Abstract:
Low-cost and energy-efficient C0 2  and H 2 S capture is provided obtaining greater than 99.9% capture efficiency from pressurized gas. The acid species are captured in an ammonia solution, which is then regenerated by stripping the absorbed species. The solution can capture as much as 330 grams of C0 2  and H 2 S per 1000 gram of water and when regenerated it produces pure pressurized acid gas containing more than 99.7% C0 2  and H2S. The absorption of the acid species is accomplished in two absorbers in-series, each having multiple stages. More than 95% of the acid species are captured in the first absorber and the balance is captured in the second absorber to below 10 ppm concentration in the outlet gas. The two absorbers operate at temperatures ranging from 20-70 degrees Celsius. The two absorbers and the main stripper of the alkaline solution operate at similar pressures ranging from 5-200 bara.

Description:
FIELD OF THE INVENTION 
       [0001]    The invention relates to methods and systems for high efficiency capture of acid species, mainly CO 2  and H 2 S, from pressurized gases in alkaline solution containing ammonia and regeneration of the solution by stripping the absorbed acid species. 
       BACKGROUND OF THE INVENTION 
       [0002]    Capturing H 2 S and CO 2  from gases is important in many industrial processes where the H 2 S and CO 2  are contaminants and have to be removed before further processing. Capturing the H 2 S and CO 2  is also important for environmental reasons where H 2 S, before or after its combustion, contributes to the formation of acid rain and CO 2  is associated with global warming. 
         [0003]    There are many commercial chemical and physical processes for capturing H 2 S and CO 2  from pressurized gas. Chemical processes include absorbents such as amine based processes, the Benfield process using potassium carbonate and many more. Physical processes include the Selexol process, the methanol based Rectisol process and more. These processes are typically expensive and require significant input of heat and electricity. In addition, most of available commercial processes can only capture small amount of CO 2  and H 2 S per unit volume of absorbent, typically in the 30-60 grams/liter and thus requiring the pumping and circulation of large volumes of solutions and making the reactors, pumps, pipes, heat exchangers large and expensive. In many of these processes the capture efficiency of CO 2  and H 2 S is relatively low and requires polishing steps downstream. Furthermore, another concern is that the stripped acid species are at low pressure and require high cost and energy intensive compression. Also, higher pressure gas results in higher solubility of non-acidic species such as H 2 , CO and CH 4  in the absorber outlet solution. As a result, the stripped acid gas contains H 2 , CO and CH 4  in concentrations that require further cleaning treatment and also results in loses of valuable matter. 
         [0004]    There is a need in the art for a dramatically improved system and process for capturing H 2 S and CO 2  and reduce its cost. The present invention addresses this need. 
       SUMMARY OF THE INVENTION 
       [0005]    The present invention provides a process and system integrating very high capture efficiency typically greater than 99% and potentially as high as 99.99%, high acid gas loading in the range of 100-330 grams per 1000 grams of water (3-7 times higher loading than commercially available technologies), produce acid gas at pressure in the range of 5-200 bara and containing more than 99.7% CO 2  and H 2 S, less than 0.3% moisture and practically no non-acidic species such as H 2 , CO and CH 4 . The process consumes less than half the energy, combined heat and power than any commercial process. It utilizes a low cost ammonia reagent which is non-degradable and produces no harmful waste stream. The integrated system of the invention reduces the cost of CO 2  and H 2 S capture to less than half the cost of state of the art technologies. The high efficiency and low cost of the process is enabled by multi-stage absorption system with multiple stages each designed and operated under conditions that optimized the system performance. 
         [0006]    Embodiments of the invention capture acid gases, mainly CO 2  and H 2 S, from pressurized gas streams into an absorbing solution and thermally strips the CO 2  and H 2 S from the absorbing solution to produce pressurized acid gas stream. The absorbing solution is a concentrated ammonia solution containing NH 3 —CO 2 —H 2 O—H 2 S. In addition to NH 3  the absorbing solution may contain alkaline cations such as Na + , K +  and Li + . 
         [0007]    Embodiments of the invention include the following units. 
         [0008]    1. A multi-stage absorber where 5-15 molal ammoniated solution captures most of the CO 2  and H 2 S from a pressurized gas stream at net CO 2 +H 2 S loading in the range of 100-330 grams per 1000 grams of water. 
         [0009]    2. A multi-stage polishing absorber weak ammonia water solution containing 0-0.2 molal ammonia is used to capture the residual CO 2 +H 2 S in the gas. In addition, the wash solution captures ammonia entrained from the absorber. 
         [0010]    3. A main CO 2 +H 2 S stripper where the CO 2  and H 2 S are stripped from the solution at 5-200 bara pressure to generate pure acid gas stream containing more than 99.7% CO 2 +H 2 S less than 0.3% H 2 O and practically no H 2 , CO, CH 4  and NH 3 . 
         [0011]    4. A sour water stripper where NH 3 , CO 2  and H 2 S species captured in the polishing absorber are stripped from the water. 
         [0012]    Advantages of embodiments of the invention result in much lower capital costs, energy consumption and overall operating costs than any state of the art technology for CO 2  and H 2 S captures and it could reduce the cost of unit CO 2  and H 2 S captured by more than 50%. 
     
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
         [0013]      FIG. 1  shows a schematic of a process and system for the simultaneous high efficiency capture of CO 2  and H 2 S from pressurized gas according to an embodiment of the invention. 
           [0014]      FIG. 2  shows according to an embodiment of the invention data of the emission of CO 2  from the first stage absorber and the second stage absorber when the solution is at equilibrium with the gas. 
           [0015]      FIG. 3  shows according to an embodiment of the invention a schematic of a three-stage first multistage absorber for the capture of CO 2  and H 2 S. 
           [0016]      FIG. 4  shows according to an embodiment of the invention a schematic of a two-stage second multistage absorber for the capture of CO 2 , H 2 S and NH 3 . 
           [0017]      FIG. 5  shows according to an embodiment of the invention a schematic of typical CO 2  and H 2 S stripper with top stage washer for NH 3  capture. 
       
    
    
     DETAILED DESCRIPTION 
       [0018]    The present invention is a system and process for the high efficient capture of acid gases mainly CO 2  and H 2 S. The acid gas species are captured simultaneously in an alkaline solution containing ammonia or a combination of ammonia and cations such as Na+, K+ and Li+. 
         [0019]    A schematic of the system for high efficiency capture of CO 2  and H 2 S from pressurized gas stream is shown in  FIG. 1 . Stream  120  is a gas stream at a pressure of 5-200 bara, temperature in the range of 10-75 degrees Celsius and containing CO 2  or CO 2 +H 2 S. Stream  120  can be syngas from coal or petcoke gasification, syngas from fuel gas steam or auto thermal reformer, natural gas from gas wells, refinery process gas and more. The gas is typically water-saturated and its CO 2  concentration is 1-50% mole and H 25  concentration is 0-7%. Stream  120  flows through the first stage CO 2  and H 2 S absorber  300  where more than 95% of the CO 2  and H 2 S are captured by the absorbing solution. Stream  122  contains the residual of the acid species not captured in absorber  300  and in addition, stream  122  contains NH 3  derived from the vapor pressure of the absorbing solution inlet to the first stage absorber  300 . 
         [0020]    Stream  122  flows to second stage absorber  400  for further cleaning. The feed absorbing solution in the second stage absorber  400  is water from sour water stripper unit  106  containing low concentration of NH 3 , typically in the range of 0-0.2 molal. The absorbing solution captures the residual CO 2  and H 2 S to very low levels and in addition it captures entrained ammonia from the first stage absorber. The system is designed in such a way that the gas stream from the first stage absorber  122  contains ammonia to CO 2  plus H 25  mole ratio smaller than 0.4. The clean gas outlet from the second stage absorber  400 , stream  124 , contains less than 10 ppm ammonia, less than 10 ppm CO2 and less than 1 ppm H2S. 
         [0021]    The inlet absorption solution in the absorber, stream  130 , is an ammoniated solution containing 8-15 molal NH 3  and with a mole ratio of CO 2 /NH 3  in the range of 0.2-0.4. The concentration of CO 2  and H 25  depends on the stripper operation and it can vary depending on the specific application. Stream  130  is fed to the top of a multistage first stage absorber vessel and after absorbing the bulk of the CO 2  and H 2 S, e.g. more than 95%, it is discharged at the bottom as a CO 2 +H 2 S rich solution, stream  132 , with mole ratio of CO 2 /NH 3  in the range of 0.6-0.7. 
         [0022]      FIG. 2  shows an example of the CO 2  emission limit at equilibrium from the top of the first and second stage absorbers operating at 50 Bara pressure and at 35 degrees Celsius solution feed temperature. It also shows that the first stage absorber CO 2  emission can be as low as 100 ppm at CO 2 /NH 3  mole ratio below 0.45 and that the CO 2  emission from the send stage absorber is nil. 
         [0023]    Referring back to  FIG. 1 , the system is typically designed with the absorber  300  and the stripper  500  operating at similar pressures and in the range of 5-200 Bara. In  FIG. 1 , the stripper pressure is only 1-3 bar higher than that of the absorber and low pressure drop pump  100  is required to pump the rich solution, stream  132 , from the absorber through a flash chamber  101  and recuperating heat exchanger  102  to the stripper. In the flash chamber the dissolved H 2 , CO, CH 4  and other non-acidic gas species are stripped from the rich solution to form gas stream  136  that flows back to the absorber. Gas stream  136  also contains small amount of CO 2 , H 2 S and NH 3 . The solubility of H 2 , CO and CH 4  in the solution at the bottom of the absorber is low and the flash chamber reduces the concentration to single digit ppm levels. As a result, the stripper gas outlet, stream  144 , is practically free of H 2 , CO and CH 4 . The solution feed stream to the stripper  500 , stream  138 , is a heated solution with heat content that is recovered in the recuperating heat exchanger  102  from the lean solution stream  140 . 
         [0024]    Heat input to the stripper is typically in the range of 40-60 KJ per mole of acid gas stripped is delivered to the reboiler  104  by heating recycle stream  142 . The ammoniated solution is chemically stable and does not degrade under the operating conditions of the stripper. As a result, the heat source in the reboiler is not limited to using condensing steam, but it can also use other sources of heat such as hot syngas, hot flue gas, hot oil from solar collectors, hot brines etc. 
         [0025]    Stream  140  is a hot, typically in the range of 150-250 degrees Celsius, and CO 2 /H 2 S lean solution from the stripper. It is cooled in the recuperating heat exchanger  102  while heating the rich solution stream  134 . Further cooling of the lean solution is provided in heat exchanger  103 . Typical temperature of the feed to the absorber, stream  130 , when using cooling tower water for heat sink in heat exchanger  103 , is 20-40 degrees Celsius. 
         [0026]    The stripper  500  is designed in such a way that its temperate at the gas outlet is lower than 40 degrees Celsius and typically in the range of 20-40 degrees Celsius. As a result water and ammonia concentration in the outlet gas stream is low, corresponding to their vapor pressure over the inlet solution to stage  506  in  FIG. 5 . For example, when the temperature at the stripper gas outlet is 40 degrees Celsius and the stripper operates at 50 Bara the moisture content of the acid gas from the stripper is about 0.15% and ammonia concentration is about 1 ppm. The acid gas stream  144  from the stripper contains more than 99.7% CO 2  and H 2 S, less than 0.3% water vapor and practically no NH 3 , H 2 , CO and CH 4 . 
         [0027]    The water from second stage absorber  400  and from the top of the stripper  500  contains NH 3 , CO 2  and H 2 S captured from the pressurized gas stream  122  and from the product acid gas stream  144 . The water is sent to a sour water treatment system where heat is provided through reboiler  114  to generate water containing low concentration of ammonia in the range of 0-0.2 molal. The treated water from stripper  106  is re-used and is sent back to the second stage absorber, stream  146 , and to the top of the main stripper, stream  148 . The gas from the sour water stripper  106  containing CO 2 , H 2 S, NH 3  and water vapor, stream  150 , is sent to the bottom of the stripper  500 . Depending on the relative operating pressures of the sour stripper  106  and the main stripper  500  a compressor may be used to push gas stream  150  to the main stripper. 
         [0028]    First Stage Absorber Vessel and System 
         [0029]    The first stage absorber and system is a multistage vessel  300  with at least two absorption stages each designed to achieve optimal results. A schematic of a three-stage absorber designed for high efficiency capture of CO 2  and H 2 S is shown in  FIG. 3 . 
         [0030]    Feed gas, stream  122 , containing CO 2  and H 2 S is injected to the bottom of the absorber, stage  306 , and it flows upwards through the absorber stages  304  and  302  to exit as clean outlet gas at the top, stream  122 . The solution fed to the top of the absorber, stream  326 , is a mix of lean solution from the stripper, stream  130 , and semi-rich solution from the second stage, stream  324 . The resultant stream has CO 2 /NH 3  mole ratio of 0.3-0.4 a ratio which is designed to optimize the capture of CO 2  and H 2 S while minimizing ammonia emission from the absorber. The mixed gas stream is cooled in heat exchanger  103  to below 40 degrees Celsius before it is fed to the top of the absorber. 
         [0031]    For example, the equilibrium gas concentration above 12 molal NH 3  solution at 35 degrees Celsius and 50 bara and containing CO 2 /NH 3  mole ratio of 0.33 is about 4,000 ppm NH 3  and less than 100 ppm each for CO 2  and H 2 S. The absorber gas outlet can be designed to achieve ammonia equilibrium concentration of 4000 ppm, CO 2  concentration above equilibrium and less than 1,500 ppm and H 2 S concentration of 100 ppm. It is important to keep the acid to ammonia mole ratio in the gas at above 0.4 so that after capturing all the residual species from the gas the second stage absorber solution is highly alkaline with CO 2 /NH 3  smaller than 0.4 so that it can capture all the residual CO 2  and H 2 S. 
         [0032]    The top stage absorber in  FIG. 3  is shown as a tray tower, but it can also be a packed tower or other gas-liquid contacting device. Also, the stage can operate as a once through liquid as shown in  FIG. 3  or it can have a recycle. Due to the low CO 2 /NH 3  mole ratio of the solution, 0.3-0.4 in the top stage  302 , the absorption rate of CO 2  and H 2 S into the solution is high. At the liquid outlet from the top absorber stage the solution is 3-15 degrees Celsius warmer than the feed solution  326  due to the heat of reaction of the CO 2  and H 2 S absorption. 
         [0033]    The solution from the top stage is fed to the middle stage  304  where it is mixed with cooled recycle solution, stream  322 .  FIG. 3  shows the middle absorber stage  304  as a packed tower with recycle. Other gas-liquid contacting devices can be applied as well. The recycle stream in the middle stage, stream  322 , is designed to increase liquid flow in the stage, to increase the rate of mass transfer and to control the stage temperature. Heat exchanger  310  removes excess heat of reaction from the recycled solution and it prevents overheating of the solution maintaining the temperature at 40-60 degrees Celsius. The recycle of the solution is done by recycle pump  308 . Excess solution from the middle stage flows to the bottom stage. 
         [0034]    The bottom stage  306  absorber is designed to produce CO 2  and H 2 S rich solution and to maximize the CO 2  and H 2 S loading of the solution. Depending on the partial pressure of CO 2  and H 2 S in the gas feed, stream  120 , and on the design characteristics of the stage, i.e. height of the stage, gas velocity, type of packing and operating temperature the outlet solution from the absorber can have as high as 0.7 CO 2  to NH 3  mole ratio and net loading, the difference in CO 2  and H 2 S content between the solution inlet to the absorber, stream  130 , and the solution outlet from the absorber, stream  134 , as high as 330 grams per 1000 grams of water or 7.5 molal of CO 2 . 
         [0035]    The high ionic strength and the high CO 2  loading of the solution at the bottom stage of the absorber may result in the precipitation of crystals of ammonium bicarbonate. For example, solution containing 12 molal of ammonia and having acid-to-NH 3  mole ratio of 0.7 should be at temperature greater than 60 degrees Celsius to prevent solids precipitation. As a result, the cooling of the middle stage absorber in heat exchanger  310  is controlled in such a way that the temperature in the bottom stage is 3-5 degrees Celsius higher than the precipitation temperature of solids. 
         [0036]    Depending on the absorber pressure, H 2 , CO, CH 4  can be physically absorbed in the solution. To eliminate practically all physically dissolved species from the solution so that loss of valuable species is eliminated and the acid gas from the main stripper contains only CO 2  and H 2 S a flash chamber  101  is installed at the solution outlet from the absorber. The outlet solution from the absorber, stream  132 , is flashed into the flash vessel optionally after heating the solution by 3-10 degrees Celsius. The physically dissolved species in the solution are flashed out of the solution and is sent back, stream  136 , to the bottom of the absorber. 
         [0037]    Second Stage Absorber Vessel and System 
         [0038]    The second stage absorber and system is designed to produce gas containing low concentration of CO 2 , H 2 S and NH 3  all in the less than 10 ppm level. The second stage absorber, Vessel  400  has at least two absorption stages. A schematic of a two-stage second absorber is shown in  FIG. 4 . 
         [0039]    The inlet gas stream to the second stage absorber is stream  122 , which is the outlet stream from the first stage absorber. It contains residual CO 2  and H 25  and in addition it contains NH 3  that evaporated from the ammoniated solution in the absorber. The first stage absorber is controlled in such a way that the CO2 plus H 25  to NH 3  mole ratio in gas stream  122  is less than 0.4 and as a result, the solution in the second stage absorber is highly alkaline and capable of removing residual CO 2  and H 25  from the gas. 
         [0040]    In addition to low CO 2 , H 2 S and NH 3  emission the second stage absorber is designed to minimize the use of water which is achieved by minimizing the NH3 emission from the first stage absorber and by producing high ammonia concentration bleed stream in the range of 1-6 molal. In the example shown in  FIG. 4 , the absorber has 2 stages. The bottom stage  404  is a packed tower and it utilizes recycle pump  406  to recycle solution within the absorbing stage, stream  414 , and to discharge excess solution, stream  412 , from the system and sending it to the sour water stripper. The recycle solution, stream  414 , is cooled in heat exchanger  408  to 3-10 degrees Celsius above the cooling water temperature to produce cooled solution, stream  416 . Stream  416  is fed to the top of the stage  404  and is mixed with solution from the top stage  402 . The bottom stage operates at molality in the range of 1-6 and it captures most of the ammonia from the gas as well as the most of the residual CO 2  and H 2 S in the gas stream. 
         [0041]    The top stage of the second stage absorber  402  in  FIG. 4  is a counter flow tray tower where water containing ammonia concentration in the range of 0-0.2 molal from the sour water stripper, stream  410 , is cooled to 3-10 degrees Celsius above cooling water temperature. The cooled water is fed to the top of the absorber and flow downwards through the trays and it captures residual NH 3 , H 2 S and CO 2  from the gas to below 10 ppm mole concentrations. Different type of packing or trays may be used in stages  402  and  404 . 
         [0042]    Main CO 2  and H 75  Stripper 
         [0043]    The main CO 2  and H 25  stripper is designed to strip CO 2  and H 2 S from the rich solution produced in the absorber (CO 2 /NH 3 =0.6-0.7 mole ratio) and to convert it to lean solution (CO 2 /NH 3 =0.2-0.3 mole ratio). The stripper operates at pressure in the range of 5-200 Bara and typically at a pressure close to the pressure of the absorber. The CO 2  and H 2 S stripping is done with practically no loss of NH 3  from the system. 
         [0044]    A schematic of typical main CO 2  and H 2 S stripper  500  is shown in  FIG. 5 . Rich solution from the absorber, stream  138 , is split to two. The smaller stream  510  is relatively cold and CO 2 -rich solution and it flows to the top of the stage  502 . The temperature of the stream  510  solution is typically at 60-80 degrees Celsius, to avoid the precipitation of ammonium bicarbonate. The solution flows downwards in stage  502  through a series of trays cooling the rising acid gas and capturing and condensing the ammonia and the water vapor in the gas. The heat of reaction and condensation and the sensible heat of the rising acid gas increase the solution temperature on its way down and it recovers the heat that otherwise would be lost. 
         [0045]    The solution from stage  502  liquid outlet is mixed with stream  512 , the main rich stream from the absorber, which is heated in a recuperating heat exchanger before entering the stripper. 
         [0046]    The bottom stage of the stripper, stage  504 , is typically a packed tower where hot gas from the reboiler  104 , at typical temperature in the range of 150-200 degrees Celsius or higher and containing CO 2 , H 2 S, NH 3  and H 2 O, flows upwards counter-currently to the rich feed solution. Heat and mass transfer occurs in the packed section of the stripper where the less volatile species from the gas, H 2 O and NH 3  vapor are cooled and condensed in the solution, while the more volatile species in the solution, CO 2  and H 2 S, evaporate into the gas phase. As a result, the rising gas becomes richer in CO 2  and H 2 S and leaner in NH 3  and H 2 O. Further enrichment of the gas in CO 2  and H 2 S occurs in stage  502  of the stripper. 
         [0047]    Heat is provided to the stripper in the reboiler  104 . The heat source to the reboiler can be any hot stream such as steam, syngas, flue gas and even heated oil from solar collectors. Stream  142  is a feed solution to the reboiler and stream  516  is a two phase stream from the reboiler. The gas phase in stream  516  contains the gas species that evaporated in the reboiler as well as gas species from the sour water stripper stream  150 . 
         [0048]    Hot lean solution, stream  140 , is withdrawn from the bottom of the stripper and sent to the recuperating heat exchanger  102  to cool the solution, stream  514 , and to recover its heat. In a system where the stripper pressure is higher than the absorber pressure a pump is installed to pump the rich solution to the stripper. In a system where the stripper pressure is lower than the absorber pressure the pump is installed to pump the lean solution to the absorber. 
         [0049]    A wash stage  506  is installed at the top of the stripper and is designed to capture all the ammonia from the gas stream and to further reduce the moisture content of the gas stream. Stage  506  is a packed or tray tower where cooled water from the sour water stripper, stream  518 , is fed to the top and flows downwards counter currently to the rising acid gas. The high partial pressure of the CO 2  in the acid gas results in high concentration of dissolved CO 2  in the solution and enhances the capture of NH 3 . The outlet solution from stage  506 , stream  520  contains practically all the ammonia that enters the stage in the gas phase. 
         [0050]    The outlet gas stream  144  from the top of the main stripper is pure CO 2  and H 2 S stream except for 0.1-0.3% of water vapor that can be easily removed downstream. Stream  144  contains practically no H 2 , CO, CH 4  and other physically absorbed species from the absorber. 
         [0051]    Sour Water Stripper 
         [0052]    The sour water stripper collects water containing NH 3 , CO 2  and H 25  from the second stage absorber  400  in  FIG. 4  and from the top stage of the main stripper  506  in  FIG. 5 . The sour water stripper is a conventional thermal stripper preferably operating at high pressure in such a way that the stripped gas  150  flows to the main acid gas stripper  500  by the pressure difference between the two vessels. Otherwise, a compressor is used to send the gas to the main stripper.