Technoeconomic and Dynamical Analysis of a CO2 Capture Pilot

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Techno-Economic and Dynamical Analysis of a CO2 Capture Pilot-Scale Plant Using Ionic Liquids Darinel Valencia-Marquez, Antonio FLores-Tlacuahuac, and Luis Ricardez-Sandoval Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b02544 • Publication Date (Web): 23 Oct 2015 Downloaded from http://pubs.acs.org on October 30, 2015

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Industrial & Engineering Chemistry Research

Techno-Economic and Dynamical Analysis of a CO2 Capture Pilot-Scale Plant Using Ionic Liquids Darinel Valencia-Marquez1 , Antonio Flores-Tlacuahuac1 ∗, Luis Ricardez-Sandoval2 1 Departamento

de Ingenier´ıa y Ciencias Qu´ımicas, Universidad Iberoamericana,

Prolongación Paseo de la Reforma 880, México D.F., 01219, México. 2 Department

of Chemical Engineering, University of Waterloo,

Waterloo, Ontario Canada N2L 3G1.

October 23, 2015

Abstract Carbon capture has been recognized as an attractive alternative to reduce CO2 emissions. The most feasible technology that can be developed at a commercial-scale in a short term period is CO2 capture by absorption since it is an end-pipe technology that can be installed in existing coal-based power plants and will not require retrofit of the power plant. The most studied CO2 capture process is absorption using monoethanolamine (MEA) and represents the benchmark solvent due to the favorable properties they have shown e.g. fast kinetics, high absorption capacity, good solubility in water and low price. On the other hand, this solvent is susceptible to thermal and chemical degradation and it is also corrosive. Nevertheless, the main drawback of this solvent is the energy consumption needed for solvent recovery (almost >90% of the plant’s operating cost). Ionic liquids (IL) are new alternative solvents for CO2 capture. Experimental results have shown that IL feature chemical and thermal stability, and good CO2 absorption capacity. In this work, a theoretical IL is used as physical solvent for developing a new flow-sheet of a CO2 capture plant. A techno-economic analysis was carried out to evaluate the feasibility of the proposed design. The results show that the IL-based plant features lower energy demand compared to a traditional MEA-based plant. Moreover, the dynamic analysis performed in this study provides insight on the degree of nonlinearity and the dynamics of the process, which are essential tools to design suitable control schemes. The results show that the plant can accommodate perturbations in the flue gas flow rate up to ± 10% while meeting CO2 recovery and purity targets. ∗

To whom correspondence should be addressed. E-mail: [email protected]

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Keywords: ionic liquids; carbon capture; dynamic simulation; sustainability

1

Introduction

Carbon capture and storage (CCS) is a promising technology to achieve the reduction of emissions of CO2 from large stationary sources such as fossil-fired power plants. Developing an economically feasible carbon dioxide capture processes (CCP) is becoming an important route to reduce greenhouse gas emissions and have therefore received attention over the last decades. While several technologies for CCP have being proposed, oxy-fuel, pre-combustion, post-combustion and chemical looping are thought to be the more attractive technologies for commercial deployment. In the oxy-fuel technology, the fuel is burned in de-nitrified air ,i.e. oxygen, to produce CO2 and water [1, 2, 3, 4]. In pre-combustion, the coal is chemically broken to produce syngas, the CO2 is removed from syngas whereas hydrogen is used for power production [5, 6, 7, 8]. In the case of post-combustion, the fuel is burned and CO2 from the combustion process is recovered, commonly by scrubbing using solvents [8, 9, 10, 11, 12]. In Chemical looping technology, calcium oxide (CaO) and CO2 are reversibly reacted to form calcium carbonate (CaCO3 ). This process is analogous to the conventional aminebased absorption process since the CO2 capture occurs in one vessel, then the sorbent material (CaCO3 ) is sent to a second vessel, where the sorbent material is regenerated producing an almost pure stream of CO2 [13, 14, 15, 16]. CO2 capture using aqueous amine absorption is currently considered the most feasible option in the post-combustion process. Using simple absorption and stripping configurations, mono-ethanolamine (MEA) has been commercially demonstrated to effectively scrub CO2 from the post-combustion flue gas [10, 17, 18, 19]. Post-combustion capture is the most attractive technology for commercial-scale implementation since it can be potentially installed next to existing fossil-fired power plants. Moreover, the operating conditions of the CO2 capture plant can be adjusted to satisfy energy demand during peak hours or reduce the cost of energy production. In addition, maintenance, replacement ACS Paragon 2Plus Environment

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and upgrade of components of CO2 capture plants can be done with minimal impact on power plant [13, 20, 11]. The main drawback of this CO2 capture alternative is its large consumption of energy (around 4.2 MJ/ton CO2 [21]). Also, undesired effects appear under certain operation conditions, i.e. corrosion [19, 22] and solvent degradation [23]. These aspects diminish the performance of post-combustion CO2 capture plants [12]. Several works have made use of optimization tools to show improvements in energy consumption as well as in the operating and capital costs for stand-alone equipments [24], full process flow sheet [25, 26], optimal process design under uncertainty [27] and integration of power plant and carbon capture process (CCP) [28]. Also, researchers have studied the dynamic behaviour of this process by analyzing the dynamic performance of stand-alone units (absorber [29], striping[30]) and the interactions between absorber and the regeneration sections [9, 31, 32]. Other studies have proposed basic and advanced control schemes for CCP [10, 33, 17, 34, 35, 36, 37]. On the other hand, several studies have also analyzed the economics and performance of the post-combustion process using different solvents e.g, amines [38, 39, 40], glycol-based solvents [41], mix MEA-triethylene glycol [42], piperazine [43, 44], ionic liquids (ILs) [45, 46, 47, 48, 49, 50], and mix amine-based solvent with IL [51, 52, 53, 54]. ILs are alternative solvents for CO2 capture featuring low volatility [55, 56], therefore solvent loses are negligible. Commonly, ILs are formed by proper combination of cations and anions. Also, IL are chemically and thermally stable for a wide range of operating conditions thus avoiding solvent degradation, which is an issue in amine-based solvents. Research studies on ILs for CO2 capture has been performed by academics [45, 48, 49] and industrial practitioners [57] resulting in ILs featuring high CO2 solubility. Therefore, they are attractive solvents for CO2 separation from flue gases. Most of the studies concerning the behavior of the CO2 -IL system have been performed at the laboratory scale, with pressures in the 1-200 bar range and temperatures in the 280-450 K range [58, 59, 60]. Those studies indicate that improved performance can be achieved at high pressures and low temperatures. Additionally, energy for solvent regeneration can be reduced compared to the MEA process, and even after several cycles of adsorption-desorption, the performance of ILs ACS Paragon 3Plus Environment


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for CO2 capture does not diminish [45, 49]. On the other hand, ILs have technical and economical disadvantages like high viscosities and purchase cost (up to $5000/kg). The combination of these drawbacks result in higher operating and capital costs [58, 61, 62]. However, other studies have demonstrated the technical feasibility of the CO2 capture process using ILs [63, 64, 65, 66]. It should be stressed that only a couple of these works have shown that this process can have better economic and technical characteristics in comparison to the traditional MEA-based CO2 capture processes. It is important to mention that these processes did not feature optimal operating conditions and that the selection of the IL was based on experimental results and process heuristics. Although a few process simulations at steady state of an IL-based scrubber compared to a MEA-based process are available [10, 26, 28, 63, 64, 65], a study that provides insight on the dynamic behavior of a complete CO2 capture process that make use of ILs as solvents is not available in the literature. In this work, based on the results of a previous work [67], the ionic liquid 1-decyl-3-methyl imidazolium trifluoromethanesulfonate ([C10mim][TfO]) was used for the CO2 capture process. Also, a new process flow sheet is proposed and compared against the MEA process flow sheet adapted from Nittaya et al. [10]. The IL was obtained by computer aided molecular design (CAMD) techniques as detailed in [67] and it was specifically tailored for the CO2 capture process. The CAMD methodology consists in selecting a combination of a set of functional groups to achieve a feasible molecule that mets an objective function (i.e. maximum solubility of CO2 in IL). Group contribution methods were used to predict the physical and thermodynamic properties of the new molecule since it was a novel compound for which no experimental pure component information was available. A summary of the key properties of the IL are presented in Table 1. The operating scenarios of fossil-fired power plants are known to vary daily and seasonally, .e.g. the electricity demands are typically low (high) during the night (morning) and are different from summer to winter. Accordingly, the dynamics of a power plant are continuously changing and subject to electricity demands coming from the power grid. It is expected that the flue gas flowrate, which is one of the key loads in a post-combustion CO2 capture plant, will follow the changes experienced by the power plant. Hence, the CO2 capture plant dynamics will play a key role and


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Table 1: Properties of the 1-decyl-3-methyl imidazolium trifluoromethanesulfonate ([C10mim][TfO]) Ionic Liquid estimated from group contribution methods. Properties Unit Value Source Reference Viscosity Pa s 0.3116 [68, 69] −1 −1 5 CP J mol K 5.9249×10 [70] Melting Temperature K 268.41 [71] Critical Temperature K 1115.00 [72, 73] Critical Pressure Bar 21.02 [72, 73] 3 Density kg m 1217.90 [72, 73] Boiling Temperature K 821.42 [72, 73] Acentric factor 0.31 [72, 73] determine the performance of this plant. Accordingly, the aim of this study is to provide insight on the transient behavior of the plant and analyze the potential challenges that may be faced when operating this plant in the dynamic mode. The objectives of this work are: (a) To present a feasible techno-economic study of a CO2 capture process IL-based plant; (b) To compare the operating and capital costs; (c) To compare the process performance of the IL for CO2 capture against the common MEA process used for the same aim; (d) To carry out a sensitivity analysis on the steady state performance to display the achievable limits for CO2 recovery and purity; and (e) To provide a dynamic analysis of IL-based plant under three different operating scenarios comprising step changes in flue gas flow rate, ramp changes in composition of CO2 and sinusoidal changes in the flow rate of the same stream. Once a steady-state design and the economics of a process have been analyzed, the next step is to perform an open-loop dynamic analysis to gain insight on the transient behaviour of the process in the presence of common disturbances and set-point changes. If the dynamic operability of the process is considered to be acceptable, then a controllability analysis can be performed to evaluate the process performance in closed-loop. Accordingly, we decided to perform both studies in the present manuscript. This study is organized as follows: Section 2 describes the process flowsheet proposed in this work the IL-based CO2 capture process. Section 3, the Results section, presents first a techno-economic analysis of the proposed flowsheets presented in Section 2. This is followed by a sensitivity analysis



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on the key process parameters. Furthermore, the last part of Section 3 discusses the dynamic characteristics of the proposed IL-based CO2 capture plant. Concluding remarks are presented at the end of this article.

2

Description of the Process flow sheet

This section provides a description of the two process flow sheets used in this work and that were implemented in the Aspen Plus and Aspen Dynamic simulation environments. The feed-stream data used for the CO2 capture processes simulation were taken from [9] and are presented in Table 2. Table 2: Flue gas stream conditions used Temperature (K) Molar flowrate (mol/s) Mole fraction CO2 H2 O N2

in the present analysis. 319.71 4.0125 0.175 0.025 0.8

These data were used for both the MEA-based plant and the IL-based plant simulation and analysis. As a way to measure the amount of CO2 captured by these processes, the CO2 removal and CO2 purity indicators were used. The CO2 removal and CO2 purity targets at the outlet stream were set at 90% and 95%, respectively. The CO2 removal was calculated as follows: CO2 removal = 1 −

FCO2 vented FCO2 fed

(1)

where FCO2 vented is the flow rate of CO2 in the Vented stream (VENTGAS) and FCO2 fed is the flow rate of CO2 in the flue gas (FLUEGAS) main feed stream (see notation in Figures 1 and 2). Moreover, CO2 purity is defined as the mol fraction of CO2 in the main product stream (CO2) in Figures 1 and 2. The MEA-based plant considered in this work (Figure 1) was taken from Harun et al.[9]. As shown in Figure 1, the CO2 capture MEA-based plant consists of an absorption step (ABSORBER), then the CO2 rich stream enters the stripping section (STRIPPER) where ACS Paragon 6Plus Environment

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the solvent is recovered and recirculated to the absorber. The heat exchanger EX-1 is used for energy recovery and the heat exchanger EX-2 reduces the temperature to feed stream conditions. A makeup stream of MEA and water is needed due to losses, i.e., by degradation. The nominal operating conditions for this plant are specified in Table 3, while mass and energy balances steadystate results are shown in Table 4. EX-2

WATER MIXER

LEAN-3

MEA

CO2 LEAN-MEA LEAN-2

VENTGAS

EX-1 STRIPPER

RICH-2 ABSORBER

FLUEGAS

RICH-1

LEAN-1

Figure 1: Typical CO2 capture process flow sheet using MEA [9]. CO2 removal = 0.9349, Flow (kgCO2 captured/h) = 101.55, CO2 purity = 0.9740 and Heat consumption (kW) = 155. EX-2

C-1 LEANIL CO2N2W-1

LEAN4

N2

VENTGAS1

CO2N2W-2

CO2N2

EX-1 ABSORBER RICHIL-2

FLASH-1 FLASH-2

C-COLUMN

LEAN3 P-1

CO2 WATER LEAN1

RICHIL-1 FLUEGAS

Figure 2: Flowsheet of the IL-based CO2 capture plant. CO2 removal = 0.9999, Flow (kgCO2 captured/h) = 106.47, CO2 purity = 0.9514 and Heat consumption (kW) = 45. The aim of the steady-state design of the CO2 capture IL-based plant was to achieve a similar ACS Paragon 7Plus Environment


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Table 3: Design specifications for the MEA process [9]. Equipment Number of stages = 7 Absorber Pressure = 1.035 Bar Pressure drop = 0.025 Bar Number of stages = 7 Pressure = 1.5 Bar Pressure drop = 0.0005 Bar Feed tray = 2 Stripper Condenser Heat duty = -11.5 kW Condenser Pressure = 1.6 Bar Reboiler Heat duty = 155 kW Reboiler Pressure = 1.6 Bar Flow direction: counter-current Ex-1 Hot stream outlet temperature decrease = 20 K Temperature = 314 K Ex-2 Pressure 1.01325 Table 4: Steady-state simulation results of the MEA process. Flow Pressure Temperature Molar fraction Stream [mol/s] [K] [Bar] CO2 N2 H2O CO2 0.6597 296.20 1.6 0.9716 0.0104 0.0180 FLUEGAS 4.0125 319.71 1.035 0.1750 0.8000 0.0250 LEAN-1 30.0266 387.89 1.6 0.0361 0.0000 0.8423 LEAN-2 30.0266 314.00 1.01325 0.0361 0.0000 0.8423 LEAN-3 30.3740 314.01 1.01325 0.0357 0.0000 0.8439 LEANMEA 30.3740 314.00 1.01325 0.0338 0.0000 0.8461 MEA 0.0003 314.00 1.01325 0.0000 0.0000 0.0000 RICH-1 30.6862 329.67 1.03825 0.0562 0.0002 0.8246 RICH-2 30.6862 351.10 1.7 0.0562 0.0002 0.8246 VENTGAS 3.7401 331.09 1.01325 0.0015 0.8564 0.1419 WATER 0.3072 314.00 1.01325 0.0000 0.0000 1.0000

MEA 0.0000 0.0000 0.1216 0.1216 0.1204 0.1201 1.0000 0.1190 0.1190 0.0001 0.0000

performance to that accomplished with the MEA-based plant. This means that the CO2 recovery has to be at least 90% and with a 95% purity of CO2 in the product stream. The initial flow-sheet of the IL-based plant was the same than the typical MEA-based plant shown in Figure 1. However, with this configuration the IL-based plant is not capable to reach the minimal CO2 purity (95%). Similarly, it was found that a flash unit was enough to recover the IL-based solvent. In order for the


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process to achieve the desired CO2 purity, a cryogenic column was deployed. However, it becomes imperative to remove the water in the product stream to avoid freezing water in the cryogenic column. Therefore, another flash before the cryogenic column unit was added to the flow sheet. In addition, a compressor was included helping to establish suitable operating conditions for the cryogenic column , i.e. to increase the pressure and reduce the energy for dehydrating the CO2 and N2. The proposed flowsheet used in this work for the IL-based CO2 capture plant is presented in Figure 2. The corresponding steady-state operating conditions for each unit are shown in Table 5 . Table 5: Design specifications for the IL CO2 capture plant. Equipment Number of stages = 14 Absorber Pressure = 7.91 Bar Pressure drop = 1.1 Bar Heat duty = 45 kW Flash-1 Pressure = 1.01325 Bar Temperature = 298 K Flash-2 Pressure = 2.3 Bar Number of stages = 3 Pressure =2 Bar Pressure drop = 0.2 Bar Partial condenser C-column Kettle reboiler Reflux ratio = 2 Distillate rate = 0.15 mol/s Feed tray = 3 Flow direction: counter-current Ex-1 Hot stream outlet temperature decrease = 10 K Temperature = 314 K Ex-2 Pressure 8.01 Type isentropic Compressor Discharge pressure = 2.3 Bar Regarding the ionic liquid deployed for CO2 capture purposes, a stream composed of 3.6 mol/s of the pure ionic liquid was added to the system. Because of the assumption made that the IL does not evaporate, full recirculation of the IL is possible; therefore, a make up stream for the IL is not needed. However, this assumption is valid only for steady state operation and for the open-loop ACS Paragon 9Plus Environment


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dynamic simulation analysis when the disturbances are only related to the flue gas flow rate. Under other simulation environments (i.e. different set of upsets and process control) a IL make-up stream will be required. The selection of the cryogenic system is based on its ability to achieve large throughputs and high purities compared against non-cryogenic systems, such as pressure-swing adsorption (PSA) and membrane technologies, which are typically employed for low production scale and generally for low product purities [74, 75]. In the cryogenic distillation column, the number of stages was calculated using the McCabe method. The system to be separated features a large average relative volatility (α = 26.86) thus requiring a small number of trays. The number of stages in the absorber column was obtained using the Kremser method [76]. Phase equilibrium models were used to predict the behaviour of this unit even when it is well known that rate-based models can provide a better representation of the system behaviour [77]. However, some reports indicate that the CPU time required to simulate an absorber column using rate-based models is at least one order of magnitude higher than that required by equilibrium-based models [78]. On the other hand, some studies have shown good agreement between non-equilibrium and equilibrium models when correct Murphree efficiencies were used in the simulations [79, 80, 81, 82, 83]. We should stress that in this work we assume that physical absorption is the mechanism for CO2 capture using IL. Therefore, using the equilibrium model may be suitable to predict the behavior of the IL. Moreover, data regarding the diffusivity of gases in ILs are lacking in the open literature. Therefore, an UNIFAC equilibrium-based model was employed in this work to estimate the thermodynamic equilibrium properties; an empirical model was used for computing the Henry law constant [84]. The model for the Henry law constant is influenced by the molecular weight of the ILs, since it indirectly determines the size and free space of the IL for CO2 capture. On the other hand, the gas phase was modeled using the Peng Robinson equation of state. Steady-state simulation results for the IL CO2 capture process are shown in Table 6. The pressure driven simulation option in Aspen PLUS was used to obtain a dynamic model for predicting the process dynamics of the CO2 capture process using IL. Required information for ACS Paragon10 Plus Environment

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obtaining the pressure-driven dynamic models is shown in Table 7. Table 6: Steady-state simulation results of the Flow Pressure Temperature Stream [mol/s] [K] [Bar] CO2 0.7143 192.17 2.2 CO2N2-2 0.8587 298.00 2.3 CO2N2W-1 0.9367 343.58 1.01325 CO2N2W-2 0.9367 435.73 2.3 FLUEGAS 4.0125 319.71 9.01 LEAN1 3.7680 343.58 1.01325 LEAN3 3.7680 345.44 10 LEAN4 3.7680 335.44 9 LEANIL 3.7680 314.00 7.91 N2 0.1444 166.43 2 RICHIL-1 4.7046 318.54 8.91 RICHIL-2 4.7046 327.61 2 VENTGAS 3.0758 314.68 7.91 WATER 0.0780 298.00 2.3

3

CO2 capture process using IL. Molar fraction CO2 N2 H2O IL 0.9532 0.0300 0.0169 0.0000 0.8161 0.1699 0.0140 0.0000 0.7482 0.1557 0.0961 0.0000 0.7482 0.1557 0.0961 0.0000 0.1750 0.8000 0.0250 0.0000 0.0003 0.0005 0.0437 0.9554 0.0003 0.0005 0.0437 0.9554 0.0003 0.0005 0.0437 0.9554 0.0003 0.0005 0.0437 0.9554 0.1383 0.8617 0.0000 0.0000 0.1492 0.0314 0.0541 0.7652 0.1492 0.0314 0.0541 0.7652 0.0005 0.9962 0.0033 0.0000 0.0001 0.0000 0.9999 0.0000

Results

This section presents the simulation results obtained while using the proposed IL-based CO2 capture process described in the previous section. These results are compared against similar results obtained using the conventional MEA CO2 capture process.

3.1

Techno-economic analysis

This section provides information regarding the cost of the processing units as well as the utilities costs of the MEA and IL CO2 capture plants featuring similar CO2 capture removal and CO2 purity specifications in the product stream. The performance of the IL-based plant is compared against the performance of the MEA-based plant. Nowadays, the MEA process is considered to be the most suitable process for CO2 capture. This comparison is expected to show the technical and economical feasibility of the IL-based plant. To achieve this goal, the estimation of the annualized capital cost ACS Paragon11 Plus Environment


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Table 7: Hydraulics information to obtain the pressure-driven dynamic model in Aspen Dynamics. Unit feature value Tray type bubble cap Absorber diameter 0.38 m spacing 0.6096 m Vessel type horizontal head type elliptical Sump tank height 2.34 m diameter 1.75 m Hold up 0.5 Vessel type vertical head type hemispherical Flash1 height 1.68 m diameter 0.84 m Vessel type vertical head type hemispherical Flash2 height 0.164 m diameter 0.082 m Tray type bubble cap Cryogenic column diameter 0.1 m spacing 0.3 m head type elliptical height 0.31 m Reflux drum (Condenser) diameter 0.16 m Hold up 0.5 Vessel type horizontal head type elliptical Sump tank (Reboiler) height 0.42455 m diameter 0.2123 m Hold up 0.5 for the MEA-based plant, and IL-based plant without cryogenic column was calculated using the following equation:

Ca = Cb

Aa Ab

n

(2)

where C is the purchase cost, A is the equipment cost attribute, n is the cost exponent. The a and b subindexes refer to equipment with the required attribute and equipment with the base attribute, ACS Paragon12 Plus Environment

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respectively. The base attributes were taken from [63, 19, 85, 26]. The capital costs associated with the cryogenic distillation column cost were estimated as follows [86, 87]:

0.8 0.8 CrC = CC · N T · CD1.245 + CR · AR + C C · AC

(3)

where CrC is the cryogenic column cost, CC is the column cost coefficient, C is the cost coefficient, A is the area, N T is the total number of stages, CD is the column diameter. The C and R subscripts refer to the condenser and reboiler, respectively. The utility cost was estimated considering only the cost (Cost) of some of the services (i = heat, electric power and cryogenic cooling): utility costi = F lowi · Costi

(4)

where F low stands for the flowrate value of the utility. Table 8 shows the values of the estimated capital cost for both MEA and IL CO2 capture processes. From this table it is clear that the capital costs for the IL CO2 capture plant are 71% higher than similar costs related to the MEA plant. The cost of the absorbent was also considered because ILs are expensive solvents compared to aqueous amines. The basis of the MEA price is $1.25/kg [12, 26]. The cost of ILs is highly dependent upon the choice of cation, anion, purity, and manufacturer. The cost of the IL used in this work with a purity of 95% was estimated to be $20/kg [88, 63]. The annualized cost results were obtained using a 15 year basis. On the other hand, because of the negligible vapour pressure of the ionic liquid we can assume no solvent losses. Furthermore, the capability of the IL for CO2 capture does not diminish with time. Therefore, it can be deployed even for several cycles of absorption-desorption [45,49].

Table 9 shows the corresponding estimated values of the utilities operating costs for both the MEA and the IL-based CO2 capture plants. As shown in this table, the total utilities cost for the proposed IL CO2 capture plant are 60 % lower than similar utility costs for the MEA-based plant. This is due ACS Paragon13 Plus Environment


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Table 8: Capital cost for the equipments (All in $1000). US MEA absorber 390 stripper 423.15 exchanger 65.65 precooler 150.15 pump 10.4 cryogenic column 0 solvent 22.76 Annualized cost (MM$ /year) 0.1593

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dollars are referred to 2013. IL 492.8 66.5 51.8 77.5 17.5 552.5 240 0.2248

to the fact that the IL CO2 capture plant features smaller flowrates of the heating utility (29%). In fact, in the MEA process the main heating utility consumption occurs in the stripping section. On the other hand, the demand of the heating utility in the IL process takes place in the Flash-1 unit (see Figure 1) where only a modest amount of heating utility is required to meet the plant performance specifications. Energy consumption, CO2 removal, CO2 purity and amount of CO2 captured are shown in the captions of Figures 1 and 2 for both CO2 capture process. The capital costs for the MEA and IL CO2 capture plants are 0.1593 and 0.2248 MM$/yr, respectively. This result indicates that, from a capital costs point of view, the IL CO2 capture plant is 74% less expensive than the MEA CO2 capture plant for the flow sheets and operating conditions that were considered for these plants. It should be noted that the flowsheet of the MEA-based plant was obtained from real pilot plant [89] and the flowsheet and operating conditions for IL-based plant were obtained from process heuristics and engineering judgment. Optimal operating process designs of the IL plant can be obtained using formal algorithmic tools based on mathematical programing [90]. From the results shown in Tables 8 and 9, the total annual cost (TAC) obtained by adding both capital and operating costs are 892.78 M$/year and 660.11 M$/year for the MEA and IL CO2 capture processes, respectively. Hence, the IL CO2 capture process is 26% less expensive than the corresponding MEA process. Also from Tables 8 and 9 it is clear that the highest cost for both process are related to utilities consumption. In fact, the operating cost for the MEA process is 82% of the TAC, whereas for the IL process this reduces to 66%. In addition, the cost of the IL represents 6% of the TAC of the IL CO2 capture processes. If the price of the IL increases twice, ACS Paragon14 Plus Environment

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then the TAC value increase 5.5% representing 10% of the TAC. In this work, the price of the IL was estimated assuming an industrial production with a purity of 95%. Further, a simplified cost analysis was performed since we have proposed a preliminary design analysis. Our aim was to carry out a basic cost analysis to compare the economics of the our proposed CO2 capture plant design and the typical MEA-based plant design.

Power costs Heating costs Cryogenic costs Total operating costs

3.2

Table 9: Operating costs MEA IL MEA $1000/year (2013 US$) $/kg CO2 0.25 5.20 0.0003 733.21 212.87 0.8243 0.00 217.25 0.0000 733.47 435.32 0.8246

IL captured 0.0056 0.2282 0.2329 0.4667

Individual costs $/kWh 0.16 0.54 9.70

Steady state sensitivity analysis of the IL CO2 capture plant

The aim of this section is to present a steady-state sensitivity analysis of the IL CO2 capture plant that would provide insight on the effect of the flow rate and CO2 mol fraction of the Flue Gas stream on the CO2 capture removal, purity and mass flow rate of the CO2 product stream. Previous studies on post-combustion CO2 capture have identified that the flue gas conditions are the key variables affecting the performance of this plant [27, 91, 92]. Therefore, it becomes relevant to analyze how the normal operation of the plant is affected when changes in the flue gas stream conditions enter into the process. The results shown in Figure 3 indicate how the CO2 capture removal and the CO2 mass flow rate of the main product stream are affected when the molar flow rate of the Flue Gas stream is modified. As shown in this figure, when the Flue Gas flow rate is increased beyond the nominal base-case value considered in this study, the amount of CO2 captured by the IL increases until it reaches an upper limit in the Flue Gas flow rate (around 6 mol/s). At this point the IL becomes saturated and the amount of additional CO2 captured becomes negligible. This is the reason why the CO2 mass flow rate in the rich stream curve turns out to be asymptotic. This also explains the behaviour observed ACS Paragon15 Plus Environment


for the CO2 capture removal. Since the IL becomes saturated, the CO2 that is not captured by the IL needs to be removed in the venting gas stream of the absorber shown in Figure 2. Figure 4 shows the changes in CO2 capture removal and CO2 product purity when the mol fraction of CO2 in the Flue Gas stream is modified. As shown in Figure 4, small changes were observed in the output variables due to the relatively small amount of CO2 fed in the Flue Gas stream. Figure 4 does not show the saturation of the IL because the change in the CO2 mol flowrate of the feed stream turns out to be smaller (0.2 mol/s) when compared to the same variable displayed in Figure 3 (0.79 mol/s). 40

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20 0.75

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Figure 3: CO2 mass flow rate and CO2 removal at different Flue Gas flow rates. Another operating variable whose impact should be clarified is the flowrate of the ionic liquid. Accordingly, Figure 5 shows the effects of the IL flowrate on the CO2 removal rate and CO2 purity. As illustrated in this Figure, a linear decrease in the CO2 removal rate is observed when the IL flowrate falls below 2.75 mol/s; this phenomenon occurs because of the limited availability of solvent in the system and the fact that a saturation condition is reached at that operating point (0.25 mol CO2 /mol IL). Note that the same behavior is observed from Figure 3 when the flue gas flowrate is increased. On the other hand, the IL also removes (captures) N2 from the system; thus, an increase ACS Paragon16 Plus Environment

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0.9982

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Figure 4: CO2 removal and CO2 purity product varying CO2 mol fraction in Flue Gas stream. in the lL flowrate also increases the capture rate of both CO2 and N2 at the expense of reducing the purity in the CO2 product stream, i.e. more N2 is present in the product stream at large IL flowrates. 1

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Figure 5: CO2 removal and CO2 purity at different Ionic Liquid flow rates.

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3.3

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Transient response of the IL-based CO2 capture plant

Process dynamics are not typically considered at the design stage; which may return economically attractive processes that are dynamically inoperable [93], [94]. Ideally, since design and operability are highly related objectives, they should be taken into account simultaneously, so interactions between these two activities render optimal and operable process designs. However, considering simultaneously economic and operable designs is challenging, specially for large-scale chemical processes like that proposed in this study. When economics is deployed as the only design objective, it is important to run an a posteriori operability analysis such that the designer gains insight on potential controllability problems. This is the aim of the open-loop dynamic analysis presented in this section of our study. Moreover, the cases considered to evaluate the plant’s dynamic behaviour were selected since these are the more likely scenarios that are expected to occur during the actual operation of a CO2 capture plant. In addition, some other previous studies have also used similar tests to evaluate the dynamic performance of this plant [9], [10], [17], [35]. The steady-state flowsheet shown in Figure 2 was implemented in Aspen Dynamics with the purpose of providing insight on the dynamic characteristics of the IL-based CO2 capture plant. To the authors’ knowledge, this is the first study that presents the dynamic behavior of an IL-based CO2 capture plant. To analyze the dynamic behavior of this process, the following scenarios were considered: (a) step changes in the flue gas flow rate, (b) ramp changes in the CO2 composition of the feed stream and, (c) sinusoidal changes in the flue gas flow rate. In all the case studies, changes of ± 10% with respect to the nominal (base-case) condition in the flue gas stream reported in Table 2 was considered. Scenario 1: Step change in flue gas flow rate As shown in Figure 6, positive step changes of 10% in the flue gas flow rate give rise to inverse response in the CO2 capture removal. For a positive step change in the flue gas flowrate, the cause of the inverse response behaviour has to do with the fact that a sudden increase in the amount of ACS Paragon18 Plus Environment

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CO2 of the absorber’s feed stream (see Figure 2) initially exceeds the IL equilibrium capability for absorbing such an amount of excess CO2 . For this reason, after the disturbance hits the process, there is a quick and sudden reduction in the CO2 removal characteristics meaning that the amount of non-absorbed CO2 increases in the VENTGAS stream of Figure 2. However, from here on, the equilibrium absorption characteristics are slowly established again until a new steady-state is reached, which explains why the CO2 removal dynamic response has now a different response direction. Following similar arguments, the inverse response observed for a negative step change in the flue gas flowrate can also be explained. Once the response reaches its settling time, the amount of CO2 recovered increases since the CO2 removal index increases when the flue gas flow rate increases. Because of the increase in the amount of CO2 fed to the plant, the IL can recover larger amounts of CO2 . This behavior was also observed for the CO2 purity since for both positive and negative step changes in the flue gas flowrate the CO2 purity shows inverse response. 1 0.968 0.998 0.966

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Figure 6: Response in CO2 removal (- - -) and CO2 purity (—) to step changes in the flue gas flow rate. As shown in Figure 7, the equilibrium temperature in the first flash of the flow sheet (FLASH-1 in Figure 2) decreases when the flue gas flow rate is increased. This is because the vaporization ratio (for a two-phases flash unit featuring a feed stream flowrate F , vapor flowrate stream V and a liquid ACS Paragon19 Plus Environment


flowrate L, the vaporazation ratio is defined as ratio vapor flowrate/feed stream flowrate) in the flash decreases leading to a lower equilibrium temperature since the flow rate of the product gas stream in this equipment increases. Therefore, the liquid level in the Flash-1 decreases for positive changes; similarly, the absorber the liquid level decreases when positive changes hit the process. As shown in Figure 7, the same disturbance applied in different directions results in different dynamics. In fact, for a positive step change in the flue gas flow rate, the flash temperature shows a larger deviation, and also a more significant inverse response. These results are a clear indication of the degree of nonlinearity of the plant. For a positive step change in the flue gas flowrate, the inverse response observed in Figure 6 can be explained on the following grounds. Initially, after the disturbance hits the process, the temperature increase in the equilibrium flash temperature is due to a reduction in the concentration of the light components other than the IL. After that small temperature increase reaches its peak, the concentration of the light components starts increasing, which reduces the equilibrium flash temperature. The implications of inverse response are well known. Inverse response can impose some stability and performance constraints when PID controllers are employed [95]. 349

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Figure 7: Temperature in the first Flash unit (see Figure 2) during step changes in the flue gas flowrate.


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Scenario 2: Ramp change in the CO2 composition of the flue gas stream Figure 8 shows the dynamic response when ramp changes in the CO2 composition of the flue gas enter the process. After 24 hrs of steady-state operation, the ramp was implemented by linearly increasing/decreasing the flue gas flowrate for 10 h until +/- 10% of the nominal (base-case) value is reached. Contrary to the previous scenario, the CO2 removal shows a somewhat overdamped response up to the end of the disturbance. The response for CO2 purity presents an oscillatory behaviour up to 30 hr. For negative changes the response quickly decreases. However, for a positive change there is a sudden variation in CO2 purity which occurs around 34 hr. As shown in Figure 8, a negative ramp change in CO2 composition produces a drop of 90K in the temperature of the condenser. Therefore, the amount of N2 increases in the cryogenic column reducing the temperature profile of the cryogenic column (see Figure 9). On the other hand, when the CO2 mol fraction increases in the flue gas, the temperature profiles in the cryogenic column increases up to 215 K. In this case, the residence time does not change because the flow is constant and the flow of CO2 in the other stream is constant as well. 0.999

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Figure 8: Response in CO2 removal (- - -) and CO2 Purity (—) to ramp in CO2 composition.



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160

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Figure 9: Steady-state temperature profile in cryogenic column for ramp disturbance in CO2 composition. Condenser is tray 1. Scenario 3: Sinusoidal perturbation in flue gas flow rate In most of the power plants, the demand of electricity varies with respect to time (i.e. higher demands are expected at peak loads) giving rise to time changes in the processing conditions of the flue gas stream. Therefore, it becomes important to realize how the CO2 capture efficiency changes during a normal operating cycle. The aim of this scenario is to analyze the dynamic performance of the IL-based CO2 capture plant for a 24 hr operating cycle assuming that the flue gas flowrate follows a sinusoidal-like change with a frequency of 12 h and amplitude of 10%. In Figure 10 the CO2 capture removal and CO2 purity responses to this change are presented. This figure shows that the magnitude of the output response is larger when the CO2 purity is analyzed. Figure 10 also presents the responses observed in CO2 capture removal and CO2 purity. As shown in this Figure, there is a larger variability in the CO2 purity than in CO2 capture removal, i.e. the amplitude of the CO2 purity signal is +0.41% and -2.22% with respect to its nominal value. On the other hand, the variability observed in the CO2 capture removal is almost insignificant, i.e. +0.23% and -0.37% with respect to its nominal value. In summary, the output responses suggest that during power


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peak loads the plant can accommodate sinusoidal-like variations in the flue gas flow rate upsets. Of course, we expect superior performance (i.e. good disturbance rejection) when a control system is used. 0.99

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Figure 10: Response in CO2 removal to sinusoidal upset .

4

Conclusions

A study on the economic impact and the dynamic performance of the CO2 capture plant using IL has been presented. The techno-economic analysis for an IL CO2 capture plant shows higher capital cost than a typical MEA CO2 capture plant. However, the CO2 capture process with IL can save energy and utility costs even though they require additional units compared to the MEA-based CO2 capture process. It is important to mention that even when the proposed IL achieves larger CO2 capture removal, the amount of CO2 that was captured before the cryogenic unit is far from its target specification (i.e. around 85%). Therefore, adding the cryogenic unit allowed the process to reach the CO2 purity specification in the product stream. From the sensitivity analysis carried out in this work it should be remarked that the sensitivity of CO2 removal only becomes important for flue gas flow rates above 6 mol/s. Moreover, the sensitivity of CO2 purity is negligible for positive ACS Paragon23 Plus Environment


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changes in the Flue gas flow rate and composition. In all cases, the CO2 recovery is over 99% and the CO2 purity is around 90%. The open-loop dynamic analysis carried out suggests that the CO2 capture IL-plant is a highly nonlinear process. The perturbation in the flow rate of the flue gas has a larger impact than the composition, due to the fact that the flow of IL cannot capture larger amounts of CO2 .

Acknowledgments The authors are grateful to the Emerging Leaders in the Americas Program from Canada (ELAP) and Economic and Development Patronage of Universidad Iberoamericana (FICSAC) for the financial support provided for this research.


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