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Techno-economic Optimization of a Green-Field Post- Combustion CO2 Capture Process using Superstructure and Rate-Based Models Ung Lee, Jannik Burre, Adrian Caspari, Johanna Kleinekorte, Artur Schweidtmann, and Alexander Mitsos Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b01668 • Publication Date (Web): 28 Oct 2016 Downloaded from http://pubs.acs.org on November 1, 2016

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Industrial & Engineering Chemistry Research

Manuscript Prepared for Industrial & Engineering Chemical Research

Techno-economic Optimization of a Green-Field PostCombustion CO 2 Capture Process using Superstructure and Rate-Based Models

Ung Leea, Jannik Burrea, Adrian Casparia, Johanna Kleinekortea, Artur Schweidtmanna and Alexander Mitsosa* a

AVT Process Systems Engineering (SVT), RWTH Aachen University, Aachen 52064, Germany

Apr. 27th, 2016, Revised October 17th, 2016

*Author to whom correspondence should be addressed. Tel.: (+49) 2418094704;

E-mail: [email protected] 1

ACS Paragon Plus Environment


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Manuscript Prepared for Industrial & Engineering Chemical Research

Abstract A techno-economic optimization of a commercial-scale amine-based post combustion CO2 capture process is carried out. The most economically favorable process configuration, sizing and operating conditions are identified using a superstructure formulation. The superstructure has 12,288 possible process configurations and unit operations in the superstructure are described using rigorous, rate-based models. In order to simplify the optimization problem, the problem is decomposed and process simulations are explicitly handled in the process simulator. Optimization is performed externally using a genetic algorithm. The best found process configuration includes the absorber intercooling, the rich vapor recompression, and the cold solvent split. The result of this study is compared in terms of the cost of capture basis and shows 38% reduction on the annual operation cost compared to the conventional amine-based CO2 capture process. Moreover, the savings on the total annualized cost is approximately 13%; a 30% of increase on the annualized investment cost results from additional unit operations.

Keywords Post-combustion carbon capture process; Genetic algorithm; Rate based reactive distillation, Techno-economic optimization; Superstructure; MINLP

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Introduction

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Manuscript Prepared for Industrial & Engineering Chemical Research Post-combustion carbon capture (PCC) is an effective technology to reduce greenhouse gases from existing large-scale sources which are responsible for more than 70% of the worldwide CO2 emissions1, 2. Amongst the available PCC technologies, the absorption process using aqueous amine is most commonly used for treating flue gases from power plants3. Also, it is the most mature carbon capture technology and its operability has been demonstrated in many industrial projects2, 4. However, PCC using aqueous amine consumes a considerable amount of energy to regenerate amine solvent. In addition, low loading capacity and degradation have been pointed out as obstacles of commercialization of this technology. Consequently, technical and economical improvements are still required to overcome weaknesses of the PCC process2. The efficiency of PCC highly depends on the solvent selection. Capacity, degradation, cost, and regeneration energy of solvents differ significantly, which makes it challenging to choose the most suitable one. Monoethanolamine (MEA) has received most attention as a solvent for CO2 separation because of its fast reaction rate5. However, the high corrosiveness of the solution requires to use diluted solution (30~40 wt%) and the low concentration together with its low loading capacity result in high thermal regeneration energy. Many researchers propose alternative solvents with high loading capacity. Secondary and tertiary amines such as diethanolamine, methyldiethanolamine, and triethanolamine6-9 are extensively studied and compared with MEA. Since higher absorption capacities also lead to a more efficient separation process, piperazine (PZ) was found to be a promising solvent, which absorption capacity is more than twice the capacity of MEA.10, 11 Amine mixtures are also proposed for replacing MEA solvent. Mitsubishi Heavy Industries, Ltd. reported the development of amine-based mixture solvents, whose application may save around 20% 3



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Manuscript Prepared for Industrial & Engineering Chemical Research regeneration energy compared to MEA12. Korea Electric Power Corporation developed a solvent called KoSol which is a mixture of amines. This solvent can reduce both regeneration temperature and regeneration energy13. Li et al. 14 studied mixtures of PZ and 2aminomethylpropanol. These mixtures have lower viscosity than PZ having comparable CO2 capture capability. Sherman et al.15 and Li et al.16 conduct similar researches. They investigated PZ and 2 methyl piperazine mixture and MDEA and methyl diethanolamine mixture. Recently, amine and ionic liquid mixtures are also received attention. Lu et al. investigated kinetics of CO2 absorption in a MEA and ionic liquid mixture. They used a MEA and 1 buthyl-3-methylimidayolium tetrafluoroborate mixture and compare its performance with MEA. Comprehensive review of recent CO2 capture technology is available in Dutcher et al.17 Even though these amine based solvents are promising to reduce the solvent regeneration energy, only limited numbers of industrial scale PCC projects have been carried out because economic incentives and regulatory requirements are not enough. Ammonia has also gained attention as a substitute for amine-based solvents, since it offers higher CO2 loading capacity, lower regeneration energy, less absorbent cost, less corrosive impact, and higher thermal and chemical stability 2, 18. However, the slow reaction kinetics, and the high volatility are issues need to be addressed for ammonia19. While significant efforts have been made for developing solvents, many studies have focused on advanced process configurations of the absorption process to reduce the overall capture cost rather than solvent replacement. These process configurations include: absorber intercooling20, 21, condensate heating22, condensate evaporation23, stripper overhead compression20, 23, 24, lean amine flash20, multi pressure stripping25, heat integration26-28, cold solvent split20, 23, 29, 30, flue gas/lean solvent precooling20, stripper interheating25 31, lean vapor 4


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Manuscript Prepared for Industrial & Engineering Chemical Research recompression20, 24, and rich amine recompression29. Each of them may offer specific process improvements and they have been evaluated in terms of solvent regeneration energy. Recently, Lin and Rochelle32 proposed advanced flash stripper configuration. In this study, a warm rich bypass stream and a cold bypass stream are introduced in order to minimize lost energy of the condenser and the advanced configuration can save about 16% of the solvent regeneration energy. Comprehensive review of these processes is provided in Le Moullec et. al.33 In spite of the capability to reduce solvent regeneration energy, process alternatives proposed in these studies should be evaluated further in order to capture economic penalties of including additional unit operations. Recent studies, therefore, focus more on the economics of PCC alternatives rather than energy consumption. Different cost estimations for carbon capture processes have been published in literature. Al-Juaied et al.34 summarized different costs estimations given by industry and compared the effect of first-of-a-kind capture plants against more mature technologies on the cost estimation. Herzog 35 discussed future opportunities to lower carbon capture costs based on a detailed analysis of costs associated on technology for CO2 separation and capture. Hence a composite model for costs from several types of power plant is developed. Davison36 presented cost data for coal and natural gas-fired power plants based on information from studies carried out recently for the International Energy Agency (IEA) Greenhouse Gas Programme. The recent estimation methodology is given by Rubin et al.37 They presented a review of carbon capture costs in 2015 in comparison to the Intergovernmental Panel on Climate Change (IPCC) special report1 of 2005 and summarized the results of recent cost studies. Subsequently, they proposed a new cost estimation method for combustion-based power plants with and with-out post-combustion capture systems based 5



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Manuscript Prepared for Industrial & Engineering Chemical Research on their approach presented in Rubin et al.38 Herein the authors recommend a consistent method of aggregating intermediate cost items to obtain the total capital and operating costs of a project based on the methods 39-42. In conclusion they calculated averaged costs of 74 USD/t CO2 captured for natural gas fired power plants and 46 USD/t CO2 captured for coalfired power plants. Vaccarellia et al.43 performed an economic evaluation of a postcombustion carbon capture process integrated in a power plant cycle. Gupta et al.44 presented a techno-economic analysis of absorber intercooling and multi pressure stripper. Similar work was done by Li et al.45. Four process configurations (i.e., Absorber intercooling, rich- amine split, modified rich amine split and stripper interheating) are assessed considering capital and operating cost. Several studies also performed process optimization in terms of cost. Fernandez et al.46 carried out an optimization of the lean vapor recompression process in terms of cost. Similarly, Oi et al.47 optimized three single process configurations considering operational and capital cost. However, these studies used simplified process models for optimization by assuming phase equilibrium in columns. Arias et al.48 optimized an aminebased carbon capture process with respect to the specific total annualized cost consisting of capital and operational cost. Although these studies have determined CO2 capture cost for certain advanced configurations quantitatively, the number of considered configurations was limited. Consequently, the optimum process configuration cannot be identified. The optimal process configuration for the PCC process can be obtained from superstructure optimization. Since a superstructure contains most or all of the potential structural options, the optimal design is decided through optimization. Several studies proposed an optimum process flowsheet using superstructure optimization. Yee and Grossmann49 presented a multi-stage heat exchanger network superstructure optimization. 6


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Manuscript Prepared for Industrial & Engineering Chemical Research The optimum process of micro power generation50 and thermal desalination51 are also proposed using superstructures among hundreds of different process alternatives. Trespalacios and Grossmann52 provided comprehensive reviewed superstructure optimization approaches. Superstructure optimization was also carried out in the past in the field of carbon capture and storage. Hasan et al.3 published a superstructure optimization of an adsorptionbased carbon capture process. Damartzis et al.53 presented a superstructure optimization of an absorption-based carbon capture process considering five process configurations. Therein, the observed processes have been optimized with respect to the capital cost and operational expenses of the respective process configuration. However, information about the rate-based models that have been used is not given. In addition, the superstructure includes limited process options, thus the global optimized solution may be overlooked. This article builds upon our previously published work54. In the previous work, we built a rigorous process model based on the pilot plant operation in South Korea and proposed the optimum retrofit process configuration using superstructure. Limited numbers of process alternatives were considered because the currently operating pilot plant does not allow implementing some of process options without major unit replacement. In addition, the optimization was carried out in terms of energy saving, thus it does not guarantee the lowest CO2 treating cost. Herein, an optimum PCC process for a 500 MWe pulverized coal-fired power plant is proposed using a superstructure. We include more than 10,000 process configurations in a superstructure, thus most of feasible options identified in the literature are included in the superstructure. In order to evaluate effects of process alternatives precisely, a rigorous rate-based distillation model including thermodynamics, kinetic, transport, and hydraulics models is optimized in terms of cost herein. In this way, the most economical 7



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Manuscript Prepared for Industrial & Engineering Chemical Research process configuration and operation condition can be obtained and the CO2 avoidance cost for the power plant can be evaluated more accurately. The optimization formulation is thus quite challenging computationally. The genetic algorithm (GA)55 is an attractive solution algorithm to solve complex optimization problems because it does not necessarily have derivatives or mathematical information of the models56. Thus, herein, the complex superstructure model is optimized using a GA interfaced with Aspen Plus. In the way, the complex optimization model can be decomposed into process simulation and optimization parts. Although the global optimum of the problem is not guaranteed with the stochastic solver, the optimization problem can be more easily handled with the decomposition.

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Process description 2.1

Conventional process

The base case of the previous study of Lee et al.54 scaled up to model the conventional commercial scale PCC process. In the conventional process, 90% of flue gas generated from a 500 MWe pulverized coal-fired power plant is captured using a 30 wt% MEA solvent. The flue gas includes 72% N2, 4% O2, 11% H2O, and 13% CO2 by mole for its constituents. As shown in Figure 1, the conventional PCC process mainly consists of an absorber, a stripper, and lean/rich amine heat exchanger. The SOx free flue gas enters the absorber from the bottom and CO2 selectively reacts with the MEA solvent. Then, the scrubbed gas is vented to the atmosphere. Herein, amine scrubbing of the absorber outlet is not considered because of the low amine concentration of the stream. However, the amine scrubbing may be required in many PCC processes when amine concentration of the emission stream is beyond the permissible level. The cold rich amine solvent leaves the absorber from the bottom. Two

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Manuscript Prepared for Industrial & Engineering Chemical Research coolers are used to increase the CO2 loading capacity of the solvent by lowering temperature of the lean amine solvent and the flue gas. The rich amine stream enters the top of the stripper and thermally regenerated. The overhead stream of the stripper is cooled to recover high purity CO2 and the hot lean amine stream is sent back to the absorber through a heat exchanger. As desorption in the stripper is endothermic, a high stripper temperature, which is limited to 120°C to prevent solvent degradation, leverages solvent capacity. Waste heat of hot lean amine solvent stream is partially recovered in the lean/rich amine heat exchanger in order to reduce the reboiler duty. For the process operability and preventing solvent degradation, two constraints i.e., minimum temperature approach (5°C) and the hot rich amine solvent temperature (99°C) are introduced. In this study, a rate-based absorption model, RadFrac, in Aspen Plus is used to simulate the absorber and the stripper. This approach is rigorous and offers higher model accuracy over the traditional equilibrium-stage modeling57. To successfully simulate the process model with accuracy, we include all thermodynamics, kinetic, transport, and hydraulics as well as their parameters. This model is originally developed in the work of Lim et al.58 and applied in several studies29, 59. In order to scale up pilot scale PCC process to commercial scale, diameters of columns are calculated using D = where V, ρ , and v are the volumetric

flow rate of the flue gas, the gas density, and the superficial gas velocity, respectively. We assume 16.8m and 11.25m of absorber and stripper heights, 1.1 ft/sec of superficial velocity, and 1.5 kg/m3 of gas density according to pilot scale PCC process model. The complete treatment of flue gas generated from 500 MWe coal fired power plant require six trains of PCC processes. Less numbers of PCC process trains may require larger than the maximum

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Manuscript Prepared for Industrial & Engineering Chemical Research column diameter (12.2m)60. Note that the scaled up PCC model is not validated. Column diameters are 11.5 and 10.3m for the absorber and stripper, respectively. More detail information can be found in the Aspen Plus input file provided in supporting information. Electrolyte NRTL model and Redlich-Kwong equation of state are used to predict thermodynamic properties. Solution chemistries such as water dissociation, bicarbonate dissociation, and MEA protonation are employed using Aspen chemistry model. Binary interaction parameters and ion interactions parameters are adapted from Austgen et al.61 and from the Aspen Plus database. The reaction kinetics and parameters for the rate controlled reactions are adopted from Pinsent et al.62 and Hikita et al.63 In the rate-based model, mass and heat transfer resistance are calculated according to the film theory with film discretization for nonhomogeneous liquid phase. The film model is calculated together with the reaction kinetics and electrolyte thermodynamics. The liquid phase MEA solution is discretized into six films and separate balance equations are solved for each film. The column hydraulics is accounted for via mass transfer, heat transfer, and interfacial area correlations. Onda’s method64 and Henley IMTP65 are used for both the interfacial area and the mass transfer coefficient correlation for the absorber and the stripper, respectively. The heat transfer coefficient is calculated using the Chilton and Colburn method66. With the interfacial area factors of 1.2 and 1 for the absorber and stripper, the model agrees with experimental data with high precision58. In the simulation model, several equality constraints and assumptions are embedded in order to simplify the optimization problem. The CO2 capture rate is assumed to be 90% and it is controlled by adjusting MEA lean solvent flowrate. Four and three random packings are used for absorber and stripper and solvent withdrawn and reinjection are possible between 10


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Manuscript Prepared for Industrial & Engineering Chemical Research packings. The stripper operating pressure is fixed at 1.8 bar and the reboiler temperature resigns approximately 120°C. The cooling water temperature is assumed to be 35°C. As a result, hot stream outlet temperatures of heat exchangers (e.g., condensers of absorber and stripper) are set to 40°C. Detailed input of the process model is summarized in table 1.

2.2

Process alternatives

The main disadvantage of the conventional process is the high reboiler duty for solvent regeneration. Furthermore, low loading capacity and degradation of the solvents may cause increasing cost. In order to overcome these disadvantages, various process configurations have been published. Based on comprehensive literature reviews, we include most of promising configurations: intercooling, flue gas/lean amine precooling20, 21, interheating25, 31, lean vapor recompression14, 18, rich vapor recompression29, stripper overhead compression14, 17, 18

, and cold solvent split14, 17, 23, 24. These are discussed in detail hereinafter.

Absorber intercooling increases the solvent loading capacity and mass transfer rate by lowering gas and solvent temperature in columns, and it eventually leads to a reduction of the amount of circulating solvent. Damartzis et al.53 indicates that absorber intercooling may reduce the reboiler duty by 11.6%. However, a too low absorber operating temperature may not be economically attractable by requiring expensive cold utilities. In addition, solvent regeneration energy may increase with lower solvent temperature, because the cold rich solvent temperature should be compensated in the stripper and may disturb reaction kinetics. Thus, the operating temperature of the absorber should be carefully optimized with a rigorous model embodying reaction kinetics. The flue gas and lean amine precooling also increase the solvent capacity, thus can be beneficial to reducing the reboiler duty. An important difference

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Manuscript Prepared for Industrial & Engineering Chemical Research between the intercooling and precooling is utility usage. In general, the intercooling is carried out using cooling water while the precooling requires chilled water. As a result, the precooling process can be less economically favorable due to the expensive utility cost. The stripper interheating targets to optimize the temperature profile in the stripper. Therefore, a semi-lean solvent stream at the middle of the column is extracted, reheated, and injected into the stripper again. Three vapor recompression processes, i.e., lean vapor recompression, rich vapor recompression, and overhead vapor recompression processes are also considered in this study. In the lean vapor recompression, the lean solvent stream is flashed under low pressure. The gas phase exiting the flash is recompressed and injected back to the stripper. In rich vapor recompression, cold rich amine stream is heated up without pressurization. A fraction of the rich amine solvent is evaporated and introduced to the stripper following by adequate pressurization. Both the lean vapor recompression and rich vapor recompression process recover heat of the hot lean amine stream not only by direct heat exchange but also by vapor generation. Consequently, the rich amine stream temperature can be lowered. The reduced rich amine stream temperature decreases condenser duty as well as reflux ratio, and eventually reduces reboiler heat duty. In addition, the reduced temperature makes further heat recovery possible in lean/rich amine heat exchanger. Overhead vapor recompression is a type of direct heat integration. In order to recover waste heat of hot overhead stream of the stripper, several process alternatives are suggested. Condensate heating22 and condensate evaporation23 are proposed, but heat recovery using these processes is very marginal and can cause solvent thermal degradation problems. Overhead vapor recompression20, 23, 24, 67 is suggested in many studies to recover waste heat from overhead steam by generating steam or 12


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Manuscript Prepared for Industrial & Engineering Chemical Research direct heat integration. Overhead vapor recompression is capable of reducing thermal energy consumption through external heat integration between the column overhead stream and the bottom stream and producing a higher pressure product stream. The high-pressure product stream makes it possible to avoid CO2 conditioning processes that consume about 20% of energy within the entire CCS chain. Woodhouse et al. and Ahn et al. claimed that the steam requirement can be reduced by using overhead vapor recompression and it can be reached about 32%23, 68. In the cold solvent split configuration, the cold rich solvent stream is split into two parts: preheated and cold. The cold stream enters the stripper at the top, while the preheated stream is injected at the middle of the stripper. Due to the additional cooling at the top, the reflux ratio and condenser duty of the stripper can be reduced29. This may lead to a decrease of the reboiler duty of about 10%33. There exist more process configurations than these which have been considered in this work. Multi pressure stripping follows the idea to separate the stripper into different sections operated at different pressures. Thus, the vapor flow going up is compressed and the liquid flowing down is expanded and flashed. Therefore, additional steam is supplied and the reboiler duty is reduced23. However, simulation studies have shown that the negative effect of additional auxiliary power outweighs the reduction in reboiler duty. Another configuration is the split-amine flow. In contrast to the cold solvent split, the rich amine flow is split after the preheater and a semi-lean amine stream flows back from the stripper to the absorber. Kun Bae et al. 69 and Oi et al.47 have found that this configuration has a great impact on the reboiler duty for high CO2 concentrations but not for the low CO2 concentration considered in this work. Advanced flash stripper configuration proposed by Lin and Rochelle32 shows the 13



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Manuscript Prepared for Industrial & Engineering Chemical Research solvent regeneration energy can be further reduced by introducing a warm rich bypass and a cold rich bypass at the same time. Although only cold amine split is considered in this study, the advance flash stripper configuration can be a promising option and should be considered in future studies.

3

Superstructure optimization 3.1

Superstructure configuration

Superstructure optimization is a powerful tool for the structural design of chemical processes. A superstructure contains all considered possible process alternatives and allows the user to simultaneously optimize both design and operating variables. Figure 2 illustrates the superstructure constructed in this study. The process options mentioned in Section 2.3 are implemented in one flowsheet in Aspen Plus, where the use of splitters and mixers gives possibilities of by-passing certain configurations. In the superstructure, the absorber intercooling and the stripper interheating are implemented using heaters and coolers option in the RadFrac model. In this way, convergence of the model is improved compared with using side streams by describing process alternatives using only continuous variables. It is worth to mention that intercooling and interheating are only possible in between packings. Three simple intercooling loops (located at 4.2, 8.4, and 12.6m) and two interheating loops (located at 4.2 and 8.4m) are employed in the absorber and the stripper, respectively. Flue gas and lean amine precooling are employed using HE-101 and HE-201 by adjusting their outlet temperatures. We assume that these heat exchangers are always selected based on the conventional process

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Manuscript Prepared for Industrial & Engineering Chemical Research configuration, thus only continuous variables are required to implement these options in the superstructure. The rich amine stream can be introduced to the stripper with three different configurations: simple heat recovery, lean vapor recompression and rich vapor recompression. Lean and rich vapor recompressions recover more waste heat as compared with simple heat recovery. They, however, require additional unit operations and higher quality utilities. Accordingly, the best option should be decided by using optimization solvers based on cost. The first splitter directs the rich amine stream either to the direct heat exchanger and lean vapor recompression (stream 9) or the rich amine recompression (stream 9-1). The cold solvent split configuration is embedded for both streams and the cold split flow is introduced to the top of the stripper. Meantime, the hot split stream can be injected to the stripper through any stages. Integer variables are used for deciding optimum feed location of the hot split stream, split fraction of hot lean amine stream (stream 12, 12-1, 12-2), and cold rich amine stream (stream 9 and 9-1). An overhead vapor recompression is also implemented in the superstructure. The stripper overhead stream can flow either to the CO2 compression process or to the overhead vapor recompression. In both options, the output pressure of the last stage of the compressor is fixed at 51.8 bar, which is the typical CO2 liquefaction pressure without using external coolant70. In overhead vapor recompression, a three stage compressor is used and inter-stage steams are directly integrated with the stripper bottom (M26) using a multi stream heat exchanger. Pressurized reflux streams are, then, expanded, flashed and recycled back to the stripper. The optimum feed stage of the flashed vapor and liquid streams are also decided using integer variables. In many cases, the water content in the last stage of the compressor 15



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Manuscript Prepared for Industrial & Engineering Chemical Research inlet stream is negligible. Thus, inter-stage cooling to the low temperature (40°C) is often advantageous rather than recovering a small amount of latent heat of water. However, it should be carefully decided because the optimal structure can differ by other operation condition. In this study, an additional binary variable is used to choose between inter-stage cooling and heat integration.

3.2

Optimization formulation

The superstructure optimization of the PCC process can be formulated as the following MINLP problem. ,

!" , ≤ 0, = 1,2,3, … . . , *

ℎ, , = 0, = 1,2,3, … . . , - ∈ / ⊂ 1 2 , ∈ 3 ⊂ 4 5

(P)

where x and y are vectors consisting of n continuous and m integer variables. , ,

!, , and ℎ , represent objective function, inequality, and equality constraints, respectively. As nonlinear constraints related to process modeling are explicitly calculated in Aspen Plus, the original problem can be reformulated using penalty function without having equality constraints. min 9:; = :; + γ*, @

!" , ≤ 0

A ≤ ≤ *, ∈ 3 ⊂ 4 5

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Manuscript Prepared for Industrial & Engineering Chemical Research 0, ℎ, , = 0 CD !",2E2F"2GHI , ≤ 0 *, = B 1, ℎ, , ≠ 0 2, ℎ, , = 0 CD !",2E2F"2GHI , > 0

(1)

The *, in equation 1 are penalty terms checking the convergence of the process

simulator and nonlinear constraints where γ indicates appropriately selected large penalty

parameters ( γ = $ 6.66 × 10O ). If the process simulator converges without any problem, the

penalty function, *, , returns zero. When process simulator reports violation of the

equality constraints (e.g., flash calculation, mass and energy balance) the penalty term becomes 1 assigning a large value to the fitness function. Similarly, the penalty term becomes two when nonlinear inequality constraints, !",2E2F"2GHI , , are violated (e.g., minimum

temperature violation). The objective function of the optimization problem is the minimization of expected total annualized cost, TAC, of the process and it is explicitly calculated using separate m-file. Note that following equations (2, 3, 4, 5, 7, and 9) are included in the objective function m-file, thus the GA solver can avoid using equality constraints. The TAC is composed of annualized investment cost, AIC, and averaged annualized operating cost, AOC. We assumed 25 years of plant life time and 0.1543 for capital recovery factor for annualizing total capital investment cost. The cost of the process are treated following the procedure in Hasna et al.3. The AIC comprises of the total plant cost (TPC) and the annual maintenance cost (AMC): :

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