The present disclosure relates to methods for removing carbon dioxide from a flowing gas stream, such as the flue gas of a coal fired power plant. The present disclosure also relates to apparatus' for removing carbon dioxide from such flowing gas streams.

Background

Much of the energy used in the world today is derived from the combustion of hydrocarbons, such as coal, oil, and natural gas. In addition to carbon and hydrogen, these fuels often contain numerous other contaminants such as sulfur, mercury, and nitrogenous compounds. Combustion of these fuels produces waste gases that include acidic components, such as sulfur and nitrogen oxides. Due to the harmful effects of these so-called "acid gases" on the environment, stringent regulations mandate the use of pollution control equipment by power plants producing energy from the combustion of hydrocarbon fuels. These regulations have prompted the development of numerous processes for capturing acid gases from the flue gas stream of power plants, before they are emitted into the atmosphere.

Due to its contribution to global warming, carbon dioxide (CO2) emissions have recently been targeted for similar regulation as other acidic gas emissions. CO ₂ emissions from the electric power sector account for approximately 40 percent of the total energy related CO ₂ emissions in the United States. As roughly 50 percent of U.S. electricity is generated from coal, it is becoming increasingly important that CO ₂ capture solutions are developed which are suitable for use in existing coal-burning plants, as well as for planned, new capacity. Currently, several CO ₂ capture technologies have been proposed for use on coal fired power plants. However, few or none of those technologies have been deployed at the scale of a commercial power plant. Moreover, existing CO ₂ capture solutions have significant capital and operating costs, which may be uneconomical for use with a traditional coal-fired plant. One known method for capturing CO ₂ from a gas stream is thermal swing absorption ("TSA"). In a conventional TSA process, CO ₂ is absorbed from a gas stream by a CO ₂ lean solvent in an absorber column. The resulting CO ₂ rich solvent is routed to a regenerator column, where it is heated to evolve the absorbed CO ₂ and regenerate the CO ₂ lean solvent. The regenerated CO ₂ lean solvent is then circulated back to the absorber column for reuse in the absorption process.

FIG. 1A schematically illustrates a conventional TSA process for removing carbon dioxide from a gas stream, such as flue gas from a power plant. In this process, gas stream 101 (e.g., flue gas) enters a bottom portion of absorber column 102. Within absorber column 102, gas stream 101 comes into contact with CO ₂ lean solvent 104. During this contact, CO ₂ lean solvent absorbs CO ₂ from gas stream 101, thereby producing CO ₂ rich solvent 105 and scrubbed gas stream 103.

CO ₂ rich solvent 105 exits absorber column 102 and is conveyed to a liquid entrance 106 of regenerator column 108. Within regenerator column 108, CO ₂ rich solvent 105 is heated to evolve offgas 112, thereby producing regenerated CO ₂ lean solvent 104'. If the process is implemented to capture CO2 from a coal-fired power plant, the heat required to for the regeneration process is typically supplied by a steam feed 109 from the plant. The steam feed either provides "stripping" steam to the regenerator column (i.e., steam directly injected into the column), or is used to produce steam from the liquid at the bottom of the regenerator column, e.g., via reboiling.

Offgas 1 12 comprises water vapor and carbon dioxide, and exits regenerator column 108 via vapor exit 1 11. Offgas 1 12 is then dried (e.g., via a condenser) and compressed for use in other processes (not shown). Regenerated CO ₂ lean solvent 104' exits regenerator column 108 via liquid exit 1 10, and is recirculated back to absorber column 102 for reuse in the absorption process. Conventional TSA processes such as the one shown in FIG.1 A are energy inefficient. One factor contributing to this inefficiency is the loss of much of the heat inputted to the system during the regeneration process. The CO ₂ rich solvent is heated as it travels down the regenerator column. During heating, the CO ₂ rich solvent evolves offgas and regenerates CO ₂ lean solvent. Near the bottom of the regenerator, the solvent reaches its "regeneration temperature," which is determined by the solvent composition near the bottom of the regenerator column and the regenerator operating pressure.

The resulting regenerated CO ₂ lean solvent exits the regenerator at an elevated temperature, typically the regeneration temperature. In many cases, this elevated temperature reduces or limits the capacity of the regenerated CO ₂ lean solvent to absorb CO ₂ .

To improve CO ₂ absorption capacity, the regenerated CO ₂ lean solvent is typically cooled before it is reused in the absorption process. As a result, much or all of the heat added during the regeneration process is lost as sensible heat. Further, additional energy must be expended in order to drive the cooling process.

Various modifications have been proposed to improve the energy efficiency of conventional TSA processes. For example, one method for cooling a hot regenerated lean solvent stream in a TSA process utilizes a heat exchanger to pre-heat the CO ₂ rich solvent prior to its entry into a regenerator column. Such a process is illustrated in FIG. IB, where cross flow heat exchanger 107 functions to transfer heat from hot regenerated CO ₂ lean solvent 104' to CO ₂ rich solvent 105. This heat transfer generally provides two benefits. First, it preheats CO ₂ rich solvent 105 prior to its entry to the regenerator column, thereby lowering the amount of heat required to produce offgas 1 12 and regenerated CO ₂ lean solvent 104' in regenerator column 108. Second, it cools regenerated CO ₂ lean solvent 104' before its reintroduction into absorber column 102, which increases the CO ₂ absorption capacity of the regenerated CO ₂ lean solvent prior to its reintroduction into absorber column 102. Economic and functional limitations often limit the amount of heat that may be cost- effectively transferred by cross flow heat exchanger 107 from regenerated CO ₂ lean solvent 104' to CO ₂ rich solvent 105. Thus, even if cross flow heat exchanger 107 is utilized, further cooling of regenerated CO ₂ lean solvent 104' is often still necessary.

U.S. Patent No. 3,962,404 ("the '404 patent") describes another method for improving the efficiency of a TSA process. Specifically, the '404 patent describes a TSA process using primary and secondary regeneration columns that are operated at different pressures to facilitate the dissociation of absorbed CO ₂ from a CO ₂ rich solvent and the regeneration of a CO ₂ lean solvent. The '404 patent's process uses heat from the regenerated CO ₂ lean solvent exiting the primary regeneration column to aid the regeneration of lean solvent in the secondary column. As a result, at least some of the heat added to the primary column during regeneration is recovered in the secondary column.

Another method for improving the efficiency of a TSA process is described in U.S. pre-grant publication no. 2009/0205946 ("the '946 publication"). The '946 publication utilizes mechanical vapor recompression ("MVR") to recover at least some of the heat added during the regeneration process. Specifically, the '946 publication describes a process wherein a portion of the regenerated lean solvent is flashed to generate flashed steam, which is compressed and routed to the regenerator column for reuse in the regeneration process. No cross flow heat exchanger is used in the processes described by the '946 publication.

While these previously proposed modifications can improve the energy efficiency of conventional TSA processes, they focus only on recapturing and/or recycling heat added to the system during the regeneration process. They do not recognize the opportunity for improving the economics and energy efficiency of the TSA process by adjusting operational parameters at other points in the process. Accordingly, there remains a need in the art for other methods and configurations for improving the efficiency of a TSA process for capturing CO ₂ from a gas stream, such as flue gas. The present disclosure is directed at satisfying this need.

Summary

One aspect of the present disclosure relates to methods for removing carbon dioxide from a gas stream. In some embodiments, the methods described herein include providing a CO ₂ rich solvent to a liquid entrance of a regenerator column at a CO ₂ rich solvent inlet temperature. Within the regenerator column, the CO ₂ rich solvent is heated to produce a regenerated CO ₂ lean solvent and an offgas comprising carbon dioxide and water vapor. The offgas and regenerated CO ₂ lean solvent are removed from the regenerator column via a vapor exit and a liquid exit, respectively. The amount of water vapor in the offgas leaving the regenerator column is controlled by adjusting the temperature of the regenerated CO ₂ lean solvent after the liquid exit.

In some non-limiting embodiments, the water vapor content of the offgas is adjusted by independently controlling the CO ₂ rich solvent inlet temperature and the temperature of the regenerated CO2 lean solvent exiting the liquid exit.

In further non-limiting embodiments, the methods according to the present disclosure include conveying a CO ₂ rich solvent at a temperature T3 to a liquid entrance of a regenerator column. The regenerator column further includes a vapor exit and a liquid exit. The CO ₂ rich solvent is heated in the regenerator column to produce a CO ₂ lean solvent at a fourth temperature T4 and an offgas comprising carbon dioxide and water vapor. The regenerated CO ₂ lean solvent at T4 and the offgas are removed from the regenerator column via the liquid exit and vapor exit, respectively. T4 is then adjusted to a fifth temperature T5. The amount of water vapor in the offgas is controlled by T5. Another aspect of the present disclosure relates to apparatus' for removing CO ₂ from a gas stream. In one non-limiting embodiment, an apparatus according to the present disclosure includes a regenerator column having a liquid entrance for receiving a CO ₂ rich solvent, a vapor exit for discharging an offgas comprising CO ₂ and water vapor, and a liquid exit for discharging a regenerated CO ₂ lean solvent. The apparatus further includes a heat source for directly or indirectly heating (e.g., via direct steam injection from a boiler, or via reboiling) the CO ₂ rich solvent to a regeneration temperature in the regenerator column, thereby producing the regenerated CO ₂ lean solvent and the offgas. The apparatus further includes a heat recovery apparatus operatively coupled to receive regenerated CO ₂ lean solvent from the liquid exit of the regenerator column. A heat exchanger is operatively coupled to the heat recovery apparatus, and transfers heat from the regenerated CO ₂ lean solvent to the CO ₂ rich solvent. The heat recovery apparatus adjusts the amount of water vapor in the offgas discharged from the vapor exit by controlling the temperature of the regenerated CO ₂ lean solvent.

Additional objects and advantages of the present disclosure will be set forth in part in the description which follows, and in part will be obvious from the description, or may be learned by practice of the present disclosure. The objects and advantages of the present disclosure will be realized and attained by means of the elements and combinations particularly pointed out in the appended claims.

It is to be understood that both the foregoing general description and the following detailed description are exemplary and explanatory only, and are not restrictive of the claims. BRIEF DESCRIPTION OF THE DRAWINGS

The advantages, nature, and various additional features of the present disclosure will appear more fully upon consideration of the illustrative embodiments described below in detail in connection with the accompanying drawings. In the drawings: FIG. 1A is a schematic illustration of a conventional TSA process.

FIG. IB is a schematic illustration of a conventional TSA process including a cross flow heat exchanger for transferring heat from a regenerated CO ₂ lean solvent to a CO ₂ rich solvent.

FIG. 2A is a schematic illustration of a TSA process according to the present disclosure.

FIG. 2B is a schematic illustration of a TSA process according to the present disclosure, including a mechanical vapor recompression ("MVR") process and a cross flow heat exchanger for transferring heat from regenerated CO ₂ lean solvent leaving the MVR process to a CO ₂ rich solvent.

FIG. 2C is a magnified view of an MVR process/apparatus coupled to a regenerator column and a cross flow heat exchanger.

FIG. 2D is a schematic illustration of a TSA process according to the present disclosure, which includes means for transferring heat from a hot regenerated CO ₂ lean solvent to a boiler feed water stream, and a cross flow heat exchanger for transferring heat from a cooled regenerated CO ₂ lean solvent stream leaving the means for transferring heat to a CO ₂ rich solvent stream.

FIG. 3 is a plot of regenerator outlet H ₂ O and CO ₂ concentration vs. ΔΤ, based on measured data using monoethanolamine as a CO ₂ absorbing solvent.

FIG. 4 is a plot of regenerator outlet H ₂ O concentration vs. regenerator inlet temperature, based on measured data points using a proprietary amine solvent produced by Powerspan Corp., and predicted data points.

DETAILED DESCRIPTION

Conventional TSA processes generally produce an offgas comprising carbon dioxide and a significant amount of water vapor, e.g., 70 mol% or more of the offgas. As a result, offgas produced by conventional TSA processes requires a substantial quantity of energy to convert water from the liquid phase to the vapor phase, particularly when compared to the reaction energy needed to release CO _2. The offgas produced by conventional TSA processes also often requires energy intensive drying (e.g., with a condenser) to remove excess water vapor, in order to produce a concentrated CO ₂ gas stream for sequestration. Moreover, the high water vapor content typically dictates the use of larger, more efficient, and/or more sophisticated drying equipment, leading to an increase in capital cost.

Significant efficiencies and cost savings may be realized by controlling the water vapor content of the offgas exiting the regeneration column in a TSA process for capturing CO ₂ from a gas stream, i.e., prior to any drying process. In particular, lowering the water vapor content of the offgas may lead to a reduction in the amount of heating energy (e.g., stripping steam) supplied to the regenerator column, as well as a reduction in the size (and thus, the cost) of the drying equipment need to dry the offgas and produce a sequestration ready CO ₂ gas stream. Further, the methods and apparatus described herein may lower the operating and capital cost of a TSA process, e.g., by reducing heat inputs due to water vapor evaporation, lowering equipment requirements at other points in the process, and/or recapturing and reusing heat inputted to the system during the regeneration process.

Accordingly, one aspect of the present disclosure relates to methods for capturing CO ₂ from a gas stream via a TSA process, wherein the amount of water vapor present in the offgas produced during the regeneration of a CO ₂ lean solvent is controlled. For example, some methods according to the present disclosure produce, prior to any drying process, an offgas comprising carbon dioxide and water vapor, wherein the water vapor accounts for less than about 70 mol% of the offgas. In some embodiments, the methods described herein can produce, prior to a drying process, an offgas comprising water vapor in an amount ranging from about 69 mol%, 65 mol%, 60 mol%, 58 mol%, 55 mol%, 54 mol%, 50 mol%, 45 mol%, 40 mol%, 38 mol%, 35 mol%, 31 mol%, 30 mol%, 29 mol%, 25 mol%, 24 mol% 20 mol%, 15 mol%, 10 mol%, 5 mol%, and/or 1 mol% of the offgas. Of course, the water vapor content of the offgas may fall within, above, or below these endpoints, without limitation.

In accordance with the present disclosure, the water vapor content of the offgas produced during a TSA process may be controlled in various ways, such as by adjusting the CO ₂ rich solvent temperature at an entrance to the regenerator column (hereafter, the "CO ₂ rich solvent inlet temperature"), relative to the regeneration temperature. Specifically, the inventors have found that at a set regeneration temperature/pressure, the amount of water vapor in the offgas decreases as the CO ₂ solvent inlet temperature decreases. Conversely, as the CO ₂ solvent inlet temperature increases relative to a set regeneration

temperature/pressure, the amount of water vapor in the offgas increases.

The water vapor content in the offgas may also be controlled by adjusting the difference between the CO ₂ rich solvent inlet temperature and the regeneration temperature. Hereafter, this temperature differential is referred to as the "delta T across the regenerator," or simply, "ΔΤ." In other terms, some methods according to the present disclosure can obtain desired offgas water vapor content over a range of CO2 rich solvent inlet temperatures and regeneration temperatures/pressures by maintaining ΔΤ at a constant or relatively constant value. As implied by the foregoing, ΔΤ may be controlled by adjusting the CO ₂ rich solvent inlet temperature alone, the regeneration temperature alone, or by adjusting both the CO ₂ rich solvent inlet temperature and the regeneration temperature, relative to one another.

Accordingly, another aspect of the present disclosure relates to methods of controlling the water vapor content of the offgas exiting a regenerator column in a TSA process for removing CO ₂ from a gas stream. In some embodiments, the water vapor content of the offgas is controlled by adjusting the CO ₂ rich solvent inlet temperature relative to a set regeneration temperature/pressure. In other embodiments, the water vapor content of the offgas is controlled by adjusting the regeneration temperature/pressure, relative to a set CO ₂ rich solvent inlet temperature. In additional embodiments, the water vapor content of the offgas is controlled by adjusting ΔΤ, i.e., by adjusting both the CO ₂ rich solvent inlet temperature and the regeneration temperature/pressure, relative to one another.

The CO ₂ rich solvent inlet temperature may range, for example, from about 190 to about 250 °F, such as from about 195 to about 245 °F, more specifically from about 200 to about 240 °F. In some embodiments, the methods according to the present disclosure use a C0 ₂ rich solvent inlet temperature of about 200 °F, 205 °F, 210 °F, 21 1 °F, 215 °F, 217 °F 220 °F, 225 °F, 230 °F, 235 °F , 240 °F, and 245 °F. Of course, the C0 ₂ rich solvent may be set within, above, or below these ranges, as the solvent composition, regeneration temperature/pressure, economics, and operating parameters of the process permit.

The regeneration temperature used in the methods according to the present disclosure may vary widely. Generally, the lower limit of the regeneration temperature is set by the boiling point of the regenerated CO ₂ lean solvent near the bottom of the regenerator column at a particular operating pressure of the regenerator column. That is, if the regenerator column is run under atmospheric pressure, the regeneration temperature is approximately the boiling point of the regenerated CO ₂ lean solvent (typically governed by composition). As the operating pressure of the regenerator column increases, the boiling point of the regenerated CO ₂ lean solvent (and hence, the regeneration temperature) increases. The upper limit of the regeneration temperature is governed by various practical considerations, such as the capital cost of high pressure equipment, and the temperature at which the regenerated CO2 lean solvent degrades.

The methods according to the present disclosure may, for example, employ a regeneration temperature ranging from about 250 to about 310 °F, from about 260 to about 300 °F, from about 270 to about 290 °F, or even about 270 to about 280 °F. In some embodiments, the methods according to the present disclosure employ a regeneration temperature of about 250 °F, 251 °F, 255 °F, 266 °F, 260 °F, 265 °F, 270 °F, 272 °F, 275 °F, 280 °F, 285 °F, 286 °F, 290 °F, 295 °F, 300 °F, or higher. Of course, other regeneration temperature ranges and endpoints are possible as permitted by the composition of the regenerated CO ₂ lean solvent and other practical/economic limitations on the system.

As previously explained, the regeneration temperature can be dictated by the operating pressure of the regenerator column ("regeneration pressure"). A wide range of regeneration pressures may be used in accordance with the present disclosure. For example, the methods according to the present disclosure may use a regeneration pressure of less than about 60 psia, such as about 50 psia, about 40psia, about 30 psia, or about 20 psia. In some embodiments, the methods according to the present disclosure employ a regeneration pressure of about 15, 20, 25, 28, 36, 40, 45, 50, 55, and 60 psia. Of course, regeneration pressures that are above, below, and within these endpoints are possible, as permitted by the composition of the regenerated CO ₂ lean solvent and other practical/economic limitations on the system.

ΔΤ may vary considerably in the methods according to the present disclosure. For example, ΔΤ may be set to about 10, 15, 20, 21 25, 26, 30, 31, 35, 36, 40, 41, 45, 46, 47, 49, 50, 51, 52, 55, 56, 57, 60, 61, 65, 65.6, 67, 70, 71, 75, 80, 81, or 85 °F or more, by controlling one or both of the CO ₂ rich solvent inlet temperature and the regeneration temperature. Of course, ΔΤ may be set to any value within, above, or below the aforementioned endpoints, as permitted by the CO2 absorbing solvent and other practical/economic constraints on the system.

As noted above, certain parameters of the methods described herein (e.g., CO ₂ rich solvent inlet temperature, regeneration temperature/pressure etc.) may be impacted by the composition of the CO ₂ absorbing solvent. For example, monoethanolamine (MEA) or an amine solvent other than MEA may be used as the CO ₂ absorbing solvent. While solvent composition may determine the precise values of various operating conditions, the inventors emphasize that the methods described herein may be practiced with any CO ₂ absorbing solvent having similar CO ₂ absorbing behavior as MEA and other amine solvents.

In one non-limiting example of a method in accordance with the present disclosure, an aqueous CO ₂ absorbing solvent comprising an amine is used to absorb CO ₂ from a gas stream in an absorber column. The resulting CO ₂ rich solvent is regenerated in a regenerator column by heating the CO ₂ rich solvent to 252 °F at a pressure of about 28 psia. At a CO ₂ rich solvent inlet temperature of 243 °F, ΔΤ is 9 °F and the offgas produced during regeneration comprises about 30 mol% CO ₂ and about 70 mol% water vapor. Lowering the CO ₂ rich solvent inlet temperature from 243 °F to 225 °F while holding the regeneration temperature constant at 252 °F (ΔΤ = 27 °F) reduces the water vapor content of the offgas to about 52 mol%, while increasing its CO ₂ content to about 48 mol%.

In this non-limiting example, the offgas produced by the second set of operating conditions (ΔΤ = 27 °F) requires substantially less input energy to the regenerator and less drying than the offgas produced by the first set of operating conditions (ΔΤ = 9 °F). This opens up the possibility for substantial cost savings at the regeneration and/or drying stage, e.g., by reducing the size or operating efficiency requirements of the heating and/or drying equipment.

Another aspect of the present disclosure relates to methods for removing CO ₂ from a gas stream in a TSA process, wherein a heat recovery apparatus is used to adjust the water vapor content of the offgas produced during the regeneration of CO ₂ lean solvent, by directly or indirectly adjusting the temperature of a CO ₂ rich solvent at an inlet to a regenerator column. As a non-limiting example of such a method, reference is made to FIG. 2A. As shown, gas stream 201 (e.g., flue gas) comes into contact with CO ₂ lean solvent 204 in absorber column 202. CO ₂ lean solvent 204 absorbs CO ₂ from gas stream 201, producing scrubbed gas stream 203 and CO ₂ rich solvent 205. CO ₂ rich solvent 205 is conveyed through cross flow heat exchanger 207, where it is warmed by regenerated CO ₂ lean solvent 204'. CO ₂ rich solvent 205 is then conveyed to liquid entrance 206 of regenerator column 208. Within regenerator column 208, CO ₂ rich solvent is heated (e.g., by contact with stripping or reboiled steam produced by steam feed 209) to evolve offgas 212 and produce regenerated CO ₂ lean solvent 204'. Offgas 212 comprises CO ₂ and water vapor, and exits regenerator column 208 via vapor exit 211. Regenerated CO ₂ lean solvent 204' exits regenerator column 208 via liquid exit 210, and is routed through cross flow heat exchanger 207. Ultimately, regenerated CO ₂ lean solvent 204' is returned to absorber column 202 for reuse in the absorption process.

The method shown in FIG. 2A further includes at least one of heat recovery apparatus' 213, 213', and 213" which is/ are configured to control the water vapor content of offgas 212 by directly or indirectly controlling the temperature of CO2 rich solvent 205 at liquid entrance 206 of regenerator column 208. For example, heat recovery apparatus 213, 213', and/or 213" may be configured to control the temperature of the CO ₂ rich solvent 205 at liquid entrance 206 by directly or indirectly adjusting (e.g., via evaporation, heat transfer, etc.) the temperature of at least one of regenerated CO ₂ lean solvent 204', CO ₂ lean solvent 204, or CO2 rich solvent 205, respectively, at a constant regeneration temperature/pressure.

Alternatively, heat recovery apparatus 213, 213', and/or 213" may be configured to maintain, directly or indirectly, the temperature of CO ₂ rich solvent 205 at liquid entrance 206, while the regeneration temperature/pressure is increased. Under either of these methods, one or more of heat recovery apparatus 213, 213', 213" can control the water vapor content of offgas 212 by controlling ΔΤ across regenerator column 208. Of course, a specific ΔΤ can be attained by configuring one or more of heat recovery apparatus 213, 213', 213" to lower, directly or indirectly, the temperature of CO ₂ rich solvent 205 at liquid entrance 206, while the regeneration temperature/pressure is increased.

In some embodiments, the methods according to the present disclosure control the CO ₂ rich solvent inlet temperature by adjusting the temperature of the regenerated CO ₂ lean solvent at a point between the liquid exit of the regenerator column and the cross flow heat exchanger. A non-limiting example of such a method is illustrated in FIG. 2B. As shown, heat recovery apparatus 213 is configured as a mechanical vapor recompression (MVR) system operatively coupled to liquid exit 210 of regeneration column 208. The MVR system includes a flash tank 214 and a compressor 215.

In operation, regenerated CO ₂ lean solvent 204' leaves liquid exit 210 at a set temperature and pressure, typically the regeneration temperature/pressure. It is then conveyed to MVR system 213. Within MVR system 213, at least a portion of regenerated CO ₂ lean solvent 204' is flashed at a flashing pressure lower than the vapor pressure of regenerated CO ₂ lean solvent 204'. Rapid vaporization of at least a portion of regenerated CO ₂ lean solvent 204' occurs; producing cooled regenerated CO ₂ lean solvent 204" and vapor stream 216.

Vapor stream 216 comprises steam and optionally other components (e.g., CO2), and is conveyed to compressor 215 (e.g., a dynamic compressor, positive-displacement compressor, or thermocompressor), where its pressure is increased sufficiently to elevate its temperature above the regeneration temperature used in the regeneration process. The resulting pressurized vapor stream 217 is then returned to regenerator column 208, where its heat contributes to the regeneration of CO ₂ lean solvent 204. Cooled regenerated lean solvent 204" is conveyed to cross flow heat exchanger 207, and ultimately is routed back to absorber column 202 for reuse in the absorption process.

FIG. 2C provides an expanded view of one MVR process according to the present disclosure. As shown in this non-limiting example, regenerated CO ₂ lean solvent 204' exits regenerator column 208 via liquid exit 210 at a regeneration temperature (Tl) of 254 °F, and a pressure of 30 psia. Upon entering flash tank 214, regenerated CO ₂ lean solvent 204' is flashed at a pressure of 21 psia, thereby forming vapor stream 216 and cooled regenerated CO ₂ lean solvent 204", each of which has a temperature T2 of 235 °F and a pressure of 21 psia. Vapor stream 216 is compressed in compressor 215 to a pressure of 32 psia and a temperature of 333 °F, and conveyed back to regenerator column 208 for reuse in the regeneration process. Cooled CO ₂ lean solvent 204" is conveyed to cross flow heat exchanger 207 (not shown), and ultimately back to absorber column 202.

The inventors have found that at a given set of absorber and regenerator operating conditions, the temperature of CO ₂ rich solvent 205 at liquid entrance 206 of regenerator column 208 (i.e., the CO2 rich solvent entry temperature) may be controlled by adjusting the flashing pressure within flash tank 214. Specifically, as the flashing pressure is lowered, the temperature of cooled regenerated lean solvent 204" is reduced. This reduction in temperature results in a corresponding (though not necessarily a 1 : 1) reduction in the CO ₂ rich solvent inlet temperature.

One explanation for the reduction in CO ₂ rich solvent inlet temperature is that because cooled regenerated CO ₂ lean solvent 204" is at lower temperature than regenerated CO ₂ lean solvent 204', it transfers less heat (relative to CO ₂ lean solvent 204') to CO ₂ rich solvent 205 in cross flow heat exchanger 207. As a result, CO ₂ rich solvent 205 is "warmed less" by regenerated CO ₂ lean solvent 204" than it would have been had regenerated CO ₂ lean solvent 204' been directly conveyed to cross flow heat exchanger 207. From this, the inventors determined that at a particular set of operating conditions, both ΔΤ across regenerator column 208 and the water vapor content of offgas 212 may be controlled by adjusting the flashing pressure in flash tank 214.

Another non-limiting embodiment of the present disclosure is illustrated in FIG. 2D.

As shown, heat recovery apparatus 213 includes heat exchanger 218, which is configured to recover at least some of the heat from regenerated CO ₂ lean solvent 204' prior to delivery of regenerated CO ₂ lean solvent 204' to cross flow heat exchanger 207. For example, heat exchanger 218 may be configured to transfer heat from regenerated CO ₂ lean solvent 204' to a boiler feed water stream 219. Boiler feed water stream 219 may, for example, feed water to the boiler of a power plant. The resulting cooled regenerated lean solvent 204" is conveyed from heat exchanger 219 to cross flow heat exchanger 207, and ultimately is routed back to absorber column 202 for reuse in the absorption process.

The method illustrated in FIG. 2D is believed to operate in much the same way as the method described in FIG. 2C. That is, cooled regenerated lean solvent 204" transfers less heat to CO ₂ rich solvent 205 in cross flow heat exchanger 207, resulting in a corresponding (though not necessarily 1 : 1) reduction in the temperature of CO ₂ rich solvent 205 at liquid entrance 206. Thus, at a given set of operating conditions, both ΔΤ across regenerator column 208 and the water vapor content of offgas 212 may be controlled by adjusting the temperature of cooled regenerated solvent 204" with heat exchanger 218.

Economic and/or practical considerations may limit the size, efficiency, etc. of MVR systems and/or heat exchangers that may be economically used as heat recovery apparatus 213. In such circumstances, the amount of heat removed by heat recovery apparatus 213 may be insufficient to cool the regenerated CO ₂ lean solvent 204' to a desired temperature. As a result CO ₂ rich solvent 205 may be warmed more than is optimum by cooled regenerated CO ₂ lean solvent 204" in cross flow heat exchanger 207, thus raising the CO ₂ rich solvent inlet temperature and potentially lowering ΔΤ to a less than optimum value. Thus, additional active cooling or another mechanism of removing, capturing, or otherwise dealing with excess heat in cooled regenerated CO ₂ lean solvent 204" may be desirable.

In some non-limiting embodiments, the methods of the present disclosure address the issue of excess heat in cooled regenerated CO ₂ lean solvent 204" by bypassing at least a portion of the regenerated solvent flow around cross flow heat exchanger 207. Because the bypassed flow never enters cross flow heat exchanger 207, its heat is not available for transfer to CO ₂ rich solvent 205. Thus, the amount of heat transferred to CO ₂ rich solvent 205 can be controlled by adjusting the flow of cooled regenerated CO ₂ lean solvent 204" though cross flow heat exchanger 207. Of course, other methods of further removing heat from cooled regenerated CO ₂ lean solvent 204", or of limiting heat transfer between cooled regenerated CO ₂ lean solvent 204" and CO ₂ rich solvent 205, may be used.

Another aspect of the present disclosure relates to apparatus' for removing CO ₂ from a gas stream. In some embodiments, such apparatus' include a regenerator column comprising a liquid entrance for receiving a CO ₂ rich solvent, a vapor exit for discharging an offgas comprising CO ₂ and water vapor, and a liquid exit for discharging a regenerated CO ₂ lean solvent. The apparatus further include a heat source for directly or indirectly heating the CO ₂ rich solvent within the regenerator column to produce the regenerated CO ₂ lean solvent and the offgas. A heat recovery apparatus is operatively coupled to the exit of the regenerator column and configured to receive the regenerated CO2 lean solvent. A heat exchanger is operatively coupled to receive regenerated CO ₂ lean solvent from the heat recovery apparatus, and configured to transfer heat from the regenerated CO ₂ lean solvent to the CO ₂ rich solvent. The heat recovery apparatus adjusts the amount of water vapor in the offgas discharged from the vapor exit by directly or indirectly controlling the temperature of the regenerated CO ₂ lean solvent discharged from the liquid exit.

One non-limiting example of a heat source that may be used in accordance with the apparatus and methods of the present disclosure is a steam feed. The steam feed may be provided by the boiler of a power plant, a reboiler, or another source of steam. The apparatus may be configured such that the heat source directly or indirectly heats the CO ₂ rich solvent in the regenerator column. For example, where the heat source is a steam feed, the steam feed may directly heat the CO ₂ rich solvent by providing a source of stripping steam that is injected into the regenerator column. Alternatively, the heat source may be configured to indirectly heat the CO ₂ rich solvent by producing steam from liquid at the bottom of the regenerator column, which ultimately contacts the CO ₂ rich solvent. Of course, other heat sources are useable in the apparatus and methods according to the present disclosure.

The heat recovery apparatus according to the present disclosure may, for example, be in the form of an MVR system or a heat exchanger for pre-heating a boiler water feed, as previously described. The heat recovery apparatus may also be configured to transfer heat via an intermediate transfer fluid. The heat recovery apparatus functions to adjust the water vapor content of the offgas leaving the regenerator column, e.g., by controlling the CO ₂ rich solvent inlet temperature and hence, ΔΤ. A combination of two or more heat recovery apparatus' may also be used.

In some embodiments, the methods and apparatus' described herein are configured such that the net energy usage of the system is reduced, relative to a conventional TSA process. For example, the methods and apparatus may be configured to recover at least a portion of the sensible heat that would otherwise be lost in a conventional TSA process, while simultaneously controlling the water vapor content of the offgas produced during

regeneration. In some embodiments, the methods and apparatus of the present disclosure recover at least some of the sensible heat produced during the regeneration process with MVR, boiler water pre-heating, or another process for the beneficial use of sensible heat. Such processes and apparatus' are further configured to reduce the water vapor content of the offgas produced during regeneration by controlling ΔΤ across the regenerator. Non-limiting examples of such methods and apparatus are provided in the examples below.

Other than in the examples, or where otherwise indicated, all numbers expressing endpoints of ranges, and so forth used in the specification and claims are to be understood as being modified in all instances by the term "about." Accordingly, unless indicated to the contrary, the numerical parameters set forth in the specification and attached claims are approximations that may vary depending upon the desired properties sought to be obtained by the present disclosure. At the very least, and not as an attempt to limit the application of the doctrine of equivalents to the scope of the claims, each numerical parameter should be construed in light of the number of significant digits and ordinary rounding approaches.

Notwithstanding that the numerical ranges and parameters setting forth the broad scope of the present disclosure are approximations, unless otherwise indicated the numerical values set forth in the specific examples are reported as precisely as possible. Any numerical value, however, inherently contains certain errors necessarily resulting from the standard deviation found in their respective testing measurements.

Examples

Laboratory scale experiments were conducted to determine the impact of ΔΤ and regenerator liquid inlet temperature (i.e., CO ₂ rich solvent inlet temperature in a TSA process for removing CO ₂ from a gas stream) on the water vapor and CO ₂ content of offgas produced during a regeneration process. In these experiments, a CO ₂ lean solvent comprising 30 weight % monoethanolamine (MEA) was introduced to the top of a laboratory scale mass transfer absorption column. Simultaneously a simulated flue gas stream was introduced into a bottom portion of the absorption column. During the resulting counter current contact, the CO ₂ lean solvent absorbed CO ₂ from the simulated flue gas, resulting in the production of a CO ₂ rich solvent.

The CO ₂ rich solvent was introduced into the top of a packed regenerator column at a controlled temperature. As it travelled down the packed column the CO ₂ rich solvent was heated with reboiled steam, resulting in the release of offgas comprising CO ₂ and water vapor. The pressure of the packed column was controlled using a pressure control valve to regulate the offgas flow out of the column.

Over a series of runs, the H ₂ 0 and CO ₂ concentrations in the offgas were measured using Fourier Transform Infrared Spectroscopy as the ΔΤ across the column was increased. The resulting data is provided in Table 1 below. The measured ΔΤ vs. offgas H ₂ 0 content is plotted in FIG. 3.

Table 1: 30 weight % MEA

RILT: Regenerator Inlet Liquid Temperature (CO ₂ rich solvent inlet temperature) RP: Regenerator Pressure

AROLT: Actual regenerator outlet liquid temperature

To further illustrate the impact of ΔΤ on the water vapor and CO ₂ content of the offgas produced during a TSA process, and to demonstrate the practice of the methods described with an alternative solvent, additional sample runs were performed using an aqueous amine solvent other than MEA, namely a proprietary solvent produced by Powerspan Corp. under the name, "ECO ₂ Solvent." Further data points were generated by calculating ΔΤ and offgas water vapor concentration using assumed regenerator column operating parameters, boiling point elevation, and steam table values. The results are reported in Table 2, with samples 4-9, 13, and 18 corresponding to measured data, and examples 10-12, 14-17, and 19-25 corresponding to calculated data.

Table 2

18 225 28 251 26 17.20 61.4 58

19 225 36 266 41 17.20 47.8 **

20 225 40 272 47 17.20 43.0 **

21 225 50 286 61 17.20 34.4 **

22 230 28 251 21 18.93 67.6 **

23 230 36 266 36 18.93 52.6 **

24 230 40 272 42 18.93 47.3 **

25 230 50 286 56 18.93 37.9 **

RILT: Regenerator Inlet Liquid Temperature (CO ₂ rich solvent inlet temperature) RP: Regenerator Pressure

AROLT: Actual regenerator outlet liquid temperature

* Based on steam tables using AROLT and a 5 F BPE

** Not measured

As demonstrated in Table 1, the measured offgas H ₂ 0 content decreased with increasing Delta T, using an MEA solvent. This is confirmed by the data in Table 2, which shows a similar trend in the measured offgas FLO content for both the ECO ₂ solvent and the calculated examples.

As also shown in Table 2, when ΔΤ was decreased by increasing regenerator inlet temperature while holding regenerator temperature/pressure constant, the actual and predicted water vapor content of the offgas increased. Further, the data demonstrates that when ΔΤ was held constant (or relatively constant), the water vapor content of the offgas remained at roughly the same level over a range of regenerator inlet temperatures and regeneration temperatures/pressures.

Further insight is provided by FIG. 4, which plots regenerator inlet temperature of the actual and calculated examples in Table 2 vs. the actual/predicted water vapor concentration of the offgas at four different regenerator pressures. As illustrated by each plot, the water vapor content of the offgas decreased as regenerator inlet temperature decreased (i.e., as ΔΤ increased), regardless of pressure. Further, comparison of these plots to one another revealed that by adjusting ΔΤ (e.g., by controlling regenerator inlet liquid temperature), roughly the same over head vapor water vapor content can be obtained over a range of regeneration pressures. For example, an offgas water vapor content of approximately 40 mol% can be obtained by running a regenerator inlet temperature of 205 °F at a regeneration pressure of 28 psia, or by running a regenerator inlet liquid temperature of about 217 °F, 221 °F, or -232 °F, at a regeneration pressure of 36 psia, 40 psia, and 50 psia, respectively.

The precise relationship of temperature to pressure reported is specific to the solvent utilized in the process as demonstrated by the differences in offgas water vapor content at similar ΔΤ values in the MEA, alternative amine, and calculated examples. However, it is expected (and the data suggests) that the relationship of ΔΤ to water will be generally true for all CO ₂ absorbing solvents that exhibit similar, CO ₂ absorption/desorption behavior.

Other embodiments of the present disclosure will be apparent to those skilled in the art from consideration of the specification and practice of the present disclosure as described herein. It is intended that the specification and examples be considered as exemplary only, with a true scope and spirit of the present disclosure being indicated by the following claims.