US20110041740A1

US20110041740A1 - Recuperative combustion system

Info

Publication number: US20110041740A1
Application number: US12/860,520
Authority: US
Inventors: Timothy J. Reilly
Original assignee: Individual
Current assignee: Individual
Priority date: 2009-08-20
Filing date: 2010-08-20
Publication date: 2011-02-24
Also published as: WO2011022653A2; WO2011022653A3

Abstract

The methods and systems described herein relate to a recuperative combustion system that recuperates energy from fuel combustion that would otherwise be lost. The recuperative combustion system minimizes or eliminates the need for an air separator unit through the use of a clean water splitter section, consisting of a thermochemical cycle or high-temperature electrolysis. Water is split into its component hydrogen and oxygen, primarily with process heat from the combustion process. The oxygen produced by the water splitter provides oxygen necessary for oxy-fuel combustion, thereby reducing or eliminating the need for the power intensive air separator unit and/or external oxygen source, significantly increasing the efficiency of the oxy-fuel combustion cycle. Hydrogen produced by the water splitter may be used for a variety of industrial uses, or combined with carbon dioxide (captured from the flue gases produced by said combustion process) to produce methanol. Methanol can further be refined in a methanol to gasoline reactor to produce dimethyl ether, olefins or high grade gasoline. Described herein are methods and systems that 1) increase oxy-fuel combustion efficiency, 2) produce hydrogen for a suite of industrial/energy uses, and 3) capture carbon dioxide and convert it to high value hydrocarbons.

Description

RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application No. 61/274,745, filed Aug. 20, 2009, and U.S. Provisional Application No. 61/345,541, filed May 17, 2010. The entire contents of these patent applications are hereby incorporated herein by reference.

FIELD OF THE INVENTION

BACKGROUND

Energy supply and concerns over unmitigated greenhouse gas emissions are two critical issues of the 21^stcentury. The use of fossil fuels (coal, oil and natural gas) shall continue to play a central role in electricity production for decades to come, and is projected to significantly increase before they are phased out. Hydrogen, a clean energy carrier, is anticipated to play a significant role in energy production in the future.
While carbon capture and sequestration results in a reduction in atmospheric inputs of carbon from industrial sources, it will require large-scale construction of a pipeline distribution and storage system, which will be extremely costly to build. Additionally, the effectiveness of long-term subterranean CO₂storage on a scale of hundreds to thousands of years (required for carbon dioxide mineralization) is presently unproven. Rather than simply disposing of purified, captured and pressurized CO₂from oxy-fuel combustion and purification systems at great expense, consideration for use of carbon dioxide as an industrial feedstock for developing reconstituted high-value carbon-based compounds (e.g., hydrocarbons and oxygenated hydrocarbons) may be seen as an economically and environmentally attractive alternative. Reconstituting waste carbon dioxide into useful materials turns a significant liability into an asset which may 1) reduce carbon emissions and 2) yield fuels, fuel precursors and other beneficial industrial compounds and products.

SUMMARY OF THE INVENTION

In one aspect, provided herein is a method for oxy-fuel combustion, comprising:
providing a system comprising a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat;
capturing heat produced by the oxy-fuel combustion;
using a portion of the heat to power a water splitter, thereby generating hydrogen gas and oxygen gas; and
transferring the oxygen gas from the water splitter to the combustion chamber for use in said oxy-fuel combustion.
In another aspect, provided herein is a method for oxy-fuel combustion, comprising:
providing a system comprising a combustion chamber arranged and disposed to receive coal/water slurry, oxygen and recycled flue gas, wherein the chamber is arranged and disposed to receive said oxygen from an air separator unit and/or external oxygen source, and/or a water splitter, and combust said coal/water slurry, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide; one or more heat exchangers arranged and disposed to capture heat from said heated flue gas and transfer a portion of the captured heat to a water splitter; a water splitter using a 4-step hybrid copper-chlorine thermochemical cycle for the conversion of heat and/or electricity into hydrogen gas and oxygen gas, arranged and disposed to transfer the oxygen gas to the combustion chamber for use in said oxy-fuel combustion;
combusting the coal/water slurry and oxygen to produce heat and heated flue gas containing carbon dioxide;
capturing heat from heated flue gas and transferring captured said heat to the water splitter;
using a portion of the heat to power the water splitter using a 4-step hybrid copper-chlorine thermochemical cycle, thereby producing hydrogen gas and oxygen gas;
transferring the oxygen gas to the combustion chamber for use in said oxy-fuel combustion; and
reducing or eliminating the amount of oxygen that the combustion chamber requires from an air separator unit and/or external oxygen supply, in proportion to the amount of oxygen received from the water splitter.
In yet another aspect, provided herein is a method for the reaction of hydrogen produced by a water splitter and carbon dioxide obtained from combustion flue gas to form methanol and water.
In still aspect, provided herein is an oxy-fuel combustion system, comprising:
a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide;
one or more heat exchangers arranged and disposed to capture heat produced by the oxy-fuel combustion and transfer said heat to a water splitter; and
a water splitter, for the conversion of heat and electricity into hydrogen gas and oxygen gas.
In another aspect, provided herein is an oxy-fuel combustion system, comprising:
a combustion chamber arranged and disposed to receive coal/water slurry, oxygen and recycled flue gas, wherein the chamber is arranged and disposed to receive said oxygen from an air separator unit and/or external oxygen source, and/or a water splitter, and combust said coal/water slurry, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide;
one or more heat exchangers arranged and disposed to capture heat from said heated flue gas and transfer a portion of the captured heat to a water splitter; and
a water splitter using a 4-step hybrid copper-chlorine thermochemical cycle for the conversion of heat and electricity into hydrogen gas and oxygen gas, arranged and disposed to transfer the oxygen gas to the combustion chamber for use in said oxy-fuel combustion.
As described herein, these methods and systems of oxy-fuel combustion are recuperative.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 depicts the recuperative combustion process described herein, further comprising a methanol reactor, means for separating methanol and water, and means for conversion of methanol to downstream products.

FIG. 2 depicts a recuperative combustion system Integrated with an 875 MWTH (HHV) ISOTHERM PWR System (modified from Hong et al. 2008).

FIG. 3 depicts a flow diagram of a methanol synthesis plant (modified from Ushikoshi et al. 2000).

FIG. 4 depicts a recuperative combustion system compared to an archetypal combustion system.

FIG. 5 depicts the steps of the four-step copper chlorine thermochemical cycle (Wang et al. 2009).

FIG. 6 depicts an overview of the sulfur-iodine thermochemical cycle (Brown et al. 2003).

FIG. 7 depicts an archetypal convective recuperator (Bureau of Energy Efficiency 2004).

FIG. 8 depicts an archetypal radiation/convective recuperator (Bureau of Energy Efficiency 2004).

FIG. 9 depicts an archetypal regenerator (Bureau of Energy Efficiency 2004).

DETAILED DESCRIPTION OF THE INVENTION

In one aspect, provided herein is a method for oxy-fuel combustion, comprising: providing a system comprising a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat; capturing heat produced by the oxy-fuel combustion; using a portion of the heat to power a water splitter, thereby generating hydrogen gas and oxygen gas; and transferring the oxygen gas from the water splitter to the combustion chamber for use in said oxy-fuel combustion.

Combustion Chamber and Fuel

Oxy-fuel combustion for the production of electricity (FIG. 1-1) can occur in a variety of combustion systems. Non-limiting examples of combustion systems include: circulating fluidized bed boiler, pulverized coal boiler, and combustors. Combustible materials, including, but not limited to coal, coal/water slurry, petroleum products including oil, methane and natural gas, biomass, and plasma fuels (FIG. 1-3), may undergo oxy-fuel combustion. The primary products of oxy-fuel combustion include heat and flue gas rich in carbon dioxide (FIG. 1-23). A portion of the flue gas is recycled, as described herein (FIG. 1-13).
In one embodiment of the method, the oxy-fuel comprises any combustible material. In another embodiment, the oxy-fuel comprises any hydrocarbon-based fuel. In yet another embodiment, the oxy-fuel comprises coal/water slurry. In still another embodiment, the oxy-fuel comprises oil.
Pressurized oxy-fuel combustion systems have the potential for better performance when compared to conventional atmospheric oxy-fuel combustion power cycles such as the ITEA ISOTHERM® pressurized oxy-fuel system (Hong et al, 2008). For example, oxy-fuel combustion at high pressures may increase the burning rate of char and the heat transfer rates in the convective sections of the heat transfer equipment. Further, because of the raised dew point and the corresponding available latent enthalpy in the raw flue gases, the pressurized oxy-fuel system can recover more thermal energy from the flue gases and eliminate the bleeding from the high-pressure and the low-pressure steam turbines. Consequently, the cycle efficiency for the pressurized oxy-fuel system may be superior to the atmospheric system. To operate this high combustion pressure system, a high pressure deaerator and a flue gas acid condenser can be used. The acid condenser may be modified to work at a high pressure level with flue gas composition seen in oxy-combustion. (Hong et al. 2008).

Air Separator Unit and Water Splitter

In one embodiment of the method, the system comprises a water splitter arranged and disposed to provide oxygen to the combustion chamber for use in oxy-fuel combustion.
In one embodiment of the method, the system further comprises an air separator unit arranged and disposed to provide oxygen to the combustion chamber for use in oxy-fuel combustion.
The normal source of oxygen for oxy-fuel combustion is an air separator unit (ASU) (FIG. 1-5). The selection and implementation of air separator units will be well known to those skilled in the art. Commonly, oxygen is produced in the ASU through a cryogenic distillation process. However, the production of oxygen from an ASU requires 15-20% of the gross power output of an industrial facility. The need for an ASU for oxygen generation may be reduced or eliminated through the use of a water splitter. Candidate water splitters include a hybrid copper-chlorine thermochemical Cycle (CuCl Cycle), sulfur-iodine thermochemical cycle, hybrid sulfur thermochemical cycle (HyS Cycle) and/or high temperature electrolysis (HTE). Other candidate water splitters include volatile metal oxide thermochemical cycles (e.g.,zinc/zinc oxide, hybrid calcium), and non-volatile metal oxide cycles (e.g., iron oxide, cerium oxide) thermochemical cycles. Other water-splitting thermochemical cycles are available, and known to those skilled in the art. Water splitters use thermal or electrical energy in order to convert water (FIG. 1-17) to hydrogen gas (FIG. 1-15) and oxygen gas (FIG. 1-11). In the recuperative combustion process described herein, the water splitter is powered by excess process heat provided by oxy-fuel combustion (FIG. 1-7), and by electricity (FIG. 1-17). Oxygen produced by a thermochemical cycle or HTE offsets, reduces or eliminates the need for an ASU, or external oxygen source, for oxy-fuel combustion, thereby increasing combustion efficiency. The thermochemical cycle or HTE unit employed in the system also produces hydrogen as a product of water splitting. The hydrogen may be used for a variety of purposes, including reaction with the carbon dioxide produced by oxy-fuel combustion to yield methanol and water (FIG. 1-25).
In another embodiment of the method, the system further comprises an air separator unit, wherein the combustion chamber is arranged and disposed to receive oxygen from the air separator unit and/or the water splitter.
Use of a water splitter to produce oxygen results in a reduction, or elimination, of the need for an air separation unit (ASU), or other oxygen source, to supply oxygen for oxy-fuel combustion. The scaling of the water splitter determines if an ASU output is either reduced or eliminated. The employment of a water splitter may significantly reduce, or eliminate, the cost of an ASU unit and, correspondingly, reduce the power penalty resulting from the operation of an ASU. Captured process/waste heat, some of which may be upgraded by chemical heat pumps, may be used to supply additional power needs by the water splitter as well.
In one embodiment of the method, the combustion chamber is arranged and disposed to receive oxygen from the air separator unit and/or external oxygen supply, and/or the water splitter.
The selection and implementation of suitable water splitters, apparatus and procedures for use in the methods and systems described herein will be well known to those skilled in the art.
In a pressurized combustor (Hong et al. 2008), the oxygen purity delivered from the water splitter and/or ASU may be 95% (mol %) and delivered at 200° C.; oxygen content in the recycled flue gas may be about 3% (mol %). The oxygen delivery temperature to the combustor is controlled to prevent the acid condensation when mixed with the recycled flue gases. The flue gases contain acid gases, such as SO₃, SO₂, nitrogen oxides, and HCl produced during combustion. As shown in FIG. 2, the recycled flue gases, state 2-19, are mixed with the oxygen stream from the water splitter, state 2-13, which may be colder, depending on the ASU or water splitter technology employed. To avoid corrosion due to the condensation of these acid gases when mixed with the oxygen stream, the oxygen delivery temperature needs to be carefully controlled to nearly 200° C. The 200° C. delivery temperature target is achieved by using a two-stage oxygen compressor with an intercooler. The mass flow rate of the oxygen stream is determined such that the raw flue gases exiting the combustor have 3% oxygen on a molar basis (Hong et al. 2008)
In one embodiment of the method, the amount of oxygen required from the air separator unit and/or external oxygen supply is reduced or eliminated in proportion to the amount of oxygen provided by the water splitter. The amount of oxygen that is required from the air separator unit, and/or another oxygen source, and/or the water splitter, may be determined by methods taught herein (see, for example, the section of Algorithms and Formulae), or by methods known to those of skill in the art.

Thermochemical Cycles

In another embodiment of the method, the water splitter produces hydrogen gas and oxygen gas by means of a thermochemical cycle. In still another embodiment, the water splitter produces hydrogen and oxygen gas by means of a hybrid thermochemical/electrochemical cycle.
The selection and implementation of suitable thermochemical or hybrid thermochemical/electrochemical cycles, apparatus and procedures for use in the methods and systems described herein will be well known to those skilled in the art.
Non-limiting examples of thermochemical cycles and hybrid thermochemical/electrochemical cycles include sulfur cycles (e.g., hybrid sulfur and sulfur-iodine thermochemical cycles), low temperature cycles (e.g., hybrid copper-chlorine thermochemical cycle), volatile metal oxide cycles (e.g.,zinc/zinc oxide, hybrid calcium), and non-volatile metal oxide cycles (e.g., iron oxide, cerium oxide) thermochemical cycles. Other suitable thermochemical water-splitting cycles are available, and known to those skilled in the art.
In one embodiment, the thermochemical cycle is selected from: a hybrid copper-chlorine cycle; a sulfur-iodine cycle; and a hybrid sulfur cycle.
In one embodiment, the thermochemical cycle is a hybrid copper-chlorine cycle. In another embodiment, the copper-chlorine cycle is selected from: a 3-step cycle, a 4-step cycle, and a 5-step cycle. In yet another embodiment, the hybrid copper-chlorine cycle is the 4-step cycle. In still another embodiment, the 4-step cycle of the hybrid copper chlorine cycle may be represented by the following steps:

- Step I: Cu(s)+2HCl(g)→2CuCl(molten)+H₂(g)
- Step II: 4CuCl(s)→2Cu(s)+2CuCl₂(aq)+HCl(aq)
- Step III: CuCl₂(aq)+n_fH₂O(l)→CuOCuCl₂(s)+2HCl(g)+(n_f−1)H₂O(g)
- Step IV: CuOCuCl₂(s)→2CuCl(molten)+0.5O₂(g)

In certain embodiments, n_fis 5-30.
Three-step, four-step and five-step versions of the copper-chlorine (Cu—Cl) hybrid electrochemical-thermochemical cycle are described in Wang et al. (2009). While the five-step version requires less heat than the three-step version of the hybrid Cu—Cl cycle, it is more complex from equipment and process engineering standpoints. A hybrid Cu—Cl Cycle using the four-step process is described by Chukwu et al. (2008). Step III of the four-step process is an electrochemical step, requiring the input of electricity. Chukwu et al. provides reaction-specific thermodynamic data and Aspen Plus heat/mass balance modeling of the four-step hybrid Cu—Cl cycle.
The four-step process (FIG. 5) consists of three thermal reactions in which H₂, O₂and HCl are generated, and an electrochemical step in which CuCl is disproportionated to yield copper metal and CuCl₂. The oxygen is released from the copper oxychloride between 450° C. and 530° C., which is the highest temperature limit for this cycle (Sattler 2010).
In one embodiment of the method, the water splitter operates at a temperature 450° C. for Step I; 30-80° C. for Step II; 375° C. for Step III; and 530° C. for Step IV. another embodiment, the heat required to produce 1 mole of O and 1 mole of H₂is about 554.7 kJ/mol.
The hybrid Cu—Cl four-step cycle receives process heat in adequate supply from oxy-fuel combustion through a heat exchanger system. The hybrid Cu—Cl four-step cycle is sized to accommodate oxygen requirements for oxy-fuel combustion within the fossil fuel combustor or boiler. This sizing is based on either a partial or complete replacement of the air separator unit (ASU), depending on a number of pre-existing, technical and economic factors unique to an industrial/power plant, or other site of combustion. Hydrogen production rates for various downstream hydrogen uses are also an important aspect of sizing the four-step hybrid Cu—Cl thermochemical cycle reactor.
In one embodiment of the method, the thermochemical cycle is a sulfur-iodine (S—I) cycle. In another embodiment, the sulfur-iodine cycle may be represented by the following steps:

- Step I: 2H₂O+SO₂+I₂→H₂SO₄+2HI
- Step II: H₂SO4→H₂O+SO₂+½O₂
- Step III: 2HI→H₂+I₂

Heat and mass balances associated with the S—I Cycle can be found in Brown et al. (2003). The gross heat needed per mole H₂(and 0.5 mole O₂) is 674.9 kJ/mole H₂with net heat requirements is 391.3 kJ/mol H₂(full HI gasification), 432.9 kJ/mol H₂(no HI gasification), as compared with 554.7 kJ/H₂gross heat and 322.7 kJ/H₂net heat required for the Cu—Cl cycle (Wang et al. 2009). A major difference between the Cu—Cl and S—I thermochemical cycles is the substantially lower temperatures that the Cu—Cl cycle operates at relative to the S—I cycle, allowing almost 30% of the heat needed for the Cu—Cl process to come from low grade heat (i.e., heat at temperatures lower than 343 K).
The hybrid sulfur (HyS) cycle is a hybrid electrochemical—thermochemical cycle (Bilgen 1988). It consists of splitting sulfuric acid into water and sulfur trixode (endothermic), followed by further decomposition of sulfur trioxide to sulfur dioxide (highly endothermic). As a final step, sulfur dioxide is electrochemically oxidized to sulfuric acid with concomitant production of hydrogen. The steps of the HyS cycle are as follows:

- Step I: H₂SO₄→H₂O+SO₃(>450° C.)
- Step II: SO₃→SO₂+½O₂(>850° C. with catalyst; 1150° C. without catalyst)
- Step III: 2H₂O+SO₂→H₂SO₄+H₂(electrolysis, 80° C.)

Electrical power is required for the electrolysis, but the electrochemical oxidation of SO₂is far more efficient than the electrolytic splitting of water. The overall efficiency of the process is calculated to be about 40%. Carbon-supported platinum electrodes are used for the SO₂oxidation. Cells made from ceramics such as silicon carbide, silicon nitrite, and cermets possess excellent resistance to corrosion by sulfuric acid at ambient temperature and at low acid concentration. Catalysts mainly based on iron oxide are available for accelerating the reaction rate of the SO₃reduction at “low” temperature (850° C.). The kinetics of the reaction are much faster if higher temperatures are available as in solar tower installations. Therefore, the use of catalysts might be reduced or even unnecessary if the sulfuric acid splitting is coupled to concentrated solar radiation. It has to be evaluated whether the higher temperatures are more efficient than the catalyzed reaction on an annual basis. Reactors used in laboratory tests have been made of glass or fused silica; solar reactors are mostly constructed from ceramics such as silicon carbide, but gold-coated steel has also been used (Noglik et al. 2009; Sattler 2010). Like the S—I Cycle, the HyS cycle may be integrated with the recuperative combustion process described herein. However, the high grade heat (>850 C with catalyst, or 1150 C without catalyst) limits the application for combustion facilities to those applications with very high heat output (i.e., oxy-fuel combustion with high oxygen content/lower recirculated flue gases, or high temperature industrial processes such as, for example, steel milling or glass production).

High Temperature Electrolysis

In one embodiment of the method, the water splitter produces hydrogen gas and oxygen gas by means of high-temperature electrolysis.
The selection and implementation of suitable high temperature electrolysis apparatus and procedures for use in the methods and systems described herein will be well known to those skilled in the art.
Hydrogen and Oxygen can be produced via the classical electrolysis of water at low temperature or, alternatively, by using the different fuel cell technologies. These technologies are based on (i) proton-exchange membrane fuel cells (PEMFCs) (referring to the solid polymeric electrolyte membrane), (ii) fuel cells using solid oxide proton conductors, and (iii) fuel cells with a solid oxide ion (O²⁻) conductor (SOFCs). In a fuel cell, electrical energy is generated by the exothermic oxidation of hydrogen. In the reverse electrolysis operation of such a cell, steam is reduced in an endothermic reaction using electrical energy.
The operating temperatures of fuels cells vary widely, from around 80-120° C. for PEMFCs to 700-1000° C. for SOFCs. The free energy required for the reaction (ΔG) decreases with increasing temperature whereas the free enthalpy (ΔH) remains almost constant. This thermodynamic relation, in principle unfavorable for the fuel cell mode at high temperatures, explains the particular interest in performing electrolysis at high temperatures. Since the SOFCs achieve competitive (chemical-to-electrical) energy conversion efficiencies despite the less favorable thermodynamic conditions, one can a priori expect that high-temperature electrolysis (HTE) cells achieve much higher (electrical to-chemical) energy conversion efficiencies (the term energy conversion efficiency for the HTE refers to the electrical-to-chemical energy conversion). Because ionic transfer numbers are close to one for both cell types, the difference in cell voltage translates linearly to the energy consumption for the reaction.
The primary motivation for HTE is the above-mentioned potential of a reduced demand for electrical energy compared with electrolysis at low temperature. This may allow electrical-to-chemical energy conversion efficiencies even exceeding 100%, as already recognized in early work (Isenberg 1981). The free energy of the reaction ΔG decreases from ˜1.23 eV (237 kj mol⁻¹) at ambient temperature to −0.95 eV (183 kj mol⁻¹) at 900° C., while the free enthalpy term remains essentially unchanged (ΔH≈1.3 eV or 249 kj mol⁻¹at 900° C.). Part of the energy required for an ideal (loss-free) HTE can thus be provided by heat. Increasing ohmic and/or reaction losses in a real HTE system increase the demand for electrical energy and decrease the demand for an external heat supply until, finally, the reaction becomes exothermic. Hence three modes of operation are distinguishable in HTE: thermoneutral, endothermic, and exothermic. HTE operates at thermal equilibrium (the thermoneutral mode) when the electrical energy input equals the enthalpy of the reaction. In that case, the entropy necessary for water splitting equals the heat generated by the loss reactions, and the energy conversion efficiency is 100%. In the exothermic mode, on the other hand, the electric energy input exceeds the ΔH term, which corresponds to efficiency below 100%. Finally, in the endothermic mode, the electric energy input remains below the enthalpy term. Therefore, heat must be supplied to maintain the cell temperature. This mode means that energy conversion efficiencies of the cell or the stacks are above 100%.
An HTE system can be operated with and without an external heat supply. This is different to low-temperature electrolyzers, which run in the exothermic mode, because the energy losses, which arise mainly from the electrochemical reactions, exceed the small difference between ΔH and ΔG at low temperature. The availability of an external heat source influences the design of an HTE system.
Without a heat source, the goal is to approach the thermoneutral mode, that is, to limit the thermal losses to a value required to compensate for the endothermic reaction. This leaves a wide margin for cell overvoltages and, therefore, for an increase in the current density or a lowering of the temperature. Operating temperatures in the range 600-700° C., known from the SOFC development, may therefore also be accessible for electrolysis.
With an external heat source of high temperature, on the other hand, the goal is to reduce the overvoltages as far as possible to allow for a significant uptake of heat. This implies, at least with present cell technology, operation under higher temperatures (800° C. or above) and lower electrode overvoltages (i.e., current densities somewhat lower than those achieved in thermoneutral operation).
The operation of SOFCs in electrolysis mode has been demonstrated in several research projects since 2004 (EIFER 2010). Cells of both commercial and research types and including the common designs were tested (electrolyte, hydrogen electrode, and metal substrate-supported). As for fuel cell operation, the hydrogen electrode-supported cells showed the highest performance owing to the low resistance of the thin electrolyte layer. A high current density of −3.6 A cm⁻²at a cell voltage of 1.48 V and a cell temperature of 950° C., for example, was reached with such a cell at DTU-Risoe (Denmark) (Mogensen et al. 2006; Zahid et al. 2010).

Heat Exchangers

In one embodiment of the method, the system further comprises one or more heat exchangers for the capture and transfer of heat from the oxy-fuel combustion to the water splitter.
Heat exchangers are required for transferring process heat from the heated flue gas to endothermic reactions within the water splitter section. For example, in the Cu—Cl Cycle, sufficient heat must be transferred from the flue gas coming from the boiler/combustor to the chlorination step (Step I), oxychlorination step (Step (III) and decomposition step (Step IV) (FIG. 5).
Numerous types of heat exchangers are available for the adequate transfer of heat to the water splitter. Specific choice of heat exchanger is dependent on a number of factors, including, but not limited to: the nature or quality of the flue gases, the temperature of primary process flue gases, matching heat demand of the secondary (i.e., water splitting) process with the heat supply from the primary process, matching timing of the heat supply for the primary process and the heat demand in the secondary process, and placement of primary and secondary heating equipment.
The selection and implementation of suitable heat exchanger apparatus and procedures for use in the methods and systems described herein will be well known to those skilled in the art.
In one embodiment of the method, the system further comprises one or more heat exchangers for the capture and transfer of heat from the heated flue gas to the water splitter. In another embodiment, said one or more heat exchangers are selected from: a convective recuperator, a radiation/convective recuperator, a ceramic recuperator and a regenerator. In yet another embodiment, said one or more heat exchangers are selected from appropriately scaled heat exchangers which function within specified heat ranges (i.e., heat ranges specified herein or known to those of skill in the art). Other suitable heat exchangers are available, and known to those of skill in the art.
Convective Recuperator. In a recuperator, heat exchange takes place between the flue gases and the air through metallic or ceramic walls. Ducts or tubes carry the air or other gas to be heated; the other side contains the waste heat stream. A common configuration for recuperators is called the tube type or convective recuperator. As seen in the FIG. 7, the hot gases are carried through a number of parallel small diameter tubes, while the incoming air/gas to be heated enters a shell surrounding the tubes and passes over the hot tubes one or more times in a direction normal to their axes. If the tubes are baffled to allow the gas to pass over them twice, the heat exchanger is termed a two-pass recuperator; if two baffles are used, a three-pass recuperator, etc. Although baffling increases both the cost of the exchanger and the pressure drop in the combustion air path, it increases the effectiveness of heat exchange. Shell and tube type recuperators are generally more compact and have a higher effectiveness than radiation recuperators, because of the larger heat transfer area made possible through the use of multiple tubes and multiple passes of the gases.
Radiation/convective Recuperator. For maximum effectiveness of heat transfer, combinations of radiation and convective designs are used, with the high-temperature radiation recuperator being first followed by convection type. These are more expensive than simple metallic radiation recuperators, but are less bulky. A Convective/radiative Hybrid recuperator is shown in FIG. 8.
Ceramic Recuperator. The principal limitation on the heat recovery of metal recuperators is the reduced life of the liner at inlet temperatures exceeding 1100° C. In order to overcome the temperature limitations of metal recuperators, ceramic tube recuperators have been developed whose materials allow operation on the gas side to 1550° C. and on the preheated air side to 815° C. on a more or less practical basis. Early ceramic recuperators were built of tile and joined with furnace cement, and thermal cycling caused cracking of joints and rapid deterioration of the tubes. Later developments introduced various kinds of short silicon carbide tubes which can be joined by flexible seals located in the air headers. Earlier designs had experienced leakage rates from 8 to 60 percent. The new designs are reported to last two years with air preheat temperatures as high as 700° C., with much lower leakage rates.
Regenerator. Regenerators are rechargeable storage batteries for heat. A regenerator (FIG. 9) is an insulated container filled with metal or ceramic shapes that can absorb and store relatively large amounts of thermal energy. During the operating cycle, process exhaust gases flow through the regenerator, heating the storage medium. After a while, the medium becomes fully heated (charged). The exhaust flow is shut off and cold combustion air enters the unit. As it passes through, the air extracts heat from the storage medium, increasing in temperature before it enters the burners. Eventually, the heat stored in the medium is drawn down to the point where the regenerator requires recharging. At that point, the combustion air flow is shut off and the exhaust gases return to the unit. This cycle repeats as long as the process continues to operate. For a continuous operation, at least two regenerators and their associated burners are required. One regenerator provides energy to the combustion air, while the other recharges. An alternate design of regenerator uses a continuously rotating wheel containing metal or ceramic matrix. The flue gases and combustion air pass through different parts of the wheel during its rotation to receive heat from flue gases and release heat to the combustion air. Regenerators may be preferable for large capacities and have been very widely used in glass and steel melting furnaces. Important relations exist between the size of the regenerator, time between reversals, thickness of brick, conductivity of brick and heat storage ratio of the brick. In a regenerator, the time between the reversals is an important aspect. Long periods would mean higher thermal storage and hence higher cost. Also long periods of reversal result in lower average temperature of preheat and consequently reduce fuel economy.
In one aspect, provided herein is a method for oxy-fuel combustion, comprising:
providing a system comprising a combustion chamber arranged and disposed to receive coal/water slurry, oxygen and recycle flue gas, wherein the chamber is arranged and disposed to receive said oxygen from an air separator unit and/or a water splitter, and combust said coal/water slurry, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide; one or more heat exchangers arranged and disposed to capture heat from said heated flue gas and transfer a portion of the captured heat to a water splitter; a water splitter using a 4-step hybrid copper-chlorine thermochemical cycle for the conversion of heat and/or electricity into hydrogen gas and oxygen gas, arranged and disposed to transfer the oxygen gas to the combustion chamber for use in said oxy-fuel combustion;
combusting coal/water slurry, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide; capturing heat from heated flue gas and transferring captured heat to the water splitter; Using a portion of the heat to power the water splitter using a 4-step hybrid copper-chlorine thermochemical cycle, thereby producing hydrogen gas and oxygen gas; transferring the oxygen gas to the combustion chamber for use in said oxy-fuel combustion; reducing or eliminating the amount of oxygen that the combustion chamber requires from an air separator unit and/or external oxygen supply, in proportion to the amount of oxygen received from the water splitter.

Efficiency

In one embodiment of the method, the amount of oxygen that the combustion chamber requires from an air separator unit and/or external oxygen supply, is reduced or eliminated in proportion to the amount of oxygen received from the water splitter results in increased efficiency.
In one embodiment of the method, the increased efficiency is measured in terms of increased gross power output of the combustion process. In one embodiment, the gross power output of the combustion process is increased by 1-20%. In another embodiment, the gross power output of the combustion process is increased by 1-10%. In yet another embodiment, the gross power output of the combustion process is increased by 10-20%. In still another embodiment, the gross power output of the combustion process is increased by 15-20%.

Primary and Secondary Products

In one embodiment of the method, the oxy-fuel combustion products further comprise heated flue gas containing carbon dioxide.
Oxy-fuel combustion yields flue gases consisting of predominantly carbon dioxide (FIG. 1-23) and condensable water, whereas conventional air-fired combustion flue gases are nitrogen-rich with only about 15% (by volume) of carbon dioxide (Hu et al. 2008; IEA, 2008). The high carbon dioxide concentration (up to 95%) and the significantly lower nitrogen concentration in the oxy-fuel raw flue gases is a unique feature that lowers the energy and capital costs of oxy-fuel carbon dioxide capture when compared to alternatives (Buhre et al 2005). Further, decreasing relative recycled flue gas input increases combustion, and flue gas, temperatures, and may be a mode for increasing heat grade for secondary water splitting processes. For example, oxy-fuel combustion, using a pressurized coal combustor, occurs at 1400 to 1600° C.; however, stoichiometric combustion of coal in pure oxygen reaches up to 3500° C. Optimization of oxygen/recycled gas content with respect to high grade heat production and oxygen/hydrogen yields requires additional research.
Oxy-fuel combustion involves the burning of fuel in an oxygen-rich, nitrogen-lean and carbon dioxide-rich environment, which is achieved by feeding the combustor or boiler with an oxygen-rich stream and recycled flue gases. Oxy-fuel combustion produces a flue gas stream containing mostly CO₂, which can be directly purified and compressed for conversion to useful materials or for carbon sequestration purposes. In the process shown in FIG. 1, the CO₂concentration in the flue gas is greatly increased by using a mixture of recirculated flue gas and pure oxygen instead of air for firing coal. Recirculation of flue gas is necessary to provide sufficient mass flow of flue gas for cooling the flame and also heat capacity and flue gas velocity for convective heat transfer in the boiler. In the oxy-fuel process, CO₂purity is mainly influenced by (a) where the flue gas is recycled in the process (the cleaning that has been done up to this point), (b) the sealing of boiler and other components to prevent air ingress, (c) the purity of the oxygen from the Air Separation Unit (ASU) (or alternative oxygen source), (d) the performance of all air quality control system equipment (e.g., SCR, FGD, and ESP), and (e) additional CO₂purification during/after compression (Wu et al. 2009). Oxy-fuel combustion involves using a mixture of recirculated flue gas and pure oxygen, instead of air, for firing the fuel. Oxy-fuel combustion may take place at 1400 to 1600° C. The stoichiometric combustion of coal in pure oxygen may reach up to 3500° C. While oxy-fuel combustion is close to stoichiometric, lower combustion temperature are achieved by using the appropriate amount of the recycled flue gases (Hong et al 2008). Further, certain additional efficiencies are realized when oxy-fuel combustion takes place under pressurized conditions (Hong et al. 2008).
The primary outputs of the recuperative combustion process include the following:
CO₂(and small amounts of CO) in the flue gas, following post combustion treatment;

H₂and O₂from the water splitter (i.e., Cu—Cl, S—I, or HyS thermochemical cycles) or HTE. The secondary outputs from the downstream portions of the process include: methanol from the methanol reactor (FIG. 1-25), dimethyl ether (DME), olefins, and gasoline (mainly C5-C9) from methanol to gasoline conversion (FIG. 1-35), and water, as a by-product of methanol and other hydrocarbon production. This water may be treated, as necessary, and recycled to the water splitter (FIGS. 1-31 and 1-41). The secondary outputs of the recuperative combustion process described herein are valuable products on the global market (FIG. 1-39).

In one embodiment of the method, the hydrogen from the water splitter is used directly or indirectly in a subsequent process. In another embodiment of the method, the hydrogen from the water splitter is used directly in a subsequent process.
In another embodiment of the method, the hydrogen from the water splitter and the carbon dioxide from the combustion flue gas are reacted to form methanol and water.
In one aspect, provided herein is a method for the reaction of hydrogen produced by a water splitter and carbon dioxide obtained from combustion flue gas to form methanol and water. In one embodiment, the carbon dioxide is purified and compressed prior to reacting with hydrogen. In another embodiment, the hydrogen and carbon dioxide are both produced by the recuperative combustion system. In yet another embodiment of the method, the products of the reaction further comprise waste heat.
Using carbon dioxide or carbon monoxide for downstream conversion to hydrocarbons requires carbon dioxide separation and removal of the impurities in the high-concentration carbon dioxide flue gases resulting from oxy-fuel combustion. Carbon dioxide purification involves the removal of contaminants from the flue gas, including nitrogen oxides, sulfur oxides, and mercury, generally under pressurized conditions. While numerous strategies for removal of these contaminants already exist at modern fossil fuel power plants, including flue gas desulfurization (FGD) for sulfur oxides, selective catalytic reduction (SCR) for nitrogen oxides compounds and activated carbon or sorbents for mercury, these are capital-intensive technologies that do not result in a comprehensive removal of impurities from, and separation of, carbon dioxide from flue gas. A host of developing technologies for increasing carbon dioxide purity are currently under development (e.g., see White and Fogash 2009; Hong et al., 2008; and Shah 2006). Specific technologies employed for contaminant removal in the carbon dioxide stream depends on the end use of carbon dioxide (i.e., for methanol production, enhanced oil recovery (EOR) or sequestration).
The high-concentration carbon dioxide flue gas that is produced by the oxy-fuel combustion process may be hydrogenated in a fluidized bed reactor with hydrogen gas at 200-300° C. at a pressure of 50-100 bar in a catalytically-mediated reaction (heterogeneous catalyst includes, but is not limited to: Cu/ZnO/ZrO₂/Al₂O₃/SiO₂), yielding methanol, water and substantial heat (FIGS. 1-25, 1-21 and 1-31). The following three reactions are controlled in the methanol reactor to maximize methanol synthesis (Hirotani et al. 1998):
CO+2H₂→CH₃OH (ΔH=−90.6 kJ/mol),
CO₂+3H₂→CH₃OH+H₂O (ΔH=−49.4 kJ/mol), and
CO₂+H₂→CO+H₂O (ΔH=+41.2 kJ/mol).
The selection and implementation of suitable apparatus and procedures for the production of methanol for use in the methods and systems described herein will be well known to those skilled in the art.
The methanol reactor may be of several types, including a modified test methanol synthesis reactor from Ushikoshi et al. (2000). FIG. 3 shows a flow diagram of the test plant, which is designed with facilities for recycling unreacted gases. The gases (mixture of CO₂, CO and H₂) supplied from the CO₂purification and compression unit (mainly CO₂) and the water splitter (H₂) (FIGS. 3-1, 3-2, 3-3, 3-4) are compressed (FIGS. 3-5, 3-6, 3-7) along with recycled gases (FIG. 3-23), and then fed into the reaction tube (FIGS. 3-8, 3-13) through a pre-heater (FIG. 3-9). The reactor (FIG. 3-14) is surrounded by a heat exchanger divided into four parts to facilitate isothermal operation, and capture of waste heat. The temperature profile along the bed is measured by means of eight thermocouples situated at the central axis of the reactor. The temperature difference along the reactor is less than 2° K. The pressure is controlled within 0.1 MPa by changing the total flow rate of the make-up gas, in which the H₂/CO/CO₂ratio is adjusted with the flow controllers.
The flow rate of the inlet gas to the reactor is controlled by the flow controller placed just after the recycle gas compressor. Reaction products are cooled down (FIGS. 3-18, 3-19), and then the mixture of methanol and water is separated at the gas-liquid separator (FIGS. 3-20, 3-21) from unreacted gases. Unreacted gases and gaseous products, such as CO, methane and so on, excluding small amounts of purge gas, are recycled back to the reactor (FIGS. 3-22, 3-23)). The mixture of methanol and water is subjected to separation (see below) then stored in a container (FIGS. 3-25, 3-26, 3-27). Control of temperature in the methanol reactor for pre-reaction reduction of the catalyst (in the presence of heat and hydrogen and nitrogen gases) as well as during methanol synthesis is controlled by a preheater, oil heater and oil cooler system with (FIGS. 3-9, 3-10, 3-11, 3-12, 3-13, 3-14, 3-15, 3-16 and 3-17).
The make-up gas, the inlet and outlet gases of the reactor and the recycle gas are analyzed with an on-line gas chromatograph. Gas chromatography is employed for analysis of the reaction products; H₂, CO and CO₂are analyzed by thermal-conductivity detector; methanol, dimethyl ether, methyl formate and hydrocarbons are analyzed by the flame ionization detector. Excess heat produced by exothermic methanol reaction (FIG. 3-30) may be recuperated and recirculated for use in the water splitter (FIG. 3-32) (through heat upgrading using chemical heat pumps) (FIG. 3-31).
The methanol and water are separated through a solvent dehydration process (e.g., by energetically attractive pervaporation using a hydrophilic ZeoSep A membrane or equivalent hydrophobic membrane for concentration of organics, or a distillation column). Final selection of appropriate pervaporation membrane or distillation column design shall depend on optimized methanol and water concentrations in the methanol/water mixture.
The methanol may be used as is, or used as an industrial feedstock (e.g., for conversion to dimethyl ether, olefins, gasoline and a spectrum of other industrial applications) using a variety of industrial processes such as the Mobil Methanol to Gasoline process. The byproduct water is filtered and re-used in the in the water splitter.

Chemical Heat Pumps

Chemical Heat Pumps (CHPs) are systems that use coupled exothermic and endothermic reactors to store thermal energy and transform it to another temperature, including waste heat whose thermal energy at low temperatures can be upgraded to higher temperatures (Naterer 2008). CHPs may be useful in the conversion of waste heat captured from exothermic portions of the Water Splitter Section or downstream Methanol Reactor for Methanol to Gasoline Sections and upgraded for use in heat-intensive, endothermic portions of the Water Splitter Section. Gainful use of CHPs result in significant increase in hydrogen and oxygen production efficiency, and overall oxy-fuel combustion system efficiencies.
The selection and implementation of suitable chemical heat pump apparatus and procedures for use in the methods and systems described herein will be well known to those skilled in the art.
Two specific solid-gas CHPs, namely salt/ammonia and MgO/water systems, are particularly useful in application to waste heat upgrading for thermochemical hydrogen production, especially when configured in series (Naterer 2008). These are described below.
A salt/ammonia chemical heat pump consists of salts that are able to absorb/desorb ammonia vapor at different operating temperatures. The ammonia vapor pressure is a function of temperature for two different salts, designated by LTS (low-temperature salt) and HTS (high-temperature salt). The desorption reaction is endothermic. Heat must be supplied to the gas/solid reactor to release ammonia vapor from the LTS. When this ammonia vapor flows to the HTS, it is absorbed and heat is released in an exothermic reaction. The pair of salts is MnSO₄/NH₃(LTS) and NiCl₂/NH₃(HTS). The operation of the chemical heat pump is described below:
MnSO₄.6NH₃ +Q _waste
4MnSO₄.2NH₃+4NH₃; ΔH=+57.6 kJ/mol (NH₃)
NiCl₂.2NH₃+4NH₃
NiCl₂.6NH₃ +Q _out; ΔH=−55.3 kJ/mol (NH₃)
An integrated closed cycle of a salt/ammonia chemical heat pump was presented and analyzed by Spoelstra et al. (2002). In their analysis, 5000 kW of low-temperature heat at 140° C. was upgraded to 2051 kW of high-temperature heat at 240° C. Shell-and-tube reactors were used with finned tubes to achieve this operating capacity. Each reactor vessel was about 6 m in height, with a diameter of 2-3 m. The total weight of one vessel was about 50 tons, including the salt and heat exchanger tubes. A very high coefficient of performance for this CHP was reported by the authors (COP=97), since electrical power is only required to pump around liquid streams. This chemical heat pump could be used as a “bottoming cycle” to upgrade waste heat to an intermediate stage, before another CHP upgrades further to higher temperatures. It is anticipated that equipment performance and reaction kinetics would become unfavorable if a single CHP attempts to operate over an excessively large temperature range. Therefore, a magnesium oxide (MgO)/vapor chemical heat pump, which operates at temperatures above the salt/ammonia CHP may be useful in increasing the grade of heat further.
The MgO/Vapor CHP is described by the following chemical reaction:
MgO(s)+H₂O(g)
4Mg(OH)2; ΔH=_—−81.02 kJ/mol
The rightward reaction is exothermic MgO hydration. The kinetics of the reaction have been reported by Kato et al. (1996). The operation consists of heat storage and heat supply modes, with solid products from each reactor supplied as solid feed to the other. Magnesium hydroxide (Mg(OH)₂) is initially charged into a gas/solid reactor. Heat is added, after which solid MgO and water vapor are formed. The heat of condensation is recovered from the steam and the resulting water is stored as a liquid. In the heat supply mode, the stored water is then vaporized by another separate heat input. The vapor is supplied to an exothermic solid/gas reactor for hydration of MgO. Scientific feasibility of the MgO/vapor chemical heat pump has been demonstrated by Kato et al. (1996). A lab-scale demonstration was performed by the authors, with an average heat output rate of 349 W per kg of Mg(OH)₂solid feed. The experimental apparatus consisted of an evaporator, gas/solid reactor, heating supply with an electric furnace, condenser, water trap and vacuum pump. Future research and development are still needed to scale up this system to larger heating capacities (Naterer 2008).
Using a sequence of chemical heat pumps in series, a conceptual framework of coupled CHPs and a Cu—Cl thermochemical cycle may be advantageous. Specifically, an exothermic step within the Cu—Cl cycle (or Methanol Reactor Section) could supply heat into the MgO/vapor CHP, to be subsequently upgraded to a higher temperature that is then used by the endothermic hydrolysis step in the Cu—Cl cycle. The lower temperature endothermic step of copper oxychloride decomposition could then be supplied separately from the salt/ammonia CHP. Input power is needed to drive compressors in the CHPs. With existing heat recovery technology available in commercial systems, electricity generated from waste heat could be supplied directly to the CHPs, thereby potentially making the CHPs and Cu—Cl cycle solely driven by process/waste heat from industrial or power plants (Naterer 2008).
Naterer (2008) presents a thermodynamic analysis of combined chemical heat pumps for a thermochemical cycle of hydrogen production, demonstrating that low-grade waste heat can be upgraded to higher temperatures via salt/ammonia and MgO/vapor chemical heat pumps, which release heat at successively higher temperatures through exothermic reactions. Using this new approach, waste heat industrial sources can be transformed to a useful energy supply for thermochemical hydrogen and oxygen production. Naterer (2008) further provides an example the application of salt/ammonia and MgO/vapor chemical heat pumps to the Cu—Cl thermochemical cycle production of oxygen and hydrogen.
In one embodiment, thermal energy from exothermic processes of the method is captured by one or more chemical heat pumps. In another embodiment, said thermal energy is transformed to another temperature. In yet another embodiment, said thermal energy is used in endothermic processes of the method. In certain embodiments, thermal energy from the water splitter cycle and/or the methanol reactor is captured, transformed to another temperature, and used in endothermic processes of the method.

System

In one aspect, provided herein is an oxy-fuel combustion system, comprising:
a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide; one or more heat exchangers for capturing heat produced by the oxy-fuel combustion and transferring said heat to a water splitter; and a water splitter, for the conversion of heat and electricity into hydrogen gas and oxygen gas.
In one embodiment of the system, the water splitter is arranged and disposed to provide oxygen to the combustion chamber for use in oxy-fuel combustion.
In one embodiment, the system further comprises an air separator unit arranged and disposed to provide oxygen to the combustion chamber for use in oxy-fuel combustion. In another embodiment, the system further comprises an external oxygen supply.
In one embodiment, the system further comprises an air separator unit, wherein the combustion chamber is arranged and disposed to receive oxygen from the air separator unit and/or external oxygen supply and/or the water splitter.
In one embodiment of the system, one or more heat exchangers serve as the means to capture the heat produced by the oxy-fuel combustion and to transfer said heat to the water splitter. In another embodiment, said one or more heat exchangers captures heat from the heated flue gas. In yet another embodiment, said one or more heat exchangers are selected from: a convective recuperator, a radiation/convective recuperator, a ceramic recuperator and a regenerator. In still another embodiment, said one or more heat exchangers are selected from appropriately scaled heat exchangers which function within specified heat ranges (i.e., heat ranges specified herein or known to those of skill in the art).
In one embodiment of the system, the amount of oxygen required from the air separator unit and/or external oxygen supply is reduced or eliminated in proportion to the amount of oxygen provided by the water splitter. The amount of oxygen that is required from the air separator unit, and/or another oxygen source, and/or the water splitter, may be determined by methods taught herein (see, for example, the section of Algorithms and Formulae), or by methods known to those of skill in the art.
In one embodiment of the system, the oxy-fuel combustion products further comprise heated flue gas containing carbon dioxide.
In one embodiment of the system, the oxy-fuel comprises any combustible material. In another embodiment, the oxy-fuel comprises any hydrocarbon-based fuel. In yet another embodiment, the oxy-fuel comprises coal/water slurry. In still another embodiment, the oxy-fuel comprises oil.
In one embodiment of the system, the water splitter produces hydrogen gas and oxygen gas by means of high-temperature electrolysis.
In one embodiment of the system, the water splitter produces hydrogen gas and oxygen gas by means of a thermochemical cycle. In another embodiment, the thermochemical cycle is selected from: a hybrid copper-chlorine cycle; a sulfur-iodine cycle; and a hybrid sulfur cycle. In yet another embodiment, the thermochemical cycle is a sulfur-iodine cycle. In still another embodiment, the sulfur-iodine cycle may be represented by the following steps:

- Step I: 2H₂O+SO₂+I₂→H₂SO₄+2HI
- Step II: H₂SO4→H₂O+SO₂+½O₂
- Step III: 2HI→H₂→I₂

In one embodiment of the system, the thermochemical cycle is a hybrid copper-chlorine cycle. In another embodiment, the hybrid copper-chlorine cycle is selected from: a 3-step cycle, a 4-step cycle, and a 5-step cycle. In yet another embodiment, the hybrid copper-chlorine cycle is the 4-step cycle. In still another embodiment, the 4-step hybrid copper chlorine cycle may be represented by the following steps:

In certain embodiments, n_fis 5-30.
In one embodiment of the system, the water splitter operates at a temperature 450° C. for Step I; 30-80° C. for Step II; 375° C. for Step III; and 530° C. for Step IV. In another embodiment, the heat required to produce 1 mole of O and 1 mole of H₂is about 554.7 kJ/mol.
In one embodiment of the system, the hydrogen from the water splitter is used directly or indirectly in a subsequent process. In another embodiment of the system, the hydrogen from the water splitter is used directly in a subsequent process.
In one embodiment of the system, the hydrogen from the water splitter and the carbon dioxide from the combustion flue gas are reacted to form methanol and water.
In one aspect, provided herein is an oxy-fuel combustion system, comprising:
a combustion chamber arranged and disposed to receive coal/water slurry, oxygen and recycled flue gas, wherein the chamber is arranged and disposed to receive said oxygen from an air separator unit and/or external oxygen source, and/or a water splitter, and combust said coal/water slurry, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide; one or more heat exchangers arranged and disposed to capture heat from said heated flue gas and transfer a portion of the captured heat to a water splitter; a water splitter using a 4-step hybrid copper-chlorine thermochemical cycle for the conversion of heat and electricity into hydrogen gas and oxygen gas, arranged and disposed to transfer the oxygen gas to the combustion chamber for use in said oxy-fuel combustion.
Integration of the unit processes of the system described herein will be well known to those of skill in the art. An example of an integrated system is depicted in FIG. 2 and described in Table 1.

Module

In one aspect, provided herein is a method for converting a non-oxy-fuel combustion system into a recuperative oxy-fuel combustion system, said method comprising:
Providing a system comprising an air separator unit and/or external oxygen source; one or more heat exchangers; a thermochemical and/or electrochemical water splitter; and a flue gas converter; and converting a non-oxy-fuel combustion system into a recuperative oxy-fuel combustion system.
In one embodiment of the method, the non-oxy-fuel combustion system comprises a combustion chamber.
In one embodiment of the method, the water splitter is arranged and disposed to provide oxygen to the combustion chamber for use in oxy-fuel combustion.

Algorithms and Formulae

The following algorithms and formulae pertain to system scaling. Underlying chemical formulae pertaining to water splitters may be found in the Thermochemical Cycles section. System scaling is based on the maximum mass flow rate of the fuel (e.g., coal) used in the boiler or combustor, which in turn is based on the gross cumulative power produced by the entire power plant (MWg), not considering the power penalty incurred by equipment at the plant. The following are specific scaling considerations for the combustion system described herein. Specifically, system scaling is based on the following method (modified from Rubin et al. 2007):

- i. Calculate maximum fuel flow rate;
- ii. Calculate oxygen requirement to meet this fuel flow rate;
- iii. Considerations for sizing the water splitter to meet the oxygen flow rate
- iv. Calculate CO₂flow rate from combustion
- v. Calculate hydrogen flow rate for converting carbon dioxide to methanol; and
- vi. Considerations for sizing the methanol to gasoline section.
  These considerations are described below.
  i. Calculate Maximum Fuel Flow Rate

Fuel flow rate is calculated based on a plant's gross cumulative power (MWg), heat rate and fuel properties (heating value). The relationship in the equation below can be used to determine the fuel flow rate required to generate the desired (or actual) gross power, given the fuel properties and gross heat rate.
$M_{coal} = \frac{{MW}_{g} \times {HR}_{steam}}{2 \times η_{boiler} \times {HHV}_{coal}}$
wherein,

M_coal=mass flow rate of fuel (ton/hr)
MW_g=gross cumulative power produced by the entire power plant; this does not consider power used by equipment in the power plant (MW)
HR_steam=heat rate of the steam cycle, which excludes the effects of the boiler efficiency (Btu/kWh)
η_boiler=boiler efficiency (fraction)
HHV_Coal=higher heating value of the coal on a wet basis (Btu/lb)
ii. Calculate Oxygen Requirement to Meet this Fuel Flow Rate

The maximum rate of oxygen produced by the water splitter is determined as follows:

- a) Calculate the stoichiometric O₂requirement based on the fuel flow rate, fuel composition, and emission factors for incomplete combustion reactants;
- b) Calculate the total O₂requirement based on the excess oxygen specified (approximately 3-5% excess); and
- c) Calculate the total oxygen product (i.e., oxidant) flow rate based on the oxygen purity (≧95%) and total O₂requirement.

This oxygen flow should replace the oxygen that would have been produced by the Air Separation Unit (ASU) in a normal oxy-fuel system.
iii. Considerations for Sizing the Water Splitter to Meet the Oxygen Flow Rate
The water splitter (i.e., Cu—Cl or S—I or HyS thermochemical cycles; or HTE) is sized to produce the maximum oxygen flow rate, thereby replacing the ASU. Sizing the water splitter such that oxygen flow rate is lower than the maximum flow rate for the facility may result in the need to supplement oxygen production with another source (e.g., an ASU).
The water splitter is sized based on stoichiometric oxygen output from particular water splitter reactions to meet the O₂flow rate determined above.
iv. Calculate CO₂Flow Rate from Combustion
The carbon dioxide mass flow rate, m_FG, can be derived based on the following equation:
m _CO2 =[m _COAL(1−% ash)+m _RFG +m _O2 −m _IMPURITIES](CO2_CAPTURE)(CO2_PURITY)

Wherein:

m_FG=flue gas mass flow rate
% ash=fuel ash content, mass fraction
m_RFG=recycle flue gas mass flow rate
m_Impurities=mass of impurities removed during CO2 purification (i.e., NOx, SOx and Hg)
CO_2Capture=capture efficiency of CO2 in the flue gas
CO_2pUrity=carbon dioxide purity requirement (generally≧95%)

Zhou et al. (2010) found that oxy-fuel combustion in a conventional utility boiler had an ideal flue gas recycle (FGR) ratio generally around 0.7-0.75; whereas Hong et al. (2008) found a flue gas recycle ratio in a pressurized coal combustor to be about 0.78. Flue gas ratios depends on boiler exit O₂and fuel properties such that flue gas recycle ratio is a linear function of the boiler exit O₂and increases slightly with air-to-fuel ratio.
v. Calculate Hydrogen Flow Rate for Converting Carbon Dioxide to Methanol
The hydrogen flow rate is determined stoichiometrically based on the following equation:
CO₂+3H₂→CH₃OH+2H₂O
Thus, the flow rate of hydrogen is three times the flow rate of carbon dioxide on a molar basis. This likely will be offset somewhat by the overall efficiency of carbon dioxide to methanol conversion, which is dependent on specific methanol reactor employed. The methanol reactor shall be sized to accommodate maximum carbon dioxide flow, based on maximum fuel flow rates.
vi. Considerations for Sizing the Methanol to Gasoline Section
The methanol to gasoline reactor system shall be sized to accommodate maximum methanol flow, based, in turn, on maximum fuel and carbon dioxide flow rates, respectively. Methanol to gasoline reactor sizing is based on stoichiometric considerations of the following overall reactions:
The dimethyl ether product is then further dehydrated over a zeolite catalyst (preferred zeolites may include ZSM-5, ZSM-11, ZSM-12, ZSM-35, and ZSM-48) to give a gasoline with 80% C5+ hydrocarbon products. Conversion efficiencies for methanol to gasoline conversion shall also be considered when sizing the reaction system.
Like the Methanol Section, the Methanol to Gasoline Section produces substantial amounts of heat which may be recuperated and transferred through chemical heat pumps and/or other heat exchangers to the Water Splitter Section to power endothermic reactions. Therefore, these Sections should be sized and positioned in a manner which facilitates waste heat recuperation and transfer.

REFERENCES

The following publications are incorporated herein by reference:

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EXEMPLIFICATION

An example of how the recuperative combustion system integrates with the ISOTHERM® pressurized coal combustion follows. The ISOTHERM® system, as designed, makes use of an ASU, which requires nearly 20% of gross facility power output to operate. The recuperative combustion system described herein is integrated with this system, replacing the ASU with a water splitter, and resulting in significant reductions in power penalty for the plant. FIG. 2 provides a system schematic of the integration of the recuperative combustion system with the ISOTHERM® oxy-fuel (coal) combustion process. Table 1 provides the temperature and mass data from each process, and includes a description of each step numerically keyed to the process step numbers in FIG. 2. System efficiencies, heat requirements and electrical requirements may vary with the choice of water splitter employed (i.e., Cu—Cl Cycle, S—I Cycle, HyS Cycle, high-temperature electrolysis or other suitable water splitter known to those in the art). In all cases ASU (and corresponding power requirements) are significantly reduced, or eliminated, due to the production of oxygen from the water splitter, hence increasing system efficiency. Additional heat and mass balance tests are required to further quantify power penalty reductions and corresponding increases in system efficiencies with the recuperative combustion system relative to the ISOTHERM® base case (i.e., where oxygen is supplied for combustion from an ASU).

TABLE 1

Recuperative Combustion System Integration with ISOTHERM PWR System for a 875 MW_TH
Pressurized Coal Fired Power Plant

				Mass Flow
		Pressure	Temperature	Rate
State #	Process Stream State/Description	(bar)	(C.°)	(kg/s)

2-1	Condensate leaves the condenser and is compressed	0.41	32.6	198.1
	by the first feedwater pump.
2-2	The pressurized condensate leaves the first feedwater	11.2	32.9	198.1
	pump and enters the Acid Condenser Unit where
	most of the latent enthalpy in the flue gases is
	recovered.
2-3	Condensate leaves the Acid Condenser Unit and picks	11.2	158.7	198.1
	up more thermal energy from the Coal Combustor
	Unit walls.
2-4	Condensate leaves the Coal Combustor Unit walls and	11.2	177.2	198.1
	enters the Deaerator Unit.
2-5	According to the saturation condition of the deaerator,	11.2	179.9	210
	the design point pressure level fixes the exit
	temperature of the water leaving the deaerator. After
	the deaerator, the feedwater stream at state 2-5 is
	pumped to the supercritical state, by the second
	feedwater pump.
2-6	After leaving the second feedwater pump, the	250	215	210
	supercritical steam feedwater is heated regeneratively
	and enters the Heat Recovery Steam Generator
	(HRSG).
2-7,	The supercritical steam feedwater leaves the Heat	250	600	210
2-9 to 2-12	Recovery Steam Generator (HRSG) where it was
(process	heated to 600° C. at 250 bars. It then enters the High
stream/	Pressure Turbine (HPT) to generate electricity.
state	Resultant subcritical steam leaves the HPT and is
description	circulated for 1) steam injection into the Coal
only)	Combustor Unit (State 2-8); 2) reheat to
	HRSG/Intermediate-Pressure Turbine (IPT; States 2-9
	and 2-10), and reheat to HRSG/Low-Pressure Turbine
	(LPT; States 2-11 and 2-12); and 3) regeneratively
	reheated to supercritical steam as in State 2-6 for entry
	to HRSG followed by the HPT.
2-8	Steam is bled from the high pressure steam turbine to	70	316	3
	be injected into the pressurized combustor in order to
	atomize the slurry particles.
2-13	An oxygen stream is produced from the Water	10	201.2	73.52
	Splitter Section. Produced oxygen is mixed with the
	recycled flue gases, state 2-19.
2-14	The recycled flue gas/oxygen mixture is injected	10	256.5	335.8
	into the pressurized Coal Combustor.
2-15	The Coal Combustor Unit yields flue gases at about	10	1549.7	382.8
	1550° C. Some of the thermal energy in the flue gas
	is transferred through one or more heat
	exchangers to the Water Spltter Section. The flue
	gas temperature downstream of the heat
	exchangers is at least 800° C.
2-16	The flue gas stream leaves the Water Splitter	10	800	1004.7
	Section and enters the HRSG and transfers
	thermal energy to the steam while being cooled
	down to state 2-17.
2-17	The flue gas stream leaves the HRSG and is either 1)	9.351	259.7	1004.7
	recycled, or 2) passed through the Acid Condenser
	Unit.
2-18	The flue gas stream is cooled down to 800° C. - as	10	268.7	621.9
	necessary - by recycled flue gases.
2-19	Recycled gases leaving the HRSG are either used	10	268.7	262.2
	for 1) mixing with the oxygen stream for injection
	into the Coal Combustor Unit, or 2) cooling flue
	gas stream downstream from the Water Splitter
	Section to 800° C.
2-20	A portion of the flue gas exhaust stream goes to acid	9.351	259.7	120.6
	condenser for cooling/heat recovery
2-21	Cooled flue gas leaves acid condenser and enters	9.351	60.51	87.7
	Carbon Dioxide Purification/Compression Unit.
2-22	CO₂is pumped to the Methanol Reactor Section.	110	30	72.5
2-23	Exhaust gases are monitored and vented.	1.2	30	16.9
2-24	Hydrogen gas (H₂) is produced in the Water	50	25	9.96
	Splitter Section and pumped to the Methanol
	Reactor.
2-25	A methanol/water mixture is produced in the	75	250	82.46 (Methanol
	Methanol Reactor and separated in the			and Water);
	Methanol/Water Separation Unit through either			52.78
	pervaporation or distillation.			(Methanol);
				29.68 Water)
2-26	Water is separated from methanol through	1.38	26.7	29.68 (water
	distillation in the Methanol/Water Separation Unit,			from Methanol
	and is treated, as necessary, and recirculated to the			Reactor);
	Water Splitter Section). Additional water is			29.56 (water
	provided from State 2-29 (water separated from			from MTG
	petrochemical distillates in MTG Unit) and			Unit); 82.78
	external water sources, as necessary.			(total water
				needed to run
				Water Splitter
				Section to
				produce
				adequate H₂
				and O₂for
				Combustor
				System)
2-27	Methanol resulting from the Methanol/Water	1	25	52.78
	Separation Unit is pumped to bulk storage tanks for
	temporary storage prior to transport and sale, or
	pumped to the Methanol to Gasoline Reactor Unit.
2-28	Hydrocarbon distillate products of the Methanol to	≦55 (MTG	200-540 (MTG	23.22
	Gasoline Reactor/Distillation System are pumped to	Reactor Unit);	Reactor Unit);
	bulk storage tanks for temporary storage prior to	1 (storage)	25 (storage)
	transport and sale.
2-29	Water resulting from the Methanol to Gasoline	1	25	29.56
	Reactor/Distillation System is treated and recirculated
	to the Water-Splitting Section via State 2-26.
2-30	Coal/Water Slurry injection to Coal Combuster. Coal			30 - Coal Only
	is supplied in the form of a coal-water slurry stream
	which contains 0.35 kg water per 1 kg of its total
	weight.

Claims

1. A method for oxy-fuel combustion, comprising:

providing a system comprising a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat and heated flue gas;

capturing heat produced by the oxy-fuel combustion;

using a portion of the heat to power a water splitter, thereby generating hydrogen gas and oxygen gas; and

transferring the oxygen gas from the water splitter to the combustion chamber for use in said oxy-fuel combustion.

2. The method of claim 1, further comprising providing an air separator unit, or another external oxygen supply, wherein the combustion chamber is arranged and disposed to receive oxygen from the air separator unit and/or external oxygen supply and/or the water splitter.

3. The method of claim 1, further comprising one or more heat exchangers for the capture and transfer of heat from the heated flue gas to the water splitter.

4-5. (canceled)

6. The method of claim 1, wherein the amount of oxygen required from the air separator unit and/or external oxygen supply, is reduced or eliminated in proportion to the amount of oxygen provided by the water splitter.

7. (canceled)

8. The method of claim 1, wherein the oxy-fuel comprises any hydrocarbon-based fuel.

9. The method of claim 8, wherein the oxy-fuel comprises a coal/water slurry.

10. (canceled)

11. The method of claim 1, wherein the water splitter produces hydrogen gas and oxygen gas by means of high-temperature electrolysis.

12. The method of claim 1, wherein the water splitter produces hydrogen gas and oxygen gas by means of a thermochemical cycle.

13. The method of claim 12, wherein the thermochemical cycle is selected from: a hybrid copper-chlorine cycle; a sulfur-iodine cycle; and a hybrid sulfur cycle.

14-15. (canceled)

16. The method of claim 13, wherein the thermochemical cycle is a hybrid copper-chlorine cycle.

17. The method of claim 16, wherein the hybrid copper-chlorine cycle is selected from: a 3-step cycle, a 4-step cycle, and a 5-step cycle.

18. The method of claim 17, wherein the hybrid copper-chlorine cycle is the 4-step cycle.

19. The method of claim 18, wherein the 4-step cycle of the hybrid copper chlorine cycle may be represented by the following steps:

Step I: Cu(s)+2HCl(g)→2CuCl(molten)+H₂(g)

Step II: 4CuCl(s)→2Cu(s)+2CuCl₂(aq)+HCl(aq)

Step III: CuCl₂(aq)+n_fH₂O(l)→CuOCuCl₂(s)+2HCl(g)+(n_f−1)H₂O(g)

Step IV: CuOCuCl₂(s)→2CuCl(molten)+0.50₂(g)

20. (canceled)

21. The method of claim 19, wherein n_fis 5-30.

22. (canceled)

23. The method of claim 1, wherein the hydrogen from the water splitter is used directly or indirectly in a subsequent process.

24. (canceled)

25. The method of claim 1, wherein the hydrogen from the water splitter and the carbon dioxide from the combustion flue gas are reacted to form methanol and water.

26-29. (canceled)

30. The method of claim 1, wherein the amount of oxygen that the combustion chamber requires from an air separator unit and/or external oxygen supply, is reduced or eliminated in proportion to the amount of oxygen received from the water splitter, resulting in increased efficiency.

31. The method of claim 30, wherein the increased efficiency is measured in terms of increased gross power output of the combustion process.

32. The method of claim 31, wherein the gross power output of the combustion process is increased by 1-20%.

33-35. (canceled)

36. An oxy-fuel combustion system, comprising:

a combustion chamber arranged and disposed to receive fuel, oxygen and recycled flue gas and combust said fuel, oxygen and recycled flue gas to produce heat and heated flue gas containing carbon dioxide;

one or more heat exchangers arranged and disposed to capture heat produced by the oxy-fuel combustion and transfer said heat to a water splitter; and

a water splitter, for the conversion of heat and electricity into hydrogen gas and oxygen gas.

37. The system of claim 36, further comprising an air separator unit, or another external oxygen supply, wherein the combustion chamber is arranged and disposed to receive oxygen from the air separator unit, or external oxygen supply, and/or the water splitter.

38. The system of claim 36, wherein one or more heat exchangers serve as the means to capture the heat produced by the oxy-fuel combustion and to transfer said heat to the water splitter.

39-62. (canceled)