CO2 Capture Modeling, Energy Savings, and Heat ... - ACS Publications

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CO2 Capture Modeling, Energy Savings, and Heat Pump Integration Stuart J. Higgins and Y. A. Liu* SINOPEC/AspenTech Center of Excellence in Process System Engineering, Department of Chemical Engineering, Virginia Polytechnic Institute and State University, Blacksburg, Virginia 24061, United States S Supporting Information *

ABSTRACT: While chemical engineers have designed amine-based CO2 capture systems since 1927, concern over anthropomorphic climate change sparked renewed interest in the 1980s. Subsequent research has led to significant advances such as well-fit property and thermodynamic models, rigorous models for unit operations, and improved process designs. The aim of this work is to summarize proposed process improvements and energy-saving schemes, including absorber intercooling, stripper interheating, stripper condensate rerouting, distributed cross heat exchanger, flash stripper, multipressure stripper, and heat pump integration. We present simulation examples demonstrating that energy-saving schemes must be considered together because of strong and complex interactions. We report an optimal process design integrating all of these improvements in Aspen Plus V8.5, and then evolutionally simplify the design through simulation results. In particular, we propose an optimum energysaving design that uses an absorption-driven heat pump together with a distributed cross heat exchanger and a stripper vapor condensate rerouting to reduce both the cooling and the heating utility consumptions. The resulting predicted solvent regeneration energy is an absorption-driven heat pump for waste heat recovery with a predicted solvent regeneration energy of 1.67 (GJt/tonne CO2 captured). To our knowledge, this is the lowest solvent regeneration energy yet reported in the literature or patents. 1.3. Quantifying System Performance. Authors vary widely in how they report the performance of their CO2 capture unit designs. The most commonly reported values include thermal regeneration energy, energy penalty, and equivalent work. 1.3.1. Regeneration Energies. Most CO2 capture units use both heat (usually from utility steam) and electricity. Authors report these usages in terms of energy per unit of CO2 captured. Thermal regeneration energy Eregen quantifies heat usage as t

1. INTRODUCTION 1.1. History. Carbon capture is well-established as part of large-scale industrial processes such as that used in the Haber process for producing ammonia. Catching carbon dioxide in these processes is relatively cheap and easy to due to the high pressures at which they operate, typically between 15 and 25 bar. Figure 1 shows how physical absorption tends to be favored over chemical absorption at higher pressures. In the rest of this Article, we refer to chemical absorption only. This supply has long satisfied the demand for CO2 for purposes such as producing food or chemicals, although recent concerns over anthropomorphic contributions to climate change and increasing interest in enhanced oil recovery have sparked a drive to capture carbon dioxide more economically. The first known CO2 capture process patent was filed in 1927, and it was essentially a series of absorption towers.2 The more familiar flowsheet, which includes solvent regeneration, was filed in 1930,3 again in 1961,4 and again in 2005.5 1.2. CO2 Capture System. Figure 2 shows a flowsheet for the baseline CO2-capture system first described in U.S. Patent 1,783,901 in 1930.3 Figure 3 shows the CO2 capture cycle. Flue gas from a major point source, for example, a coal-fired power plant, loses CO2 to the solvent as they flow countercurrent through the absorption tower. The CO2-depleted flue gas is typically released to atmosphere, while the solvent goes to the regeneration tower. The regeneration tower strips CO2 from the solvent through heating. Captured CO2 is generally compressed to somewhere between 110 and 150 bar.6−9 In the figure, we typically refer to the solvent stream leaving the bottom of the absorber and going to the top of the stripper that is loaded with captured CO2 as “rich solvent”. We refer to the solvent stream leaving the bottom of the stripper (regenerator) and returning to the top of the absorber as the “lean solvent” because it contains relatively little carbon dioxide as compared to the rich solvent. © 2015 American Chemical Society

Etregen

∑∀ utility heat i Di ≡

consumers

Fcaptured CO

2

(1)

where ∀ represents “for all”, Di is the heat duty of unit i, and Fcaptured CO2 is the flow rate (usually mass flow rate) of captured CO2. Simple designs like that shown in Figure 2 only employ a single heater, the reboiler, so for such designs: Etregen =

DSTR‐REB Fcaptured CO

(2)

2

where DSTR‑REB is the heat duty of the stripper’s reboiler. Similarly, electrical regeneration energy Eeregen quantifies electricity usage as Eeregen

∑∀ electricity i Bi ≡

consumers

Fcaptured CO

2

(3)

where Bi is the electrical duty of unit i. Typically, the large majority of Eregen will come from the electricity requirement for e Received: Revised: Accepted: Published: 2526

November 25, 2014 February 9, 2015 February 10, 2015 February 26, 2015 DOI: 10.1021/ie504617w Ind. Eng. Chem. Res. 2015, 54, 2526−2553

Article

Industrial & Engineering Chemistry Research

Figure 1. Chemical absorption tends to dominate at lower pressures, while physical absorption tends to dominate at higher pressures. Graph from DOE/NETL.1

Figure 2. Flowsheet representation of a baseline CO2 capture unit.

Figure 3. CO2 cycle in a basic CO2 capture unit. Flue gas loaded with CO2 enters the process, has some portion (known as the capture rate R) of its CO2 absorbed, and then leaves, usually to atmosphere. The captured CO2 is stripped from the solvent in the regeneration tower (stripper) through heating. 2527

DOI: 10.1021/ie504617w Ind. Eng. Chem. Res. 2015, 54, 2526−2553

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The size-up factor refers to how many times larger (less one) a power plant implementing CCS would need to be to produce the same amount of output as one not implementing CCS:

the CO2 compression train, while the balance will come from solvent pumps. Some CO2 capture designs incorporate compressors in other locations such as on the flue gas stream, and these designs will suffer much higher electrical regeneration energies as a result. Section 3.7 discusses a correlation for estimating CO2 compression work from stripper pressure provided by Van Wagener.7 1.3.2. Equivalent Work. We are interested in the amount of energy production lost due to implementing CCS (carbon capture and storage) in a power plant. CCS systems generally use both utility steam for heating and electricity. The electricity used to capture a unit of carbon, Eregen , is taken directly out of the t plant’s production. The utility steam used by the CCS system could have otherwise been used in the turbine to produce electricity, so it reduces electrical output by ηiEregen where ηi is the t,i conversion factor from steam heat to the electrical energy it could have produced in the turbine. Total equivalent work of the CCS system is then CCS W eq = Eeregen +

∑

ηisteamEt,regen i

∀ steam sinks i

EPsize‐up ≡ =

(4)

(5)

where εCarnot = 0.75 is the Carnot efficiency and Tturbine is the temperature at which the turbine condenses the steam, for example, Ttrubine ≈ 40 °C. In practical terms, the Carnot efficiency factor εCarnot can vary with the plant’s metrics such as the efficiency of its coal-fired boiler. Alternatively, ηsteam can be estimated as i ηi

steam

≈

plant ηw/o CCS

η boiler

(6)

where ηplant w/o CCS is the ratio of net electrical energy output of a plant without CCS to the thermal energy it obtains by burning fuel, and ηboiler is the thermal efficiency of the coal-fired boiler. This estimation does not consider steam temperature Tsteam,i, and ηsteam is the same for all i. i 1.3.3. Energy Penalties. There are two common definitions for energy penalty: the production-loss factor EPprod‑loss and the size-up factor EPsize‑up.6,13 The production-loss factor is the portion by which energy production is reduced due to implementing CCS: EPprod‐loss ≡ 1 −

plant ηw/o CCS

⎞ ⎟ ηisteamEt,regen i ⎟ ⎠ ∀ steam sinks i

1 1

−1

(8)

2. ENERGY-SAVING SCHEMES We can modify the basic process in Figure 2 to save energy. These energy-saving schemes generally focus on reducing the regeneration tower’s reboiler duty, DSTR‑REB. We will first look at the simplified case of individual energysaving schemes used in isolation. Section 2.4 will explain that, in practice, energy-saving schemes must be considered together because of their significant interactions. To illustrate these interactions, we include some figures and tables based on simulation results for a demonstration plant being built in the Asia Pacific for capturing one million tonnes of CO2 per year. We discuss the details of our simulation in sections 3 and 4. 2.1. Larger Unit Sizes. Using larger units is the simplest way to increase energy efficiency. The primary adjustable unit sizes are the packed heights of absorption and regeneration towers, and heat-transfer areas of cross heat exchangers. We find that the regeneration tower’s packed height has a modest effect on the regeneration energy, so we are primarily concerned with the absorption tower’s packed height and the area of the cross heat exchangers.

plant ηwith CCS

⎛ = RεeCCS⎜⎜Eeregen + ⎝

−1

The size-up factor will always be larger than the production-loss factor. Energy penalty is a poor metric for comparing reported CO2 capture units because it heavily depends on three factors: (i) the efficiency of the CO2 capture unit; (ii) the efficiency of the reference plant without CCS (e.g., turbine efficiency); and (iii) the carbon capture rate assumed. An author might reasonably refer to (i) a reference plant without CCS efficiency of 35%14 to 44.4%15 or an efficiency determined from equivalent work; (ii) an electricity-based, CO2 emission rate from 0.21 (tonne CO2/GJe)6 to 0.2833 (tonne CO2/GJe);16 (iii) a carbon capture rate between 80% and 95%; and (iv) either the production-loss energy penalty or the size-up energy penalty. For a process with a thermal regeneration energy of Eregen = t 1.67 (GJt/tonne CO2) and an electrical regeneration energy of Eregen = 0.393 (GJe/tonne CO2), the reported energy penalty e then could range from 10.9% to 43.9% (see Table 1). This range neglects the fact that regeneration energy tends to depend on carbon capture rate as exemplified in Figure 23, which could cause the range to vary further. Additional ambiguity exists in the referenced flue gas composition (see Figure 23). Table 1 illustrates this ambiguity resulting from the different assumptions of the efficiency of the CO2 capture unit, the efficiency of the reference plant without CCS, and the carbon capture rate assumed. For comparing CO2 capture systems between authors or works, we recommend avoiding the use of energy penalty as a basis. Our principal concern is that reports of energy penalty percentage typically do not describe their assumed values of turbine efficiency, emission rate, or/and capture rate.

Tsteam, i − Tturbine Tsteam, i

plant ηwith CCS

RεeCCS(Eeregen + ∑∀ steam sinks i ηisteamE t,regen i )

Babatunde and Rochelle10−12 suggest estimating ηsteam assuming i that the power plant’s turbine functions at 75% Carnot efficiency, such that ηisteam ≈ εCarnot

plant ηw/o CCS

∑

(7)

where ηplant (i = reference plant with or without CCS) is the ratio i of net electrical energy output of plant i to the thermal energy it obtains by burning fuel; ηsteam is the ratio of electrical energy i obtainable to the thermal energy used for steam i; R is the carbon CO capture rate; and εe 2 is the thermal emission rate: the mass flow rate of CO2 emitted per unit of thermal energy produced by burning fuel in the base case plant. 2528


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Table 1. An Illustration of the Variations of the Energy Penalty Measures Resulting from Different Assumptions of Carbon Capture Rate, Turbine Efficiency, and Carbon Emission Ratea energy penalty ηsteam from eq 5 i

ηsteam from eq 6 i

capture rate R

steam-to-power efficiency (ηplant w/o CCS)/ ηboiler [MWe/MWt]

2 CO2 emission rate εCO e [tonne CO2/GJe]

size-up penalty EPsize‑up

production-loss penalty EPprod‑loss

size-up penalty EPsize‑up

production-loss penalty EPprod‑loss

80% 80% 80% 80% 95% 95% 95% 95%

35.0% 35.0% 44.4% 44.4% 35.0% 35.0% 44.4% 44.4%

0.210 0.283 0.210 0.283 0.210 0.283 0.210 0.283

12.2% 17.2% 12.2% 17.2% 14.8% 21.1% 14.8% 21.1%

10.9% 14.7% 10.9% 14.7% 12.9% 17.4% 12.9% 17.4%

19.6% 28.4% 23.5% 34.6% 24.2% 35.7% 29.2% 43.9%

16.4% 22.1% 19.1% 25.7% 19.5% 26.3% 22.6% 30.5%

a

This table uses a thermal regeneration energy of Eregen = 1.67 GJt/tonne CO2, an electrical regeneration energy of Eregen = 0.393 GJt/tonne CO2 t t (for a stripper pressure of 1.4 bar), a reboiler temperature of TREB = 110 °C, a reboiler temperature approach of ΔTREB = 10 K, and an assumed condenser temperature of 40 °C.

Figure 4. Increasing the absorber’s packed height tends to decrease regeneration energy.

which distribution of the cross heat exchanger area was optimized for the case Atotal ≅ 27.4 (m2/tonne CO2). 2.2. Simple Process Modifications. Several process modifications are relatively simple, involving the change or addition of two or less unit operations. These modifications do not significantly change the fundamental way in which the process operates. 2.2.1. Absorber Intercooling. Figure 6 shows the absorber intercooling scheme in which a side cooler is attached to the absorption tower. Absorber intercooling is meant to drive CO2 absorption by making it more thermodynamically favorable. Absorber intercooling was filed for patent in 1931 and then again in 1967.17,18 It was reported to be effective in several large diglycol amine (DGA) plants in 1981.19,20 Absorber

Figure 4 shows the decrease in solvent regeneration energy as the packed height of the absorber is increased. All other factors being equal, a taller packed section in the absorption tower will result in a higher rich solvent loading and thus a lower regeneration energy. Also, because CO2 absorption is exothermic, the increased net extent of absorption will cause the rich solvent to be warmer. Figure 5 shows the decrease in solvent regeneration energy with increasing cross heat exchanger area. Generally, larger cross heat exchangers will lead to greater energy efficiency up until a point where further increase in size begins to have a negligible effect on heat transfer because of approaching a minimal temperature approach. The case plotted in Figure 5 uses the distributed cross heat exchanger configuration (discussed in section 2.2.4) in 2529


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Figure 5. Increasing total cross heat exchanger size tends to decrease thermal regeneration energy. Heat exchanger distribution fractions are optimized for the 27.4 (m2/tonne CO2) case with a packed height of 30 m.

Figure 7. Flowsheet representation of a CO2 capture unit with generalized absorber intercooling. Multiple intercoolers are placed at different points on the absorption tower.

Figure 6. Flowsheet representation of a CO2 capture unit with absorber intercooling. The side heat exchanger on the absorber provides cooling to a liquid draw from the tower.

found appearance of this explanation was by Ragatz on the cover of his 1931 absorber intercooling patent.17 Figure 7 generalizes the single-intercooler configuration shown in Figure 6 by further distributing the cooling across the column. An implementation of this general scheme was patented in 2012 by Stoever et al.25 While generalized absorber intercooling should yield greater energy efficiency, it is unusual to see it in practice. The disadvantages of absorber intercooling include: increased capital cost; increased process complexity; increased utility cost of cooling; reduced kinetic reaction rate; and loss of heat from the system. These disadvantages can outweigh benefits even when the increase in capital and cooling utility costs are neglected. We recommend case-specific simulations to determine what degree of absorber intercooling, if any, is optimal for a given process.

intercooling became part of the Fluor Econamine process between 1992 and 2003, and was further explored by Freguia and Rochelle.21−23 Recently, Alstom and Dow24 filed a patent for a series of generalizations of absorber intercooling in which other streams are combined with the draw-off stream from the absorber or combined with the cooled output from the intercooler. Conceptually, absorber intercooling simply distributes the cooling ordinarily done in the lean solvent cooler downstream to positions on the absorption tower. The advantage is that higher temperatures are maintained at the top of the absorber where absorption is kinetically driven, while cooling still brings down the temperature lower on the column where absorption into the loaded solvent becomes thermodynamically driven. The earliest 2530


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heat available for the cross heat exchangers. While the interheater delivery location is superior, this opportunity cost further mitigates its overall effect. Ignoring capital costs and pressure drops, stripper interheating tends to provide a slight benefit to an otherwise optimized system. Supporting Information Table S2 shows our optimization approach for stripper interheating. In the table, Nstage is the number of rate-based stages for the stripper, and HSTR is the height of the stripper. 2.2.3. Stripper Condensate Rerouting. Figure 9 shows stripper condensate rerouting, the idea that returning the stripper

Table S1 in the Supporting Information shows our optimization approach for absorber intercooling. In the table, NABS is the number of ideal stages for the absorber, HABS is the height of the absorber, DSTR‑REB is the heat duty of the stripper’s reboiler, and kscale is a scaling constant as a ratio of plant size to simulation size that we will detail in section 3.2. We model the absorber intercooler in terms of its height on the absorber HABS‑IC at which it applies its heat duty DABS‑IC. Optimization determines the ideal height and duty for this side cooler. If optimization determines that DABS‑IC = 0, then we do not use absorber intercooling. 2.2.2. Stripper Interheating. Figure 8 shows stripper interheating, the idea that adding a heat exchanger between the

Figure 8. Flowsheet representation of a CO2 capture unit with stripper interheating. The stripper has a side heat exchanger to recover heat from the freshly regenerated solvent.

Figure 9. Flowsheet representation of a CO2 capture unit with stripper condensate rerouting. The condensate coming from the stripper’s condenser is usually returned to the top of the stripper despite the need for the stripper to be hot. This scheme redirects the condensate to the lean stream cooler, avoiding unnecessarily cooling the stripper.

freshly regenerated lean solvent and the side of the stripper will improve the process. This side heat exchanger is very limited in how much heat it can recover, but the heat that it does recover is more effective than that recovered by the central cross heat exchanger because it is delivered to an optimized intermediate point in the stripper. The first found appearance of stripper interheating comes from Cabanaw and Mann in their 1982 patent on an alkylation process.26 In 2003, Leites et al. reported the implementation of stripper interheating from 198627 wherein the interheater is constructed into the regeneration column.28 In 2006, Oyenekan and Rochelle coined internal stripper interheating as internal exchange.11,29 The primary difference between conventional stripper interheating and internal exchange is in the construction the column and interheater. In practice, whichever is designed to transfer more heat should be more efficient. We model the stripper’s interheater in terms of its height on the stripper HSTR‑IH and its heat-exchanging surface area ASTR‑IH. Optimization determines the ideal height and size for this side heater. If optimization determines that ASTR‑IH = 0, then there is no side heater. Stripper interheating tends to suffer from a low thermal driving force because much of the stripper usually operates near the temperature of the stripper’s bottoms. In some cases, such as in a process using the multipressure stripper discussed in section 2.3.2, the stripper bottom stream can be cooler than the inside of the column. Further use of the stripper interheater limits the

condenser’s condensate to a location other than the top of the stripping column will improve the process. The first appearance of stripper condensate redirection comes from Bresler et al. (1963).4 Supporting Information Table S3 shows our optimization approach for stripper condensate rerouting. We set this location to be the top of the absorption column and define the split fraction to it as XSTR‑CON. Optimization determines the ideal split fraction. If optimization determines that XSTR‑CON = 0, then there is no condensate rerouting. Figure 10 shows the effect of stripper condensate rerouting in the reported design (section 4). In this case, stripper condensate rerouting reduces the total steam consumption by 0.5%. This relatively modest potential for savings follows from the significantly reduced vapor flow rate induced by the distributed cross heat exchanger (see Figure 13). 2.2.4. Distributed Cross Heat Exchanger. Figure 11 shows a simplified version of central cross heat exchanger distribution, the idea that splitting a portion of the rich solvent in the central cross heat exchanger off for further heating before entering the stripper at a midpoint will improve the process. Figure 12 expands in Figure 11 by adding vapor/liquid separations on the streams entering the stripper. The liquid splits enter the stripper at their usual locations, as in Figure 11. The vapor splits enter the bottom of the stripper along with the vapor from the reboiler. Vapor splits can be removed from the design 2531


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Figure 10. Effect of stripper condensate rerouting in the reported design. We report a full implementation of rerouting with XSTR‑CON = 1, although reducing XSTR‑CON linearly reduces performance. This effect is limited in the reported design due to an interaction of stripper condensate rerouting with the distributed cross heat exchanger configuration, which significantly reduces vapor condensate flow rate and thus the effect of rerouting it.

Figure 12. Flowsheet with a full distributed cross heat exchanger configuration. This configuration splits steam generated by the central cross heat exchanger into the bottom of the stripper rather than putting the steam into the top of or midpoint on the stripper.

Figure 11. Flowsheet representation of a CO2 capture unit with a distributed central cross heat exchanger. The central cross heat exchanger has part of its cold stream, the rich solvent stream, extracted partway through and inserted into the top of the stripper. The remainder of the cold stream is heated to completion and then inserted into a midpoint on the stripper.

Table 2. Composition of Vapor Streams Generated by the Reboiler and Distributed Cross Heat Exchanger

when a stream’s vapor fraction is negligible. For example, the final design reported in this Article anticipates a negligible vapor fraction in the stream entering the top of the stripper, so the phase separator on that stream is not part of our final design. However, if that stream does have a significant vapor fraction, we strongly recommend splitting it. Table 2 shows the composition of the vapor streams from the reboiler and the vapor split from the stream entering the midpoint on the stripper. The vapor from the cross heat exchanger contains more CO2 because it comes from a richer solvent, although it is otherwise qualitatively similar to the vapor from the reboiler. Supporting Information Table S4 shows our optimization approach for the distributed cross heat exchanger configuration.

vapor composition (mol %) component

distributed cross heat exchanger

H2O CO2 MEA N2 + O 2

85.7% 13.9% 0.31% 0.02%

reboiler 91.8% 7.9% 0.34% 0.00%

We model the distributed cross heat exchanger in terms of: (1) the split fraction to the distributed side XDHEX; (2) the height of the distributed feed into the stripper HDHEX; (3) the total cross heat exchanger area AHEX available; and (4) the heat exchanger area distributed ADHEX. 2532


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Figure 13. Profiles of the stripper’s internal temperatures and flow rates. The left of the plot is the top of the stripper; the right of the plot is the bottom of the stripper. Temperatures and water vapor both peak toward the right side of the graph because the reboiler is at the right axis.

We can select a total cross heat exchanger area AHEX to be distributed between the central cross heat exchanger CHEX and the distributed cross heat exchanger DHEX. Initial optimizations can often assume that the distributed feed enters the stripper’s packed sections at the midpoint, HDHEX ≅ HSTR/2. We then determine XDHEX, ADHEX, and optionally HDHEX through optimization. If optimization finds that XDHEX = 0, then we do not use the distributed cross heat exchanger configuration. No full implementation of cross heat exchanger distribution has yet appeared in the literature, although Leites et al.28 and Lin et al.8 have both reported partial implementations. Section 5.2 discusses the implementation given by Lin et al. Conceptually, the distributed cross heat exchanger functions by delivering more of the heat to the bottom section of the stripper than to the top section. The cooler top has a lower solvent vapor fraction, which reduces useful heat lost to the condenser. Figure 13 shows the vapor flow rates and temperatures within the column for a simple stripper and one in a simulation using a distributed cross heat exchanger. In the distributed case, the temperature at the top of the column (left axis of the plot) is a mere 54 °C as compared to the 82 °C found in the simple stripper case. This results in an 85% reduction in water vapor reaching the condenser for significant energy savings. The identical values for CO2 lost to the condenser show that the same amount of carbon is captured in both cases. 2.3. Major Process Modifications. Major process modifications involve extensive changes to the system and may have a strong impact on how the system is designed and operates. 2.3.1. Flash Stripper. Figure 14 shows the flash stripper configuration in which the regeneration tower is replaced by a simple flash drum. While this configuration is generally less energy efficient than the standard configuration, it has the advantages of being easier to simulate and potentially cheaper to build.

Figure 14. Flowsheet representation of the flash stripper configuration in which the regeneration tower is replaced with a flash drum.

2.3.2. Multipressure Stripper. Figure 15 shows a multipressure stripper, the idea that separating the stripper into multiple sections of differing pressures will improve the process. The bottom section has the lowest pressure, and vapor is brought from the top of each section to the bottom of the one above it through a compressor. Each stage is considered as its own pressurized section with a compressor on its outgoing vapor to increase pressure by ΔPSTRi. The first found report of the multistripper configuration comes from a series of patents by Rochelle30 in 2004. Rochelle later published analyses of the multipressure stripper configuration.31 Supporting Information Table S5 shows our approach for optimizing a multipressure stripper. Optimization determines all pressure increases ΔPSTRi. If optimization determines that any 2533


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2.3.3. Heat Pump Integration. Heat pumps are used to move thermal energy against a temperature gradient in units like refrigerators and air conditioners. We use a heat pump to move waste heat produced by the CO2 capture unit to its reboiler. For example, we can recover waste heat from a condenser rather than simply cool the condenser with cooling water, and then use the recovered heat to help power the reboiler. 2.3.3.1. Basic Heat Pump Operation. Figure 17 shows a block diagram for an absorption-driven heat pump. The heat pump is composed of six major unit operationsHP: (1) steam generatorHP; (2) absorberHP; (3) condenserHP; (4) evaporatorHP; (5) pumpHP; and (6) cross heat exchangerHP. The heat pump and CO2 capture unit share several names in common, for example, “absorber”, so we will specify blocks belonging to the heat pump with the subscript “HP” to avoid ambiguity. The working solution of the heat pump is aqueous lithium bromide (LiBr), which is like a dense salt water solution. The chemical species in the working solution are H2O, H3O+, OH−, LiBr, Li+, Br−, LiOH, and HBr. This working solution is constantly looping back and forth between the generatorHP and the absorberHP. Part of its water is boiled off in the generatorHP, but this separated water rejoins the working solution in the absorberHP. The generatorHP shown in Figure 16a concentrates the working solution by boiling off some water from the working

Figure 15. Flowsheet representation of a CO2 capture unit with a multipressure stripper. The compressor on the side of the stripper moves vapor from the lower pressure section below to the higher pressure section above.

particular pressure increase ΔPSTRi is zero, then there is no side compressor immediately above stage i. If all pressure increases ΔPSTRi are zero, then we do not use a multipressure stripper.

Figure 16. Four primary unit operations inside of an absorption-driven heat pump. 2534


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Figure 17. Block diagram for an absorption-driven heat pump.

2.3.3.2. Coefficient of Performance (COP). We quantify the effectiveness of the heat pump by its coefficient of performance (COP):

solution, producing concentrated aqueous LiBr and steam. The generated steam goes to the condenser, while the concentrated working solution heads back toward the absorberHP. Eventually, the steam and the concentrated working solution will meet again in the absorberHP, and then be sent back to the generatorHP to be separated again as a new cycle begins. The next major unit operation, the absorberHP shown in Figure 16b, recombines the concentrated working solution with the boiled-off water, producing heat. This recombines the material separated by the generatorHP. The heat produced by reabsorption is used to warm a heating target. We will use it to help power the stripper’s reboiler (see section 2.3.3.4). Physically, the absorberHP sprays the concentrated working solution onto a heating tube inside which is the rich solvent stream to be heated. The steam combines with the concentrated working solution, diluting it, producing heat, and then exiting the absorber. The third major unit operation, the condenserHP shown in Figure 16c, receives steam boiled off from the working solution by the generatorHP. This stream is condensed at high pressure. Condensation releases heat to a heating tube inside which is the rich solvent stream to be heated. The rich solvent stream is first heated by the absorberHP, and now it receives additional heat from the condenserHP. The liquid water leaving the condenserHP is sprayed into the evaporatorHP shown in Figure 16d with a significant pressure drop, for example, from 2.1 to 0.1 bar. Any unvaporized liquid water is boiled by waste heat received from heating tubes containing the waste heat streams such as the vapor streams from the absorber and stripper as shown in Figure 18. The stream generated in the evaporatorHP goes to the absorberHP, where it will recombine with the concentrated working solution. The heat pump is able to recover waste heat through its evaporatorHP, which operates at low temperature. By the conservation of energy, the heat pump must output both that waste heat and the heat it takes from utility steam in the generatorHP through the condenserHP and absorberHP. We select process conditions such that the condenserHP and absorberHP are hot enough to be useful in a given process, for example, to power the CO2-capture unit’s reboiler. This selection usually requires the generatorHP to recover heat at a higher temperature, which requires higher pressure steam.

COP ≡

usefulness of energy received from heat pump usefulness of energy lost to heat pump (9)

which can be approximated as COP =

+ DHP‐ABS DHP‐CON + DHP‐ABS D ≅ HP‐CON DHP‐GEN + ηref WHP‐pump DHP‐GEN (10)

All of the notations in this equation refer to the heat pump (HP), with DHP‑CON being the heat duty of the condenser, DHP‑ABS being the heat duty of the absorber, DHP‑GEN being the duty of the generator, and WHP‑pump being the work for the pump. This approximation neglects the solution pump’s work, WHP‑pump, because its magnitude is generally negligible. Assuming ηref = 0.4 (MWe/MWt) (see section 1.3.3), then we have generally found results bounded by (WHP‑pump/ηref)/DHP‑gen ≤ 0.003, such that the error due to this approximation is less than 0.3%. Alternatively, we can use the economic coefficient of performance (ECOP): ECOP ≡

utility costs mitigated by heat pump utility costs incurred by heat pump

(11)

or ECOP = ζSTR‐REB(DHP‐CON + DHP‐ABS) + ζcooling waterDHP‐VAP ζHP‐GENDHP‐GEN + ζelectricityWHP‐pump (12)

where ζi is the cost of utility for i. We can assume that steam costs are proportional to their equivalent work: ζsteam for i ∝

Ti + ΔTapp − Tturbine Ti + ΔTapp

which is explained in section 1.3.2 in which Ti is the temperature of unit i, ΔTapp is the temperature approach, and Tturbine ≅ 40 °C is the turbine’s condenser temperature typically assumed to be 2535


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Figure 18. CO2-capture process implementing an absorption-driven heat pump to help power its reboiler.

about 40 °C. Using this relationship, we can estimate the ECOP without the need for actual utility costs:

2.3.3.4. Heat Pump Integration. A single-stage, absorptiondriven heat pump can exchange heat from each of its four major components shown in Figure 17. For applications such as ours: (i) The generatorHP will require utility heating from steam. (ii) The evaporatorHP will recover waste heat. (iii) The absorberHP and condenserHP will provide potentially useful heating. Erickson34 listed several of these possibilities. Figure 18 shows our preferred configuration in which: (i) The potentially useful heating provided by the absorberHP and condenserHP is used to help power the stripper’s reboiler. (ii) Waste heat recovered by the evaporatorHP comes from the vapor streams from the top of the absorption tower and regeneration tower. Computationally, we first simulate the CO2-capture process without a heat pump. This provides us with the stripper reboiler duty required if no heat pump is used, DSTR‑REB. From Figure 18, we then know that

ECOP ζSTR‐REBDHP‐CON + ζSTR‐REBDHP‐ABS + ζcooling waterDHP‐VAP ≅ ζSTR‐REBDHP‐GEN 1−

Tturbine TSTR‐REB + ΔTapp

1−

Tturbine THP‐GEN + ΔTapp

×

(13)

2.3.3.3. History. Ferdinand Carré patented the first absorption-driven heat pump in 1860.32 Kasley patented the modern design for a single-stage absorption-driven heat pump in 1924,33 six years before the modern design for a basic absorptiondriven CO2-capture process.3 Growing interest in waste heat recovery led to a flurry of patents and other publications starting in the late 1970s. In 1980, Erickson34 filed a patent for using a single-stage absorptiondriven heat pump to recover waste heat for a distillation column’s reboiler. Also in 1980, Wilkinson and Hanna35 filed a patent for heat recovery using one-, two-, three-, and four-stage absorptiondriven heat pumps, although a three-stage absorption-driven heat pump was patented again in 1993.36 In 2013, Zhang et al. reported an absorption-driven heat pump to a regeneration column in a CO2 capture process like that shown in Figure 2.37 Their design is discussed in section 5.3, although we note that the reported energy savings were just 2.62%, because of their misunderstanding of how to effectively apply heat pump to minimize the energy consumption in a CO2 capture process.

conventional DSTR‐REB = DSTR + DHP‐ABS + DHP‐CON ‐REB

(14)

This equation says that the combination of the heat duty of a conventional reboiler Dconventional and duties of the heat pump STR‑REB absorber and condenser (DHP‑ABS + DHP‑CON) will supply the heat that otherwise would have been provided solely by the stripper reboiler duty required when no heat pump is used, DSTR‑REB. We define the heat pump extent χSTR‑REB as the portion of the total stripper reboiler duty provided by the heat pump such that χSTR‐REB ≡ 2536

DHP‐ABS + DHP‐CON DSTR‐REB

(15) DOI: 10.1021/ie504617w Ind. Eng. Chem. Res. 2015, 54, 2526−2553

Article


Figure 19. Waste heat available to the heat pump. Only a fraction of the total waste heat is recovered for use by the heat pump.

2.3.3.6. Determining Waste Heat Temperature Twaste. Figure 19 shows the fractions of waste heat drawn from each unit. We determine these fractions to maximize the quality of the recovered waste heat. We find that the full waste heat requirement could be met from waste heat above 58.9 °C, so all waste heat below this temperature is unrecovered. This temperature above which waste heat is sufficient is Twaste, defined such that

Ideally, the heat pump is used to provide the entire stripper reboiler duty such that χSTR‑REB = 1. However, this may require more waste heat than is efficiently available from the rest of the process, so optimization may report designs with χSTR‑REB < 1. If optimization reports χSTR‑REB = 0, then we do not use the heat pump at all. Our final design discussed in section 4.2 does have a heat pump extent of unity, χSTR‑REB = 1, so the conventional portion of the reboiler duty Dconventional drops out of the final STR‑REB design. In general, the heating savings ΔEheating due to the heat pump will be ΔE heating = 1 −

DHP‐VAP =

(16)

ΔEcooling =

(17)

In the equation, ∀ represents “for all”, DABS‑CON is the heat duty of the heat pump condenser, DSTR‑CON is the heat duty of the stripper condenser, and DLSC is the heat duty of the lean solvent cooler. In section 4.2, we report cooling savings of 40.9% because the full heat recovery requirement of DHP‐VAP = (COP − 1)DSTR‐REB

∑ ∀ waste heat sources i

Di|T dT

(19)

where Di|T is the available heat duty of waste source i at temperature T. Figure 20 shows the waste heat available from each of the three sources i. DHP‑VAP is the total solid area under the curve above Twaste = 58.9 °C. The waste heat curves shown in Figure 20 are bound on the left by the cooling target for the full-waste heat’s source and on the right by the source’s initial temperature before cooling. For example, the lean solvent enters the lean solvent cooler (LSC) at 51.83 °C and is cooled to 45 °C, so waste heat from the lean solvent cooler is available from 45 to 51.83 °C. Unused waste heat is removed from the process as normal, for example, with utility cooling water. For example, none of the waste heat from the lean solvent cooler is used, so its entire cooling duty is provided from cooling water. We note that strippers usually have a higher condenser duty than that shown in Figures 19 and 20, although our stripper’s condenser duty is much lower due to the distributed cross heat exchanger discussed above in section 2.2.4. This effect of the distributed cross heat exchanger on heat pump implementation is an example of the strong, complex interactions discussed below in section 2.4. 2.3.3.7. Determining Steam Temperature THP‑steam for the Heat Pump’s Generator. In this Article, our technique is to be applied in a 1 (Mtonne/yr) demonstration plant by an Asia Pacific power plant. The turbines in the power plant enable the extraction of multiple levels of steam. We choose steam at 236 °C and 31.2 bar that was readily available. In the general case, we can vary the heat pump’s parameters such as operating pressures, working solution salt concentrations, and heat exchanger sizes to fit the available waste heat temperature Twaste and the available utility steam temperatures THP‑steam for the generatorHP. We can adjust the heat pump’s salt

where the coefficient of performance (COP) is as defined above in section 2.3.3.2. When there is a sufficient amount of waste heat available, maximum energy savings are obtained when χSTR‑REB = 1 such that ΔEheating = 1 − [COP]−1. 2.3.3.5. Cooling Provided by the Heat Pump. The heat pump provides useful cooling because it recovers waste heat from streams that would otherwise have been cooled by utility cooling water. For example, utility cooling water, which would have been needed for the condensers shown in Figure 18, is reduced by the total amount of waste heat recovered through the heat pump vaporizer, DHP‑VAP. In general, cooling savings ΔEcooling due to the heat pump will be DHP‐VAP ∑∀ coolers i |Di| DHP‐VAP = |DABS‐CON| + |DSTR‐CON| + |DLSC|

∞

waste

χSTR‐REB COP

∫T

(18)

is met by 40.9% of the net cooling duty of the CO2-capture process. 2537


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Figure 20. Total waste heat available from the process by temperature. The warmest waste heat, above Twaste, is recovered by the heat pump. In this case, Twaste is 58.9 °C, although in general Twaste is selected such that the total waste heat available above it (the solid area under the curve) is equal to the heat pump’s waste heat requirement.

Figure 21. Comparison between the effectiveness of using semiwaste heat versus full-waste heat to power the heat pump. This early simulation found a COP of approximately 1.8 due to optimistic assumptions for temperature approaches; more recent work would find somewhat (∼12.5%) reduced energy savings for both series.

In the case of full-waste heat, heat consumption is reduced by (1 − (1/COP))DSTR‑REB. This is still true in the case of semiwaste heat; however, the siphoning of semiwaste heat raises DSTR‑REB. Figure 21 shows a comparison between using semiwaste heat and full-waste heat. In this particular case, the parasitic effect of drawing semiwaste heat increases DSTR‑REB by 38.7%. The overlap between semiwaste heat and full-waste heat at low heat pump extents shows that some parasitic effect does not increase DSTR‑REB, so it is likely that the central cross heat exchanger is oversized in that base case. In the nonoverlapping region, increasing the heat pump extent saves only about 42.6% of the energy in the semiwaste heat case as in the full-waste heat case.

concentrations and internal pressures to match available steam temperature and pressure. 2.3.3.8. Recovering Semiwaste Heat from the Stripper’s Bottom Stream. Some waste heat that needs to be removed from a stream can still be somewhat useful to the process. We call this type of waste heat semiwaste heat. Our primary source of semiwaste heat is the freshly regenerated solvent coming out of the bottom of the stripper. While the high temperature of the stream must be corrected before it can enter the absorption tower, making it waste heat, it is still somewhat useful in that a large portion of it will be transferred in the central cross heat exchanger. Unlike full-waste heat, drawing semiwaste heat to power the heat pump comes at a cost to other parts of the system. 2538


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Figure 22. Effect of lean solvent loading on regeneration energy in two different cases. These cases overlap for lower loadings, but diverge above about 0.1725 mol CO2/mol MEA.

incorporates several energy-saving schemes: absorber intercooling; distributed cross heat exchanger; stripper interheating; and heat pump integration. 3.1.1. Solvent. The lean solvent stream is 30 wt % MEA excluding CO2. CO2 is added as to meet the specified lean solvent loading, for example, 0.1825 (mol CO2/mol MEA), for which sensitivity analyses such as that shown in Figure 22 have found to be optimal. While the rich solvent loading is calculated and tends to increase with the absorption tower’s packed height, it is usually close to 0.5 (mol CO2/mol MEA). 3.1.2. Flue Gas. Our dry flue gas typically contains 15 mol % CO2 with the balance being 90 mol % N2 and 10 mol % O2. We then saturate the flue gas to 100% relative humidity because the flue gas is generally washed just before the capture process. We model the flue gas at 40 °C and 1.06 bar, although several authors have pointed out that heat can be recovered from higher temperature flue streams.38,39 The first known reference to recovering heat from the flue gas in a CO2 capture system appears in a patent from 1961.4 3.1.3. Absorption Tower (Absorber). Figure 26 shows a common implementation of the absorption tower. CO2 is captured from the flue gas in the bottom section, beneath the inflow of lean solvent. The upper section washes the treated flue gas to reduce the loss of MEA and other compounds to atmosphere. Figure 27 shows a simplified absorption model in which the top wash section is replaced by a condenser. We model the condenser with a single flash calculation, which significantly reduces simulation time. We model the packed capture section of the absorber with 20 stages. Each stage is rate-based, following the model shown in Figure 28. The liquid film has a geometric discretization such

Our earlier work recovered semiwaste heat when insufficient full-waste heat was available for a full heat pump extent of χHP = 1, although this is unnecessary in the final design reported in section 4.2. 2.4. Energy-Saving Scheme Interactions. Literature presentations of energy-saving schemes often consider each scheme in isolation. While this approach can be useful for formulating ideas, our more computationally demanding approach is necessary as energy-saving schemes have strong and complex interactions. We present examples below. 2.4.1. Shifts of Optimal Lean Solvent Loading. The effectiveness of tactics as trivial as parameter optimization varies depending on the rest of the process. Figure 22 shows how, depending on the process, the optimal lean solvent loading can vary. Lin et al. report an optimal lean solvent loading for a similar process at 0.38 (mol CO2/mol MEA), discussed below in section 5.8 2.4.2. Effect of Flue Gas Composition. Figure 23 shows the effect of capture rate on the same CO2 capture process operating on flue streams with a minor difference in CO2 concentration. In both cases, the regeneration energy increases with capture rate, but at different rates when the CO2 concentration is varied. 2.4.3. Effect of Stripper Pressure. Generally, the regeneration column is under pressure. Figure 24 illustrates how the stripper pressure affects the thermal regeneration energy in three different cases. Lin et al.8 recommend pressures ranging from 3.0 to 3.3 bar for their energy-saving regeneration design using rich solvent bypass.

3. SIMULATION APPROACH 3.1. Flowsheet Description. Figure 2 shows our basic flowsheet. Figure 25 shows a more complete flowsheet, which 2539


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Figure 23. Effect of capture rate on normalized regeneration energy for two otherwise identical cases with different flue gas CO2 concentrations.

Figure 24. Effect of stripper pressure on regeneration energy.

that the thinnest section of liquid film borders the vapor film and each section of film is 5 times thicker than the prior section of film.

The rate-based model’s bulk liquid and bulk vapor phases are at their own individual equilibria except that their kinetic reaction rates are finite. Additionally, mass and heat transport between the 2540


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Figure 25. Flowsheet incorporating multiple energy-saving schemes.

Figure 27. Simplified absorption tower model. Figure 26. Common absorption tower implementation.

Our models use rate-based stages instead. A rate-based approach calculates mass and heat transfer through the liquid and vapor films rather than assuming instantaneous transmission between the liquid and vapor. Ideally, a rate-based model should employ infinitely many stages, although in practice, the number of stages is limited by available computational power. Figure 29 shows the effect of stage count on the absorber’s profile. We use 20 rate-based stages per column, packed with 30 m of Mellapak 750Y with a diameter of 1.32 m. The packing diameter is determined during each simulation such that the fractional approach to flooding is 70%. 3.1.4. Regeneration Column (Stripper). We model the regeneration column using 20 rate-based stages, a flash calculation for the condenser, and a flash calculation for the reboiler. We choose the reboiler heat duty such that the amount of CO2

liquid and vapor phases is mediated by film, so the bulk phases are not at equilibrium with each other despite being at their own individual equilibria. Kale et al.40 discuss the rate-based approach to modeling in detail. Traditional equilibrium-based stages assume that all reactions are at equilibrium in every section of the column, and the effective size of an equilibrium-based column is Hpacking =

∑ Nstages

HETPi = Nstages × HETPaverage (20)

where HETPi is the height equivalent to a theoretical plate (stage) i and Nstages is the number of equilibrium stages. 2541


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Figure 28. Comparison of the equilibrium model (top) and rate-based model (bottom) for reactive distillation column stages.

3.2. Scaling. We employ scale-invariant models. All simulations are designed to capture exactly 1 (tonne CO2/h), so we report our results in terms of the scaling constant: kscale ≡

2

1 (tonne CO2 )/h

(22)

All material flow rates, heat duties, and heat exchanger surface areas are scaled by kscale. Column packing diameters are scaled by (k scale ) 1/2 because their crosssectional areas scale with kscale. All other parameters, such as temperatures, pressures, compositions, split fractions, and column heights, are scale-invariant. This scaling process is not an estimation method, but rather mathematically identical because of the underlying assumptions in the physical models. 3.3. Simulation and Optimization Approach for Baseline Flowsheet. Table 3 shows our simulation and optimization approach for the baseline flowsheet. In actual applications, we add other energy-saving schemes’ variables and design specifications to this baseline table. 3.4. Heat Pump Optimization Approach. We optimize our heat pump using sequential quadratic programming (SQP) for a combination of maximizing the coefficient of performance (COP from section 2.3.3.2) and maximizing the three temperature approaches of the interfaces between the heat pump and CO2 capture unit. We seek to maximize

Figure 29. CO2 flow rates inside a rate-based absorption tower. Higher stage counts lead to greater accuracy although they require more time to compute.

released from the top of the condenser is equal to the amount to be captured, R × Fflue CO2, or 1 (tonne CO2/h) by scaling convention (see section 3.2). The regeneration column has a packed height of 10 m, and its diameter is selected such that the maximum approach to flooding is 70%. 3.1.5. Heat Exchangers. We describe cross heat exchangers by their heat exchanging surface area, Ai, where i is the heat exchanger. We avoid specifications such as log-mean temperature difference (LMTD) because they overestimate the relative performance of designs with higher solvent flow rates or, equivalently, a larger loading difference ΔL: ΔL ≡ Lrich − L lean

Fcaptured CO

-HP = -HP1Dwaste − -HP2

(23)

where Dwaste is the corrected recovered waste heat and both -HP1 and -HP2 are optimization functions. The corrected recovered waste heat is the total waste heat consumed by the heat pump:

(21)

where Lrich is the loading of the rich solvent in (mol CO2/mol MEA) and Llean is the loading of the lean solvent in (mol CO2/mol MEA). For example, doubling the solvent flow rate (or halving ΔL) implicitly doubles the cross heat exchanger size when LMTD is specified. In some cases, this might lead to misleading results; for example, an optimizer might prefer a design with twice the flow rate for a 5% energy savings, when using the same double-sized heat exchanger with a lower solvent flow rate may have saved 10%.

Dwaste =

∑ ∀ waste sources i

|Di| + β

∑ ∀ semi‐waste sources i

|Di|

(24)

except semiwaste heat is weighted by the semiwaste heat penalty β to account for the decreased desirability of using semiwaste heat. We typically use β ≅ 1/3 following from the observation that semiwaste heat recovery tends to be 1/3 as desirable as fullwaste heat recovery. 2542


optimization

stripper packed diameter temperature at top of absorber temperature of stripper condenser lean solvent loading rich solvent loading central cross heat exchanger heat exchanging surface area central cross heat exchanger temperature approach stripper bottom pressure absorber packing height stripper packing height lean solvent cooler temperature

Lrich

ACHEX ΔTCHEX PSTR HABS HSTR TLSC

6

7

9 10

8

5

4

5−100 5 1.0−2.0 5−40 5−30 40−55

0.35−0.55

0.55 40 40 0.15−0.25

absorber packed diameter stripper flooding factor

3

70−75 1.10 70−75

absorber flooding factor

2

design practices

range 80−95 55 0.75

description CO2 capture rate lean solvent flow rate stripper reboiler duty

symbol

R Flean DSTR‑REB

- ABS dABS -STR dSTR TABS‑TOP TSTR‑CON Llean

1

parameter set

emission target for plant

selection criteria

parameters

Table 3. Process Parameters Needed To Specify the Baseline Flowsheet

1.5 30 25 50

100

0.50

40 40 0.18

70

70

80

suggested

m2 × kscale K bar m m °C

(mol CO2)/(mol MEA)

m × kscale1/2 °C °C (mol CO2)/(mol MEA)

taller absorbers tend to lead to significantly improved energy efficiency; taller strippers tend to lead to slightly improved energy efficiency

if Llean is too low, the stripper will have to spend much more energy per mole of CO2 stripped taller absorption towers will tend to yield higher Lrich, which generally leads to greater energy efficiency as the loosely bound CO2 is easier to strip along with having a tall absorber, increasing ACHEX tends to be one of the easiest ways to increase energy efficiency

if the flooding factors are too high, flooding will occur higher flooding factors will allow for smaller packed diameters

% m × kscale1/2 %

notes design specification determines Flean such that the effective capture rate in the absorber 9ABS ABS is R; design specification determines DSTR‑REB such that the effective capture rate in the stripper 9 STR is R if the flooding factors are too low, weeping will occur

units % (kmol/h) × kscale MW × kscale

typical values

Industrial & Engineering Chemistry Research Article

2543


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Industrial & Engineering Chemistry Research We multiply Dwaste by the first optimization factor: -HP1 ≡

∏

3.6. Physical Models. We use the physical models described in AspenTech’s documentation for their MEA-based modeling package.41 AspenTech’s package employs the electrolytic nonrandom two-liquid (ENRTL) model,42−44 which extends the earlier nonrandom two-liquid (NRTL) model45 to handle solutions containing ions. We use true electrolyte components (including ions) in our simulation. For our stream tables below, we report on the molar flow rate and mole fractions of apparent components, which represent properties within Aspen Plus’ electrolyte simulation results. Zhang et al. explain rate-based modeling of CO2 capture in detail.46 Vivier et al. independently examine use of the NRTL model for aqueous MEA with CO2 through use of differential evolution algorithms.47 Kamalpour and Mehablia conduct a similar study for MDEA using the fuller ENRTL model.48 AspenTech regressed the MEA model parameters to fit the pilot plant data reported by Notz et al.49 3.7. Compressor Work Correlation. Generally, captured CO2 must be compressed above 100 bar after capture. Compression work is significant and can vary with the pressure at which CO2 is captured. For example, a stripper operating at 2.5 bar might require about 22% less in postcapture compression energy than a stripper operating at 1.0 bar. Wagener7 reports a correlation for calculating the equivalent work (section 1.3.2) necessary to compress captured CO2 to 150 bar. This correlation assumes a train of compressors, each with (i) a compression ratio, * ≡ (Vin /Vout), no greater than 2; (ii) a polytropic efficiency of 72%; and (iii) an intercooler exiting at 40 °C after each compressor. The correlation is

exp(− exp(η1 − ΔTi , j))

∀ heat exchangers i

(25)

∀ heat exchanger sides j

where η1 = −1 K is a shift factor. -HP1 is essentially a series of factors for each temperature approach between the heat pump and CO2 capture unit. Each factor is between zero and one. Very small temperature approaches have the optimization factor near zero, so the optimization function is strongly incentivized to improve small temperature approaches. Very large temperature approaches have an optimization factor near one, so while the optimizer is still somewhat rewarded for making large temperature approaches even larger, it will tend to focus on improving smaller temperature approaches or the COP instead. In some cases, we have simulations that would initialize the optimization problem with strongly negative temperature approaches. While very negative temperature approaches do not present a mathematical problem, limits on computer precision caused -HP = -HP1Dwaste to be so close to zero that the optimizer could not figure out what to do, causing optimization to fail. To correct for this, we introduce the second optimization factor: -HP2 ≡

∑

exp(η2 − ΔTi , j)

∀ heat exchangers i

(26)

∀ heat exchanger sides j

where η2 = −17.5 K is a second shift factor, which is only relevant for negative temperature approaches. This second optimization factor -HP2 subtracts a large amount from the objective function for each negative temperature approach, guiding the optimizer toward realistic, physical solutions. 3.5. Flowsheet Convergence Considerations. We build our flowsheet using an object-oriented approach in which more complex units are derived from simpler ones. We model the absorption tower as a rate-based distillation packed section with a condenser, and set its packed diameter such that the closest approach to flooding is 70%. We determine the lean solvent flow rate such that the effective capture rate in the absorption tower, 9ABS ≡ 1 − (Flost CO2 /Fflue CO2), meets the desired capture rate R. The base object for the absorption tower is the packed section, and each additional feature is a step on the inheritance tree. We also model the regeneration tower as a rate-based distillation packed section with a condenser, although it also includes a reboiler. As with the absorption tower, we set the regeneration tower’s packing diameter such that the closest approach to flooding is 70%. Its effective capture rate,

(

⎧ ⎛P ⎞ ⎪ 0.4275 − 0.1039 × ln⎜ in ⎟ ⎝ bar ⎠ if Wcompression ⎪ =⎨ GJe /tonne CO2 ⎛ Pin ⎞ if ⎪ ⎪ 0.4085 − 0.0914 × ln⎜⎝ bar ⎟⎠ ⎩

Pin > 4.56 bar

(27)

where Wcompression is the equivalent work of the CO2 compression train and Pin is the pressure of the CO2 before it enters the compression train.

4. SIMULATION RESULTS 4.1. Column Packing Effectiveness. Figure 30 shows the results for an absorption column with various packings with correlation data in Aspen Plus. Packings are characterized by the lean solvent flow rate necessary to achieve the target capture rate R and the column diameter necessary to achieve the target maximum approach to flooding -target when used to pack the absorption tower. 4.2. Optimized Flowsheet Results. We find an optimized solution with a thermal regeneration energy Eregen of 1.67 t (GJt/tonne CO2 captured) with a heat pump COP of 1.71, or 2.86 (GJt/tonne CO2 captured) without the heat pump. We compare these results to other reported thermal regeneration energies below in section 5.1. The central cross heat exchanger area ACHEX was found to be 35 k × m2 with an LMTD ≈ 9.56 K, and the distributed cross heat exchanger area ADHEX was found to be 100 k × m2 with an LMTD ≈ 4.43 K. Figure 31 shows the initial flowsheet for our complete CO2 capture system excluding heat pumps. This design implements: (i) absorber intercooling; (ii) stripper interheating;

)

9 STR ≡ Fcaptured CO /Fflue CO2 , is set to the desired capture rate 2

Pin ≤ 4.56 bar

R by varying the reboiler duty DSTR‑REB. As with the absorption tower, the packed section is the base object, and each additional feature is a step on the inheritance tree. Our design employs a single-stage, absorption-driven heat pump. Its low pressure PHP‑low and high pressure PHP‑high are set such that the absorberHP and condenserHP can provide useful heating to the stripper’s reboiler, while the evaporatorHP can recover waste heat and, if necessary, semiwaste heat from the CO2 capture process. We determine its flow rate such that the heat pump’s useful heating meets the stripper’s reboiler duty, or DHP‑ABS + DHP‑CON = DSTR‑REB. The details will appear in future work. 2544


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Figure 30. Analysis of column packings for use in the absorption column. Koch’s Flexipac 700Y results in a near-minimum solvent flow rate with a smaller column size.

Figure 31. Complete flowsheet for a CO2 capture unit implementing many energy-saving methods. This flowsheet does not include a heat pump.

condensed. We can see both phenomena in Figure 33, which shows the vapor composition profile. We can see that most of the CO2 absorption happens at the top of the column, which causes the high temperature there. Water vapor evaporates more at the top as well, but water continues to evaporate at a meaningful rate throughout the column, cooling the liquid as it descends. This dynamic relies on the flue gas being saturated with water. Experiments and simulations that use a dry or partially saturated flue gas will fail to account for the full effect of water’s condensation. Industrial plants that do not wash their flue gas are likely to have a very different temperature profile and thus a different result for absorber intercooling. Finally, we simplify vapor condensate rerouting by always using its full extent, XSTR‑CON = 1, meaning that the entire condensate is redirected to the top of the absorber. This simplification is suggested by our early sensitivity analyses, which reveal that full vapor condensate redirection is generally helpful

(iii) multipressure stripping; (iv) vapor recompression; (v) distributed cross heat exchanger; and (vi) stripper condensate redirection. Early in the optimization process, we reduce the complexity of the complete flowsheet shown in Figure 31. First, we decide to focus attention on designs excluding additional compressors. This eliminates both multipressure stripping and vapor recompression. In our flowsheet, this is ΔPSTRi = 0∀i and ΔPSTR‑REC = 0. Next, we note that absorber intercooling is ineffective in our examined cases. As discussed in section 2.2.1, absorber intercooling is intended to reduce the solvent temperature in the lower part of the absorption tower to shift the equilibrium to favor further absorption. However, as shown in Figure 32, the bulk liquid phase inside of our absorption tower tends to reach approximately 40 °C without an intercooler. The temperature profile shown in Figure 32 results from heating when CO2 is absorbed and cooling when H2O is 2545


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in our examined cases; by removing an optimization parameter, we greatly simplify the computational burden, because the stripper’s condenser is entirely downstream of the rest of the stripper rather than being in a closed loop with it. Figure 34 shows the new flowsheet, which implements these specifications: ΔPSTRi = 0∀i; ΔPSTR‑REC = 0; DABS‑IC = 0; and XSTR‑CON = 1. Figure 35 shows the simplified flowsheet given in Figure 34 with the addition of an absorption-driven heat pump augmenting the reboiler using waste heat from the condensers. Our simulation approach first solves for the flowsheet without the heat pump, as shown in Figure 34, before solving for the flowsheet with a heat pump shown in Figure 35. We report both solutions. Figure 36 and Table 4 show our optimized results for the CO2 capture process without a heat pump. These stream results remain exactly the same when the heat pump is applied (discussed in section 2.3.3) as the heat pump integration essentially modifies how the process heaters and coolers operate, but not their effect on the process streams. Figure 37 shows our optimized process design for the heat pump integration. The heat pump’s internal streams contain only aqueous lithium bromide (LiBr). The regenerator’s liquid bottom stream represents an external stream for the heat pump because the heat pump is operating as the stripper’s reboiler. Table 5 shows the stream data for the heat pump’s internal streams. In general, there are only two pressures, three flow rates, and three LiBr mass fractions. We determine the two heat pump pressures by optimization to be Plow ≅ 0.129 bar and Phigh ≅ 2.141 bar. Table 6 shows the heat duties for units in the optimized design. If we use the heat pump, then the absorber’s condenser and lean solvent cooler both have part of their cooling provided by the heat pump as it siphons off waste heat. If we do not use the heat pump, then absorber’s condenser and lean solvent cooler draw their total cooling from utility cooling water. The coefficient of performance (COP) discussed in section 2.3.3.2 is 1.71, and the economic COP (ECOP) is 1.83. Figure 38 shows the flow of useful heat through the process when the heat pump is used. Both waste heat and useful heat enter the heat pump, but the entire amount of heat is converted to useful heating for the reboiler.

Figure 32. Temperature profile for the bulk liquid phase in the absorption tower. The top of the absorber reaches nearly 70 °C, while the bottom is close to the flue gas temperature of 40 °C.

Figure 33. Vapor mass flow rate profiles for the vapor phase in the absorption tower. The top of the absorber releases 0.25 tonne/h of CO2 due to the 80% capture rate and our scaling convention, which has 1 tonne/h of CO2 captured.

Figure 34. Simplified flowsheet used for the later stages in our optimization process. Absorber intercooling, multipressure stripping, and vapor compression have been removed. Stripper condensate rerouting has been simplified to its complete implementation. 2546


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Figure 35. Simplified flowsheet with a hybrid reboiler powered by an absorption-driven heat pump reboiler and a conventional heater.

Figure 36. Flowsheet with stream table labels.

paying attention only to the relative performance trends, not to the exact values. For example, Lin et al.9 use the so-called “Phoenix” thermodynamic model reported in July 2012, which includes some changes to the standard electrolyte NRTL (ENRTL) model. By contrast, our work uses the updated ENRTL thermodynamic property model in Aspen Plus V8.5 from Aspen Tech41 reported in November 2013. Additionally, some authors do not report full closed-loop flowsheets, but rather either just the absorber or just the stripper. We discuss several specific cases below. 5.2. Rich Solvent Bypass Scheme, Lin et al. Recently, Lin et al.8 have presented a rich solvent bypass scheme largely equivalent to the distributed cross heat exchanger scheme shown in Figure 11. Here, we compare their work to the designs presented in this Article.

We can view the coefficient of performance as a size-up factor for useful heating. In this case, the useful heat provided to the heat pump’s generator, DHP‑GEN ≅ 0.46 MW, is multiplied by the COP ≅ 1.71 to provide the necessary heat DSTR‑REB ≅ 0.79 MW to the reboiler.

5. COMPARISONS TO OTHER WORK 5.1. Survey of Reported Thermal Regeneration Energies. Our focus has been on reducing the thermal regeneration energy Eregen presented in section 1.3.1. Figure 39 presents a t comparison of other reported values of Eregen throughout the t literature. Authors vary widely in their modeling approaches, thermodynamic property methods, choice of flue gas, choice of solvent, etc. We should view the comparison in Figure 39 with caution, 2547


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Industrial & Engineering Chemistry Research Table 4. Stream Table for optimized CO2 capture design flow rates mole (kmol/h) stream

temp (°C)

pressure (bar)

vapor mass fraction

mass (tonne/h)

total

H2O

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23

40.00 40.00 40.00 42.91 40.00 69.78 42.48 77.29 67.48 67.48 67.48 108.70 67.48 108.70 109.01 110.20 110.20 110.20 77.29 51.83 45.00 70.20 40.00

1.06 1.06 1.40 1.06 1.06 1.06 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.40 1.06 1.40 1.40

1.00 1.00 1.00 0.00 0.00 1.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00 1.00 0.00 1.00 0.00 0.00 0.00 0.00 0.00 1.00 0.00

6.09 5.05 1.02 15.12 0.88 5.94 15.12 14.02 12.10 3.02 12.10 2.45 0.00 0.58 15.25 1.23 14.02 14.02 14.02 14.02 14.09 1.11 0.08

203.90 179.33 24.06 594.99 48.80 228.17 594.94 589.13 476.00 119.00 476.00 96.19 0.02 26.48 645.30 60.96 589.14 589.14 589.13 589.13 593.14 28.52 4.47

14.40 12.65 1.30 528.33 48.77 61.42 528.33 522.57 422.66 105.67 422.66 82.96 22.71 578.51 55.94 522.57 522.57 522.57 522.57 526.58 5.76 4.46

MEA

66.56 0.03 0.03 66.56 66.56 53.25 13.31 53.25 13.23 0.08 66.77 0.21 66.56 66.56 66.56 66.56 66.56

CO2

SO2

N2

O2

28.40 5.68 22.73 34.87

0.15 0.08

144.85 144.83 0.03 0.03

16.09 16.09 0.01 0.01

144.83 0.03

16.09 0.01

5.68 34.87 12.14 27.89 6.97 27.89 3.28 0.01 3.69 16.95 4.81 12.14 12.14 12.14 12.15 12.15 22.73

0.07 0.03 0.11 0.07 0.07 0.05 0.01 0.05 0.01

0.02 0.02 0.01

0.07 0.07 0.07 0.07 0.07 0.03

0.01

Figure 37. Flowsheet for optimized heat pump results.

We suggest putting the vapor condensate into the lean solvent cooler. Lin et al. disregard the vapor condensate as a waste stream, although this may be because their models do not include the absorption tower or its lean solvent cooler. 5.2.2. Distributing the Central Cross Heat Exchanger. The rich solvent bypass scheme distributes the central cross heat exchanger as does the distributed cross heat exchanger configuration presented in Figure 11.

Lin et al. describe various regeneration columns assuming a rich solvent input and a lean solvent output. They use the “Phoenix” property model described by Plaza50 for their MEAbased simulations. 5.2.1. Vapor Condensate. We can deal with the vapor condensate in one of three ways: (1) Return it to the regeneration column as in Figure 2. (2) Put it into the lean solvent cooler as in Figure 9. (3) Have it exit the process as a waste stream. 2548


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Industrial & Engineering Chemistry Research Table 5. Stream Table for Optimized CO2 Capture Design’s Heat Pumpa stream

temp (°C)

pressure (bar)

vapor mass fraction

mass (tonne/h)

LiBr mass fraction

1 2 3 4 5 6 7 8 9 10 11

117.50 117.65 212.47 231.14 122.65 122.65 131.22 231.14 122.50 50.86 54.89

0.129 2.141 2.141 2.141 2.141 0.129 0.129 2.141 2.141 0.129 0.129

0.000 0.000 0.009 0.000 0.000 0.000 0.098 1.000 0.000 0.127 1.000

8.22 8.22 8.22 7.65 7.65 7.65 8.22 0.57 0.57 0.57 0.57

67.4% 67.4% 67.0% 72.3% 72.5% 72.5% 67.4% 0.0% 0.0% 0.0% 0.0%

a

Table 6. Heat Duties for the Optimized CO2 Capture Process heat exchanger generatorHP (HP-GEN) condenserHP (HP-CON) absorberHP (HP-ABS) evaporatorHP (HP-VAP) absorber’s condenser (ABS-CON)

stripper’s condenser (STR-CON) stripper’s reboiler (STR-REB) lean solvent cooler (LSC)

heat duty (MW)

from heat pumpa from cooling water total

from heat pumpa from cooling water total

0.465 −0.381 −0.413 0.329 −0.311 −0.336 −0.647 −0.063 0.795 −0.028 −0.066 −0.094

a

Cooling provided by the heat pump if the heat pump is implemented. If not, the total cooling duty is provided by utility cooling water.

Internal heat pump streams only.

5.2.3. Solvent Concentration. We use 30 wt % MEA while they use 9 m MEA (∼35.5 wt %). The higher MEA concentration is energetically favorable, although it can be more difficult to work with in practice, because of increased viscosity and corrosiveness. Literature reports51 have presented inhibitors, which could potentially allow up to 40 wt % MEA, although we are unaware of any demonstrations of this technology. 5.2.4. Regeneration Tower Pressure. We find that the optimal regeneration tower tends to vary with the exact design as shown in Figure 24. Here, we present a design employing 1.4 bar, which is approximately optimal for the full design incorporating multiple energy-saving schemes. Lin et al. present designs using at least 3.0 bar for their designs based on the distributed cross heat exchanger.

Higher regeneration tower pressures require a more expensive column capable of pressurization and raise the operating temperature of the reboiler, although the higher pressure captured CO2 stream requires less compression. This reduced need for compression results in a lower capital cost and electrical requirement for the compressor. The Wagener correlation (section 3.7) predicts that the Lin et al. design saves 20.2% on compression energy as compared to our reported design by operating the stripper at 3.0 bar as opposed to 1.4 bar. 5.2.5. Lean Solvent Loading. We use a lean solvent loading of about 0.1825 (mol CO2/mol MEA) following from the data shown in Figure 22. Lin et al. use 0.38 (mol CO2/mol MEA), based on their analysis that shows this lean solvent loading to be optimal for a stripper pressure of 3.0 bar as shown in Figure 10 of

Figure 38. Both waste heat and useful heat flow through the heat pump to become useful heating for the stripper’s reboiler. 2549


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Figure 39. Comparison of thermal regeneration energies from literature.

their recent publication.8 Because Lin et al. recommend a smaller loading difference, they require a far higher solvent flow rate to achieve the same amount of carbon capture. The solvent flow rate increases by a factor equal to the ratio of the working capacities. This higher solvent flow rate requires proportionally larger heat exchangers. If we compare using the same solvent concentration, then the necessary solvent flow rate increases to

according to Van Wagener’s correlation for compressor efficiency,7 results in an equivalent work of 0.313 (GJe/tonne CO2). We report a regeneration column operating at 1.4 bar, which, according to these same correlations, results in an equivalent work of 0.3933 (GJe/tonne CO2). Without the heat pump, our thermal regeneration is approximately 2.86 GJt/tonne with a reboiler operating temperature of 110 °C; by contrast, Lin et al. report 2.96 GJt/tonne with a reboiler operating temperature of 120 °C. 5.2.7.2. With Heat Pump. Lin et al.8 do not include a heat pump in their design, although use of one in our design reduces thermal regeneration energy to 1.67 GJt/tonne. Their 10 °C increase in reboiler temperature would make it difficult for an aqueous LiBr heat pump to deliver waste heat from the absorberHP and reduce overall heat pump efficiency because of the wider difference between waste heat and delivery temperature. 5.2.7.3. Overall. Lin et al. report a flowsheet based on the distributed cross heat exchanger design employing a higher solvent concentration and regeneration tower pressure. This design results in doubling the solvent flow rate, raising the reboiler’s operating temperature by 10 °C, increased solvent degradation, and an absorption tower size, which is smaller at lower capture rates while larger at higher capture rates. We have been unable to directly compare heat exchanger sizing as we cannot meet their LMTD specifications with single-shell heat exchangers. 5.3. Absorption-Driven Heat Pump, Zhang et al. Recently, Zhang et al.37 reported an absorption-driven heat pump augmenting the reboiler in a CO2-capture process. They focus on using heat from the heat pump’s absorberHP alone. Their waste heat comes from both cooling water through the condenserHP and utility steam condensate through the evaporatorHP and generatorHP. Zhang et al. report a heat reduction of just 2.62% in moving from 3.873 to 3.772 (GJ/tonne CO2). This design also increases utility cooling water consumption.

0.5239 (mol CO2/mol MEA) − 0.1825 (mol CO2/mol MEA) = 284.5% 0.50 (mol CO2 /mol MEA) − 0.38 (mol CO2 /mol MEA)

(28)

Because Lin et al. report their design based on a higher solvent concentration, the reported flow rate would only be 284.5% ×

30.0% ≅ 240.4% 35.5%

(29)

of what we report. Doubling the solvent flow rate requires doubling the size of cross heat exchangers to achieve the same temperature approaches. Lin et al. do not penalize their higher solvent flow rate models for the larger cross heat exchangers, which they would require despite the fact that the rejected, lower solvent flow rate models may perform better given the same capital investment. 5.2.6. Absorption Tower Packed Height. We report an absorption column packing height of 30 m for the absorber, although we can readily modify this with consequences like those shown in Figure 4. For example, decreasing the absorption tower’s packed height by 10−20 m would increase regeneration energy by about 2.7%. Our packing cross section area is 1.18 m2/tonne. Lin et al.8 do not simulate an absorption tower, but rather they simply specify lean and rich solvent loadings. In our reproduction of their work, we find that an absorption tower would need to have a packed height of 15.6 m for 80% capture or 22.5 m for 90% capture. Their cross section areas are 1.39 m2/tonne for 80% capture and 1.52 m2/tonne for 90% capture. In terms of volume, their design requires a 27.4% smaller absorption tower for 80% capture, but 24.4% larger absorption tower for 90% capture. 5.2.7. Regeneration Energy. 5.2.7.1. Without Heat Pump. Lin et al. report a regeneration column operating at 3 bar, which,

6. CONCLUSIONS As discussed in section 1.3, the commonly reported metric of energy penalty can vary widely depending on the assumed values for emission rate, turbine efficiency, and carbon capture rate. Equivalent energy is an improvement, although it still assumes a 2550



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case-specific turbine efficiency based on an arbitrary efficiency of the Rankine cycle. On the basis of our analysis, we recommend that CO2 capture units should be evaluated in terms of their and electrical regeneration thermal regeneration energy Eregen t . energy Eregen e We employ a production-scale-independent approach as detailed in section 3.2. Our convergence scheme employs an object-oriented approach as discussed in section 3.5. We have reviewed many simple and major process modifications including: absorber intercooling, stripper interheating, stripper condensate rerouting, distributed cross heat exchanger, flash stripper, multipressure stripper, and heat pump integration. In section 2.4, we present examples demonstrating that process modifications must be considered together because of strong and complex interactions. For example, the high pressure stripper recommended by Lin et al.8 would result in a higher reboiler temperature, which would significantly reduce the effectiveness of a heat pump integration. We have presented a highly effective approach to heat pump integration, finding heating savings of ΔE = 1 − (1/COP) ≅ 41.2%, in contrast to the 2.62% reported by Zhang et al.37 The latter study actually increases the cooling utility requirement by using cooling water as a source of waste heat, although our application reduces cooling water consumption by drawing waste heat from areas in the process that would have otherwise required utility cooling water. We present a CO2-capture unit design with a thermal regeneration energy of 1.67 (GJt/tonne CO2) employing several energy-saving schemes: (i) distributed cross heat exchanger; (ii) stripper vapor condensate rerouting; and (iii) heat pump integration. 6.1. Optimized Distributed Cross Heat Exchanger. The distributed cross heat exchanger is optimized when 80% of the rich solvent flow rate is distributed along with 74% of the cross heat-exchanging surface area. This energy-saving scheme results in significantly reduced solvent loss as vapor to the condenser, conserving energy for solvent regeneration. 6.2. Stripper Vapor Condensate Rerouting. The stripper’s vapor condensate is optimized when fully redirected to the lean solvent cooler. An additional benefit of this configuration is the computational efficiency in that it removes an inner convergence loop from the regeneration column. 6.3. Heat Pump Integration. We report use of an absorption-driven heat pump to power the stripper’s reboiler, resulting in heating savings of 1 − (1/COP). Waste heat is primarily recovered from vapor leaving the top of the absorption tower, although we have also explored using semiwaste heat from the regeneration tower’s liquid bottom stream.

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Article

AUTHOR INFORMATION

Corresponding Author

*Tel.: (540) 231-7800. Fax: (540) 231-5022. E-mail: design@ vt.edu. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This Article is an invited presentation in the symposium, “How Computing Has Changed Chemical Engineering − Session in Honor of Professor Larry Evans’ 80th Birthday”, at the AIChE Annual Meeting, Atlanta, GA, November (2014). We are most thankful to Professor Evans for his strong support and encouragement over the years. We thank Aspen Technology, China Petroleum and Chemical Corp. (SINOPEC), BAE Systems, and Mid-Atlantic Technology, Research and Innovation Center for supporting our educational programs in computer-aided design and process system engineering at Virginia Tech. We are grateful to Professor Gary T. Rochelle of the University of Texas at Austin for his valuable comments on our manuscript. We thank Andy Lui, Ashok Bhakta and Selim Anavi of Aspen Technology, Inc. for their reviews of the manuscript.

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SYMBOLS the “for all” symbol heat exchanging surface area of heat exchanger i [kscale × m2] Di heat duty of unit operation i [kscale × MWt] Di|T incremental heat duty of unit operation i at temperature T [kscale × (MWt/K)] DABS‑IC cooling duty of the only absorber intercooler in a single intercooler configuration [kscale × MWt] DABS‑ICi cooling duty of absorber intercooler i [kscale × MWt] DSTR‑REB heat duty of the stripper’s reboiler [kscale × MWt] Eregen electric regeneration energy [kscale × GJe] e Eregen thermal regeneration energy [kscale × GJt] t Fi total flow rate of stream i [kscale × (tonne/h)] Fij flow rate of component j in stream i [kscale × (tonne/h)] Hi total packed height in unit operation i [m] kscale scaling constant as ratio of plant size to simulation size [−] Keq equilibrium constant [−] Li solvent loading for stream i [mol CO2/mol MEA] LMTD log-mean temperature difference [K] Ni number of stages in packed unit operation i [stage] P pressure [bar] Pi pressure of flow sheet element i [bar] Q heat flow [kscale × (GJt/h)] R capture rate [−] effective capture rate in the absorber [−] 9ABS effective capture rate in the stripper [−] 9 STR Ti temperature of i [K] Weq total equivalent work for a process [MWeq/tonne CO2] equivalent work contribution of unit operation Weqi i [MWeq/tonne CO2] xi liquid-phase mole fraction of chemical species i [mol i/mol solution] Xi material split fraction to unit operation i [tonne i/tonne total] ∀ Ai

ASSOCIATED CONTENT

S Supporting Information *

Table S1: Process parameters needed to specify the absorber intercooling energy-saving scheme. Table S2: Process parameters needed to specify the stripper interheating energy-saving scheme. Table S3: Process parameters needed to specify the stripper condensation rerouting energy-saving scheme. Table S4: Process parameters needed to specify the distributed cross heat exchanger. Table S5: Process parameters needed to specify the multipressure stripper energy-saving scheme. This material is available free of charge via the Internet at http://pubs.acs.org. 2551


Article

Industrial & Engineering Chemistry Research yi XSTR‑CON εCarnot γi

the 7th International Conference on Greenhouse Gas Control Technologies (GHGT-7), 2004; p 9. (14) Geisbrecht, R.; Dipietro, P. Evaluating Options for U.S. Coal Fired Power Plants in the Face of Uncertainties and Greenhouse Gas Caps: The Economics of Refurbishing, Retrofitting, and Repowering. GHGT-9 2009. (15) E. E. Agency. Air Pollution from Electricity-Generating Large Combustion Plants, 2008. (16) U.S. EPA Clean Energy: Coal; http://www.epagov/cleanenergy/ energy-and-you/affect/coal.html (accessed Sep. 10). (17) Ragatz, E. G. Method for the Absorption of Gases. U.S. Patent US1987267, 1935. (18) William B. Borst, J. Absorption Process. U.S. Patent US1987267, 1967. (19) Huval, M.; Venne, H. v. d. DGA Proves out as a Low Pressure Gas Sweetener in Saudi Arabia. Oil Gas J. 1981. (20) Kohl, A.; Nielsen, R. Gas Purification, 5th ed.; Gulf Publishing Co.: Houston, TX, 1997. (21) Freguia, S. Modeling of CO2 Removal from Flue Gases with Monoethanolamine. M.S. Thesis, The University of Texas at Austin, 2002. (22) Fluor’s Econamine FG Plus Technology. NETL Carbon Sequestration Conference, 2003. (23) Scherffius, J. R.; Reddy, S.; Klumpyan, J. P.; Armpriester, A. LargeScale CO2 Capture Demonstration Plant using Fluor’s Econamine FG Plus Technology. Energy Procedia 2013, 37, 9. (24) Baburao, B.; Schubert, C. Advanced Intercooling and Recycling in CO2 Absorption. U.S. Patent US8,460,436 B2, 2013. (25) Stoever, B.; Konig, D.; Bergins, C.; Schonwalder, M.; Buddenberg, T. Coal Power Plant Having an Associated CO2 Scrubbing Station and Heat Recovery. U.S. Patent US2012/0216540 A1, August 30, 2012. (26) Cabanaw, E. J.; Mann, J. W. Alkylation Process. U.S. Patent US4404419A, September 13, 1983. (27) The Handbook of the Nitrogen Engineer; Chimia: Russia, 1986. (28) Leites, I. L.; Sama, D. A.; Lior, N. The Theory and Practice of Energy Saving in the Chemical Industry−Some Methods for Reducing Thermodynamic Irreversibility in Chemical Technology Processes. Energy 2003, 28, 43. (29) Oyenekan, B. A.; Rochelle, G. T. Alternative Stripper Configurations to Minimize Energy for CO2 Capture. 8th International Conference on Greenhouse Gas Control Technologies (GHGT-8), Trondheim, Norway, 2006; p 5. (30) Rochelle, G. T. Regeneration of an Aqueous Solution from an Acid Gas Absorption Process by Multistage Flashing and Stripping. U.S. Patent US7901487B2, March 4, 2004. (31) Figueroa, J. D.; Fout, T.; Plasynski, S.; McIlvried, H.; Srivastava, R. D. Advances in CO2 Capture Technology−The U.S. Department of Energy’s Carbon Sequestration Program. Int. J. Greenhouse Gas Control 2008, 2. (32) Carré, F. Improvement in Apparatus for Freezing Liquids. U.S. Patent US 30201, 1860. (33) Kasley, A. T. Refrigerator. U.S. Patent 1,506,530, February 4, 1924. (34) Erickson, D. C. Absorption Heat Pump Augmented Thermal Separation Process. U.S. Patent 4,350,571, September 21, 1982. (35) Wilinson, W. H.; Hanna, W. T. Process and System for Boosting the Temperature of Sensible Waste Heat Sources. U.S. Patent 4,333,515, 1982. (36) Rockenfeller, U.; Sarkisian, P. Triple Effect Absorption Cycle Apparatus. U.S. Patent 5,335,515, August 9, 1994. (37) Zhang, K.; Liu, Z.; Li, Y.; Li, Q.; Zhang, J.; Liu, H. The Improved CO2 Capture System with Heat Recovery Based on Absorption Heat Transformer and Flash Evaporator. Appl. Therm. Eng. 2014, 62, 7. (38) Zarrinehkafsh, M. T.; Sadrameli, S. M. Simulation of Fixed Bed Regenerative Heat Exchangers for Flue Gas Heat Recovery. Appl. Therm. Eng. 2004, 24, 10.

vapor-phase mole fraction of chemical species i [mol i/mol solution] material split fraction of the stripper condensate turned to the alternative location [tonne split/tonne total] fractional efficiency as compared to a Carnot heat engine [−] liquid-phase activity of chemical species i

Unit Operations

ABS-CON ABS-IC CHEX CHEX-MAIN CHEX-SIDE STR-CON STR-REB STR-IH

absorber’s condenser absorber intercooler combination of CHEX-MAIN and CHEX-SIDE main part of the central cross heat exchanger distributed part of the central cross heat exchanger stripper’s condenser stripper’s reboiler stripper interheater

Models

NRTL PDH RK RKS

nonrandom two-liquid activity coefficient model Pitzer−Debye−Huckel activity coefficient model Redlich−Kwong equation of state Redlich−Kwong−Sauve equation of state

Phases

L liquid V vapor

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REFERENCES

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NOTE ADDED AFTER ASAP PUBLICATION This paper was originally published ASAP on February 26, 2015. Corrections were made to the Supporting Information, and the paper was reposted on February 27, 2015.

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