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Technoeconomic Assessment of an Advanced Aqueous AmmoniaBased Postcombustion Capture Process Integrated with a 650-MW Coal-Fired Power Station Kangkang Li,*,†,‡ Hai Yu,*,† Shuiping Yan,§ Paul Feron,† Leigh Wardhaugh,† and Moses Tade‡ †

CSIRO Energy, 10 Murray Dwyer Circuit, Mayfield West, New South Wales 2304, Australia Department of Chemical Engineering, Curtin University of Technology Australia, GPO Box U1987, Perth, Western Australia 6845, Australia § College of Engineering, Huazhong Agricultural University, No. 1 Shizishan Street, Hongshan District, Wuhan 430070, PR China ‡

S Supporting Information *

ABSTRACT: Using a rigorous, rate-based model and a validated economic model, we investigated the technoeconomic performance of an aqueous NH3-based CO2 capture process integrated with a 650-MW coal-fired power station. First, the baseline NH3 process was explored with the process design of simultaneous capture of CO2 and SO2 to replace the conventional FGD unit. This reduced capital investment of the power station by US$425/kW (a 13.1% reduction). Integration of this NH3 baseline process with the power station takes the CO2-avoided cost advantage over the MEA process (US $67.3/tonne vs US$86.4/tonne). We then investigated process modifications of a two-stage absorption, rich-split configuration and interheating stripping to further advance the NH3 process. The modified process reduced energy consumption by 31.7 MW/h (20.2% reduction) and capital costs by US$55.4 million (6.7% reduction). As a result, the CO2-avoided cost fell to $53.2/ tonne: a savings of $14.1 and $21.9/tonne CO2 compared with the NH3 baseline and advanced MEA process, respectively. The analysis of energy breakdown and cost distribution indicates that the technoeconomic performance of the NH3 process still has great potential to be improved.

1. INTRODUCTION Postcombustion capture (PCC) by chemical absorption is the leading technology for CO2 capture and is claimed to be the most-likely dominant technology for commercial-scale CO2 capture from the coal-fired power station before 2030.1 The world’s first commercial CO2 capture plant using aqueous amines at SaskPower Boundary Dam Power Station in Estevan, Saskatchewan, Canada has taken a significant step to commercialize this technology.2 Despite its commercial application, the amine technology is still suffering from some severe technical and economic difficulties restricting its further application. Specifically, monoethanolamine (MEA), as benchmarking solvent, has struggled with its heavy energy penalty, solvent degradation, low loading capacity.3,4 Piperazine (PZ), as second-generation solvent, has an advanced technical performance more so than MEA including faster absorption, higher CO2-carrying capacity, and lower degradation.5,6 However, precipitation and slurry formation of PZ-based PCC at low temperature and lean solvent condition represent a great process challenge,7 which likely affects the process flexibility and increases the operation difficulty. The blended amine solvents, such as MEA/2-amino-2-methyl-1-propanol (AMP),8 PZ/AMP,9 etc., are extensively investigated to take their respective advantages for enhancing the overall technical © 2016 American Chemical Society

performance, though the issue of amine degradation is still unsolved resulting in solvent loss, equipment corrosion, and generation of volatile degradation compounds.10 In addition, the amine technology requires clean flue gas with preferred SO2 concentration below 10 ppmv, which requires extra facilities and equipment for deep SO2 removal, resulting in additional capital investment.11,12 Aqueous ammonia (NH3) is a promising, alternative solvent for CO2 capture with some obvious advantages over amines, such as lower solvent cost, no solvent degradation, and simultaneous capture of acid pollutant gases (CO2 and SO2, etc.).13−17 Intensive research has been conducted to evaluate the technical feasibility of the NH3 process, including the pilot and demonstration trials by industrial companies and research organizations such as Alstom, 18,19 Powerspan, 20 and CSIRO.21,22 These trials confirmed the technical feasibility of the NH3 process and demonstrated many of the expected benefits: 80−90% CO2 removal efficiency, >99% high-purity CO2 product, simultaneous capture of CO2 and SO2. Modeling Received: Revised: Accepted: Published: 10746

June 1, 2016 August 11, 2016 September 9, 2016 September 9, 2016 DOI: 10.1021/acs.est.6b02737 Environ. Sci. Technol. 2016, 50, 10746−10755

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Figure 1. Schematic of advanced aqueous ammonia-based postcombustion capture plant integrated with a greenfield coal-fired power station.

studies also highlight the technical advantages of the NH3 process with a low regeneration duty of 0.93−2.5 MJ/kg CO2,23 which is lower than 3.7 MJ/kg CO224 and 2.6 MJ/kg CO25 for MEA and PZ processes, respectively. However, NH3-based PCC processes are confronted by a few technical challenges. The most critical one is the high NH3 loss due to its intrinsic volatility. The chilled ammonia process (CAP)25 operates at low absorption temperatures to reduce NH3 vaporization significantly. However, this comes at the expense of a heavy penalty for solvent chilling duty26−28 and introduces problems in slurry handling during operation. The second challenge is the high parasitic energy penalty involved in CO2 capture. For example, integrating an aqueous NH3 process into a coal-fired power plant would reduce net power plant efficiency by 20−30%.28,29 Moreover, ammonia itself is a hazardous gas that strongly irritates the throat, eyes, and respiratory system and forms the aerosol, fine particles by the interaction with gaseous acid pollutants in the atmosphere. The NH3 emission level should be controlled to less than 25 ppmv for the sake of human health and environment protection.30 Recently, process designs and modifications were investigated to enhance the technical performance and increase the process feasibility in NH3-based PCC.30−32 Specifically, an advanced NH3 recycle process was designed to recover >99% vaporized NH3 with NH3 emission level below 25 ppmv and remove >99% SO2 with low energy penalty; a two-stage absorption was proposed to reduce NH3 emissions by >50%; rich-split and interheating processes were applied to reduce regeneration energy by around 25%. Although these modifications were technically advantageous, their economic feasibility is still uncertain, because the implementation of new process configurations will always require the extra equipment and more auxiliary power to operate, leading to an increase of capital cost and operating cost. It is therefore of great importance to evaluate the trade-off between the technical improvement and economic burden, resulting in comprehensive assessment of these novel configurations.

To date most work focuses on the technical assessment and improvement of the NH3 process, with limited attention paid to its economic performance when integrated with a coal-fired power station. Table SI-1 summarizes the cost performance of the NH3-based CO2 capture process integrated with a coalbased power station.29,33,34 The first economic assessment was a scoping study by Ciferno et al.,33 who concluded that the NH3 process has significant cost advantages over MEA. However, this was a preliminary study based on limited knowledge and understanding of the NH3 system, which led to optimistic estimated costs. More recently, Versteeg et al.29 and Valenti et al.34 independently performed a technoeconomic analysis of the CAP, but their conclusions are contradictory. The former declared that the energy and economic performance of the CAP were superior to the MEA process, while the latter indicated no such advantage, since the significant chilling loads and associated costs offset the benefits from CAP. However, none of studies discussed the validation of technoeconomic models for the NH3 process, and the reliability of the models is unknown. Moreover, their technical results were obtained using equilibrium-based models, from which the results are thermodynamically achieved. Such models are likely to underestimate energy consumption and column sizes and hence underestimate the capital costs and overall economics of the PCC plant. A rigorous, rate-based model is preferable when describing the technical performance of the NH3 process. In addition to these shortcomings, no studies to date have investigated the economic benefits of integrating SO2 removal into the NH3-based PCC process, despite this bringing the technical advantage and cost savings of simultaneously capturing SO2 and CO2 in the NH3 process. The present study performs a detailed technoeconomic analysis of the aqueous NH3-based PCC process integrated with a coal-fired power station. The novelties of this study include the following: (i) exploring the economic benefits of combined SO2 and CO2 capture in the NH3 process; (ii) technoeconomic performance of the NH3 process and its process improvements based on a validated rate-based model 10747

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secondary FGD to deeply clean the flue gas to below 10 ppmv SO2,11,12 the proposed NH3 process can replace the FGD equipment with the NH3 recycle and SO2 removal system, thereby reducing capital costs. As indicated by Cichanowicz,36 the capital investment of a wet FGD system for a 650-MW unit increased from US$320/kW in 2004−2006 to US$376/kW in 2008−2010: an increase of around $16/kW per year (in 2008 dollars). Thus, the capital cost of an FGD system in 2012− 2014 is estimated at US$433/kW (2008 dollars), which is equivalent to US$425/kW (2013 dollars) when using the conversion rate of Chemical Engineering Plant Cost Index.37 Replacing an FGD system with the NH3 process thereby reduces the capital investment for a 650-MW power station by around US$276 million (from US$3,246 to US$2,281/kW). Note that the cost benefit from FGD replacement with this NH3 process is only applied for the greenfield power station. 2.1.2. NH3-Based PCC Process. The NH3-based PCC process is designed with an 85% CO2 removal efficiency when dealing with 560 tonne/h (∼4 million tonne/year) CO2 from the power station. The rigorous, rate-based model of the NH3−CO2−SO2−H2O system, which has been validated against laboratory and pilot experimental results,30,31 is employed to simulate the technical performance of the NH3based PCC. This enables a reliable sizing of columns and equipment and adequate calculations of the energy requirements during CO2 capture. Due to the large flue gas flow rate, four parallel process trains are proposed, each with an NH3 recycle and SO2 removal unit, CO2 capture unit, and a CO2 compressor (detailed column sizing is referred to in the Supporting Information). We first evaluate the energy and technical performance of a baseline NH3 process, under the process conditions which were described in detail in a previous investigation:32 25 °C lean solvent, 6.8% NH3 concentration, 0.225 mol/mol lean CO2 loading, 10 bar stripper pressure, and 150 bar CO 2 compression. Then process modifications for both absorber and stripper, as shown in Figure 1, are integrated to improve the technical performance of the NH3 process. A two-stage absorption configuration with intermediate cooling is introduced to reduce NH3 slip during the CO2 absorption process: the bottom stage scrubs the majority of CO2, while the top stage captures vaporized NH3 leaving the bottom absorber. The rich-split configuration reduces regeneration duty by splitting the cold solvent to recover the steam heat in the stripping process. The stripper interheating configuration decreases reboiler duty by using the high-quality, high-temperature heat in the hot, lean solvent leaving the stripper. Lastly, the three advanced configurations are integrated into one process to

and economic model; (iii) insight into whether process modifications are technically and economically feasible and integrate different modifications into one process to maximize the technoeconomic benefits. We also perform a technoeconomic evaluation of the benchmark MEA process for comparison. To our knowledge, this is the first technoeconomic evaluation that combines rate-based modeling, simultaneous CO2 and SO2 removal, and process modifications to analyze an aqueous NH3-based PCC process integrated with a coal-fired power station.

2. METHODOLOGY 2.1. Process Description. Figure 1 illustrates the NH3based PCC process integrated with a coal-fired power station. The PCC plant includes an NH3 recycle and SO2 removal unit, CO2 capture unit, and CO2 compressor. 2.1.1. Power Station. The study is based on an NH3-based PCC plant that is assumed to be integrated with a greenfield coal-fired power station with a designed power output of 650 MW and a net efficiency of 38.9%. The technical design and cost estimation for the power station has been described in detail by the United States Energy Information Administration.35 Table 1 summarizes the technical and costing information for the power station. Table 1. Technical and Cost Information (2013 US Dollars) of the 650-MW Greenfield Coal-Fired Power Station with and without FGD parameter

unit

net electricity output net efficiency total flue gas flow rate CO2 flow rate SO2 content total capital investment total capital cost per kW fixed operation and maintenance (O&M) cost variable O&M cost

MW % tonne/h tonne/h ppmv $ million $/kW $/kWyear $/MWh

power station with FGD

power station without FGD

650 38.9 3,180 560 ∼30 2,110 3,246 37.8

650 38.9 3,180 560 ∼200 1,834 2,821 32.9

4.5

3.9

It should be mentioned that the power station includes wet flue gas desulfurization (FGD) equipment for SO2 removal. However, our previous investigation has shown that the NH3 process has the capacity for a high SO2 removal efficiency (>99.9%) at various SO2 concentrations in the NH3 recycle and SO2 removal unit.31 Unlike the MEA process, which requires a

Figure 2. Capital costing methodology for a CO2 capture plant (Source: United States Department of Energy). 10748

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water, NH3 solvent, and wastewater. The associated consumption is determined from the process simulation based on the heat and material balance of the entire CO2 capture process. The electricity and steam consumed in the PCC are assumed to be supplied from the power station at the expense of the net power penalty. This penalty will be reflected in the overall economic performance of the cost of electricity delivered from power station and the CO2-avoided cost. When the total capital investment and the assumptions are determined, the economic assessment is conducted based on the calculation of net present value to achieve a breakeven electricity-selling price, termed the levelized cost of electricity (LCOE). This is the minimum price sold to the grid system so that the power plant owner can receive the revenue to cover all capital investments and O&M costs during the plant life. The LCOE calculation method is in the Supporting Information. For simplification, plant income does not include sulfur product sales, even though the proposed NH3 process captures SO2 from flue gas. It is anticipated that the NH3 process would receive the economic benefits from selling sulfur-containing fertilizers. Lastly, the CO2-avoided cost (US$/tonne CO2) is calculated to quantify the cost of capturing 1 tonne CO2 from the power station and to evaluate the economic viability of the NH3-based PCC process. The calculation method is shown in eq 1

improve the technical performance of the NH3 process (For detailed descriptions of process modifications refer to ref 32.). 2.2. Economic Assessment. 2.2.1. Capital Cost and Economic Assumption. Figure 2 illustrates the capital costing method for a PCC plant proposed by the United States Department of Energy (US DOE)12,38 (The framework is referred to in the Supporting Information.). The Aspen Capital Cost Estimator (ACCE) V8.6 is employed to calculate the main equipment costs (in 2013 US dollars) involved in the NH3 process, such as columns, packings, blowers, cross heat exchangers, pumps, and compressors. ACCE uses the equipment models contained in the Icarus Evaluation Engine to generate preliminary equipment designs and simulate vendorcosting procedures to develop detailed costs for each piece of equipment through mapping, sizing, and evaluating.39 Previous investigation revealed that equipment costs estimated from the ACCE agree well with the cost results from the most detailed DOE economic study for a MEA-based CO2 capture plant,40 providing us with the confidence to reliably estimate equipment costs for the NH3 process. The total material cost includes all the equipment in the PCC plant and the associated costs for piping, civil engineering, structural steel, instrumentation, electrical wiring, insulation, and paint, while the total labor cost includes manpower for construction and installation of equipment. The process contingency and project contingency are included in cost estimates to compensate for process uncertainties and unknown costs due to a lack of complete project definition and engineering. The percentages used to calculate the total investment cost were under the guidelines of US DOE economic analysis38 and the Association for the Advancement of Cost Engineers International Recommended Practice,41 with an expected accuracy of total capital cost of ±30%. Table 2 lists the primary economic assumptions for the assessment of the PCC process integrated with the coal-fired

cos t of CO2 avoided =

where (LCOE)PCC and (LCOE)ref are the LCOE of the power plant with and without PCC, respectively, in $/MWh; (CO2 emission)ref and (CO2 emission)PCC are the CO2 mass flow rate emitted from the power plant without and with PCC, respectively, in tonne CO2/MWh. The proposed economic model has been validated by comparison with published cost results from an MEA-based CO2 capture plant.40 The capital costing methodology performed satisfactorily when predicting the total capital estimation of a MEA-based PCC plant with ∼1 million tonne CO2 capture capacity. The cost results for the MEA process agree well with the published results, with deviations of ±30% for specific capital investment and ±10% for CO2-avoided cost. Because the NH3 process adopts the same economic model as the MEA process, these satisfactory agreements indicate the rationality of employing this economic model to evaluate the economic performance of the NH3 process.

Table 2. Primary Economic Assumptions of the Postcombustion Capture Process Integrated with the 650MW Coal-Fired Power Station parameter present value plant life capital cost discounted cash flow rate construction time fuel cost budget allocated in construction year 1, year 2, and year 3 plant capacity factor plant capacity factor of first year’s operation fixed O&M cost cooling water demineralized water NH3 solvent wastewater treatment

(LCOE)PCC − (LCOE)ref (CO2 emission)ref − (CO2 emission)PCC (1)

value 2013 US dollars 30 years calculated 8% 3 years $2/GJ 40%, 30%, and 30%

3. TECHNOECONOMIC PERFORMANCE OF THE BASELINE NH3 PROCESS 3.1. Energy Performance. As indicated in Table 3, the NH3 process is energy-intensive, placing an energy burden of 157 MW on the power station and decreasing its net efficiency by 9.5%. This energy performance is better than that of baseline CAP of 10.4%−10.9% by Hanak et al.27 and 10.4%−11.6% by Linnenberg et al.28 This energy savings is mainly attributed to the advanced process design of NH3 recycling which greatly utilizes the flue gas waste heat and saves the heat requirement of NH3 regeneration. The second reason is the rise of absorption temperature to 25 °C instead of chilling solvent, which also contributes to the saving of solvent chilling duty. The heat requirement in the CO2 stripper is the largest power consumer accounting for 64% of total energy consumption, which is among the range of 60−65% in the published

85% 50% 3.5% of total capital investment $0.35/m3 $2.0/m3 $600/tonne $100/m3

power plant, which follows the criteria set out by the International Energy Agency for technical and economic assessment of power plants with low CO2 emissions.42 The fixed O&M cost is assumed to be 3.5% of total capital investment of PCC plant.43 The variable O&M cost includes the material and chemical costs of cooling water, demineralized 10749

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project contingencies, and owner’s costs. The NH3’s capital cost is greater than that of the MEA process (US$645 million excluding secondary FGD).40 As indicated in Table 4, three primary equipment items contribute to the higher capital cost: (i) larger CO2 absorber, due to the slow reaction rate; (ii) additional NH3 wash column, owing to high NH3 loss; (iii) larger cross-heat exchanger, due to the greater temperature excursion from 25 °C (rich solvent) to 145.5 °C (lean solvent) compared with 40−120 °C for the MEA process. However, the NH3 process benefits the cost savings from the smaller CO2 compressor, and a more important capital savings from FGD replacement. Specifically, high stripping pressure in the NH3 process (10 bar) results in low compression ratio and a subsequent small compressor size, which saves US$ 91.1 million in capital investment compared to the MEA process. The replacement of FGD with the advanced NH3 recycle unit, as discussed in section 2.1, leads to capital savings of US$276 million. Both of these contribute the cost savings of US$53.2 million in total capital investment, making the NH3 process economically advantageous over the MEA process. 3.3. Economic Performance. As shown in Table 3, with the integration of the NH3 process, the LCOE of the power station rises from $71.9 to $118.0/MWh (an increase of 64%). This results in a CO2-avoided cost of $67.3/tonne, which is lower than $73.2/tonne of the CAP process estimated by Versteeg et al.29 This cost savings is primarily due to the great energy reduction as aforementioned. Compared with the MEA process, the NH3 process has an obvious economic advantage of $67.3 vs $86.4/tonne: a savings of $19.1/tonne CO2 avoided. Three reasons account for this cost advantage: (i) less energy required for the NH3 process; (ii) less capital investment, due to the high stripping pressure and FGD replacement; and (III) cheap cost and lower degradation rate, which reduces the variable O&M cost.

Table 3. Energy and Economic Performance of the Baseline NH3 Process and the Baseline MEA Process with a Capture Capacity of ∼4 Million Tonne CO2/Year (Four Process Trains) baseline NH3

baseline MEA40

MW

+650

+650

%

+38.9

+38.9

MW MW MW MW MW MW MW MW % % %

−100.1 −21.5 −9.5 −7.8 −13.5 −4.6 −157 +493 −24.2 −9.5 +29.4

−101.7 −52.6 −14.3 −2.4 − −6.0 −177 +473 −27.2 −10.6 +28.3

US$/kW

2820

3,246

832.8 21.3

643 24.6

18.8

2.16

89.4

89.4

29.1

22.6

5.5

23.2

157

177

71.9 118.0

71.9 130.8

67.3

86.4

unit Energy Performance power island net electricity output without PCC net efficiency PCC island stream extraction compressor blowers pumps chiller auxiliary total energy penalty of PCC power output with PCC power output reduction net efficiency penalty net efficiency with PCC Economic Performance capital cost and O&M cost capital investment of power station capital cost of PCC plant fixed O&M cost of power station variable O&M cost of power station coal cost

US$ million US$ million/ year US$ million/ year US$ million/ year fixed O&M cost of PCC plant US$ million/ year variable O&M cost of PCC plant US$ million/ year power penalty MW economic performance of the MEA process LCOE of power station US$/MWh LCOE of power station and US$/MWh PCC plant CO2-avoided cost US$/tonne CO2

4. PROCESS IMPROVEMENT 4.1. Advanced Absorber Configuration. As shown in Table 5, the advanced absorber configuration significantly reduces NH3 emission levels from 26,500 to 12,000 ppmv and brings many technical and economic benefits. First, it reduces the circulation flow rate and chilling temperature of the NH3 wash solvent, thus decreasing the chilling duty from 13.5 to 5.5 MW. Second, it reduces the column size of the NH3 recycle system; the packing height of the NH3 wash column reduces from 15 to 10 m, while the pretreatment column decreases from 10 to 8 m. Although the two-stage absorption process increases the absorber packing height from 16 to 20 m and requires an extra intercooler for solvent cooling, the energy and capital cost savings offset this added cost. As a result, the twostage absorption reduces the total capital cost by US$32.8 million, leading to savings of $2.8/MWh in LCOE and $4.4/ tonne in CO2-avoided cost. 4.2. Advanced Stripper Configuration. Table 6 summarized the technoeconomic performance of the stripper modifications: rich-split and interheating process. It is clear that the stripper modifications bring technical improvements and economic benefits for the NH3 process. The rich-split configuration reduces the reboiler duty from 3.27 to 2.89 MJ/kg CO2 and decreases condenser duty from 1.45 to 0.39 MJ/kg CO2, leading to an energy savings of 12.7 MW. The decreasing reboiler and condenser duties benefit the size reduction and the subsequent capital costs of the stripper

results. 27,28 The second-largest contributor is the CO 2 compressor accounting for 14%, close to 9.5−14% of baseline CAP.27,28 This low proportion of energy penalty for CO2 compression owes to the high stripping pressure used in the NH3 process. This significantly reduces the compression duty due to the decreasing compression ratio and contributes to a less net efficiency penalty of 1.1% compared to the baseline MEA process. 3.2. Capital Performance. Table 4 displays the details of breakdown equipment cost and total capital investment of the baseline NH3-based PCC plant for one process train. The cost includes all the necessary equipment involved in the NH3 recycle and SO2 removal unit, CO2 capture unit, and CO2 compression unit. The equipment cost is calculated based on carbon steel, except for the column inside and packing materials, which are of SS304 stainless steel to prevent corrosion. The total capital investment of the CO2 capture plant is estimated at US$832.8 million for four process trains, including all equipment costs, indirect costs, process and 10750

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Table 4. Total Capital Investment Cost in 2013 US Dollars of Baseline Aqueous Ammonia (NH3)-Based Postcombustion Capture Plant for One Process Train (∼1 Million Tonne/Year) equipment

material

NH3 Recycle Unit/Pretreatment blower 730,930 pretreatment column 3,481,402 pretreatment packing 3,641,900 washing column 3,824,621 washing packing 7,272,900 heat exchanger 197,483 cooler 190,633 chiller 510,840 pumps 74,846 solvent tank 160,954 CO2 Capture Unit blower 730,930 absorber 6,220,677 absorber packing 6,309,520 stripper 2,312,507 stripper packing 1,166,003 stripper reboiler 2,044,219 stripper condenser 371,059 main heat exchanger 9,070,237 lean solvent cooler 401,144 pumps 2,149,592 solvent tank 491,708 CO2 Compression CO2 compressor 12,700,000 Auxiliary water separation unit 3,914,379 others 358,857 total direct cost (TDC) total indirect cost (TIC) bare erected cost (BEC) engineering, procurement, construction (EPC) total plant cost (TPC) total capital investment total capital investment including secondary FGDb

manpower

direct cost

specificationa

direct cost of MEA40

81,303 1,283,870 89,075 1,322,576 141,803 38,593 38,398 64,386 23,643 48,192

812,233 4,765,300 3,641,900 5,147,200 7,414,700 236,076 229,031 575,226 98,489 209,146

200 m3/s ⌀10 m × H13.25 m ⌀10 m × H10 m ⌀10 m × H20 m ⌀10 m × H15 m TEMA, 510 m2 TEMA, 460 m2 TEMA, 2000 m2 5.9 m3/min 66 m3

812,000 2,875,000 1,883,000 2,842,000 2,060,000 202,000 108,000 93,000

81,303 2,294,001 92,638 242,937 259,791 165,744 51,959 4,465,863 56,568 249,156 63,962

812,233 8,514,700 6,402,158 2,555,400 1,418,094 2,209,963 423,018 13,536,100 460,700 2,398,748 555,670

200 m3/s ⌀12 m × H21.5 m ⌀12 m × H16 m ⌀6.8 m × H10.75 m ⌀6.8 m × H8 m TEMA, 9,987 m2 TEMA, 1,329 m2 TEMA, 46,300 m2 TEMA, 1,970 m2 76.2 m3/min 702 m3

812,000 8,376,000 6,117,000 2,273,000 1,464,000 3,183,000 305,000 1,502,000 522,000 211,000 388,000

2,279,500

14,979,500

6 stages, 150 bar

24,028,000

906,895 91,643

4,821,274 450,500 82,714,409 16,542,882 99,257,291 126,056,759 181,045,298 208,202,093 208,202,093

1,227,000

0.2 TDC TDC + TIC 1.27 BEC 1.2 EPC + 0.3 BEC 1.15 TPC

60,586,000 12,117,000 72,703,000 92,333,000 132,611,000 152,502,000 221,502,000

a Tubular Exchanger Manufacturers Association, Inc. (TEMA) shell and tube heat exchangers are used in this study with 2000 W/m2 K. bCapital cost of secondary FGD is US$69 million for one process train.

electricity savings, due to the significant reduction of reboiler duty and chilling duty. This is a 20.2% reduction in total energy consumption compared to the baseline. Economically, the combined process cuts total capital investment by US$55.4 million, consequently reducing the CO2-avoided cost from US $67.3 to US$53.2/tonne (a savings of $14.1/tonne CO2). Compared with the advanced MEA process, the modified NH3 process has a great CO2-avoided cost savings of US$21.9/ tonne, demonstrating the technical and economic competitiveness of the proposed NH3 process. Although implementing these advanced process configurations increases the complexity of the PCC plant by appending an intercooling heat exchanger to the absorber and an interheating heat exchanger to the stripper, the added complexity is considered affordable in light of the significant savings in energy consumption and CO2avoided cost.

reboiler and condenser. The rich-split modification also reduces the heat surface area and capital cost of the lean/rich heat exchanger, since less rich solvent flows through the heat exchanger. All of these lead to a total capital reduction of US $66.4 million for the four process trains, saving $4.2/MWh of LCOE and $7.7/tonne of CO2-avoided cost. The interheating modification reduces the reboiler duty from 3.27 to 3.0 MJ/kg CO2 and the condenser duty from 1.45 to 1.16 MJ/kg CO2. This comes at the expense of extra equipment required, such as an interheating exchanger, pump, and rich/ lean cross-heat exchanger with a larger heat transfer area, which increases capital investment by US$43.9 million compared to the reference case. However, the interheating process reduces reboiler duty by 8.3% and cuts electricity consumption by 8.4 MW. The benefit from energy savings therefore outweighs the capital cost increase, reducing the LCOE by $0.4/MWh and CO2-avoided cost by $0.7/tonne CO2. 4.3. Modified Process. The proposed modified NH3 process integrates two-stage absorption, rich-split and interheating, thereby maximizing technoeconomic benefits (Table 6). Technically, the modified process enables a 31.7-MW

5. FURTHER TECHNICAL ADVANCEMENT Figure 3 illustrates the breakdown of energy consumption, capital cost, and CO2-avoided cost for the modified NH3 process. This provides the guidelines for technical improve10751

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Article

Environmental Science & Technology

The energy penalty is clearly the major barrier to commercial application of integrating the NH3 process with the power station, accounting for 54.7% of the CO2-avoided cost. As reboiler duty is the major contributor to total energy consumption (60.6%), reducing regeneration duty should be the top priority to decrease the CO2-avoided cost. As pointed out by Goto et al.,44 decreasing 1 MJ/kg CO2 of regeneration duty improves power station net efficiency by around 2%. Theoretically, the heat of CO2 desorption ranges from 0.65− 1.6 MJ/kg CO2,23,45 depending on the NH3 concentration and CO2 loading. This indicates that there is still considerable room in the NH3 process (2.47 MJ/kg CO2 in this study) to decrease reboiler duty through process improvements. Our sensitivity study shows that a 10% decrease of regeneration duty would reduce the CO2-avoided cost by 4.4%. Capital investment is the second-highest contributor, representing 27.1% of the CO2-avoided cost. This relatively low percentage is attributed to the FGD replacement by the proposed NH3 process, which reduces the capital cost of the power station, and in a sense reduces the capital investment required for the NH3-based PCC plant. Practical approaches that could further reduce capital costs include improving CO2 absorption rate to decrease the absorber size and reducing NH3 vaporization to decrease the sizes of the pretreatment and wash columns. The former can be achieved by adding rate promoters, such as MEA and piperazine (3−4 times improvement)46 or sarcosine (∼4 times improvement).47 Doubling the absorption rate would reduce CO2-avoided cost by 7.1%, owing to the capital reduction of the CO2 absorber. The latter could be realized by introducing NH3 inhibitors. These could include transition metal ions, via complexation between metal ions and NH3 ligands (20−40% reduction in NH3 vaporization48), and organic chemicals, via interaction of physical−chemical hydrogen bonding (10−40% reduction49). A 50% reduction in NH3

Table 5. Comparison of the Technical and Economic Performance (in 2013 US dollars) of the Aqueous Ammonia (NH3)-Based Postcombustion Capture Process before and after Two-Stage Absorption two-stage absorption

baseline

Main Changes in Technical Performance (Four Process Trains) 26,500 12,000 NH3 emission after absorber, ppmv NH3 emission after wash column, ppmv ∼50 ∼15 heat transfer area of intercooler, m2 − 4500 chilling temperature, °C ∼1 ∼5 chilling duty, MW −13.5 −5.5 regeneration duty, GJ/tonne CO2 3.27 3.27 reboiler duty, MW −100.1 −100.1 total energy penalty, MW −157 −150 Main Changes in Capital Cost Performance (One Process Train) wash column packing height, m 15 10 pretreatment column packing height, m 10 8 absorber column packing height, m 15 20 wash column and packing, US$ 12,561,900 8,378,600 pretreatment column and packing, US$ 8,496,275 6,797,020 absorber column and packing, US$ 14,916,858 17,160,551 absorber intercooler and pump, US$ 0 1,135,391 others, e.g. cooler, chiller, US$ 1,138,822 379,607 savings in column cost, US$ 0 3,262,686 savings in total capital investment, US$ 0 8,212,927 Economic Performance (Four Process Trains) capital cost, million US$ 832.8 800.0 levelized cost of electricity, $/MWh 118.0 115.2 CO2 avoided cost, $/tonne CO2 67.3 62.9

ments that will increase the economic viability of NH3-based PCC.

Table 6. Primary Technical Changes and Economic Improvement (in 2013 US dollars) of Different Stripper Modifications and Comparison with the Advanced Monoethanolamine (MEA) Process referencea Technical Performance (Four Process Trains) heat transfer area of main heat exchanger, m2 heat transfer area of interheater, m2 condenser duty, GJ/tonne CO2 reboiler temperature, °C reboiler duty, GJ/tonne CO2 reboiler duty, MW total energy penalty, MW Capital Cost Performance (One Process Train) stripper reboiler, US$ stripper condenser, US$ rich/lean heat exchangerc, US$ lean solvent cooler, US$ interheating heat exchanger, US$ interheating pump, US$ change in capital cost, US$ change in total capital investment, US$ Economic Performance (Four Process Trains) capital cost, US$ million levelized cost of electricity, $/MWh CO2 avoided cost, $/tonne CO2

rich-split

interheating

combined process

advanced MEA processb

46,300 − 1.45 145.7 3.27 −100.1 −150

25,300 − 0.36 145.7 2.88 −87.4 −137.3

46,300 3,100 1.16 145.7 3.00 −91.6 −141.6

37,500 1,900 0.24 145.5 2.46 −75.3 −125.3

6,100 − 0.73 123.7 3.08 −78.4 152.5

2,209,963 423,018 13,536,100 460,712 − − 0

1,533,686 106,000 7,727,700 669,700 − − −6,592,707 −16,593,843

1,586,940 338,950 17,481,600 475,100 889,400 225,800 4,367,997 10,994,248

1,310,024 106,000 11,664,300 537,800 546,000 225,800 −2,239,869 −5,637,750

2,870,000 105,000 2,498,000 720,000 313,000 75,500 −526,420 −1,325,000

800.0 115.2 62.9

733.6 110.0 55.2

843.9 114.8 62.2

777.4 109.0 53.2

639.7 123.8 75.1

a

Reference case is based on two-stage absorption. bThe advanced MEA process includes the process modifications of absorber intercooling, richsplit, and stripper interheating process.40 cTemperature difference between solvent inside and outside stripper is 10 K. 10752

DOI: 10.1021/acs.est.6b02737 Environ. Sci. Technol. 2016, 50, 10746−10755

Article

Environmental Science & Technology

Figure 3. Breakdown of energy consumption, capital cost, and CO2-avoided cost in the ammonia-based postcombustion capture (PCC) process. The energy cost in the CO2-avoided cost breakdown is based on energy consumption and LCOE after the PCC plant is integrated with the power station.

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vaporization would decrease the CO2-avoided cost by 9.5%, due to the decrease in both capital costs and chilling duty in the NH3 recycle process. In addition to technical improvements, cheaper construction materials could reduce capital costs. For instance, the column could be constructed with concrete instead of carbon steel. This has been applied and demonstrated to be practical in the Boundary dam project.2 Packing materials of stainless steel have the potential to be replaced by cheaper plastic materials in the absorber and wash column, since the temperatures along these two columns range from 10 to 35 °C in the NH3 process, compared with 40 to 70 °C in the MEA process. It is estimated that replacing stainless steel with plastic packing could reduce material costs by 90%,50 cutting the CO2-avoided cost by 10.7%. It is worth mentioning that the technical development of the NH3 process is not isolated from the process itself and is therefore likely to have associated drawbacks. For instance, process modifications to reduce regeneration duty always require extra equipment, resulting in greater capital investment. Addition of NH3 suppressants and rate promoters demands expensive solvents, raising O&M costs. Replacement of stainless steel with plastic packing will likely deteriorate the mass transfer and subsequently CO2 absorption performance. Moreover, the proposed modifications will increase the process complexity and operation difficulties and thus increase O&M cost, project, and process contingency, etc., which is likely to offset the economic benefits from the technical improvements. Therefore, an overall investigation of technoeconomic performance combined with experimental trials of the proposed process modifications is essential to evaluate the pros and cons of each process modification and determine whether the modifications would advance the process.

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AUTHOR INFORMATION

Corresponding Authors

*Phone: +61-2-49606199. E-mail: [email protected] (K.L.). *Phone: +61-2-49606210. E-mail: [email protected] (H.Y.). Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors wish to acknowledge financial assistance provided through both CSIRO Energy and Australian National Low Emissions Coal Research and Development (ANLEC R&D). ANLEC R&D is supported by Australian Coal Association Low Emissions Technology Limited and the Australian Government through the Clean Energy Initiative. The views expressed herein are not necessarily the views of the Commonwealth, and the Commonwealth does not accept responsibility for any information or advice contained herein.

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REFERENCES

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.est.6b02737. Tables SI-S1−SI-S3 and Figure SI-1 (PDF) 10753

DOI: 10.1021/acs.est.6b02737 Environ. Sci. Technol. 2016, 50, 10746−10755

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DOI: 10.1021/acs.est.6b02737 Environ. Sci. Technol. 2016, 50, 10746−10755