Patent Document

CROSS REFERENCE TO RELATED APPLICATIONS 
     This application claims the benefit of U.S. Provisional Application No. 60/890,233, filed Feb. 16, 2007, the entire contents of which are incorporated herein by reference. 
    
    
     BACKGROUND 
     It is believed that there are global warming effects that are being caused by the introduction of increased carbon dioxide into the atmosphere. One major source of carbon dioxide emission is the flue gas that is exhausted as a result of a power generation plant&#39;s combustion process. Therefore, there have been several efforts by governments and utility companies worldwide, to reduce these emissions. 
     There are two principal types of power plants that are based on combustion processes; coal combustion and natural gas combustion. Both of these processes produce carbon dioxide as a byproduct when generating power. Efforts have been made to increase the efficiency of the burner, and, therefore, the basic combustion process itself. The intent of these efforts has been to reduce carbon monoxide (the result of imperfect combustion), oxides of nitrogen, and other pollutants. However, since the production of carbon dioxide and water are the basic products of the chemical reaction of combustion, the most efficient technique to minimize the carbon dioxide emission is to capture as much of the carbon dioxide as possible as it is being created by the power plants. In order to truly maximize the efficiency of this technique, existing coal combustion plants, which represent a large portion of the power generation plants worldwide, must also be targeted. The oxy-combustion technique is very interesting, and has significant advantages, since it can be adapted to existing facilities. 
     Traditional power plants use air as the source of oxidant to combust the fuel (typically coal). Steam is generated by indirect heat exchange with the hot combustion products. The steam is then expanded in turbines to remove useful energy, and, thereby, produce power. The combustion process produces carbon dioxide as a by-product, which is mixed with the residual nitrogen of the combustion air. Due to the high content of nitrogen in the inlet air (78 mol %), the carbon dioxide is diluted in the flue gas. To insure full combustion, the power plants must also run with an excess air ratio that further dilutes the carbon dioxide in the flue gas. The concentration of carbon dioxide in the flue gas of an air combustion plant is typically about 20 mol %. 
     This dilution of the carbon dioxide increases the size and the power consumption of any carbon dioxide recovery unit. Because of this dilution, it becomes very costly and difficult to recover the carbon dioxide. Therefore, it is desirable to produce flue gas with at least about 90% to 95 mol % carbon dioxide, in order to minimize the abatement cost. The current technology for carbon dioxide recovery from flue gas utilizes amine contact tower to scrub out the carbon dioxide. However, the high amount of heat that is needed to regenerate the amine and extract the carbon dioxide, reduces the amine processes cost effectiveness. 
     In order to avoid the dilution of carbon dioxide in the nitrogen, the power generation industry is switching to an oxy-combustion process. Instead of utilizing air as an oxidant, high purity oxygen (typically about 95% purity or better) is used in the combustion process. The combustion heat is dissipated in the recycled flue gas concentrated in the carbon dioxide. This technique makes it possible to achieve a flue gas containing between about 75 mol % and 95 mol % carbon dioxide. This is a significant improvement over the previous concentration of about 20 mol %, which is obtained with air combustion. The purity of carbon dioxide in oxy-combustion&#39;s flue gas ultimately depends on the amount of air leakage into the system and the purity of oxygen being utilized. The necessary high purity oxygen is supplied by an air separation unit. 
     In one example of the traditional oxy-combustion process, the carbon dioxide removal process begins as the flue gas exiting the boiler is cooled and sent to an electrostatic precipitator. A portion of the flue gas is further cooled, the moisture is removed, and this portion of the flue gas is recycled to the coal handling section (mill, dryer, etc). Another portion of the flue gas is recycled back to the boiler, and the remaining portion is extracted as flue gas output and is sent to the carbon dioxide purification unit. One example of this type of oxy-combustion is illustrated in  FIG. 1 . 
     Since pure oxygen, hence power input and capital cost, is required in the oxy-combustion process to facilitate the capture of carbon dioxide, the whole process, including the oxygen plant and the carbon dioxide capture and purification must be very efficient to minimize the power consumption. Otherwise, the economics of the carbon dioxide recovery will become unattractive to the operator of the power generation plant. In summary, the carbon dioxide capture with oxy-combustion is appealing in terms of pollution abatement, however, in order to achieve it, the capital expenditure and the power input must be minimized to avoid a prohibitive increase in power cost. 
     As previously mentioned, carbon dioxide purities of 90% or higher (typically 95% or higher) are desirable for many subsequent carbon dioxide abatement techniques (such as deep well injection, deep sea injection or enhanced oil recovery systems). Due to air leakage and the presence of inert gases in the high purity oxygen (nitrogen and argon), in practice the flue gas can be as low as about 75% carbon dioxide. The carbon dioxide concentration must therefore be increased to 90% to 95% in some type of purification process. Common industry specifications typically require that the overall carbon dioxide recovery ratio must be about 90% and even higher than 95% in some cases. 
     On example, of such a purification system, was described in the Publication of IEA Green House R&amp;D Programme-Oxycombustion Processes for CO 2  Capture From Power Plant (Report Number 2005/9, dated July, 2005). This process is illustrated in  FIG. 2 . 
     In the process indicated in  FIG. 2 , the flue gas is washed. Its acid content is removed, it is compressed to a pressure greater than about 30 bar, then it is dried (stream  1 ). A cryogenic partial condensation process is then utilized to concentrate the carbon dioxide (stream  7  and stream  8 ). 
     The carbon dioxide is further compressed to very high pressure (between about 80 bar and about 120 bar) (stream  9 ). The off-gas leaving the process at 30 bar (stream  10 ) is generally heated to about 300° C., then it is expanded in a hot gas expander in order to more efficiently recover the potential energy. 
     In order to heat to 300° C., the gas must be heated first to about 150° C. by exchanging heat with an adiabatic compressor (i.e. the compression heat is not removed by an intercooler, and the exit temperature is allowed to rise to about 200° C.). The gas is then heated to 300° C. by heat exchange with the flue gas from the boiler. 
     As evidence of these thermal costs, it is noted that an adiabatic compressor (either feed gas or carbon dioxide compressor) consumes more power than the isothermal compressor equipped with intercoolers. Also, the hot gas expander, because of the high expansion ration, (about 30 to 1) and high operating temperature, requires a multiple stage (usually axial type) expander. The skilled artisan will recognize that this type of expander is typically quite expensive. And the heating of the off-gas from about 150° C. to about 300° C. by the flue gas consumes the valuable heat of the boiler, and, therefore, it is possible that steam production will be effected. This will then result in a lower power output from the stream turbines. This reduces the efficiency of the overall process. This also requires a gas-to-gas heat exchanger in the boiler, which, is typically, very expensive. Furthermore, utility companies involved with oxycombustion are also evaluating techniques to minimize the air leakage to further improve the CO 2  content of flue gases. This effort also reduces the flowrate of the off-gas stream, such that its recoverable energy becomes smaller, compared with the total power input. Therefore, it becomes less attractive to use less efficient adiabatic compressors to recover the reduced power content of lower off-gas flow. 
     In another example of the existing art, European patent number 0503910 presents a process scheme, wherein the compressed dry flue gas is treated in 2 distillation columns arranged in series. The first column removes the inert gases (O 2 , N 2  and Argon) and produces a bottom liquid containing CO 2 , acid gases, and less than 5 ppm O 2 . This liquid then feeds in the second column, which then yields the pure CO 2  overhead liquid and the acid gases bottom liquid. Since these products are in liquid form, this process requires intensive cooling by external refrigeration equipment and additional nitrogen expansion by the oxygen plant. The inert gas extracted from the flue gas is expanded in 3 expanders in series with intermediate reheats to keep the exhaust temperatures of the expanders above the freezing point of CO 2 . 
     For the foregoing reasons, a need exists for a more cost effective and efficient method for removing carbon dioxide from the flue gas that is generated by oxy-combustion plants. In particular, a need exists for a method that recovers energy from the expansion of the off-gas stream in a more efficient and cost effective manner. 
     SUMMARY 
     The present invention is directed to a method that satisfies the need in general for a more cost effective and efficient method for removing carbon dioxide from the flue gas that is generated by oxy-combustion plants. 
     In one aspect of the present invention, an improved carbon dioxide separation process for oxy-combustion coal power plants is provided. This process requires warming at least a portion of a waste stream, that has been separated from a flue gas stream. This waste stream is then expanded, which results in a cool vapor exhaust stream. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
       For a further understanding of the nature and objects for the present invention, reference should be made to the following detailed description, taken in conjunction with the accompanying drawings, in which like elements are given the same or analogous reference numbers and wherein: 
         FIG. 1  is a stylized diagram of an illustrative embodiment of an oxy-combustion process for a coal power plant; 
         FIG. 2  is a stylized diagram of an illustrative embodiment of a typical partial condensation process with a hot gas expander; 
         FIG. 3  is a stylized diagram of an illustrative embodiment of the present invention having two separators to remove carbon dioxide from the flue gas, and two expanders to remove energy from the off-gas stream; 
         FIG. 4  is a stylized diagram of an illustrative embodiment of the present invention having a stripping column and a separator, and two expanders to remove energy from the off-gas stream; 
         FIG. 5  is a stylized diagram of an illustrative embodiment of the present invention having distillation column and two separators, and two expanders to remove energy from the off-gas stream; 
         FIG. 6  is a stylized diagram of an illustrative embodiment of the present invention having two distillation columns and two separators, and two expanders to remove energy from the off-gas stream; and 
         FIG. 7  is a stylized diagram of another illustrative embodiment of the present invention having striping column and two separators, and two expanders to remove energy from the off-gas stream. 
     
    
    
     DESCRIPTION OF THE PREFERRED EMBODIMENTS 
     Illustrative embodiments of the invention are described below. While the invention is susceptible to various modifications and alternative forms, specific embodiments thereof have been shown by way of example in the drawings and are herein described in detail. It should be understood, however, that the description herein of specific embodiments is not intended to limit the invention to the particular forms disclosed, but on the contrary, the intention is to cover all modifications, equivalents, and alternatives falling within the spirit and scope of the invention as defined by the appended claims. 
     It will, of course, be appreciated that in the development of any such actual embodiment, numerous implementation-specific decisions must be made to achieve the developer&#39;s specific goals, such as, compliance with system-related and business-related constraints, which will vary from one implementation to another. Moreover, it will be appreciated that such a development effort might be complex and time-consuming, but would, nevertheless, be a routine undertaking for those of ordinary skill in the art having the benefit of this disclosure. 
       FIG. 3  depicts an illustrative embodiment of process  300  according to the present invention. Process  300  includes a first separator  310 , a second separator  312 , a first pressure increasing device  320 , a second pressure increasing device  323 , a first expander  315 , a second expander  318 , a first heat transfer device  331 , a second heat transfer device  332 , a first pressure reducing device  326 , a second pressure reducing device  314 , and a collective heat transfer device, which is indicated generally as  329  in  FIG. 3 . 
     Flue gas from the oxycombustion power plant is available at essentially atmospheric pressure and relatively warm temperature. After cooling to about ambient temperature, the flue gas is then compressed, the compression heat is removed in the compressor&#39;s cooler, the compressed flue gas stream is then dried in dryer  330 . Examples of such drying methods may include, but are not limited to, desiccant dehumidification system, adsorption system by activated alumina or molecular sieves, permeation dryers or solvent scrubber/dryers. The flue gas also contains some other impurities, mainly the by-products of the coal combustion, such as traces of acid, NO x  (like nitrogen oxide NO and nitric oxide NO 2 ), SO x  (like sulfur dioxide SO 2 , sulfur trioxide SO 3 ) etc. In some circumstances, it is preferable to remove some of these impurities in a scrubber system prior to cryogenic treatment. For example, NO 2  can react with water and SO 2  in the scrubber to yield sulfuric acid or, in the absence of SO 2  or if SO 2  is depleted, can react with water to yield nitric acid. With sufficient residence time, NO can react with oxygen to form NO 2 , which, is then converted to the acids, as described. The acids in the water can be neutralized with a hydroxide solution or some other chemical means. The choice of front-end removal of those impurities depends upon the final use of CO 2  and the economics of wet treatment of flue gas. Indeed, the NO 2  and SO 2  being heavier than CO 2  would concentrate in the CO 2  product. The presence of SO 2 , NO 2 , and sometimes O 2  and NO, in the CO 2  can be objectionable for sequestration or EOR applications. In this situation, these impurities can be removed in the front-end treatment so that CO 2  will not contain significant level of those impurities. 
     Once the compressed flue gas stream is cooled and dried, and its impurities optionally removed, to form compressed dry flue gas stream  301 , it is further cooled  302  and sent to a first separator  310 . The compressed dry flue gas stream  301  may be at a pressure of about 30 bar, its temperature can be between about 5° C. and about 35° C. It is possible to perform the drying of the flue gas at a lower pressure followed by further compressing the dry flue gas to the required pressure for cryogenic treatment. The further cooled flue gas stream  302  will be at least partially condensed. Within the first separator  310 , this further cooled flue gas stream  302  is separated into a first vapor stream  303  and a first liquid stream  311 . This first liquid stream  311  may be comprised of at least 90% carbon dioxide. The first vapor stream  303  is further cooled and at least partially condensed  304 , and sent to a second separator  312 . The at least partially condensed stream  304  may have a temperature of about −52° C. Within the second separator  312 , this further cooled first vapor stream  304  is separated into a second vapor stream  305  and a second liquid stream  313 . This second liquid stream  313  may be comprised of at least 90% carbon dioxide. 
     The second liquid stream  313  is warmed and vaporized  307 . This warmed and vaporized stream  307  may have a pressure of about 9 bar and a temperature as low as of about −40° C. The colder temperature lowers the compression power of the carbon dioxide compressor. The temperature is preferably warmer than the dew point of the gas, so sending liquid droplets into the compressor inlet can be avoided. The −40° C. minimum temperature allows the use of lower cost carbon steel and not higher cost stainless steel for piping and compression equipment. The second liquid stream  313  may pass through a second pressure-reducing device  314 . After passing through the second pressure-reducing device  314 , the second liquid steam  313  may have a pressure of about 9 bar. The vaporized second liquid stream  307  is compressed in a first pressure-increasing device  320 , thereby, creating a higher-pressure stream  321 . A portion of the second liquid stream  313  may remain a liquid  334 . The first liquid stream  311  may pass through a first pressure-reducing device  326 . After passing through the first pressure reducing device  326  the first liquid stream may have a pressure of about 19 bar and may have a temperature of about −6° C. The at least a portion of the first liquid stream  311  is warmed and vaporized  308 , at which point it combines with stream  321  to produce a combined stream  322 . A portion of the first liquid stream  311  may remain a liquid  333 . Combined stream  322  is further compressed in a second pressure-increasing device  323 , thereby, creating a high-pressure stream  309 . 
     The second vapor stream  305  is warmed in exchanger  329  and further warmed in first heat transfer device  331  to a temperature higher than that of the flue gas  301 , thereby, resulting in a warm third vapor stream  324 . This warm third vapor stream  324  may have a temperature that is between about 35° C. and about 80° C. This warm third vapor stream  324  is then expanded in a first expander  315 , thereby, resulting in a cool fourth vapor stream  316 . This cool fourth vapor stream  316  may have a pressure of about 6.6 bar. This cool fourth vapor stream  316  is then warmed in exchanger  329  and further warmed in exchanger  332  to a temperature higher than that of the flue gas  301 , thereby, resulting in a warm fifth vapor stream  317 . This warm fifth vapor stream  317  may have a temperature that is between about 35° C. and about 80° C. This warm fifth vapor stream  317  is then expanded to about atmospheric pressure in a second expander  318 , thereby, resulting in a cool sixth vapor stream  319 . This cool sixth vapor stream  319  is then warmed and vented. 
     Power generated by first expander  315  or second expander  318  can be used to drive electric generators to produce electricity, or can be used to partially drive the boost compressor (not shown) for the feed gas  301 , or carbon dioxide product (first or second pressure increasing devices  320  or  323 ). 
     The external heat exchanger used to heat the off-gas (first and second heat transfer devices  331  and  332 ) may be a heat recovery exchanger, wherein the hot compressed feed gas or hot compressed carbon dioxide exchanges heat with the off-gas to provide the necessary heat. These heat exchangers can be an intercooler, or aftercooler of the flue gas compressor, or carbon dioxide product compressors (first or second pressure increasing devices  320  or  323 ). In most isothermal compressors, the gas exiting a compressor stage is usually about 90° C. to about 120° C., and it can be used as heating medium, therefore, heating to the level of about 50° C. can suit very well for the isothermal compressor, which is favorable for any power saving scheme. 
     Thanks to the refrigeration supplied by the first and second expanders  315  and  318 , the carbon dioxide fractions  311  and  313  can be produced at low temperature, ranging from about −40° C. to about 3° C. Furthermore, this additional refrigeration also allows extracting the CO 2  streams  307  and  308  at higher pressures to save more compression power. 
     Since the triple point of carbon dioxide is −56.6° C., it is preferable to limit the outlet temperature of the first and second expanders  315  and  318  to about −54° C. to avoid the risk of carbon dioxide freezing at the cold end of the exchanger. This constraint can be met by using the first and second expanders  315  and  318 , with inlet temperature about 35° C. to about 70° C. and to expand from about 30 bar to about atmospheric pressure as proposed in the present application. A single expander would yield an outlet temperature that was too cold, and would require a higher expander inlet temperature, which is more difficult to achieve, as in the case of the hot gas expander. Without heating to about 35° C. to about 70° C., it is also feasible to obtain similar performance of the 2 expanders by using 3 expanders in series with inlet temperatures of about 10° C. to about 20° C. However, not only is there an additional cost for the third expander, also the heat exchanger would cost higher due to an additional passage for the third expander flow. 
     In some situations, it is desirable to produce a CO 2  product essentially free of oxygen like in applications for Enhanced Oil Recovery (EOR).  FIG. 4  depicts an illustrative embodiment of process  400  for oxygen removal according to the present invention. Process  400  includes a first separator  414 , a second separator  453 , a stripping column  440 , a first pressure increasing device  420 , a second pressure increasing device  422 , a third pressure increasing device  432 , a fourth pressure increasing device  437 , a fifth pressure increasing device  418 , first expander  425 , a second expander  428 , a first heat transfer device  451 , a second heat transfer device  452 , a first pressure reducing device  417 , a second pressure reducing device  430 , and a collective heat transfer device, which is indicated generally as  441  in  FIG. 3 . 
     Once the compressed flue gas stream  401  is cooled and dried, a portion  404  is sent to a stripping column  440  reboiler wherein it serves as the reboiler inlet stream  404 . The stripping column  440  may operate at about 10 bar. The stripping column  440  may operate at between about 10 bar and about 25 bar. This flue gas stream  404  reboils the stripping column  440  by condensing at least a portion of the flue gas stream  404  in the reboiler. This reboiler inlet stream  404  then exits the stripping column&#39;s reboiler as the reboiler outlet stream  405 . Stream  405  is sent to a second separator  453 , where it is separated into the reboiler outlet vapor stream  455  and reboiler outlet liquid stream  456 . Reboiler outlet liquid stream  456  feeds the stripping column. Reboiler outlet vapor stream  455  is then further cooled, and will be at least partially condensed, thereby, resulting in separator inlet stream  457 . The remaining portion  403  of the flue gas is cooled, partially condensed to yield stream  406 . Within the first separator  414 , streams  406  and  457  are separated into a first vapor stream  415  and a first liquid stream  416 . This first liquid stream  416  is then sent to a first pressure-reducing device  417 , thereby, resulting in a stripping feed stream  413 . This stripping feed stream  413  is then sent to the stripping  440 . 
     The stripping overhead stream  407  is warmed  402 , and then sent to a fifth pressure-increasing device  418 , thereby, creating a recycle steam  419 . Of course, the warmed stripping overhead stream can feed to a stage of the flue gas compressor thus simplifying the machine arrangement at the expense of a slightly larger drying unit. The warmed and vaporized stripping column overhead stream  402  may have a temperature that is between about 35° C. and about 40° C. This recycle stream  419  is then combined with flue gas stream  401 . 
     A portion of the stripping column bottom stream  408  is sent to a first pressure increasing device  420 , which results in a first medium pressure liquid stream  421 . The stripping column bottom stream  408  contains less than 10 ppmv of oxygen. This first medium pressure liquid stream  421  is then warmed and vaporized, then sent to a second pressure increasing device  422 , thereby, resulting in a high pressure stream  423 . This high-pressure stream  423  is then sent to the end-user. 
     The first vapor stream  415  is warmed in exchanger  441  to about ambient temperature and further warmed in exchanger  451  to a temperature higher than that of the flue gas  401 , thereby, resulting in a first warm vapor stream  424 . This first warm vapor stream  424  may have a temperature that is between about 35° C. and about 80° C. This first warm vapor stream  424  is then expanded in a first expander  425 , thereby, resulting in a cool second vapor stream  426 . This cool second vapor stream  426  is then warmed in exchanger  441  to about ambient temperature and further warmed in exchanger  452  to a temperature higher than that of the flue gas  401 , thereby, resulting in a second warm vapor stream  427 . This second warm vapor stream  427  may have a temperature that is between about 35° C. and about 80° C. This second warm vapor stream  427  is then expanded in a second expander  428 , thereby, resulting in a cool third vapor stream  429 . This cool third vapor stream  429  is then warmed and vented. 
     In another embodiment, as illustrated in both  FIG. 4  and  FIG. 4   a , a portion of the stripping column bottom stream  408  is removed prior to the first pressure-increasing device  420 . This removed portion is sent to a second pressure reducing device  430 , and warmed and vaporized, thereby, creating a low-pressure stream  431 . This low-pressure stream  431  is then compressed in a third pressure increasing device  432 , thereby, creating a second medium pressure stream  433 . 
     In another embodiment, as illustrated in both  FIG. 4  and  FIG. 4   a , a portion of the stripping column bottom stream  408  is removed after the first pressure-increasing device  420 . This removed portion is sent to a third pressure reducing device  434 , and warmed and vaporized, thereby, creating an intermediate-pressure stream  454 . This intermediate-pressure stream  454  is then compressed in a fourth pressure increasing device  437 , thereby, creating a second medium-pressure stream  439 . This second-medium pressure stream  439  is then combined with the first medium-pressure stream  421 , prior to admission into the second pressure increasing device  422 . 
     Power generated by first expander  425  or second expander  428  can be used to drive electric generators to produce electricity, or can be used to partially drive the boost compressor for the feed gas  401 , or carbon dioxide product  432 ,  437 , or  422 . 
     The external heat exchanger used to heat the off-gas  451  and  452  may be a heat recovery exchanger wherein the hot compressed feed gas or hot compressed carbon dioxide exchanges heat with the off-gas to provide the necessary heat. These heat exchangers can be an intercooler or aftercooler of the flue gas  401  or carbon dioxide product compressors  431 ,  437 , or  422 . In most isothermal compressors, the gas exiting a compressor stage is usually about 90° C. to about 120° C., and it can be used as heating medium, therefore, heating to the level of about 50° C. can suit very well for the isothermal compressor, which is favorable for any power saving scheme. 
     Thanks to the refrigeration supplied by the 2 expanders  425  and  428 , the carbon dioxide fractions can be extracted at low temperature, ranging from about −40° C. to about 3° C. This additional refrigeration also allows extracting the CO 2  product streams at higher pressures to save more compression power. 
     Since the triple point of carbon dioxide is −56.6° C., it is preferable to limit the outlet temperature of the expanders  425  and  428  to about −54° C. to avoid the risk of carbon dioxide freezing at the cold end of the exchanger. This constraint can be met by using 2 expanders  425  and  428  with inlet temperature about 35° C. to about 70° C. and to expand from about 30 bar to about atmospheric pressure as proposed in the present application. A single expander would yield an outlet temperature that was too cold, and would require a higher expander inlet temperature which is more difficult to achieve as in the case of the hot gas expander. Without heating to about 35° C. to about 70° C., it is also feasible to obtain similar performance of the 2 expanders by using 3 expanders in series with inlet temperatures of about 10° C. to about 20° C. However, not only is there an additional cost for the third expander, also, the heat exchanger would cost higher due to an additional passage for the third expander flow. 
     In another embodiment, as illustrated in  FIG. 5 , the compressed dry flue gas  560  is sent to a distillation column  580  to remove the SO 2  and NO 2  impurities. A bottom stream  570  containing the captured SO 2  and NO 2  impurities is recovered and sent to the SO 2  and NO 2  treatment units. A vapor stream  565  exiting the top of the distillation column is essentially free of SO 2  and NO 2  and is further cooled and partially condensed. The vapor and liquid fractions of the partial condensation steps then follow the similar paths as in  FIG. 3 . This type of process arrangement can be used when the CO 2  product can contain some oxygen, but only traces of SO 2  or NO 2 . 
     The embodiment of  FIG. 6  is similar to  FIG. 5 , a distillation column  680  for SO 2  and NO 2  removal is provided near the warm end of the heat exchanger  641 . The top vapor  665 , essentially free of SO 2  and NO 2 , is cooled and partially condensed in the similar paths as in  FIG. 4 . This type of process arrangement can be used when the CO 2  product contains only traces of oxygen, SO 2 , and NO 2 . 
     In another embodiment, as illustrated in  FIG. 7 , a first portion of the compressed dry flue gas  701  is sent to a first phase separation device  703 , wherein it is separated into a first vapor stream  704  and a first liquid stream  705 . A second portion of the compressed dry flue gas  702  is cooled in the condenser of a stripping column  706 , then sent to a second phase separation device  710 , wherein it is separated into a second vapor stream  711  and a second liquid stream  712 . Second liquid stream  712  is sent to stripping column  706 , wherein it is separated into a third vapor stream  707  and a third liquid stream  708 . Third vapor stream  707  is then cooled and recirculated back to the incoming flue gas line. Third liquid stream  708 , is warmed and vaporized, then compressed and sent to an end user  709 . First liquid stream  705  is heated and sent to stripping column  706 . First vapor stream  704  is warmed in exchanger  713  to a temperature higher than that of the flue gas, thereby, resulting in a warm fourth vapor stream  714 . This warm fourth vapor stream  714  may have a temperature that is between about 35° C. and about 80° C. This warm fourth vapor stream  714  is then expanded in a first expander  715 , thereby, resulting in a cool fifth vapor stream  716 . This cool fifth vapor stream  716  may have a pressure of about 6.6 bar. This cool fifth vapor stream  716  is then warmed in exchanger  717  to a temperature higher than that of the flue gas, thereby, resulting in a warm sixth vapor stream  718 . This warm sixth vapor stream  718  may have a temperature that is between about 35° C. and about 80° C. This warm sixth vapor stream  718  is then expanded to about atmospheric pressure in a second expander  719 , thereby, resulting in a cool seventh vapor stream  720 . This cool seventh vapor stream  720  is then warmed and vented.

Technology Category: f