is extracted from the bottom of the reactor. In the lower
section of the second reactor, the metal-oxide will donate oxygen
to partially oxidize the solid fuel. In certain embodiments, the
oxygen carrier formula at the exit of the reactor is
FeO.sub.dTiO.sub.2, and the governing equation for this section of
the reactor will be 0.1<=d<=c<=1.33. The syngas generated
is of high quality with the H.sub.2/CO ratio ranging from 1:1 to
4:1. The % CO.sub.2 and H.sub.2O in this configuration is below 25%
v/v and the CO/CO.sub.2 ratio is >3 with an ideal value close to
8. The overall set of governing equations for this reactor scheme
in terms of the oxygen carrier oxidation states are
1.33>=c>=b and 0.1<=d<=c<=1.33. Note that the second
reactor follows a swing in the oxidation state of the oxygen
carrier. In the top section the oxygen carrier gains oxygen and in
the bottom section the oxygen carrier loses oxygen. This oxidation
state change is different from the configuration shown in FIG. 1,
FIG. 2, FIG. 7 and FIG. 8 as the syngas is produced with the oxygen
carrier constantly reducing in oxidation states in one direction,
while the fresh fully-oxidized Fe.sub.2O.sub.3 from the top of the
reducer reactor is used to fully oxidize the enhancing gas. This
embodiment enables the enhancing gas to react with the oxygen-rich
oxidation state of iron. In certain embodiments of FIG. 10, the
solid fuel can be introduced to the second reactor from the
top.
[0095] The carbonaceous fuel in the first and second reactor can be
reacted with the metal oxide counter-currently or co-currently in a
moving bed, a rotary kiln and/or a downer reactor design. The
oxygen carrier is sent to the third reactor to react with air or an
oxygen source to re-oxidize the reduced oxygen carrier to complete
the redox cycle. The heat can be extracted inside the third reactor
and from the spent air stream through heat-exchangers for
satisfying the parasitic energy consumption. In a variation of the
scheme shown in FIG. 10, the C.sub.1-C.sub.4 hydrocarbons from the
liquid chemical complex can be used as enhancing gas for the solid
fuel and/or to supplement the carbonaceous fuel source in the first
reactor.
[0096] In the following two configurations, the general circulation
path of the metal oxide solids does not change from prior
configurations. In prior configurations, the enhancer gas, a
CO.sub.2 and H.sub.2O mixture, is recycled from different
locations. Another possibility for the source of the enhancing gas,
an oxidized carbon form and/or oxidized hydrogen form, occurs from
the oxidation of at least either a carbon-containing or
hydrogen-containing source, including hydrocarbons, in a gaseous,
liquid, or solid form, or combination thereof. Examples include
natural gas, syngas, and tail gas from the F-T reactor, waste gas,
liquid fuels, coke, coal, biomass, and solid waste. The generation
of enhancing gas can occur in the third reactor, a new fourth
reactor, or a combination of the two. By independently generating
the enhancing gas, additional electricity or high-quality heat can
be produced while removing the necessity for gas cooling,
compression, and re-heating that is necessary for recycling
enhancing gas shown in previous configurations. Further, if
oxidation of the carbon-containing and/or hydrogen-containing
source occurs solely in the fourth reactor, no temperature
restriction exists and the oxidation, either partially or fully,
can occur at temperatures higher than the third reactor, thus
allowing for electricity production or heat transfer without worry
of metal oxide degradation.
[0097] In FIG. 11, the reduced metal oxide, FeO.sub.biO.sub.2,
exiting the first reactor can be split into two streams with split
fractions m and n where m+n=1. The split stream m enters the third
reactor while split stream n continues on its normal circulation
path to the second reactor. The split stream m containing
FeO.sub.biO.sub.2 mixes with FeO.sub.ciO.sub.2 exiting the second
reactor and reacts with the oxygen-source to regenerate the metal
oxide entering the first reactor, FeO.sub.aiO.sub.2.
[0098] In the third reactor or an entirely new reactor, a
carbon-containing source, hydrogen-containing source, or
combination thereof, can be oxidized, either partially or fully, to
CO.sub.s and H.sub.2O.sub.t where 0<s<2 and 0<t<2. If
oxidation of the carbon-containing and/or hydrogen-containing
source occurs in the third reactor, the outlet temperature of the
CO.sub.s/H.sub.2O.sub.t mixture is dictated by the reaction
temperature of the third reactor. The generation of enhancing gas
in the third reactor has another added benefit over prior
configurations by removing the need for a gas-solid separation
device as the enhancing gas can also assist in pneumatically
transporting the solids into the first reactor. If a fourth reactor
is used to generate the enhancing gas, the reactor temperature has
no restrictions. For example, the fourth reactor system can be a
gas turbine where natural gas combustion occurs to produce
electricity, CO.sub.2 and H.sub.2O. The electricity can be used
either internally or distributed to the grid while the CO.sub.2 and
H.sub.2O enters the first reactor. Here, adiabatic flame
temperatures are in the range of 2000 C to 3000 C, which would
deactivate the metal oxide if such temperatures were used in the
third reactor.
[0099] The source of oxygen is unimportant so long as the
carbon-containing, hydrogen containing source or mixture thereof
can be oxidized, either partially or fully, at temperatures near
the operating temperature of the first reactor where minimal
re-heat of the gas stream is necessary. Sources of oxygen can
include ambient air, oxygen-enhanced air, or oxygen derived from an
air separation unit, with the purity of oxygen necessary being
dependent upon the desired CO.sub.2 concentration in the first
reactor system. Both air and oxygen-enhanced air have a large
volume fraction of nitrogen present, thus diluting the CO.sub.2
stream exiting the first reactor. From literature, guidelines for
CO.sub.2 purity required for sequestration are typically greater
than 90% with nitrogen less than 7%. However, should the gas stream
containing CO.sub.2 and H.sub.2O exiting the first reactor be used
for purposes other than geological sequestration such as biological
fixation or re-utilization, then the CO.sub.2 purity will not be as
important so long as it is a concentrated source of CO.sub.2 with
impurities meeting the specifications required for the purpose. The
oxidation of a carbon-containing, hydrogen-containing source or
mixture thereof with an oxygen source in a separate vessel removes
the need for the cooling, compression, and re-heating of CO.sub.2
and H.sub.2O for recycle, but rather additional heat can be
generated for additional electricity generation or heat
transfer.
[0100] In FIG. 12, the four main reactors are equivalent to FIG.
11. In FIG. 12, the syngas, with a CO:H.sub.2 ratio between 1 and 3
is used to generate chemicals through either a Fischer-Tropsch or
methanol synthesis process. A portion of the C.sub.1-C.sub.4
hydrocarbons are recycled into the second reactor, which reduces
the overall consumption of pipeline/shale natural gas going into
the second reactor as compared to the configuration shown in FIG.
11.
[0101] In yet another embodiment, as shown in FIG. 13, a 3-reactor
system is used to convert solid and gaseous fuel to high purity
CO.sub.2/H.sub.2O and high quality syngas. In certain embodiments,
the gaseous and solid fuel can be coal and natural gas. The solid
and gaseous fuels are fed into the second reactor and converted to
syngas. A portion of unreacted devolatilized solid fuel formed in
the second reactor is discharged with the oxygen carriers from the
solids outlet. This devolatilized solid fuel (e.g. char) is then
separated from the oxygen carrier via a solid-fine separation
device. The oxygen carrier is discharged to the third reactor from
the solid-fine separation device while the devolatilized solid fuel
is sent to the first reactor. In the first reactor, the
devolatilized solid fuel reacts with the fully oxidized oxygen
carrier discharged from the third reactor to produce a gas stream
predominately consisting of CO.sub.2, CO, H.sub.2, and H.sub.2O.
The first and second reactor can operate in a co-current and
counter-current gas-solid contact mode. In certain embodiments,
Fe.sub.2O.sub.3 is the primary metal oxide used as the fully
oxidized oxygen carrier. The iron-based oxygen carrier enter the
first reactor in the oxidation state of FeO.sub.x and the outlet
state of FeO.sub.y where 1.5>=x>=y>=0.75. The iron-based
oxygen carrier then enter the second reactor as FeO.sub.y and exit
as FeO.sub.x where y>=z>=0.01. In certain embodiments, a
3-reactor system can be reduced to a two-reactor system with
char-separation and re-oxidation of the oxygen carrier as the
second reactor. The oxygen carrier sent to the third reactor is
re-oxidized with air and/or an oxygen-containing source. In the
case of an iron-based oxygen carrier, the oxygen carrier enters the
third reactor as FeO.sub.z and exits as FeO.sub.x where
1.5>=z>=y>=0. In certain embodiments, a recycled product
gas stream from the first reactor is used to separate the
devolatilized solid fuel from the oxygen carrier in the solid-fine
separation device and convey the devolatilized solid fuel to the
first reactor. The recycled product gas stream can also serve to
enhance the devolatilized solid fuel gasification in this
embodiment. Yet another embodiment, the product gas stream from the
first reactor is also recycled to the second reactor to enhance the
gasification of the solid fuel and reform the gaseous fuel
introduced. The amount of devolatilized solid fuel discharged from
the second reactor and transferred to the first reactor is based on
the thermodynamic, kinetic, and hydrodynamic properties of the
solid fuel in addition to the oxygen carrier properties and reactor
contact mode used. The devolatilized solid fuel discharge amount
can be adjusted to produce a flexible syngas H.sub.2/CO ratio and
for additional heat for the system and/or electricity generation.
This exemplary embodiment produces a sequester-ready stream of
CO.sub.2 without the need for additional gas-gas separation
techniques resulting in operating and capital cost reductions.
[0102] In yet another embodiment, a two-reactor system is used to
convert gaseous fuels such as natural gas and shale gas to high
quality syngas as illustrated in FIG. 14. In the first reactor, the
gaseous fuel is introduced and reacted with the oxygen carrier in a
co-current gas solid flow to produce high quality syngas. In the
second reactor, the reduced oxygen carrier is reacted with air
and/or other gaseous oxygen-containing reactants in a fluidized
bed/entrained bed in a co-current or counter-current gas-solid
contacting mode. The gas produced from the second reactor consists
of an oxygen-depleted gas stream. The gaseous fuel is introduced to
the first reactor at varying pressures ranging from ambient to as
high as required for the downstream reactor system. The first and
second reactors operate at the same or lower pressure than the
pressure of the gaseous fuel stream. A gas compressor is used to
provide high pressure oxygen-containing gas to the second reactor.
An expander is also placed on the product gas stream for the second
reactor to recuperate a portion or all of the energy used for
compressing the gas inlet stream. A similar philosophy is used to
design the expander-compressor coupling for the gaseous fuel inlet
and the gaseous product outlet stream. The high-pressure gaseous
fuel is subjected to a Joule-Thomson expansion in an expander after
pre-heating it to a suitable temperature. The work extracted from
the expander is used to compress the gaseous product by coupling
the expander with a compressor. Down-stream of the expander, the
gaseous fuel is heated before it is injected to the reactor system.
The gaseous product has a compressor which is coupled with the
expander from the inlet stream. Down-stream of this product stream,
another compressor can be used to pressurize the product gas to the
requisite value. This method follows an optimization pathway to
reduce the operating energy requirement for a gaseous fuel to
gaseous product system.
3. METAL OXIDE
[0103] The metal oxide used in the disclosed systems and methods
includes a primary metal oxide, a secondary metal oxide, and
optionally a tertiary metal oxide. In certain embodiments, the
metal oxide composition used is MeO.sub.x (primary)-Al.sub.2O.sub.3
(secondary) or FeO.sub.x (primary)-TiO.sub.2 (secondary) (where
0<x<1.5) to ensure overall high (>85%) ratio of
(CO+H.sub.2)/(CO+H.sub.2+H.sub.2O+CO.sub.2).
[0104] The primary metal oxides can consist of Ni, Cu, Mn, Mg, Co,
Zn, Mo, or any combination thereof. The primary metal-oxide can be
chosen from any of Fe.sub.2O.sub.3, NiO, Cu, Mn, Mg, Zn, Mo, Co
etc. The primary metal-oxide should be able to donate oxygen to the
fuel mixture (e.g., the primary metal oxide must have the capacity
to donate oxygen to selectively oxidize fuels to a mixture of
syngas).
[0105] The secondary metal oxide can be chosen from any of the
previously listed primary metal oxides or other metals such as Ti,
Al, Ca, etc, that provide support and enhance reactivity to
oxidation/reduction reaction with air and/or solids fuels at the
operating temperatures. In certain embodiments, the secondary
metal-oxide can be TiO.sub.2, Al.sub.2O.sub.3, Co, Cu, Mg, Mn, Zn,
or a combination thereof. The secondary metal-oxide can serve to
strengthen the primary metal-oxide and can enhance reactivity.
[0106] A tertiary component like Ag, Au, Ca etc may be added to
impart catalytic effect to the oxygen carrier. In some embodiments,
the tertiary metal oxides like Pt, Mo, Ag, Au, and Zn serve to
catalyze the tar decomposition and char gasification reaction for
solids fuels. It can also serve to catalyze gaseous fuel
decomposition reaction for methane and higher hydrocarbons.
[0107] The oxygen-carrier metal-oxide may contain a combination of
primary, secondary and tertiary metal-oxides in varying weight
percentages. The metal oxide can have varying percentages of the
primary metal oxide ranging from greater than 0 to 100 wt %, and
the secondary metal oxide compositions can vary from 0 to 100 wt %
as well. In addition to the primary and the secondary metal oxide,
a tertiary metal oxide can be added in varying percentages up to
less than 100 wt % as a promoter to enhance the reactivity of the
primary metal oxide.
[0108] The metal oxide particles can vary in size from 250 microns
to 5000 microns, with a density varying between 500 kg/m.sup.3 to
3000 kg/m.sup.3 depending on the method of synthesis. The method of
synthesis can vary from dry mixing and pelletizing via dry or wet
mixing methods. Other techniques can include extrusion,
co-precipitation, wet or dry impregnation, spray drying and/or
sintering after mechanical compression. Techniques like sintering
the synthesized metal-oxide or adding a binder with sol-gel
combustion can be used to increase the strength of the
metal-oxide.
[0109] In certain embodiments, the metal oxide has a primary
component as iron-oxide and a secondary component as
Al.sub.2O.sub.3. In certain embodiments, the metal oxide has a
primary component as iron-oxide and secondary component as
TiO.sub.2.
4. FEEDSTOCK
[0110] In general, the feedstock used in the disclosed systems and
methods can be classified into two separate categories: (1) solids
fuels and (2) gaseous fuels. Each embodiment described below
includes gaseous and the solid fuels used in conjunction. The solid
fuel can include coal (anthracite, lignite, bituminous and
subbituminous type), biomass, pet-coke, and/or solid
hydrocarbon-based waste products including municipal solid waste.
The gaseous fuels can include natural gas, gasified coal (syngas),
and/or the light hydrocarbon off-gas stream. The light hydrocarbon
stream can be a combination produced from various units of the
refinery operation including but not limited to hydrocracking,
hydrogenation, isomerization, etc. In cases where waste naptha and
residual oil (which are essentially liquid fuels) are present, they
can be used in conjunction with or replace either solid or gaseous
fuels. The overall reaction efficiency can be increased by (1)
pre-treating the solid fuels by de-moisturizing; and (2)
pre-heating the gaseous and liquid fuels as close to the reaction
temperature as possible. The solid fuel can be pretreated to ensure
that the distribution of the fuel is uniform in the reactor to
simulate the proper reaction method and the desired solids profile.
The gaseous and/or liquid fuels can be pre-heated as close to the
reaction temperature as feasibly possible.
[0111] In certain embodiments, the feedstock for this application
can be any solid fuel like coal, pet-coke, biomass etc. The
co-injection of a higher intrinsic H.sub.2 content fuel like
methane, flare gas etc. may be used in some variations. The solid
fuel can be pre-treated and the injection mode designed to ensure
uniform distribution of fuel through the radial reactor direction.
It is generally expected that the overall reaction efficiency will
increase with increasing the pre-heat temperature of the fuel to a
value near or above the reactor operating condition.
[0112] The present invention has multiple aspects, illustrated by
the following non-limiting examples.
5. EXAMPLES
[0113] Examples 1-8 refer to the exemplary embodiment of the
two-reactor system with a gas compressor and expander for
significant increase in process efficiency. The example studies
performed herein provide a detailed analysis of the expected
efficiency gain from this disclosed exemplary embodiment.
[0114] This method can be flexibly applied to a combination of
high-pressure gaseous fuel and a high temperature reactor system
for the concept of chemical looping or otherwise. The system uses a
synergy between a compressor-expander device to recover a portion
of the energy expended in compressing the gaseous fuel or
utility.
[0115] Syngas production systems play a role in the overall cost of
a Fischer-Tropsch plant configuration for liquid fuel production
from natural gas. Typical syngas production systems include
steam-methane reforming, auto-thermal oxidation and partial direct
oxidation of the gaseous fuel. In reference to the various
configurations disclosed above and this analysis comprehensively
investigates the operating power requirements for running the
chemical looping syngas production system in an auto-thermal
operation mode. As an example, the syngas produced needs a pressure
and temperature equivalent to the downstream reactor system; out of
the syngas production system. This particular analysis explores
applications and benefits of the method disclosed.
[0116] Establishing a baseline for partial oxidation energy
requirements: The partial direct oxidation system uses an Air
Separation Unit (ASU) to supply 99% pure oxygen for partially
oxidizing the methane in a combustion chamber. The combustion
chamber is operated around 1400 C. The inlet pressure for the
oxygen as an example is expected to be around 30 atmospheres which
gives a baseline for comparing the chemical looping case
against.
[0117] The reaction chemistry targeted in direct partial oxidation
is shown in Reaction 1:
CH.sub.4+0.5O.sub.2->CO+2H.sub.2 (1)
[0118] This reaction is slightly exothermic. Reaction 1 shows the
stoichiometric ratio of O.sub.2/CH.sub.4 to be 0.5. The actual
ratios used can vary from 0.73 to 3.0 depending on the combustion
chamber design and methane conversions targeted. For this analysis
a conservative number of 0.73 is used. The air separation unit is
typically designed for two purities: 95% pure O.sub.2 and 99% pure
O.sub.2. The energy consumption for the O.sub.2 produced from an
ASU can be found in a variety of literature sources between 160 to
250 kWh/met tonne O.sub.2 (95 and 99% purity). The power
consumption for a 99% pure stream of O.sub.2 pressurized to 30 atm
is 378.34 kWh/met tonne O.sub.2. If 1 kmol/hr of CH.sub.4 is
selected as the basis flow for analysis, the amount of O.sub.2 will
be 0.73 kmol/hr. The operating power can be calculated as 9.08 kWe,
assuming the heating value of CH.sub.4 to be 52582 kJ/kg. This
value when translated on a per unit fuel input basis becomes 3.815%
kWe/(kW.sub.th CH.sub.4) processed. The comparison of the chemical
looping cases operating energy consumption is computed with this
number as the baseline.
Example 1
Chemical Looping Reactor Operated at 1 Atm ("Case 1")
[0119] The chemical looping cases are analyzed assuming the outlet
pressure of syngas produced is 30 atm. The inlet pressure of
methane is taken to be 30 atm. The overall analysis indicates that
the operating power requirement is a function of the operating
pressure of the chemical looping system. The analysis starts from a
basic set-up analysis of the chemical looping scheme being operated
at 1 atm, in two modes of heat-integration as shown in FIGS. 15 and
16. For the chemical looping case with heat-integration scheme 1,
the methane is expanded to a pressure of 1.2 atm in an expander
which recovers work. This expansion is coupled with a compressor on
the syngas outlet stream from the reducer. The compressor
compresses the syngas produced at near atmospheric conditions from
the reducer reactor. The basis of analysis is 1 kmol/hr CH.sub.4.
The reactor modeling is optimized as per the earlier discussion to
produce the syngas in accordance with performance results for an
auto-thermal operation and an optimal % of secondary metal support.
The air flow is set at 15% excess of the stoichiometric requirement
for full re-oxidation of metal oxide. An air-compressor is used to
compress the air to overcome the pressure drop in the combustor and
for transporting the solids to the reducer reactor. The reducer
reactor inlet temperature is set at 1190 C. The product gases for
the reducer and the combustor reactor are assumed to be exiting at
1190 C. The methane inlet temperature is 600 C. The methane stream
loses heat in the expander, and is heated again to 600 C. The
air-pre-heat temperature is also set to 600 C. The heat-extraction
takes place from the air-outlet and the syngas streams to satisfy
the heating requirement and produce work to off-set the compressor
power requirements. The syngas compressor for Case 1, heat
integration scheme 1 compresses the syngas to 30 atm using a
pressure ratio of 1.6. The compression efficiencies used are 0.86
for the polytropic efficiency and 0.98 for the mechanical
efficiency. The cooling water temperature for each stage is set at
69 F. The compression method used for modeling the power
requirements is polytropic using ASME method. The heat-extraction
specification after heat-integration is specified in Table 1
below.
TABLE-US-00001 TABLE 1 Available hear for Case 1, heat integration
Scheme 1, FIG. 15 Available heat T.sub.in Unit Description Q
(cal/sec) (C.) T.sub.out (C.) P (bar) B17(S36) Spent Air Stream
3972 1190 398 1 B2 Combustor (heat extrn) 69 1190 1190 1
[0120] The parasitic energy consumption is shown in Table 2
below.
TABLE-US-00002 TABLE 2 Parasitic energy consumption for Case 1,
FIG. 15 Parasitic operating energy consumption Unit Description
Energy % kWe/kWth B7 Syngas compressor 6.96 kW.sub.e 2.98 B3 Air
compressor 0.13 kW.sub.e 0.056
[0121] The last column in Table 2 converts and compares the
operating energy requirement from 1 kmol/hr CH.sub.4 to %
kWe/(kW.sub.th CH.sub.4 basis).
Example 2
Chemical Looping Reactor Operated at 1 Atm, FIG. 16 ("Case 2")
[0122] This case analyzes the chemical looping at 1 atm case
without the expander-compressor coupling from Case 1. The schematic
for the case is shown in FIG. 16. This case assumes that the
CH.sub.4 is pre-heated to 600 C. The air outlet is used to extract
heat in two stages: a primary pass reducing the temperature to 400
C and a secondary pass further reducing the stream temperature to
170 C. The syngas compressor pressurizes the syngas to 30 atm using
similar assumptions to those stated in Case 1. The overall heat
extracted is summarized in Table 3.
TABLE-US-00003 TABLE 3 Available heat for Case 2, Heat integration
Scheme 2, FIG. 15 Heat extraction source B5 Syngas outlet heat
extrn -4674 899 150 1 B17 Spent air heat extraction -4696 1190 399
1 B19 Spent air ht extrn secondary -1238 399 169 1
[0123] The total operating energy requirement is given in Table
4.
TABLE-US-00004 TABLE 4 Parasitic energy consumption for Case 2,
heat integration scheme 2. Parasitic operating energy consumption
Unit Description Energy % kWe/kWth B7 (8 stages) Syngas compressor
8.97 3.83 B3 Air compressor 0.58 0.05
[0124] The last column gives a % kWe/(kW.sub.th CH.sub.4 input)
value. If we compare the values in Table 2 and Table 4 we can see
that the heat-integration Scheme 1 gives a lower operating cost
which outweighs the energy lost in heating up the methane gas
before the expander unit. Hence, for the cases at higher pressure
the compressor-expander coupling unit is added wherever feasible.
The unit is used to couple the methane inlet stream with the syngas
outlet stream. The unit is also used to couple the air-inlet stream
with the air-outlet stream to effectively utilize the energy and
reduce the operating power requirements.
Example 3
Chemical Looping Reactor Operated at 5 Atm, FIG. 17 ("Case 3")
[0125] The schematic for this case is shown in FIG. 17. The general
schematic for this case is similar in terms of the number of units
used to analyze higher pressure cases (Case 4 to Case 7). The
methane inlet condition for this example is 30 atm. This stream is
heated to 600 C, before being subjected to a Joule-Thomson
expansion to reduce the general stream pressure to a value slightly
above the system operating pressure. This expander is coupled with
a compressor to save on the overall syngas compression cost. The
methane inlet stream is heated to 600 C as it loses temperature
after the expander. The syngas produced is compressed using the
compressor coupled with the expander. This is followed by a syngas
compressor which compresses the syngas to 30 atm. The
specifications for this compressor are the same as those specified
for Case 1. The air inlet stream is initially compressed by a
compressor which is coupled to an expander downstream of the
combustor reactor. This leads to two scenarios where the expander
supplies enough power to compress the inlet air beyond the
combustor operating pressure. In an alternative scenario, the
pressure of the air-stream at the outlet of the coupled compressor
is lower than the combustor operating pressure. In this case an
additional syngas compressor is used to compress the air to a
pressure higher than the combustor operating pressure. The
assumptions on this compressor are similar to those in Case 1
analysis. The compression typically increases the temperature of
the gas-stream. The air-stream is pre-heated to 600 C. If the
compressor outlet temperature is lower than 600 C, a heat-exchanger
is added to heat the air stream to 600 C. The air outlet stream
goes through an expander and then through two heat-extraction
passes. The primary pass cools down the stream to 400 C and the
secondary pass cools down the stream to 170 C. The expander
decompresses the gas to a pressure value slightly higher than 1 atm
(.about.1.2 atm). The syngas outlet stream also passes through two
passes with similar specifications to the air outlet stream and
passes through a condenser stream before the coupling compressor
and the syngas compressor.
[0126] A break-down of the available heat for the Case 3 is given
in Table 5.
TABLE-US-00005 TABLE 5 Available heat for Case 3, 5 atm, FIG. 17
Available heat Unit Description Q (cal/sec) T.sub.in (C) T.sub.out
(C) P (bar) B17(S36) Spent air stream 2567 1190 398 5 B2 Combustor
(heat extrn) 875 1190 1190 5
[0127] Table 6 shows the operating energy consumption for Case
3.
TABLE-US-00006 TABLE 6 Parasitic energy consumption for Case 3,
heat integration scheme 2. Parasitic operating energy consumption
Unit Description % kW.sub.e/kWth B7 Syngas compressor 3.46 kWe
1.480573 B3 Air compressor 0 kWe 0
[0128] The expander-compressor coupling for the air compressor is
sufficient to offset the amount of energy required to compress the
air inlet stream slightly above 5 atm. The compression cost is
reduced by 55% over the Case 2 scenario. As compared to the ASU
case the cost is around 60% lower on the same basis.
Example 4
Chemical Looping Reactor Operated at 10 Atm, FIG. 17 ("Case 4")
[0129] This case uses the same schematic set-up described in Case
3. The expander for syngas stream has a discharge pressure of 10.2
atm and the air compressor outlet pressure is desired to be >10
atm. The advantage of this scheme is that the syngas compressor
requirements will go down as the syngas is produced at a pressure
of 10 atm. The available heat for this case is shown in Table 7
below.
TABLE-US-00007 TABLE 7 Available heat for Case 4, 10 atm, heat
integration Scheme 2, FIG. 3 Available heat Unit Description Q
(cal/sec) T.sub.in (C) T.sub.out (C) P (bar) B17(S36) Spent air
stream -1901 1190 398 10 B2 Combustor (heat extrn) -1986 1190 1190
10
[0130] Table 8 shows the operating energy consumption.
TABLE-US-00008 TABLE 8 Parasitic energy consumption for Case 4
Parasitic operating energy consumption Unit Description %
kW.sub.e/kWth B7 Syngas compressor 2.27 kWe 0.972 B3 Air compressor
0 kWe 0
[0131] For this case the coupling is sufficient to offset the
energy requirements for the air-compressor and the syngas
compressor power requirement goes down by 35% over the 5 atm
case.
Example 5
Chemical Looping Reactor Operated at 15 Atm and FIG. 17 ("Case
5")
[0132] The syngas compressor requirement goes down further as the
base pressure after coupling is higher than 15 atm. This case shows
a higher air compressor energy demand than the Case 4. The coupling
is not sufficient to increase the pressure of the air to greater
than 10 atm. An additional single stage air-compressor is used to
compress the air to greater than 10 atm. The available heat for
Case 5 is shown in Table 9.
TABLE-US-00009 TABLE 9 Available heat for Case 5, 15 atm, FIG. 16
Available heat Unit Description Q (cal/sec) T.sub.in (C) T.sub.out
(C) P (bar) B17(S36) Spent air stream -1564 1190 398 15 B2
Combustor (heat -3378 1190 1190 15 extrn) CO.sub.2 stream heat
Syngas (heat -1224 951 150 15 extraction)
[0133] The parasitic energy consumption for Case 5 is shown in
Table 10 below.
TABLE-US-00010 TABLE 10 Parasitic energy consumption for Case 5
Parasitic operating energy consumption Unit Description %
kW.sub.e/kWth B7 Syngas compressor 1.57 kWe 0.671821 B3 Air
compressor 0.38 kWe 0.162606
[0134] As compared to Case 4 the syngas compressor power
consumption by itself goes down by 30%. The net including air
compressor reduces the cost by 15%.
Example 6
Chemical Looping Reactor Operated at 20 Atm, FIG. 17 ("Case 6")
[0135] This investigation analyzes the case when the chemical
looping reactors are operated at a pressure of 20 atm. In this
analysis the air compressor parasitic energy requirement is
expected to be higher than that in Case 5. The corresponding value
of the syngas compressor energy requirement is supposed to go
lower. The schematic is similar to the one used in FIG. 17. The
available heat and the parasitic energy requirements are shown in
Table 11 and Table 12.
TABLE-US-00011 TABLE 11 Available heat for Case 6, 20 atm, FIG. 17
Available heat Unit Description Q (cal/sec) T.sub.in (C) T.sub.out
(C) P (bar) B17(S36) Spent air stream -1347 1190 398 20 B2
Combustor (heat -4221 1190 1190 20 extrn) CO.sub.2 stream Syngas
(heat -1224 951 150 20 heat extraction)
TABLE-US-00012 TABLE 12 Parasitic energy consumption for Case 6
Parasitic operating energy consumption Unit Description %
kW.sub.e/kWth B7 Syngas compressor 0.78 kWe 0.34 B3 Air compressor
1.07 kWe 0.46
[0136] In case 6 at 20 atm, the syngas compressor requirement is
50% lower than the case 5. This is expected as a higher starting
point is given in terms of pressure for the syngas compressor. The
air power requirement is 20% higher than case 4. The overall
efficiency is around 5% reducing in the compression energy
requirement over case 5.
Example 7
Chemical Looping Reactor Operated at 25 Atm, FIG. 17 ("Case 7")
[0137] This case is expected to further reduce the syngas
compression cost, but the air compression cost increases. Operating
the chemical looping reactors at 25 atm shows the case where the
air compression cost increase outweighs the decrease in syngas
compression cost. The available heat and the parasitic energy
consumption are shown in Table 13 and Table 14.
TABLE-US-00013 TABLE 13 Available heat for Case 7, 25 atm, FIG. 17
Available heat Unit Description Q (cal/sec) T.sub.in (C) T.sub.out
(C) P (bar) B17(S36) Spent air stream -1189 1190 398 25 B2
Combustor (heat -4221 1190 1190 25 extrn) CO.sub.2 stream Syngas
(heat -4999 951 150 25 heat extraction)
TABLE-US-00014 TABLE 14 Parasitic energy consumption for Case 7
Parasitic operating energy consumption Unit Description %
kW.sub.e/kWth B7 Syngas compressor 0.77 kWe 0.328 B3 Air compressor
2.14 kWe 0.91573
[0138] The overall net compression energy requirement goes up by
50% over that required in case 6. The syngas compression cost
decreases minimally, while the air compression cost increases by
around 50% value.
Example 8
Chemical Looping Reactor Operated at 30 Atm, FIG. 18 ("Case 8")
[0139] This case investigates the operating cost at a pressure of
30 atm. This case follows a schematic as shown in FIG. 18. As the
chemical looping reactors are operated at 30 atm, the syngas
compression cost goes down to zero. The air compression cost
however will outweigh this decrease despite the coupling effect
with a high temperature spent-air stream. The available heat and
the parasitic energy consumption are shown in Table 15 and 16.
TABLE-US-00015 TABLE 15 Available heat for Case 8, 30 atm, FIG. 17
Available heat Unit Description Q (cal/sec) T.sub.in (C) T.sub.out
(C) P (bar) B17(S36) Spent air stream -1167 1190 398 30 B2
Combustor (heat extrn) -2389 1190 1190 30
TABLE-US-00016 TABLE 16 Parasitic energy consumption for Case 8,
Heat integration scheme 2 Parasitic operating energy consumption
Unit Description % kW.sub.e/kWth B7 Syngas compressor 4.65 kWe
1.989787185 B3 Air compressor 0 kWe 0
[0140] As compared to an ASU this energy requirement is still
around 48% lower. The air compressor follows similar specifications
to those defined in Case 1 and Case 2, with a polytropic efficiency
of 0.86 and a mechanical efficiency of 0.98.
[0141] Discussion Overall analysis: The pressure sensitivity in
terms of operating cost is shown in Table 17.
TABLE-US-00017 TABLE 17 An overall analysis for the example cases
is shown where the disclosed optimization method is applied Savings
(% Pressure Syngas comp Air comp over baseline (atm) (kWe) (kWe)
Total (kWe) % kWe/kWth case) Case 1 1 6.96 0.13 7.09 3.033 79.62967
Case 2 1 9.3 0.13 9.43 4.03520283 105.9108 Case 3 5 3.46 0 3.46
1.48057283 38.86018 Case 4 10 2.27 0 2.27 0.971358475 25.49497 Case
5 15 1.57 0.38 1.95 0.834426884 21.90097 Case 6 20 0.78 1.07 1.85
0.791635762 20.77784 Case 7 25 0.78 2.14 2.92 1.24950077 32.7953
Case 8 30 0 4.65 4.65 1.989787185 52.22539 Case 0 ASU-30 atm 9.08 0
9.08 3.81
[0142] The energy percentage in the last column is an energy
consumption ratio of energy consumed for case to ASU, and is
computed as a % of the overall % kWe/kWth for each case as compared
to the ASU case. As we go in blocks of 5 atm from 1 atm to 30 atm,
the overall percentage goes through a minima in-between 15-20 atm.
This same pattern can be observed in FIG. 18. The overall analysis
is a method of analysis which can be applied to a combination of
fuels and corresponding pressurized products. The optima are
obtained by a consideration of the expander-compressor coupling and
the overall heat-integration schematic.
Example 9
Specific Operating Condition
[0143] The explanation of the specific operating condition is
initially shown in an iso-thermal condition in the reducer. The
reducer system is a downward co-current moving bed. The metal-oxide
combination chosen is Fe.sub.2O.sub.3 as the primary component and
Al.sub.2O.sub.3 as the secondary component. FIG. 20 shows a sample
reducer reactor configuration investigated for the optimal reactor
operation condition, configuration and analysis. The model used is
a Gibbs-energy minimization reactor. The minimization of Gibbs free
energy gives the thermodynamic performance of the system. The Gibbs
reactor simulates the co-current downward flowing moving bed
performance. The temperature swing for the oxygen carrier in an
adiabatic reactor is limited between 1250 C and 750 C. If the
temperature swing is set between 1200 C and 900 C, the performance
is shown in FIG. 21. This condition is analyzed at a 50% HHV ratio
of CH4 and PRB Coal. The steam injection is 1 mol H.sub.2O per mole
of Carbon entering the reducer. The feasible area of operation for
satisfying the heat balance is determined by the value of heat duty
being either zero or negative. For example, when the heat duty is
negative, the temperature of the solids will be higher than 900 C.
If the heat duty is positive, the temperature will be lower than
900 C. The heat-duty may be manipulated by increasing or decreasing
the support content in the metal-oxide, increasing or decreasing
the solids flow and changing the pre-heat temperatures. A similar
scan at temperatures of 950 C and 1000 C give the values listed in
Table 18 and plotted on FIG. 22. Expectedly, it was found that
higher the temperature of analysis, higher is the heat-duty
required. If the fuel pre-heat and the oxygen carrier inlet
temperature and support weight % are kept constant, then the
neutral heat-duty condition is only satisfied when going to a
higher solids circulation rate. This leads to a higher oxygen
transfer to the syngas produced, which converts more CO to CO.sub.2
reducing the carbon utilization of the system.
TABLE-US-00018 TABLE 18 Illustration of reactor performance for
different solids outlet temperatures Solids H.sub.2/CO M % CH.sub.4
T out conversion ratio value Syngas Fe.sub.2O.sub.3/C conversion
900 33% 2.03 1.32 65% 0.47 91% 950 33% 1.94 1.23 69% 0.6 97% 1000
33% 1.83 1.08 72.10% 0.75 99%
[0144] FIGS. 23a, 23b, 24a and 24b show a sensitivity analysis of
the effect of steam addition to the reducer reactor. The steam
addition is to increase the syngas H.sub.2/CO ratio and the syngas
M value (M:(H.sub.2--CO.sub.2/CO+CO.sub.2) to the required values
for some of downstream processes. Note that there is no significant
change in the net-duty curve beyond a certain point of water
injection. Note that the H.sub.2/CO ratio increases with increasing
H.sub.2O addition while the amount of CO produced decreases
correspondingly. The point beyond where the heat-duty does not show
a significant deviation for incremental values of steam injection
is a special point of interest around which the system can be
designed to maximize the carbon utilization.
[0145] Analyzing the regions before and after this point shows two
distinct trends. Before this point, the oxygen for production of
H.sub.2 and CO from the solid fuel comes exclusively from the
oxygen carrier. Going beyond the so-called critical point of steam
injection, the amount of oxygen donated by the oxygen carrier stays
more or less constant. The amount of CO produced from the fuel is
constant. The additional steam starts donating oxygen to the CO,
converting it to CO.sub.2, while producing more H.sub.2. This
increases the H.sub.2/CO ratio to the requisite value at the cost
of converting CO to CO.sub.2. The CO.sub.2 itself can be removed in
a standard rectisol type configuration to give a M value which is
very close to the H.sub.2/CO ratio. As shown in FIG. 23, the
H.sub.2/CO ratio increases with addition of more water. Note that
the M value remains more of less the same. As shown in FIG. 24b,
the % syngas drops as predicted by the hypothesis mentioned above.
The specific operating condition is designed around the system for
this particular injection condition uniquely by a combination of an
iron based metal-oxide and the co-current downward flow and solid
fuel. This condition is unique in the sense that the operation is
designed to harness the sweet spot of maximizing the oxygen
donation from the oxygen carrier for producing syngas rather than
the water-gas shift reaction. The importance of oxygen donation
from the oxygen carrier is in the fact that it donates oxygen in a
reducing environment without oxidizing the CO to CO.sub.2. A
conventional coal-gasifier using only a water gas shift based
system gives a fuel to syngas carbon conversion values of between
25 and 35%. This is because of the fact that the water-gas shift
reaction provides an oxidizing environment which increases the
H.sub.2 content of the syngas at the cost of converting CO to
CO.sub.2. The role of oxygen carrier in providing the oxygen can be
controlled by a combination of the oxygen carrier material, the
reactor operation mode and the amount of steam injection. The
conditions claimed satisfy the heat-balance for an iso-thermal
operation.
[0146] The analysis presented so-far held a constant temperature to
evaluate the effect of various parameters for deriving a specific
condition and confirmation of the theory of maximizing the oxygen
donation by the oxygen carrier.
[0147] FIGS. 25 to 32 analyze the behavior of the looping reducer
for a commercial prototype adiabatic system. The fuel system is
maintained similar to the iso-thermal case analyzed above for ease
of understanding. The commercial reactor system is typically
designed with refractory lining of the reactor wall to have an
adiabatic operating condition. For this analysis, this will
accurately represent a realistic commercial operation. This
evaluating the GIBBS reactor system under the condition of heat
duty being near zero instead of a constant temperature. The
variable which changes with varying iron-oxide flow-rate is the
temperature coming out of the system. The temperature swing is
within the specified oxygen carrier limits of 1250 C and 750 C. The
system behavior is analyzed with steam addition as a parameter,
while the other inlet conditions remaining similar to the
isothermal cases analyzed in FIGS. 21-24.
[0148] The first condition to be isolated for an acceptable reactor
performance is the absence of Carbon deposition in the reducer. The
set of operating curves are shown as a function of the amount of
water-injected per mole of Carbon input to the reducer. The
conditions where there is no carbon-deposition are acceptable for
evaluating the system performance. FIG. 31 shows that for the same
amount of fuel injection, the minimal amount of oxygen-carrier flow
required for no-carbon deposition decreases with increase of water
flow for increasing temperature. In other words, the system can
have a lower solids flow and a lower reducer outlet temperature
with higher water injection for the same amount of carbon input to
the system.
[0149] FIG. 25 and FIG. 26 plot the CO/CO.sub.2 curves and
H.sub.2/CO value in syngas for varying iron-oxide flows as a
function of steam injection per mole of Carbon. It can be seen that
in the regions where there is no carbon deposition, the H.sub.2/CO
ratio increases with increasing steam injection. The CO/CO.sub.2
ratio decreases with increased steam injection, pointing to the
presence of the so-called critical point (claimed specific
operating condition) beyond which the oxygen donation from the
steam to the fuel occurs. The H.sub.2 produced increases as a
function of the increasing steam injection as shown in FIG. 30. The
CO.sub.2 produced also increases while the CO decreases as a
function of increased steam injection. The temperature outlet
profile does not show a significant difference as a function of
steam injection. Using a standard post-combustion rectisol-type
system downstream of the unique looping reactor helps decide the
optimal system goals for this case. The goals in accordance with
previous iso-thermal analysis are reaching a reasonable H.sub.2/CO
ratio while maximizing the CO/CO.sub.2 ratio. The specific point of
operation would be close to the point where the critical point of
transition occurs wherein the oxygen transfer from steam to the
fuel begins dominating over that from steam to the oxygen carrier
and the oxygen carrier to the fuel. This specific condition can be
achieved only with the unique combination of the oxygen carrier,
heat management, steam injection, fuel feedstock and a downward
co-current contact between the oxygen carrier and the fuel type
reactor.
[0150] In a different configuration, illustrated in FIG. 33 the
CO.sub.2 separated from the post-combustion capture can be recycled
with after moisture removal. The CO.sub.2 recycle is to improve the
condition to suppress the Carbon deposition, satisfying the
heat-balance conditions at a lower circulation rate. Another
advantage of this recycle CO.sub.2 injection is to suppress the CO
conversion to CO.sub.2. FIG. 34 and FIG. 35 show the system
performance with different levels of CO.sub.2 injection at a steam
to fuel carbon molar ratio of 0.5. The figures illustrate that the
addition of a small amount of CO.sub.2 helps suppress the Carbon
deposition to a higher extent. Higher addition has no significant
benefits as shown in FIGS. 33 and 34. A small addition of CO.sub.2
as disclosed in this configuration, in combination with the unique
operating condition disclosed earlier for maximizing the oxygen
transfer has some benefits over in cases where CO.sub.2 can be
recycled feasibly.
[0151] FIG. 36 and FIG. 37 show a case where a different amount of
natural gas and coal is added. The case essentially demonstrates
the fact that the specific operating condition can be claimed for
different combination of various fuels claimed. The metal-oxide
composition used is Fe.sub.2O.sub.3--Al.sub.2O.sub.3. In this case
the coal HHV contribution to the fuel mix is 45%. The trends shown
are similar to those analyzed for the base case above. The areas
for no carbon deposition are isolated (FIG. 36) and the
corresponding performance ratios for the systems are analyzed. A
similar behavior for the oxygen transfer from the oxygen carrier is
observed. Initially at low steam injections, the oxygen transfer
from the oxygen carrier is dominant. Later, the oxygen transfer
from steam starts to dominate. This case shows the specific
operating condition to be present at a lower steam flow-rate than
the 50% HHV case. This is attributed to the fact that the CH.sub.4
amount decreases per mole of carbon which leads to a drop in
intrinsic H.sub.2 content from the fuel mixture. A syngas
composition which is suitable for downstream application is
obtained at a higher solids flow and lower steam injection
requirement. This case shows that using a combination of any of the
oxygen carriers disclosed earlier and the unique control offered by
a co-current downward moving bed reactor can be applied with a
specific amount of steam injection to derive the claimed specific
condition for syngas production. This case shows that a similar
behavior can be expected for different combinations of fuel
injections with the critical point being present at differing
operating conditions depending on the coal type, fuel type and the
intrinsic H.sub.2 content of the fuel.
[0152] FIGS. 39, 40 and 41 show the overall performance curve using
Illinois #6 and PRB coals in combination with methane input at the
unique operating conditions for the optimal steam injection using
the Fe.sub.2O.sub.3--TiO.sub.2 based particles as example. This
case is analyzed to verify the claim of a different metal-oxide
being used for the operating condition. It should be seen that
presence of a H.sub.2-rich co-injection fuel helps improve the
carbon utilization of the system by minimizing the steam usage. A
scan of different HHV'S ratios of coal and methane fuel injection
is analyzed. Application of the given condition can improve the
carbon utilization significantly while maintaining the syngas
quality and eliminating the need for molecular oxygen transport
from an air-separation unit. The application of the specific
condition is true for all the fuel feedstock, oxygen carrier
compositions specified.
[0153] FIG. 42 shows a configuration in which the reducer is
operated with fuel injection. The general principle for operating
an oxygen carrier based looping unit is the fact that it should be
auto-thermal in operation. If carbon capture from a chemical
looping with heat-balance is considered, it leads to two design
configurations. One configuration has a part of fuel burned in the
oxidizer reactor to attain the auto-thermal operating condition.
This configuration can operate at a lower solids circulation rate,
which improves the syngas yield near the critical point of steam
injection. The configuration may require dilute CO.sub.2 capture
from the oxidizer exhaust stream. An alternative configuration
avoids this dilute CO.sub.2 capture, by increasing the solids
circulation rate for heat-balance purposes. This ensures that the
excess CO.sub.2 is obtained as a part of the CO.sub.2 rich syngas
stream, reducing the cost of separation of CO.sub.2.
[0154] FIG. 43 shows the application of the heat balance concept
and its effect on syngas capacity using
Fe.sub.2O.sub.3--Al.sub.2O.sub.3 particles as an example. The
Fe.sub.2O.sub.3--Al.sub.2O.sub.3 particles have shown equivalent
performance to the TiO.sub.2 particles used in earlier examples for
this specific application. The tables 19 and 20 show the conditions
investigated as demonstrations for applying the heat balance
concept while burning some amount of a fuel like natural gas for
electricity production.
TABLE-US-00019 TABLE 19 Carbon utilization per mole of C input to
the system shown in FIG. 43 M value Carbon % C C Coal CH.sub.4 Coal
CH.sub.4 H.sub.2O after Case CO utilization capture total reducer
reducer Oxidizer NGCC reducer rectisol 1 0.35 35 90 1 0.9 0 0 0.1 2
2.01 2 0.564 51. 81.8190 1.11 0.55 0.35 0.1 0.1 1 2.01 3 0.564
54.12 90 1 0.65 0.25 0 0.1 1.1 2.01 4 0.611 60.47 90 1 0.65 0.25 0
0.1 1.1 1.7 5 0.6 60.72 95 1 0.65 0.3 0 0.05 1 2.01 6 0.66 66.41 96
1 0.65 0.3 0 0.05 0.7 1.7
TABLE-US-00020 TABLE 20 Material balance for 1 mole of C input to
the system shown in FIG. 43 Case No Fe.sub.2O.sub.3 Al.sub.2O.sub.3
CO CO.sub.2 H.sub.2 H.sub.2O CH.sub.4 T.sub.out H.sub.2/CO M
H.sub.2O/C 3 0.848 5.31 0.528 0.369 1.065 1.111 0.0018 944.146385
2.01 0.774 1.099 4 0.806 5.05 0.594 0.300 1.028 0.742 0.0050
927.508098 1.73 0.813 0.727 5 0.848 5.31 0.5928 0.352 1.18 0.982
0.0050 919.259518 2.01 0.884 1.04 6 0.750 4.70 0.664 0.265 1.151
0.5879 0.020 889.774144 1.734 0.954 0.66
[0155] The M values and the H.sub.2/CO values are similar to those
produced by conventional gasification systems. The carbon
utilization is higher because of application of the specific
operating condition. FIG. 43 shows the increase of syngas
production per mole of total carbon in the system over the baseline
coal gasification technology. A similar performance was obtained
for combinations of other H.sub.2 deficient fuels which need some
steam injection for increasing the H.sub.2/CO ratio.
[0156] FIG. 44 shows a methodology to reduce the operating energy
requirements of a chemical looping unit. The example shown
illustrates the application of the methodology in conjunction with
the special operating condition used. The methodology is used with
the configurations disclosed above for reducing the operating
energy requirement wherever applicable. This methodology involves
coupling the air streams and the natural gas streams used for power
generation purposes. The typical downstream syngas processing units
operate at a pressure between 30 and 80 atm. The chemical looping
unit has two main reactors namely the reducer and the oxidizer. The
reducer produces the syngas which needs to be compressed to the
requisite downstream processing pressure. The following methodology
disclosed in FIGS. 44 and 45 shows a way to analyze the looping
problem and identify the optimal pressure for reducing the
operating energy. The disadvantage of operating at a pressure close
to the downstream syngas processing system is that the air has to
be compressed to that high value while the syngas compression cost
diminishes. On the other hand, operating at pressure close to
atmospheric will diminish the cost of air-compression while
significantly increasing the cost of syngas compression. It is
expected that the actual minima would lie somewhere in middle of
the pressure range scanned.
[0157] The methodology of coupling the air stream and the syngas
stream with incoming natural gas stream is unique and has certain
intrinsic advantages. The air is initially compressed and a large
volume of air remains unreacted. The options for recovery of air
from this stream include the heat extraction from a rankine-cycle
type steam cycle coupled with a Brayton-cycle type gas turbine. The
gas-turbine may have a certain let-down which can be designed to
offset some of the inlet compressor duty. The net compressor duty
would then be the excess remaining to be supplied by the air
compressor in terms of an effective net-duty. A similar analysis
for offsetting some of the load on syngas compressor will include
coupling of the high-pressure fuel stream and the syngas stream
coming out at the operating pressure of the chemical looping unit.
This will intrinsically offset some of the requirement for syngas
compression, with the net compression energy being the balance
requirement for the syngas compression. Note the intrinsic
difference in coupling on the reducer streams and the oxidizer
streams. The reducer side streams let-down the pressure on the
inlet lines. It uses that to offset some of the syngas compression
requirements on the outlet side. The let-down process itself
follows the Joule-Thomson cooling effect and requires some
pre-heating of the inlet fuel to offset for the temperature loss.
Depending on the reactor operating pressure, the oxygen carrier
composition, it may be chosen to add an additional heater to
heat-up the fuel entering in the reducer. The oxidizer works in a
reverse way as the let-down is on a high-temperature air-stream.
This intrinsically heats the inlet air-stream and depending on the
reactor operation it could be heated further.
[0158] FIG. 45 shows a model example sample of results for the net
operating energy requirement as a function of operating pressure of
the unit. It assumes that the natural gas input is at 30 atm to the
unit, while the final syngas pressure is required to be 31 atm for
a model chemical synthesis process (e.g., methanol, F-T etc). As
predicted, the net syngas compression cost becomes lower as the
operating pressure of the system increases. Correspondingly the net
air compression cost increases as the operating pressure of the
system increases. The net-effect is a total-net operating energy
curve for the chemical looping system which passes through minima
for the 10 atm to 15 atm range. This can be chosen as a suitable
operating pressure for the looping system for identifying the
critical point of steam injection beyond which the maximum oxygen
transfer from the oxygen carrier to the fuel occurs.
6. EXEMPLARY EMBODIMENTS
[0159] In one aspect, disclosed is a system for producing syngas
from one or more carbon-based fuels using oxidation-reduction of
metal oxides. The system can be configured for partially or fully
oxidizing a fuel in a first reactor, optionally to serve as an
enhancing gas in a second reactor. The metal oxide can undergo an
oxidation-reduction swing in the second reactor (e.g., a reduced
metal oxide entering the second reactor can undergo oxidation, with
CO.sub.2 or H.sub.2O for example, and subsequently undergo
reduction, with a carbon-based fuel for example). The flow pattern
in the first reactor can be counter-current or co-current,
preferably counter-current. The flow pattern in the second reactor
can be counter-current or co-current. The system can include a
third reactor to regenerate the reduced metal-oxide, using air for
example. The first reactor can use a recycled light hydrocarbon
stream, from a Fischer-Tropsch or methanol synthesis system for
example. The enhancing gas used in the second reactor can be
derived from the first reactor, the third reactor, or one or more
additional reactors through oxidation of a fuel with a metal oxide
or through a combustion process (e.g., natural gas combustion).
[0160] In another aspect, disclosed is a system for producing
syngas from one or more carbon-based fuels using
oxidation-reduction of metal oxides. The system can be configured
to oxidize a fuel in a first reactor to produce syngas. A second
reactor in communication with the first reactor can produce a
gaseous product (e.g., CO.sub.2/H.sub.2O) that can be used in the
first reactor for gasifying the fuel in the first reactor (e.g.,
gasifying a solid fuel such as coal or biomass). The system can
include a third reactor to regenerate the reduced metal-oxide,
using air for example. Gaseous product for gasifying the fuel in
the first reactor can also be derived from the third reactor, or
one or more additional reactors through oxidation of a fuel with a
metal oxide or through a combustion process (e.g., natural gas
combustion).
[0161] In another aspect, disclosed is a system for producing
sequester-ready CO.sub.2. A second reactor can process solid fuel
and optionally gaseous fuel to produce syngas. The unconverted
solid fuel can be separated in a high-temperature separation device
to recycle the char to the top of the first reactor to produce
sequester-ready CO.sub.2.
[0162] In another aspect, disclosed is a method for reducing the
operating energy and cost requirements for a chemical system
converting fuel to a product, operating at different pressures. The
method can employ expander-compressor coupling to recuperate energy
used for compression of gaseous fuel, gaseous product, or a
combination thereof.
[0163] In another aspect, disclosed is a specific unique operating
condition configuration for producing high-quality syngas from a
flexible solid fuel feedstock like coal, biomass amongst others.
The specific operating condition requires a unique combination of
the oxygen carrier, the support percentage in the oxygen carrier,
effective heat-management strategy, a specific amount of steam
injection and a co-current downward flowing moving bed system. The
operating reactor configurations can be selected to maximize the
oxygen transfer from the oxygen carrier to the fuel, maximize the
steam oxygen transfer to the oxygen carrier while minimizing the
steam oxygen transfer to the fuel to produce the exact syngas
specifications that are required for a downstream application.
[0164] In certain embodiments, disclosed is a system configuration
which includes fuel addition to the oxidizer to satisfy the
auto-thermal operating condition of the reactor system. The
trade-off is dependent on the permissible carbon emissions and the
carbon-capture mechanism used. A higher carbon utilization of the
fuel can be achieved by adhering to a lower oxygen carrier
flow-rate which will satisfy the heat balance. An alternative
configuration with a higher solids flow rate can be used for
additional carbon capture through the syngas, coupled with higher
oxygen carrier flow rates and higher carbon utilization.
[0165] In certain embodiments, disclosed is a system methodology
for minimizing the operating energy requirement of the reactor
system and choosing the reactor pressure for the configurations
disclosed. The methodology can be applied to any combination of
fuels processed, the oxygen carrier used and the downstream
standard technologies used for any of the disclosed reactor
configurations.
[0166] For reasons of completeness, various aspects of the
disclosure are set out in the following numbered clauses:
[0167] Clause 1. A system for the production of syngas, comprising:
a first reactor comprising a plurality of oxygen carrying particles
comprising a first metal oxide, wherein the first reactor is
configured to provide a counter-current contact mode between the
first metal oxide and a first fuel to reduce the first metal oxide
to a second metal oxide; a second reactor in communication with the
first reactor, the second reactor configured to oxidize the second
metal oxide to a third metal oxide, and further configured to
reduce the third metal oxide to a fourth metal oxide with a second
fuel to provide a partially or fully oxidized gaseous fuel
comprising one or more of CO, CO.sub.2, H.sub.2, and H.sub.2O,
wherein the second metal oxide is oxidized to the third metal oxide
using an enhancing gas of CO.sub.2 and H.sub.2O, the partially or
fully oxidized gaseous fuel, or a combination thereof, to generate
syngas; and a third reactor in communication with the second
reactor, the third reactor configured to regenerate the first metal
oxide by oxidizing the fourth metal oxide with an oxygen
source.
[0168] Clause 2. The system of clause 1, wherein the
counter-current contact mode between the first metal oxide and the
first fuel is such that the first metal oxide moves downward and
the first fuel moves upward.
[0169] Clause 3. The system of clause 1 or clause 2, wherein the
first metal oxide is introduced to the top of the first reactor,
and the first fuel is introduced to the bottom of the first
reactor.
[0170] Clause 4. The system of any one of clauses 1-3, wherein the
second reactor is configured to provide a counter-current contact
mode between the second metal oxide and the enhancing gas, and a
counter-current contact mode between the third metal oxide and the
second fuel.
[0171] Clause 5. The system of clause 4, wherein the second metal
oxide is introduced to the top of the second reactor, the enhancing
gas is introduced to the middle of the second reactor, and the
second fuel is introduced to the bottom of the second reactor.
[0172] Clause 6. The system of any one of clauses 1-3, wherein the
second reactor is configured to provide a co-current contact mode
between the second metal oxide and the enhancing gas, and a
co-current contact mode between the third metal oxide and the
second fuel.
[0173] Clause 7. The system of clause 6, wherein the second metal
oxide is introduced to the top of the second reactor, the enhancing
gas is introduced to the middle of the second reactor, and the
second fuel is introduced to the top of the second reactor.
[0174] Clause 8. The system of clause 6, wherein the second metal
oxide is introduced to the top of the second reactor, the enhancing
gas is introduced to the top of the second reactor, and the second
fuel is introduced to the top or the middle of the second
reactor.
[0175] Clause 9. The system of any one of clauses 1-8, wherein at
least a portion of the enhancing gas is derived from the first
reactor resulting from the reduction of the first metal oxide with
the first fuel.
[0176] Clause 10. The system of any one of clauses 1-9, wherein at
least a portion of the enhancing gas is derived from oxidation of a
carbon-containing or hydrogen-containing source in the third
reactor, a fourth reactor, or a combination thereof.
[0177] Clause 11. The system of any one of clauses 1-10, wherein
the third reactor is communication with the first reactor, wherein
at least a portion of the second metal oxide is circulated directly
to the third reactor, and oxidation of a carbon-containing or
hydrogen-containing source with an oxygen source in the third
reactor generates a stream of enhancing gas, wherein at least a
portion of the enhancing gas generated in the third reactor is used
in the second reactor as an enhancing gas.
[0178] Clause 12. The system of any one of clauses 1-11, comprising
a fourth reactor in communication with the first reactor,
configured to generate a stream of enhancing gas comprising
CO.sub.2 and H.sub.2O.
[0179] Clause 13. The system of clause 12, wherein at least a
portion of the enhancing gas generated in the fourth reactor is
used in the second reactor as an enhancing gas.
[0180] Clause 14. The system of any one of clauses 1-13, wherein
the first fuel is a solid fuel (e.g., biomass, coal, pet-coke,
solid hydrocarbon-based waste products, or a combination
thereof).
[0181] Clause 15. The system of any one of clauses 1-13, wherein
the first fuel is a gaseous fuel (e.g., natural gas, gasified coal,
a light hydrocarbon off-gas stream, or a combination thereof).
[0182] Clause 16. The system of any one of clauses 1-15, wherein
the second fuel is a solid fuel (e.g., biomass, coal, pet-coke,
solid hydrocarbon-based waste products, or a combination
thereof).
[0183] Clause 17. The system of any one of clauses 1-15, wherein
the second fuel is a gaseous fuel (e.g., natural gas, gasified
coal, a light hydrocarbon off-gas stream, or a combination
thereof).
[0184] Clause 18. The system of any one of clauses 1-17, wherein
the second reactor is in communication with a Fischer-Tropsch or
methanol synthesis system that produces a light hydrocarbon
tail-gas, wherein the second reactor provides syngas to the
Fischer-Tropsch or methanol synthesis system and the light
hydrocarbon tail-gas is optionally recycled to first reactor, the
second reactor, or a combination thereof.
[0185] Clause 19. The system of any one of clauses 1-18, wherein
the first metal oxide has formula FeO.sub.aTi.sub.x or
FeO.sub.aAl.sub.x, the second metal oxide has formula FeO.sub.bTix
or FeO.sub.bAl.sub.x, the third metal oxide has formula
FeO.sub.cTi.sub.x or FeO.sub.cAl.sub.x, and the fourth metal oxide
has formula FeO.sub.bTi.sub.x or FeO.sub.dAlx, wherein
1.5>a>b, b<c>d, 1.5>c, and x is 0.01 to 5.
[0186] Clause 20. The system of any one of clauses 1-19, wherein
the first metal oxide has formula FeO.sub.aTiO.sub.2 or
FeO.sub.aAl.sub.2O.sub.3, the second metal oxide has formula
FeO.sub.bTiO.sub.2 or FeO.sub.bAl.sub.2O.sub.3, the third metal
oxide has formula FeO.sub.cTiO.sub.2 or FeO.sub.cAl.sub.2O.sub.3,
and the fourth metal oxide has formula FeO.sub.dTiO.sub.2 or
FeO.sub.dAl.sub.2O.sub.3, wherein 1.5>a>b, b<c>d, and
1.5>c.
[0187] Clause 21. The system of any one of clauses 1-20, wherein
the third metal oxide produced in the second reactor is the same as
the first metal oxide.
[0188] Clause 22. The system of any one of clauses 1-20, wherein
the third metal oxide produced in the second reactor is different
from the first metal oxide.
[0189] Clause 23. The system of any one of clauses 1-22, wherein
the fourth metal oxide produced in the second reactor is the same
as the second metal oxide.
[0190] Clause 24. The system of any one of clauses 1-22, wherein
the fourth metal oxide produced in the second reactor is different
from the second metal oxide.
[0191] Clause 25. A system for the production of syngas,
comprising: a first reactor comprising a plurality of oxygen
carrying particles comprising a first metal oxide, wherein the
first reactor is configured to provide a co-current contact mode
between the first metal oxide and a first fuel to reduce the first
metal oxide to a second metal oxide and to generate syngas; a
second reactor in communication with the first reactor, the second
reactor configured to reduce the second metal oxide to a third
metal oxide with a second fuel; and a third reactor in
communication with the second reactor, the third reactor configured
to regenerate the first metal oxide by oxidizing the third metal
oxide with an oxygen source; wherein a gaseous product comprising
CO.sub.2 and H.sub.2O is used in the first reactor to gasify the
first fuel.
[0192] Clause 26. The system of clause 25, wherein at least a
portion of the gaseous product used in the first reactor is derived
from a gaseous product produced in the second reactor resulting
from the reduction of the second metal oxide with the second
fuel.
[0193] Clause 27. The system of clause 25 or clause 26, wherein the
co-current contact mode between the first metal oxide and the first
fuel is such that the first metal oxide moves downward and the
first fuel moves downward.
[0194] Clause 28. The system of any one of clauses 25-27, wherein
the first metal oxide is introduced to the top of the first
reactor, the first fuel is introduced to the top of the first
reactor at a level below the metal oxide introduction, and the
gaseous product is introduced to the top of the first reactor.
[0195] Clause 29. The system of any one of clauses 25-28, wherein
the second reactor is configured to provide a counter-current
contact mode between the second metal oxide and the second
fuel.
[0196] Clause 30. The system of any one of clauses 25-29, wherein
the second metal oxide is introduced to the top of the second
reactor, and the second fuel is introduced to the bottom of the
second reactor.
[0197] Clause 31. The system of any one of clauses 25-30, wherein
at least a portion of the gaseous product is derived from oxidation
of a carbon-containing or hydrogen-containing source in the third
reactor, a fourth reactor, or a combination thereof.
[0198] Clause 32. The system of any one of clauses 25-31, wherein
the third reactor is communication with the first reactor, wherein
at least a portion of the second metal oxide is circulated directly
to the third reactor, and oxidation of a carbon-containing or
hydrogen-containing source with an oxygen source in the third
reactor generates a gaseous product, wherein at least a portion of
the gaseous product generated in the third reactor is used in the
first reactor to gasify the first fuel.
[0199] Clause 33. The system of any one of clauses 25-32,
comprising a fourth reactor in communication with the first
reactor, configured to generate a stream of gaseous product
comprising CO.sub.2 and H.sub.2O.
[0200] Clause 34. The system of clause 33, wherein the gaseous
product generated in the fourth reactor is used in the first
reactor to gasify the first fuel.
[0201] Clause 35. The system of any one of clauses 25-34, wherein
the first fuel is a solid fuel (e.g., biomass, coal, pet-coke,
solid hydrocarbon-based waste products, or a combination
thereof).
[0202] Clause 36. The system of any one of clauses 25-35, wherein
the second fuel is a gaseous fuel (e.g., natural gas, gasified
coal, a light hydrocarbon off-gas stream, or a combination
thereof).
[0203] Clause 37. The system of any one of clauses 25-36, wherein
the first reactor is in communication with a Fischer-Tropsch or
methanol synthesis system that produces a light hydrocarbon
tail-gas, wherein the first reactor provides syngas to the
Fischer-Tropsch or methanol synthesis system and the light
hydrocarbon tail-gas is optionally recycled to first reactor, the
second reactor, or a combination thereof.
[0204] Clause 38. The system of any one of clauses 25-37, wherein
the first metal oxide has formula FeO.sub.aTi.sub.x or
FeO.sub.aAl.sub.x, the second metal oxide has formula FeO.sub.bTix
or FeO.sub.bAl.sub.x, and the third metal oxide has formula
FeO.sub.cTi.sub.x or FeO.sub.cAl.sub.x, wherein
1.5.gtoreq.a>b>c>0.2, and x is 0.01 to 5.
[0205] Clause 39. The system of any one of clauses 25-38, wherein
the first metal oxide has formula FeO.sub.aTiO.sub.2 or
FeO.sub.aAl.sub.2O.sub.3, the second metal oxide has formula
FeO.sub.bTiO.sub.2 or FeO.sub.bAl.sub.2O.sub.3, and the third metal
oxide has formula FeO.sub.cTiO.sub.2 or FeO.sub.cAl.sub.2O.sub.3,
and wherein 1.5.gtoreq.a>b>c>0.2.
[0206] Clause 40. A system for producing sequester-ready CO.sub.2,
comprising: a first reactor comprising a plurality of oxygen
carrying particles comprising a first metal oxide, wherein the
first reactor is configured to provide a co-current contact mode or
counter-current contact mode between the first metal oxide and a
devolatilized solid fuel to reduce the first metal oxide to a
second metal oxide and to generate product stream comprising
sequester-ready CO.sub.2; a second reactor in communication with
the first reactor, the second reactor configured to reduce the
second metal oxide to a third metal oxide with a solid fuel and a
gaseous fuel to generate syngas and a devolatilized solid fuel; and
a third reactor in communication with the second reactor, the third
reactor configured to regenerate the first metal oxide by oxidizing
the third metal oxide with an oxygen source.
[0207] Clause 41. The system of clause 40, wherein the oxygen
carrying particles are separated from devolatilized solid with a
solid-fine separation device prior to the particles entering the
third reactor.
[0208] Clause 42. The system of clause 40 or clause 41, wherein at
least a portion of the product gas stream from the first reactor is
used to separate the devolatilized fuel from the oxygen carrying
particles in the solid-fine separation device and to convey the
devolatilized fuel to the first reactor.
[0209] Clause 43. The system of any one of clauses 40-42, wherein
at least a portion of the product gas stream from the first reactor
is recycled to the second reactor.
[0210] Clause 44. The system of any one of clauses 40-43, wherein
the first metal oxide is FeOx; the second metal oxide is FeOy; and
the third metal oxide is FeOz; wherein 1.5>x>y>0.75, and
y>z>0.01.
[0211] Clause 45. The system of any one of clauses 1-44, comprising
at least one expander-compressor coupling (e.g., coupled such that
the work extracted from the expander (e.g., used to expand a
pre-heated gaseous fuel) is used to compress the syngas product
from the first reactor or second reactor; or coupled to recuperate
at least a portion of the energy used for compressing the
oxygen-containing gas in the third reactor).
[0212] Clause 46. A system for producing syngas, comprising: a
first reactor comprising a plurality of oxygen carrying particles,
wherein the first reactor is configured to reduce the oxygen
carrying particles with a gaseous fuel to generate syngas and
reduced oxygen carrying particles; a second reactor in
communication with the first reactor, the second reactor configured
to oxidize the reduced oxygen carrying particles with an
oxygen-containing gas to regenerate to the oxygen carrying
particles and produce a spent gas; and (i) a first expander
configured to subject the gaseous fuel to an expansion after
pre-heating and prior to entering the first reactor, and a first
compressor configured to compress the syngas from the first
reactor, wherein the first expander and the first compressor are
coupled such that the work extracted from the expander is used to
compress the syngas; or (ii) a second expander configured to
subject the spent gas from the second reactor to an expansion, and
a second compressor configured to provide compressed
oxygen-containing gas to the second reactor, wherein the second
expander and the second compressor are coupled such that the second
expander recuperate at least a portion of the energy used for
compressing the oxygen-containing gas.
[0213] Clause 47. The system of clause 46, wherein the system
comprises both (i) and (ii).
[0214] Clause 48. The system of clause 46 or clause 47, further
comprising a condenser between the first reactor and the first
compressor.
[0215] Clause 49. The system of any one of clauses 1-48, wherein
the metal oxide undergoes a swing of oxidation state (e.g.,
oxidation followed by reduction) within the first reactor, within
the second reactor, or both the first reactor and the second
reactor.
[0216] Clause 50. The system of any one of clauses 1-49, where the
first reactor, the second reactor, and the third reactor are each
independently selected from a packed moving bed reactor, a rotary
kiln, a down-comer, a fluidized bed reactor, a fixed bed reactor,
or any combination thereof.
[0217] Clause 51. The system of any one of clauses 1-50, wherein
the first fuel, the second fuel, or both the first fuel and the
second fuel are pre-heated prior to entering the first reactor or
the second reactor.
[0218] Clause 52. The system of any one of clauses 1-51, wherein
the first metal oxide is FeO.sub.aTiO.sub.2 with 80 wt %
TiO.sub.2.
[0219] Clause 53. The system of any one of clauses 1-52, wherein
the system has an operating pressure of 1-100 atm (e.g., 20
atm).
[0220] Clause 54. The system of any one of clauses 1-53, wherein
the air and gaseous fuels are pre-heated to about 600 C.
[0221] Clause 55. The system of any one of clauses 1-54, wherein
the first reactor has an operating temperature of 1200 C to 900
C.
[0222] Clause 56. The system of any one of clauses 1-55, wherein
the second reactor has an operating temperature of 100 C to 700
C.
[0223] Clause 57. The system of any one of clauses 1-56, wherein
the third reactor has an operating temperature of 700 C to 1300
C.
[0224] Clause 58. The system of any one of clauses 1-57, wherein
the H.sub.2/CO ratio of the generated syngas is from 1:1 to
4:1.
[0225] Clause 59. The system of any one of clauses 1-58, wherein
the carbon efficiency of the system is 90% or greater.
[0226] Clause 60. The system of any one of clauses 1-59, wherein
the generated syngas has a combined CO.sub.2, H.sub.2O content of
less than or equal to 15%.
[0227] Clause 61. The system of any one of clauses 1-60, wherein
the third reactor uses ambient air, oxygen derived from an air
separation unit, or oxygen-enhanced air (e.g., pure oxygen from an
air-separation unit, vacuum distillation unit, or oxygen tanks), or
a combination thereof to regenerate the first metal oxide.
[0228] Clause 62. The system of any one of clauses 1-61, wherein
the heat is extracted from the third reactor for satisfying
parasitic energy consumption in the system.
[0229] Clause 63. The system of any one of clauses 1-62, wherein
the first metal oxide is a fully oxidized metal oxide (e.g.,
Fe.sub.2O.sub.3).
[0230] Clause 64. The system of any one of clauses 1-62, where the
second metal oxide is a reduced metal (e.g., Fe) or reduced metal
oxide (e.g., FeO, Fe.sub.3O.sub.4).
[0231] Clause 65. The system of any one of clauses 1-24, clause 63,
or clause 64, wherein the third metal oxide is a metal oxide
intermediate (e.g., Fe.sub.2O.sub.3, Fe.sub.3O.sub.4).
[0232] Clause 66. The system of any one of clauses 1-24, or clauses
63-65, wherein the fourth metal oxide is a reduced metal (e.g., Fe)
or reduced metal oxide (e.g., FeO, Fe.sub.3O.sub.4).
[0233] Clause 67. The system of any one of clauses 25-39, wherein
the third metal oxide is a reduced metal (e.g., Fe) or reduced
metal oxide (e.g., FeO).
[0234] Clause 68. The system of any one of clauses 1-67, wherein
the third reactor generates heat, and said heat is used to satisfy
endothermic requirements of one or more of the first or second
reactors.
[0235] Clause 69. The system of any one of clauses 1-68, wherein
the system operates under autothermal conditions.
[0236] Clause 70. A method of reducing the operating energy of a
system for converting fuel, the system comprising at least one
reactor comprising a plurality of oxygen carrying particles,
wherein the reactor is configured to reduce the oxygen carrying
particles with a gaseous fuel to generate syngas and reduced oxygen
carrying particles; at least one reactor configured to oxidize the
reduced oxygen carrying particles with an oxygen-containing gas to
regenerate to the oxygen carrying particles and produce a spent
gas; and at least one expander configured to expand a gas stream,
and at least one compressor configured to compress a gas stream,
wherein the expander and the compressor are coupled, the method
comprising: using the expander to recuperate at least a portion of
the energy used for compressing a gas stream; or using the work
extracted from a high-pressure Joule-Thomson expansion in the
expander to compress a gas stream with the compressor.
[0237] It is to be understood that the first, second, third, and
fourth metal oxides referred to herein may be produced and moved
through the reactor systems in combination with other metal oxides
(e.g., FeO/Fe can be reduced metal oxides that are produced and
traverse one or more reactors). For example, conversion of a first
metal oxide to a second metal oxide may be incomplete, and thus, a
first metal oxide and a second metal oxide may be communicated to a
second reactor.
[0238] It is understood that the foregoing detailed description and
accompanying examples are merely illustrative and are not to be
taken as limitations upon the scope of the invention, which is
defined solely by the appended claims and their equivalents.
[0239] Various changes and modifications to the disclosed
embodiments will be apparent to those skilled in the art. Such
changes and modifications, including without limitation those
relating to the chemical structures, substituents, derivatives,
intermediates, syntheses, compositions, formulations, or methods of
use of the invention, may be made without departing from the spirit
and scope thereof.
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