Experimental Study on the Novel Direct Steam Stripping Process for

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Experimental Study on the Novel Direct Steam Stripping Process for Postcombustion CO2 Capture Mengxiang Fang,† Qunyang Xiang,† Tao Wang,*,† Yann Le Moullec,‡ Jiahui Lu,§ Wenmin Jiang,† Xuping Zhou,† Jinbai Zhang,§ and Guofei Chen§ †

State Key Laboratory of Clean Energy Utilization, Zhejiang University, Hangzhou, 310027, P. R. China Fluid Dynamics, Power Generation and Environment Department, EDF R&D, 6 quai Watier, Chatou Cedex 78401, France § EDF China R&D Center, EDF Asia Pacific Direction, China Division, Beijing 100005, P. R. China ‡

ABSTRACT: High energy consumption is a crucial issue for the regeneration of solvent for postcombustion CO2 capture by chemical absorption. Primary modeling results show the potential to reduce the energy consumption through a novel solvent regeneration process by direct steam stripping. This work is an extension of our previous exploration and aims to validate the direct steam stripping process by experimental studies on a lab-scale stripping platform. We investigated the direct steam stripping and the conventional stripping mode in terms of energy consumption and steam condensation. The results showed, for the direct steam stripping mode, the optimum energy consumption at 1 atm was 2.98 MJ/kg CO2, 23.2% lower than that of the conventional stripping mode. Higher solvent feeding temperature and carrier steam superheating temperature were beneficial to reduce steam condensation in the column. We found 96−99.5 °C feeding solvent temperature was suitable considering both the energy consumption and steam condensation. Besides, we proposed an improved direct steam stripping process using carrier gases such as hydrocarbon, which could further reduce the latent heat required.

1. INTRODUCTION Extensive carbon dioxide (CO2) emissions are seen as a major contributor to global warming.1 Fossil fuel power plants are considered to be the largest source of CO2 emissions.2 CO2 capture and sequestration is a viable option to significantly reduce CO2 emissions from fossil fuel power plants. Among all the CO2 capture technologies, postcombustion capture (PCC) by chemical absorption using aqueous monoethanolamine (MEA) is relatively mature and can be used on existing power plants.3,4 However, the biggest challenge for postcombustion CO2 capture by chemical absorption is the high energy consumption during solvent regeneration which results in a coal-fired power plant output loss of approximately 20−25% when coupled with CO2 compression.5−9 Several researchers have done parametric optimizations on bench-scale or pilot-scale experimental facilities to reduce the energy consumption for solvent regeneration.10−13 Meanwhile, process modifications have also been carried out to reduce the energy consumption and operation costs.14−21 The mostly discussed process modifications are solvent intercooling, solvent split-flow, flashing of hot rich solvent, vacuum stripper, multipressure stripper, and advanced heat integration of the capture process. However, most of the modifications reported in the literature and patents are conceptual designs studied only at the modeling phase. The experimental validation of the novel processes is very limited up to our knowledge.22,23 During the CO2 regeneration process, a great amount of carrier steam is needed to drive out CO2 from the liquid. In the current technologies, the carrier steam is generated through solvent boiling in a reboiler.24,25 The excessive stripping steam exiting the column is cooled down, separated from the CO2 and then sent back to the stripping or absorbing column to © 2014 American Chemical Society

maintain the water balance. In this way, the latent heat of the steam is mostly wasted. Recently, the idea of solvent regeneration through direct steam stripping in order to save the latent heat in the regeneration process has been mentioned in literature.26 The idea is to extract superheated steam from low-pressure steam turbine and inject it into the regeneration column as carrier steam. This steam is then separated from CO2 via a condensation after the regeneration column and lower pressure steam is generated by evaporation, this low pressure steam is then sent back to the low pressure turbine. The steam condensation and evaporation process is made through a heat exchanger by utilizing the heat released in condensation to evaporate the condensed water at a lower pressure. In this way, a large portion of the latent heat can be efficiently recovered. The preliminary simulation results show that the direct steam stripping process can reduce the amount of steam needed very significantly.26 This work is an extension of our previous exploration and aims to validate the direct steam stripping system by experimental methods. We have built a lab-scale stripping system and carried out series of experiments on the conventional stripping mode and the direct steam stripping mode. Through these experiments, the regeneration energy and the steam condensation at different operation conditions have been investigated in order to seeking the optimized operation conditions for the direct steam stripping process. These first experimental results show the potential of this modification and will help to improve the efficiency of the direct steam stripping process. Received: Revised: Accepted: Published: 18054

September 5, 2014 October 23, 2014 October 28, 2014 October 28, 2014 dx.doi.org/10.1021/ie503517y | Ind. Eng. Chem. Res. 2014, 53, 18054−18062

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Figure 1. Schematic view of the experimental platform.

2. EXPERIMENTAL SECTION Materials. MEA at technical grade (>99.7%) was purchased from the Oriental Petrochemical Corporation and diluted with deionized water to the desired concentration. The exact concentration of MEA solutions was determined by titration using standardized sulfuric acid (Sinopharm Chemical Reagent Co., Ltd., AR). The CO2 gas (>99.9%) used was produced by Hangzhou Jingong Gas Corporation. Experimental Platform. The schematic view of the experimental platform is shown in Figure 1. The main part of the platform is a stripper column randomly packed with 3 mm × 3 mm Dixon rings (internal diameter: 0.08 m, total packing height: 1.6 m). Other equipment include a steam generator, a steam flow meter, heating wires, electrical heating rods, richsolvent tank, a gear pump, a solvent preheater, a condenser, a water tank, and a wet gas flow meter. The direct carrier steam is supplied by a steam generator and superheated by 20−50 K above its boiling point by using heating wires before being injected into the stripper column. The CO2-rich solvent is stored in a rich-solvent tank and preheated to desired temperatures before being injected into the stripper column. The gas mixture of carrier steam and CO2 exiting from the column top is then condensed in a condenser. The condensed water is collected in the water tank, and the pure CO2 flow rate is measured by a wet-gas flow meter. Solvent samples have been taken from four extraction points along the column for analyzing the MEA concentration and CO2 loading. An automatic acid−base titrator has been used to measure the MEA concentration and the CO2 loading was measured by the same apparatus as the one reported by Zhu et al.27

Heat loss of the stripping column was measured by experiments at different room temperatures. Electrical heating rods at the bottom of the column were used to compensate the heat loss of the system. Besides, the heating rods could also provide regeneration heat in a conventional stripping mode. Experimental Conditions. The operation conditions of the experimental system are listed in Table 1. The rich solvent Table 1. Operation Conditions of the Experimental Platform parameter

value

MEA concentration (M) CO2 loading of rich solvent (mol/mol) feeding solvent flow rate (mL/min) feeding solvent temperature (°C) operation pressure (atm) heat loss of the column (W)

5 0.499−0.516 166.7 91−99.5 1 85−104

has been prepared by bubbling CO2 gas into the 5 M MEA solution until the loading reaches 0.5 mol CO2/mol MEA. An oil bath is used to preheat the rich solvent before it is injected into the column. The heat loss of the column was roomtemperature dependent. To compensate the heat loss at the room temperature range 283−298 K, the power of the heating rods was set to 85−104 W, a setting determined by experimental tests.

3. METHODOLOGY The CO2 regeneration energy required in a conventional stripping process can be calculated by eq 1. 18055

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Figure 2. Mass flow rate of CO2 regenerated (solid markers) and total energy consumption for CO2 regeneration (hollow markers) as a function of total heat duty in the conventional stripping mode. (The total heat duty includes the reboiler duty and the preheater duty; feeding solvent temperature is at 96 °C.)

Q t = (Q reboiler + Q preheater − Q loss)/qmCO

Q preheater = Q sen,0 + Q reac,0

(1)

2

= c pqmsolvent (Tfeed − T0) + ΔHabs,CO2qmCO ,0

where Qt, in J/kg, is the total heat requirement for the conventional stripping process; Qreboiler, in W, is the reboiler duty; Qpreheater, in W, is the heat supplied by the preheater; Qloss, in W, is the heat loss of the column; and qmCO2, in kg/s, is the regenerated CO2 mass flow rate. The CO2 regeneration energy required in a direct steam stripping process can be calculated by eq 2.

2

(6)

where Qsen,0, in W, is the sensible heat supplied by the solvent preheater; Qreac,0, in W, is heat of reaction supplied by the solvent preheater; cp, in J/(kg K), is the specific heat of the rich solvent obtained from literature;28,29 qmsolvent, in kg/s, is the mass flow rate of the rich solvent; Tfeed, in °C, is the feeding temperature of the rich solvent into the stripper; T0, in °C, is an assumed temperature of the rich solvent after an economizer with a temperature pinch of 15 K; we choose T0 = 85 °C for the 1 atm operation pressure; ΔHabs,CO2, in J/kg, is the CO2 absorption heat with amine obtained from available data in literature;30,31 and qmCO2,0, in kg/s, is the CO2 mass regeneration rate of the preheated solvent injected into the column without any other heat sources (reboiler or direct carrier steam) which is measured by experiments in this work. The lean CO2 loading of the regenerated solvent can be calculated by qnCO 2 αCO2,lean = αCO2,rich − camineqv solvent (7)

Q d = (Q steam + Q preheater + Q reboiler − Q loss)/qmCO

2

(2)

of all the direct steam stripping experimental cases in this work, the heat supplied by the reboiler is only used to compensate the heat loss of the system; eq 2 can therefore be simplified as Q d = (Q steam + Q preheater)/qmCO

2

(3)

where Qd, in J/kg, is the total energy requirement for the direct steam stripping process; Qsteam, in W, is the extracted energy of the carrier steam during the regeneration process, calculated by eq 4: Q steam = Hsteam,injqmsteam,inj − Hsteam,recqmsteam,rec

(4)

where αCO2,lean, in mol/mol, is the CO2 loading of lean solvent; αCO2,rich, in mol/mol, is the CO2 loading of the rich solvent; qnCO2, in mol/s, is the CO2 molar regeneration rate; camine, in mol/m3, is the amine concentration of the solvent; and qvsolvent, in m3/s, is the volume flow rate of the rich solvent.

where Hsteam,inj, in J/kg, is the enthalpy of the superheated steam before injected into the column; Hsteam,rec, in J/kg, is the enthalpy of the steam recovered after the regeneration process; qmsteam,inj, in kg/s, is the mass flow rate of the steam injected and qmsteam,rec, in kg/s, is the mass flow rate of the steam recovered which can be calculated by eq 5. qmsteam,rec = χ (qmsteam,inj − qmsteam,con )

4. RESULTS AND DISCUSSION Comparison Results between Conventional Stripping Mode and Direct Steam Stripping Mode. First, comparative experiments between the direct steam stripping mode and the conventional stripping mode have been carried out. The comparison studies have been investigated at the same rich-solvent flow rate and conditions (5 M MEA; CO2 loading, 0.5 mol/mol; solvent preheating temperature, 96 °C) and

(5)

where χ is the proportion of the carrier steam exiting the stripper column which can be recovered through a condensation and re-evaporation process; qmsteam,con is the mass flow rate of carrier steam condensed in the column. The heat supplied by the preheater can be calculated by 18056

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stripper conditions (1 atm; reboiler temperature, 100 °C). Figure 2 shows the mass flow rate of CO2 regenerated and total energy consumption for CO2 regeneration as a function of total heat duty in the conventional stripping mode. The total heat duty includes the preheater duty and the reboiler duty. The preheater duty at different feeding solvent temperatures can be calculated using eq 6. The sensible heat supplied by the preheater is calculated with the specific heat and mass flow rate of the rich solvent and the temperature difference between the feeding solvent temperature 96 °C and T0 (85 °C). The reaction heat supplied by the preheater was studied by experiments at different feeding solvent temperatures without other heat sources. The mass flow rate of CO2 regenerated with only solvent preheating is shown in Figure 3. When the feeding

total preheater duty, including the sensible heat and the reaction heat, is 229 W for the 96 °C feeding solvent temperature cases. In Figure 2, the preheater duty is fixed and the total heat duty changes with the reboiler duty. When the total heat duty is low, the CO2 regeneration rate increases rapidly with the increase of total heat duty, which results in a decrease of total energy consumption. When the reboiler duty exceeds 600 W, the increase of CO2 regeneration rate slows down with the increase of the reboiler duty. Hence, we can obtain the optimum energy consumption for CO2 regeneration of the conventional stripping mode. The value is 3.88 MJ/kg CO2 regenerated for a 406 W total heat duty. This energy consumption result is comparable to the results of literature for MEA solvent.32 Experimental results from Li et al.33 are 4.26 MJ/kg CO2 for the regeneration of a rich solution (30 wt % MEA, 0.5 mol/mol CO2 loading). Knudsen et al.22 show 3.7 MJ/kg CO2 on the industrial pilot plant of Esbjerg (30 wt % MEA, 10 K pinch economizer, 12.5% CO2 in flue gases). The modeling result from Alie et al.34 and Singh et al.35 is, respectively, 4 MJ/kg CO2 and 3.8 MJ/kg CO2. In the direct steam stripping process, the solvent preheater and the carrier steam provide the energy consumed. The preheater duty in the direct steam stripping mode is equal to that in a conventional stripping mode. The regeneration heat from the carrier steam can be calculated by eqs 4 and 5. A fraction of the carrier steam is condensed in the column through heat exchange with the solvent during regeneration, while the rest of the carrier steam exits the column with regenerated CO2. In the direct steam stripping process, the carrier steam out of the column can be recovered at lower pressure and sent back to the steam turbine. The proportion of the recovered carrier steam χ depends on the temperature pinch of the economizer and the operating parameter of the hot and cold streams. It is possible to obtain a large χ value with a small temperature pinch of the economizer. In this work, χ = 0.8 has been considered for a relatively large temperature pinch (10−12 K).

Figure 3. Mass flow rate of CO2 regenerated as a function of feeding solvent temperature with no other heat sources (0 W reboiler duty and no direct injected steam).

solvent temperature is 96 °C, the mass flow rate of CO2 regenerated is 5.13 × 10−5 kg/s, and thus the corresponding reaction heat supplied by the preheater can be calculated. The

Figure 4. Mass flow rate of CO2 regenerated (solid markers) and total energy consumption for CO2 regeneration (hollow markers) as a function of carrier steam flow rate in the direct steam stripping mode. (Feeding solvent temperature is at 96 °C; carrier steam superheated degree is 30 K and χ is 0.8.) 18057

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Figure 5. Mass flow rate of CO2 regenerated (solid markers) and condensed carrier steam percentage (hollow markers) as a function of carrier steam superheated temperature. (Feeding solvent temperature is at 96 °C, and carrier steam flow rate is 0.617 kg/h.)

Figure 6. Mass flow rate of CO2 regenerated (solid markers) and condensed carrier steam percentage (hollow markers) as a function of feeding solvent temperature. (Carrier steam superheated degree is 30 K, and carrier steam flow rate is 0.617 kg/h.)

Figure 4 shows the mass flow rate of CO2 regenerated and the total energy consumption for CO2 regeneration as a function of carrier steam flow rate in the direct steam stripping mode. The effect of carrier steam flow rate on CO 2 regeneration rate is pronounced when the carrier steam flow rate is below 0.4 kg/s and becomes not as effective above 0.8 kg/s. The minimal energy consumption for CO2 regeneration of the direct steam stripping mode can be evaluated. This value is 2.98 MJ/kg CO2 at a 0.617 kg/h carrier steam flow rate, 23.2% lower than that of the conventional process. With adequate optimization of operation conditions, the electric penalty with direct steam stripping process will be reduced by 20−30% which allows the efficiency loss on a power plant decrease to approximately 9% from the standard value of 11.95%.26 The experimental results show that the direct steam stripping process has significant advantages of regeneration energy consumption over the conventional stripping process. Furthermore, the CO2 loading of the lean solvent after

regeneration in the energy optimized case of the direct steam stripping mode is 0.291 mol/mol, lower than that of energy optimized case of the conventional stripping mode, which is 0.329 mol/mol. Lower lean CO2 loading leads to a larger cyclic capacity between rich and lean solvent indicating that less solvent is required for the CO2 capture process. Besides, lower lean CO2 loading can benefit the CO2 absorption as well.36 Effect of Carrier Steam Superheating and Solvent Preheating. In the direct steam stripping process, water balance in the stripper column is critical. The solvent is diluted by steam condensation inside the column. Hence, we chose superheated carrier steam and relatively high feeding solvent temperature to avoid most of the steam condensation. In this work, the carrier steam superheated degree ranging from 10 to 50 K and the feeding solvent temperature ranging from 91 to 99.5 °C have been tested. The effect of the carrier steam superheated degree on CO2 regeneration rate and condensed steam percentage is shown in Figure 5. The CO2 regeneration rate changes very little with the 18058

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Figure 7. Regeneration energy consumption (solid markers) and condensed carrier steam percentage (hollow markers) as a function of carrier steam flow rate at three different feeding solvent temperature: 91 (black), 96 (red), 99.5 (blue) °C. (Carrier steam superheated degree is 30 K.)

Figure 8. Titration results of MEA concentration (square markers) and CO2 loading (cross markers) at the four sample points. The dashed lines are the lean CO2 loading calculated by the CO2 regeneration rate using eq 7. (Carrier steam flow rate is 0.617 kg/h; feeding solvent temperature is at 96 °C, and carrier steam superheated degrees are 30 K (a) and 50 K (b), respectively.)

We also investigated the optimum carrier steam flow rate at three different feeding solvent temperatures. Figure 7 shows regeneration energy consumption and condensed steam percentage as a function of carrier steam flow rate at three different feeding solvent temperatures: 91, 96, and 99.5 °C. To avoid steam condensation, a higher feeding solvent temperature (99.5 °C) is preferred. However, the optimum regeneration energy at 99.5 °C is the highest among the three feeding solvent temperatures. For the selection of a suitable solvent preheating temperature for the direct steam stripping system, both the regeneration energy and the steam condensation should be considered. A 96 to 99.5 °C feeding solvent temperature can be the operation window for the direct steam stripping process coupled with the carrier steam superheating. Liquid Sample Analyses. Liquid samplings along the column allow a better understanding of the regeneration process in the column.15,37 In our experimental system, there are four sample points along the stripper column, the sample points 1, 2, 3, and 4 are respectively at 340, 740, 1140, and 1540 mm below the top of the packing layer.

increase of the carrier steam superheated degree from 10 to 50 K while the steam condensation is decreased. Since the extra energy for superheating the carrier steam is relatively small, a high carrier steam superheated temperature is preferred for the direct steam stripping process in order to reduce steam condensation. The effects of feeding solvent temperature on the CO2 regeneration rate and condensed steam percentage are presented in Figure 6. A higher feeding solvent temperature results in a small increase of CO2 regeneration rate and a significant decrease of steam condensation. At a feeding solvent temperature of 99.5 °C, the condensed steam percentage can be reduced to only 3.5%, whereas the value is approximately 40% when the feeding solvent temperature is around 90 °C. At a low feeding solvent temperature for the direct steam stripping process, a large part of the regeneration energy is supplied by carrier steam condensation which can compensate the decrease of heat supplied by the preheater. Hence, the CO2 regeneration rate is not significantly decreased at lower feeding solvent temperature. 18059

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Figure 9. Titration results of MEA concentration (a) and CO2 loading (b) at the four sample points as a function of carrier steam flow rate. The dashed line in panel b is the lean-CO2 loading calculated by the CO2 regeneration rate using eq 7. (Feeding solvent temperature is at 96 °C, and carrier steam superheated degree is 30 K.)

Figure 10. Improved direct steam stripping process.

Figure 9 displays the titration results of MEA concentration and CO2 loading at the four sample points as a function of carrier steam flow rate. The titrated MEA concentration presents a slightly decreasing trend with the increase of carrier steam flow rate as the condensed steam weight is increased. These results prove that the MEA concentration along the column is influenced by steam condensation. The titrated CO2 loading decreases with the carrier steam flow rate because of the deeper regeneration degree at higher carrier steam flow rate. The lean CO2 loading calculated by eq 7 is also shown in

Figure 8 shows the titration results of MEA concentration and CO2 loading for two experimental cases. Repeated samples were taken and analyzed at each sample point. The MEA concentrations of all samples are in the range of 4.65−4.96 M, a little lower than the MEA concentrations of the rich solutions (5 M) due to steam condensation. The titrated CO2 loading decreases from the sample point 1 to sample point 4 indicating that the regeneration degree goes deeper when the solvent drops from the top of the column to the bottom. 18060

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process at 1 atm. Besides, an improved direct steam stripping process has been proposed in order to solve the steam pollution issue and further reduce the latent heat required.

Figures 8 and 9. The values are always higher than the CO2 loading of sample from sample point 4 and even higher than that from sample point 3 especially when the carrier steam flow rate is large. This probably occurs because the solvent drops have longer residence time when accumulated at the sample points than those that drop directly to the bottom of the column. Hence, we think the regeneration process in our stripper column is not an equilibrium system. A longer stripper column or packing material with larger specific surface area will be effective to further regenerate solvent and reduce the energy consumption. Improved Direct Steam Stripping Process. One of the main issues for the direct steam stripping process is the possible pollution of the recovered steam. Owing to the evaporation of the solvent, the gas mixture may contain some solvent in the gas phase after the stripper column, and thus the recovered steam cannot be sent back to the steam turbine without prior polishing. Hence, we propose an improved direct steam stripping process as shown in Figure 10. The carrier gas is partially boiled (from 0.8 to 1 vapor fraction) through a heat exchanger by steam before being compressed prior to being injecting into the stripper. The carrier gas and CO2 mixture after regeneration is condensed for CO2 separation and re-evaporated at lower pressure. The steam condensation and evaporation process is made through a heat exchanger by utilizing the heat released in condensation to partially evaporate (80%) the condensed carrier gas. The reevaporated carrier gas can then be reused in a new cycle. Hence, the steam from the steam turbine will not be polluted and the latent heat will be reduced significantly. Moreover, the carrier gas can be other gases with a lower latent heat and highly hydrophobic, such as hydrocarbons, and thus the latent heat will be further reduced. However, the improved direct steam stripping process requires additional equipment such as an economizer, a steam−steam heat exchanger and a gas compressor which will increase the capital cost. Further study is required to understand the compromise between the reduction of the energy consumption and the increase of the capital cost with the improved direct steam stripping process.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work acknowledges much support from EDF researchers and is financially supported by EDF, National Natural Science Foundation (No. 51306161, 51276161), Natural Science Foundation of Zhejiang Province, China (No. LY13E060004), the Specialized Research Fund for the Doctoral Program of Higher Education of China (No. 20130101120143), and the Program of Introducing Talents of Discipline to University (B08026)

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REFERENCES

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5. CONCLUSIONS Absorbent development and capture system optimization are equally important aspects of research on postcombustion CO2 capture by chemical absorption. In capture system optimization, an effective way to develop our knowledge is screening flowsheet modifications by using modeling methods, followed by further experimental study of the promising potential processes. In this work, we further investigated our previous idea of the direct steam stripping process using experimental methods. A lab-scale stripper allowing injection of live steam at the bottom has been built and a series of experiments to validate the feasibility of the direct steam stripping process have been carried out. At 1 atm regeneration pressure, the optimum energy consumption of the direct steam stripping process is 2.98 MJ/kg CO2, 23.2% lower than that of a conventional stripping process. The steam condensation issue has also been investigated and can be solved by a relatively high feeding solvent temperature and large carrier steam superheating degree. We have found 96−99.5 °C feeding solvent temperature coupled with 30−50 K carrier steam superheating degree is a suitable operation window for the direct steam stripping 18061

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dx.doi.org/10.1021/ie503517y | Ind. Eng. Chem. Res. 2014, 53, 18054−18062