Carbon Dioxide Liquefaction Process for Ship Transportation

Article pubs.acs.org/IECR

Carbon Dioxide Liquefaction Process for Ship Transportation Ung Lee,† Seeyub Yang,† Yeong Su Jeong,† Youngsub Lim,† Chi Seob Lee,‡ and Chonghun Han*,† †

School of Chemical and Biological Engineering, Seoul National University, Gwanak-ro 599, Gwanak-gu, Seoul 151-742, South Korea Korea Electric Power Corporation Engineering & Construction, #512, Yeongdong-daero, Gangnam-Gu, Seoul 135-791, South Korea

‡

ABSTRACT: CO2 liquefaction is an essential process for long-distance ship transportation. The conventional CO2 liquefaction process employs either an external coolant or liquid expansion followed by multistage compression to obtain liquefied CO2 at low pressure. However, these processes consume considerable amounts of energy, which presents an obstacle to commercialization. Thus, the CO2 liquefaction process needs to be carefully researched and designed to reduce the operating energy. In this study, two alternative CO2 liquefaction processes are proposed and evaluated. These alternative processes use multistage expansion and multistream heat exchangers to lower the input stream temperature for the compressor. In addition, the system is operated in a more efficient manner by operating the process with an optimized compression ratio. Evaluation of the economic feasibility was performed in this study for a complete assessment of the alternative processes. As a result, about 98.1 kWh/t of CO2 was consumed for alternative process 2, which is only 91.8% of the total operating energy of existing CO2 liquefaction processes, and the CO2 liquefaction costs for alternative process 2 were reduced by 5.5%.

1. INTRODUCTION Recent climate change and related calamities have attracted worldwide attention and brought about efforts to reduce greenhouse gases. Among the reduction technologies for greenhouse gases, carbon capture and sequestration (CCS) has been spotlighted because of its ability to reduce a large amount of CO2, which is suspected as the main culprit of global warming. CCS technology generally consists of CO2 capture, transportation, and sequestration. Most research to date has focused on the capture and sequestration processes because the transportation process has been considered as the least technologically challenging part. However, transportation approaches such as CO2 liquefaction, boiled-off-gas (BOG) handling systems, and intermediate storage systems for ship loading also require advanced technologies and could result in the use of large amounts of energy depending on the process operation. Among transportation processes, CO2 liquefaction consumes the largest amounts of energy, which can reach up to 10% of the total energy consumption in the entire CCS process. In addition to energy consumption, CO2 liquefaction is an essential process for CO2 ship transportation because liquidphase CO2 transportation is considered to be the most economically feasible. The CO2 liquefaction process can be divided into two distinct parts: the liquefaction process with external coolants and the process without external coolants. For the former process, Hegerland et al.1 and Decarre at al.2 proposed liquefaction systems employing ammonia or light hydrocarbons as coolants. However, these processes are considered to be inadequate for application to the processing of large amounts of CO2 because they use expensive external coolants and multiple heat exchangers. Most importantly, these processes have to remove a considerable amount of water in a dehydration unit because less water is removed in the gas scrubber at low pressure.3 Aspelund and Jordal4 and Desideri and Paolucci5 proposed CO2 liquefaction processes that use CO2 itself as a coolant, but their high energy consumption is © 2012 American Chemical Society

still an obstacle for commercialization. In this study, alternative processes for CO2 liquefaction are proposed and compared with the existing process. These alternative processes use multistage expansion and multistream heat exchangers to lower the temperature of the input stream for the compressor using cold vapor streams generated from the expansion. In addition, these alternative processes operate the system in a more efficient manner despite the same or smaller number of heat exchangers. Moreover, the optimal conditions for compressor operation were achieved by optimizing the compression ratio. To evaluate the process completely, evaluation of the economic feasibility of the processes was also carried out in this study.

2. DESIGN BASIS Aspen Plus with the Soave−Redlich−Kwong (SRK) equation of state was employed as the process simulator.6 A number of previous studies have shown that the SRK equation predicts the phase behavior of CO2 and CO2 mixtures with the greatest precision at high pressure.7,8 However, the SRK equation with improper water−CO2 binary interaction parameters often fails to predict their solubility.9 In this study, the SRK equation was modified with a binary interaction parameter (kij) of 0.193 in the van der Waals mixing rule, as reported by Heggum et al.10 The conditions of the product for the liquefaction process were set near 6.5 bar and −52 °C, based on the vessel pressure of CO2 transportation carriers. Currently, small-scale semipressurized ships are operating around the North Sea. Their transport pressure is 14−20 bar, and their CO2 transportation capacity varies between 1000 and 1500 m3.1 These ships are not suitable for transporting large amounts of CO2, and the pressure needs to decreased to increase the vessel size. Commercially available large-capacity tanks are usually Received: Revised: Accepted: Published: 15122

February 17, 2012 July 20, 2012 September 24, 2012 September 24, 2012 dx.doi.org/10.1021/ie300431z | Ind. Eng. Chem. Res. 2012, 51, 15122−15131

Industrial & Engineering Chemistry Research

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operated within the pressure range of 5−7 bar,11 and several studies indicate that 6.5 bar is a feasible pressure for CO2 ship transportation.2,3,12 It is also the operating pressure of the intermediate storage terminals that link the continuous liquefaction process to discrete ship transportation.13 The feed stream in this study is assumed to be the product stream of the postcombustion capture process specified by Aspelund and Jordal.4 It is mainly composed of CO2, water, and possibly trace amounts of amine solvent. SOx and NOx might be present in the flue gas. However, the postcombustion capture process can treat these gases until only a negligible amount remains. Volatile gases such as nitrogen or argon increase the required liquefaction energy, and they should be removed during or before the CO2 liquefaction process. With experience previously gained from other industrial projects, the CO2 product stream from the postcombustion process usually contains scarce amounts of volatile gases. However, additional distillation columns might be required when the CO2 feed stream is coming from a precombustion process where the flue gas contains relatively larger amounts of volatile gases. During the CO2 liquefaction process, the water content in the liquid stream should be removed to avoid formation of hydrates, freezing of water, and corrosion. Li et al. indicated that liquid CO2 containing 100 vppm of water lies on the liquid−hydrate equilibrium line near the target pressure and temperature.8 In this study, it was assumed that the water content in the liquid CO2 stream should be less than 50 vppm. This specification is far lower than necessary from a thermodynamic point of view. Aspelund and Jordal reported that hydrate formation and ice problems do not occur until the water content reaches 500 vppm.4 However, the water solubility of the liquid CO2 stream is lower than typical for the ship transportation because of the low temperature and pressure. Therefore, a maximum water content of 50 vppm was maintained to avoid any possible operational problems in the liquefaction process. To assess the process alternatives, all process input data were assumed to be the same as in the original work by Aspelund and Jordal.4 Table 1 lists the composition of the feed stream and the specifications of the process inputs. It is also assumed that cold cooling water (10 °C) is readily available. The input feed stream is the product stream of the amine-based postcombustion CO2 capture process, and its flow rate is 323 t/h, which is equivalent to 90% of the CO2 generation of a 600-

MW coal power plant.14 A log-mean temperature difference (LMTD) correlation factor of 0.85 for the seawater heat exchanger and isoenthalpic expansion for the Joule−Thomson valves were also assumed throughout the process simulation.

3. PROCESS DESIGN 3.1. Base Case. The base case of this study was the CO2 liquefaction process proposed by Aspelund and Jordal, which is denoted as process S14 (see Figure 1). The base case included a four-stage compressor, two process heat exchangers, and two multistream heat exchangers. The CO2 feed stream was compressed using the four-stage compressor. The pressurized CO2 feed stream was subsequently sent to a condenser and liquefied. Process heat exchangers lowered the liquid CO2 stream temperature. They applied cold CO2 streams as coolant either generated from the expanding fraction of the liquefied CO2 stream or recovered as a vapor stream from the product. Multistream heat exchangers lowered the temperature of the process stream entering the compressor and the process heat exchangers. The cold liquid CO2 stream was finally sent to a Joule−Thomson (J−T) valve, and a low-pressure liquefied CO2 stream was obtained through isoenthalpic expansion. The liquefied CO2 stream from the J−T valve was recovered as a product, and the effluents of the vapor stream from the expansion were used to lower the temperature of the liquid process stream and recycled back to the compressor. In the base-case design, there are three major approaches that can be applied to reduce the operating energy. First, the vapor CO2 recycle ratio sent to the second and third stages of the compressor can be reduced. The base case employed direct expansion from a high-pressure liquefied CO2 stream to a lowpressure CO2 product stream using a single expansion. Consequently, most of the CO2 vapor stream generated from the expansion was recycled back to second stage of the compressor. The CO2 vapor stream introduced at the second stage of the compressor eventually led to increases in the compressor work duties in the second, third, and fourth stages at the same time because it had to go through the third and fourth stages of the compressor. However, by applying multistage expansion, one can reduce the recycle ratio directed to the second stage of the compressor. For both process alternatives, multistage expansion was employed instead of using the fraction of liquefied CO2 as a refrigerant. Second, the location of the multistream process heat exchanger is also an important design variable. The amount of work provided to the compressor depends on the input stream temperature. The higher the input stream temperature, the greater the amount of compressor work consumed in general. Therefore, the location of the multistream heat exchanger, which lowers the temperature of the compressor input stream, needs to be carefully examined and optimized. The location and heat duty of the heat exchanger were optimized for both process alternatives 1 and 2. Finally, the compression ratio of each compressor stage was also optimized to achieve the minimum operating energy. 3.2. Alternative Process 1. Figure 2 shows the flow diagram of alternative process 1. In contrast to the base case, this process used multistage expansion and only cold CO2 vapor streams as the coolant. Cold vapor streams were generated from each expansion stage with J−T valves and used to lower the temperature of the CO2 stream. The output pressure of J−T valves 1 and 2 was determined from the input pressures of the fourth and third stages, respectively, of the compressor. J−T valve 3 expanded the product stream to 6.5

Table 1. Gas Quality and Input Data for CO2 Liquefaction Process parameter

value

units

94.39 5.61 trace

mol % mol % mol %

82 85

% %

0.5 5

bar °C

0.1 3

bar °C

Input Stream Composition CO2 H2O volatiles Process Input Data compressor efficiency pump efficiency Seawater HX pressure drop MITA (minimum internal temperature approach) Process HX pressure drop MITA (minimum internal temperature approach)

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Figure 1. Process flow diagram of base case of CO2 liquefaction (S1).4

Figure 2. Flow diagram of alternative process 1.

bar, and the CO2 vapor from the expansion was subsequently sent to J−T valve 4, which expanded the vapor stream (24) to

the input pressure of the second stage of the compressor. The expanded cold vapor stream (27) was recycled back to flash 2 15124



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Figure 3. Flow diagram of alternative process 2.

LMTD correction factor (FT).17 The LMTD correction factor can be expressed as

though process heat exchanger 4. The cold vapor stream (22) from J−T valve 2 was also directed to process heat exchanger 4, but stream 19 from J−T valve 1 could not be directed to process heat exchanger 4 because the temperature of the CO2 vapor stream increased to nearly the cooling water temperature as it passed through process heat exchanger 1. The location of process heat exchanger 4 is an important design variable. There are two possible designs in which the heat exchanger cools the output stream leaving either the condenser or the seawater heat exchangers. In the former case, the total operating energy can be reduced because the compressor discharge pressure decreases. In the latter case, the process heat exchanger is able to lower the temperature of the feed stream entering each compressor stage. The lower the temperature of the input stream, the less compression work required to achieve the same compression ratio in general. In this process, process heat exchanger 4 was installed to lower the temperature of stream 4, which enters the second stage of the compressor, to minimize the operating energy. Several liquid−vapor separation drums were installed to ensure that no entrained liquid entered the CO2 compressor. This is the simplest way to remove components with boiling points lower than that of gaseous CO2. In this process, water is the main component removed through the flash drums. The recovered water was subsequently sent to flash 1 to recover the CO2 dissolved in the water at high pressure. For a CO2 liquefaction process with an operating pressure higher than 30 bar, the specification of the water contents suggested by Austegard et al. can be only achieved through the flash drum separation without a dehydration column.4 However, the alternative process employed a dehydration column to decrease the water content to less than 50 vppm. A triethylene glycol (TEG) drying column15 or molecular sieve16 can be employed for the dehydration unit. The amount of seawater required for cooling and condensing the CO2 gas streams was calculated from the value of the

⎡ (1 − S) ⎤ R2 + 1 ln⎣ (1 − RS) ⎦ FT = ⎡ ⎤⎫ 2 ⎧ ⎪⎢ ⎣2 − S(R + 1 − R + 1 )⎥⎦ ⎪ ⎬ (R − 1) ln⎨ ⎡ ⎤⎪ 2 ⎪ ⎩ ⎣⎢2 − S(R + 1 + R + 1 )⎦⎥ ⎭

(1)

where

R= S=

Thot,in − Thot,out Tcold,out − Tcold,in Tcold,out − Tcold,in Thot,in − Thot,out

In typical heat exchanger applications, FT values of 0.85 or higher are desirable because small errors in R and S can result in FT values much lower than anticipated.18 Thus, an FT value of 0.85 was used throughout the process design, and the corresponding amount of seawater was calculated for cooling the gaseous CO2 vapor stream. The compression ratio of each stage of the compressor was determined by solving nonlinear programs min J = W1(r1) + W2(r2) + W3(r3) + W4(r4) q (ri)Fi with Wi = i η s.t.

1 ≤ ri ≤ 4

(2)

i = 1, 2, 3, 4

4

∏ ri = 51.8 i=1

Equation 2 shows the optimization of the objective function of the compression ratio, where W and r represent the compression work and the ratio, respectively, of each 15125



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CO2 vapor stream still remains at high pressure and can be directly recycled back to either the third or fourth stage of the compressor. Using this method, the operating energy of the compressor can be reduced by reducing the flow rate entering the second and third stages of the compressor. Table 2 illustrates the effects of multistream expansion. The vapor streams recycled to the third and fourth stages were

compressor stage. According to Seider et al., the maximum compression ratio of each stage can reach up to 4.18 The optimum compression ratio was calculated using the Matlab optimization solver with automated Aspen Plus by solving the nonlinear programming problem. 3.3. Alternative Process 2. Figure 3 shows the flow diagram of alternative process 2. In comparison with alternative 1, this process did not employ process heat exchangers for cooling the liquefied CO2 streams. With no process heat exchangers for the liquid CO2 streams, the amount of vapor generated from each J−T valve was greater than the amount in alternative process 1 because of the higher stream temperature. As a result, this process caused more of the CO2 gas stream to enter the CO2 condenser and required a CO2 condenser larger than that used in alternative process 1. The larger condenser eventually led to an increase in capital expenditures (CAPEX), even though the total number of process heat exchangers decreased. The process heat exchangers that decreased the temperature of the stream flowing into flash 1 and 2 utilized the vapor streams from the J−T expansions as a refrigerant. This necessitates multistream heat exchangers for the process heat exchangers. Spiral-wound, plate-frame, or brazed aluminum heat exchangers can be used for multistream heat exchangers.19 In this study, spiral-wound heat exchangers were used for the multistream heat exchangers, which is commonly used in the liquefaction of natural gas. The heat duty of process heat exchanger 1 was determined by the temperature of vapor stream 6, and the temperature of this stream was set at 3 °C. If the stream temperature is too low to maintain the vapor or liquid phase, problematic ice or hydrate formation can occur. Song and Kobayashi reported that icing problems would occur at −8 °C and 6.9 bar with 1/10 of the water content used in this study.20 Even though the solubility of water in CO2 gas increases at lower pressure and higher temperatures,21 the temperature of the stream was kept above the freezing point of water to maintain secure process operation. Process heat exchanger 2 decreased the temperature of stream 1 entering flash 1. Cold vapor stream 20 from J−T valve 2 was sent to process heat exchanger 2 after it had decreased the temperature of stream 5 to 3 °C in process heat exchanger 1. Cold vapor stream 17 from the J−T valve 1 was also supplied to heat exchanger 2 as a coolant. Process heat exchangers could be avoided in this process by mixing the cold vapor stream with the CO2 feed stream entering each the compressor stage. However, direct mixing of cold CO2 streams might cause problems such as the local formation of ice or hydrates.

Table 2. Vapor CO2 Recycle Introduced to Each Stage of the Compressor base case

recycle stage recycle stage recycle stage

to compressor 2 to compressor 3 to compressor 4

multistream expansion

flow rate (kmol/h)

fraction

flow rate (kmol/h)

fraction

1119.32

0.18

883.83

0.14

3693.50

0.60

2972.12

0.49

6111.11

1.00

6111.11

1.00

generated from the CO2 coolants for the base case. Compared with the base case, the molar flow rates recycled to the second and third stages of the compressor were reduced, whereas that of the fourth stage was increased. As a result, the total amount of vapor CO2 passing through the second and third stages was reduced, whereas that of the fourth stage of the compressor remained the same as in the base case. The operating energy of the compressor was decreased because of the energy reductions for the second and third stages, even though the energy consumption of the fourth stage remained relatively unchanged compared with the base case. Of the two possible locations for process heat exchanger 4 in alternative process 1, using it to decrease the output stream temperature from the seawater heat exchanger utilizes less operating work than using it to decrease the output stream temperature from the condenser. Figure 4 shows the change in the operating energy ovtained by varying the cooling duty of process heat exchanger 4 installed after the condenser. As the cooling duty of heat exchanger 4 increased, the temperature of cold vapor streams 28 and 29 also increased, and the discharge pressure of the compressor decreased. Consequently, the stream temperature entering each stage of the compressor increased. In Figure 4, when more cooling duty

4. RESULTS AND DISCUSSION Three major improvements were made for both process alternatives 1 and 2: employing multistage expansion, selecting the location of the multistream heat exchanger, and optimizing the compression ratio. The main advantage of multistage expansion is reducing the vapor CO2 stream recycle ratio directed to the second or third stage of the compressor. When the liquid CO2 product stream is generated from a single expansion, the CO2 vapor generated from the expansion has the same pressure as the CO2 product. This vapor stream has to recycle back to the second stage of the compressor and increases the compressor work duties in the second, third, and fourth stages. In contrast, if the product stream is generated from multistage expansion, a fraction of the

Figure 4. Operating energy change as a function of heat-exchanger cooling duty. 15126



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compression ratios were calculated using Matlab solver with the Aspen Plus simulation model by solving eq 2. The amount of work to achieve the same compression ratio generally increases as the starting pressure rises. Therefore, the optimized compression ratio decreased as the feed stream proceeded through the compressor. Figure 6 depicts the operating lines for alternative processes 1 and 2. The major difference between the two alternatives was

was exchanged in the heat exchanger, a higher operating energy was required. This demonstrates that the contribution of decreasing the stream temperature entering the compressor is greater than that of decreasing the discharge pressure of the compressor. Therefore, multistream heat exchangers were used for lowering the compressor input stream temperature in this study. In the proposed alternative processes, the total operating energy also varied depending on the location of the multistream process heat exchanger. As shown in Figure 5, cooling the

Figure 5. Energy consumption as a function of the location of process heat exchanger 4 in alternative process 1.

second-stage compressor feed reduced the operating energy the most. Generally, it is most efficient to cool the compressor stage with the highest compression ratio. The compression ratio of the first stage was slightly higher than that of the second stage. However, the high water concentration in the first-stage CO2 feed stream (1) caused the operating energy to increase because the water unnecessarily takes up a significant amount of the cooling duty. In alternative process 1, the cold vapor streams (20, 27) could not supply sufficient cooling duty for the process by heat exchanger 4 to decrease the hot-feed stream temperature to 3 °C. Therefore, only one multistream process heat exchanger was used to decrease the temperature of the CO2 gas stream entering the second compressor stage. However, there is surplus cooling duty in alternative process 2 even though it cooled CO2 gas stream 6 to 3 °C. Therefore, one additional process heat exchanger was also installed to decrease the temperature of CO2 gas stream 1 flowing into the first compressor stage. In this manner, the operating energy was reduced compared to the base case, which did not use an additional heat exchanger. The compression ratios and discharge pressures on each compressor stage are presented in Table 3. The optimized

Figure 6. Pressure−enthalpy diagram of CO2 and process operating line for (a) alternative process 1 and (b) alternative process 2.

the amount of cold CO2 vapor stream generated. In alternative process 1, cold vapor streams from the J−T valves were used to cool the liquefied CO2 stream, and these subcooled liquid CO2 streams generated less CO2 vapor than in alternative process 2. Cold vapor streams from process heat exchangers 2 and 3 were subsequently sent to process heat exchanger 4 to decrease the temperature of the second-stage compressor feed stream. The input temperature of the second compression stage was 5.9 °C and located near the 5 °C line in Figure 6a. In alternative process 2, cold CO2 vapor streams were used to decrease the temperatures of both streams 1 and 5. As a result, the workloads for the first and second stages of the compressor were reduced, which eventually resulted in the compressor utilizing a smaller amount of operating energy than in alternative process 1. The temperature of the CO2 vapor stream leaving process heat exchanger 1 was set at 3 °C to avoid the formation of ice or hydrates and that of the process heat exchanger 2 was 5.4 °C (Figure 6b). Figure 7 presents the total operating energy of each CO2 liquefaction system. Alternative processes 1 and 2 required 7.5% and 8.2% less operating energy, respectively, than the base case, and their energy consumptions were 98.9 and 98.1 kWh/t of CO2, respectively.

Table 3. Optimized Compression Ratio and Pressure on Each Compressor Stage alternative 1 stage feed stage 1 stage 2 stage 3 compressor out

alternative 2

pressure (bar)

compression ratio

pressure (bar)

compression ratio

1.0 4.0 13.0 30.0 51.8

4.0 3.7 2.4 1.8

1.0 4.0 12.9 29.2 51.8

4.0 3.7 2.4 1.8

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The calculation of the direct costs inside the battery limits (ISBL) is given by the equation min ISBL = x1 × equipment + x 2 × installation + x3 × instrumentation + x4 × piping + x5 × electrical s.t.

(4)

0.2 ≤ x1 ≤ 0.4 0.073 ≤ x 2 ≤ 0.26 0.025 ≤ x3 ≤ 0.07 0.03 ≤ x4 ≤ 0.15 0.025 ≤ x5 ≤ 0.09

Figure 7. Total operating energy consumption of CO2 liquefaction cases.

x1 + x 2 + x3 + x4 + x5 = 0.6

5. EVALUATION OF ECONOMIC FEASIBILITY For complete evaluation of the CO2 liquefaction process, information on both capital investments and operating costs were obtained from a number of resources.22−25 Economic depreciation was assumed to be 5% over 30 years.2 The cooling water cost was estimated as $0.013/m3 from the literature.18 Also, the price of electricity, $0.065/kWh, was assumed based on a 2010 U.S. Department of Energy report. The total cost of the CO2 liquefaction process can be calculated as follows cost of CO2 liquefaction ($/tonne of CO2) annualized CAPEX + OPEX = annual liquid CO2 production

where x1 is the fraction of equipment cost to fixed capital investment (FCI), x2 is the fraction of installation cost to FCI, x3 is the fraction of instrumentation cost to FCI, x4 is the fraction of piping cost to FCI, and x5 is the fraction of electrical cost to FCI. The value of the ISBL was assumed as 60% of the fixed capital costs based on the literature. Equation 4 is a simple linear programming problem and can be calculated easily. The inside the battery limits (OSBL) costs and indirect cost can be calculated in the same way, and their results are presented in Table 5. The total capital investments for the base case, alternative process 1, and alternative process 2 were about $65.87 million, $65.15 million, and $67.97 million, respectively. These values represent almost 45% of the CO2 total capital investment of the capture process identified by Abu-Zahra et al.26 Alternative process 2 required the highest capital investment, which was about 3% higher than that of the base case. The total capital investment can be varied depending on the xi values. Therefore, it is important to identify the lower and upper bounds of the capital investment and justify profitability of both alternatives 1 and 2 within this range. The upper bound of the capital investment was calculated using the same method as the minimum, but the objective function of eq 4 changed to max ISBL. Table 6 reports the minimum and maximum total capital investments obtained from the optimization problems. The maximum cost reached more than twice the minimum value. As the capital investment reached the maximum value, the cost savings of the alternative cases decreased. However, both alternatives 1 and 2 provide total cost reductions compared with the base case throughout the specified ranges. 5.2. Operating Costs. The total operating costs include production costs and general expenses. In Table 7, detailed estimates for each category are specified. Most of the factors that were used in the OPEX calculation were obtained from Peters et al.,24 except for the operating labor, supervision, and support labor costs, which were obtained from Rao and Rubin.25 The minimum total operating costs for the base case and alternative processes 1 and 2 were $29.29 million, $27.49 million, and $27.32 million per year, respectively. Because it had the lowest electric power consumption, alternative process 2 required the lowest total production costs among the investigated cases. As in the capital cost calculation, the maximum values of the operating costs were calculated using the same method applied for the total capital investment evaluation. The results of the

(3)

Table 4 summarizes the equipment costs for each case evaluated using the Aspen Process Economic Analyzer. The Table 4. Overview of Equipment Costs (M$) type of equipment

base case

alternative 1

alternative 2

water separation flash drum compressor dehydration column multistream process HX liquid CO2 HX compressor interstage cooler product separator CO2 condenser total

1.12 18.58 0.15 0.81 0.14 0.42 0.06 1.06 22.33

1.51 18.11 0.15 0.07 0.08 0.37 0.06 2.42 22.76

1.50 18.10 0.16 0.22 0.08 0.36 0.06 2.56 23.04

equipment costs of alternative processes 1 and 2 were higher than that of the base case because of the higher cost of the CO2 condenser. The condenser cost increased because of the increases in the CO2 vapor flow rate and the LMTD of the condenser drop. The compressor is the most expensive piece of equipment and can make up to about 80% of the total equipment cost. Auxiliary equipment such as valves, splitters, or mixers was not included in this study. 5.1. Capital Costs. Table 5 presents the minimum total capital investments of the CO2 liquefaction process. The direct and indirect costs were estimated using the factors specified by Douglas.22 As indicated in the table, each item has a wide range of costs. As a result, the total capital investment can differ depending on the factors used in the calculation. The factors specified in Table 5 were calculated to achieve the minimum total capital investment. 15128



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Table 5. CO2 Liquefaction Process Total Capital Expenditures (CAPEX) cost (M$) portion of FCI (%) ISBL purchased equipment purchased equipment installation instrumentation and control piping electrical

20−40 7.3−26 2.5−7.0 3−15 2.5−9.0

OSBL building and building services yard improvements services facilities land total direct costs

6−20 1.5−5.0 8.0−35.0 1−2

engineering construction expenses contractor’s fees contingency total indirect costs

4−21 4.8−22.0 1.5−5.0 5−20

fixed capital investment (FCI) working capital startup costs total capital investment (CAPEX)

1.00 10−20 8−10

used

Direct Costs 0.60 0.40 0.07 0.03 0.08 0.03 0.20 0.08 0.02 0.10 0.01 0.80 Indirect Costs 0.05 0.06 0.03 0.06 0.20 Total Costs 1.00 0.10 0.08

minimum

maximum

65.87 67.15 67.97

145.13 147.97 149.77

alternative 1

alternative 2

33.49 22.33 4.07 1.40 4.30 1.40 11.16 4.24 1.00 5.36 0.56 44.66

34.15 22.76 4.15 1.42 4.38 1.42 11.38 4.33 1.02 5.46 0.57 45.53

34.56 23.04 4.21 1.44 4.44 1.44 11.52 4.38 1.04 5.53 0.58 46.08

2.89 3.34 1.49 3.45 11.16

2.95 3.40 1.52 3.51 11.38

2.98 3.44 1.54 3.56 11.52

55.82 5.58 4.47 65.87

56.91 5.69 4.55 67.15

57.61 5.76 4.61 67.97

calculation are presented in Table 8. The maximum total operating cost was also found to be more than twice the minimum value. 5.3. Total CO2 Liquefaction Costs. Figure 8 shows an overview of the minimum CO2 liquefaction costs for the base case and both alternative processes. The total liquefaction cost for the base case was $10.51/t of CO2, in good agreement with the results of Aspelund et al.,3 who reported a value of $10.5/t of CO2. Alternative processes 1 and 2 had liquefaction costs of

Table 6. Minima and Maxima of the Total Capital Cost Investments (CAPEX, M$) base case alternative 1 alternative 2

base case

Table 7. CO2 Liquefaction Process Total Operating Costs (OPEX) cost (M$) range

used

base case

alternative 1

alternative 2

0.56 0.22

0.57 0.23

0.58 0.23

0.00 1.78 23.20 0.56 0.59 0.25 0.08 0.09 0.70

0.00 1.75 21.46 0.57 0.59 0.25 0.09 0.09 0.71

0.00 1.74 21.30 0.58 0.59 0.25 0.09 0.09 0.71

0.09 0.59 0.59 29.29

0.09 0.55 0.55 27.49

0.09 0.55 0.55 27.32

Production Costs fixed charges local taxes insurance direct production costs raw material cooling water electricity maintenance operating labor (OL) supervision and support labor operating supplies laboratory charges plant overhead costs administrative cost distribution and marketing R&D costs total production cost (OPEX)

1−4% of FCI 0.4−1% of FCI

1.0% 0.4%

1−10% FCI 1.0% 2 jobs per shift $33.7/h 30% of total labor costs 15% of maintenance 15.0% 15% of OL 15.0% 50−70% of (M + S + OL) 50.0% General Expenses 15−20% of OL 15.0% 2−20% of production costs 2.0% 2−5% of production costs 2.0%

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Notes

Table 8. Minima and Maxima of the Total Operating Costs (OPEX, M$) minimum

maximum

29.29 27.49 27.32

66.61 65.04 65.22

base case alternative 1 alternative 2

The authors declare no competing financial interest.

■

■

Figure 8. CO2 liquefaction costs.

$9.98 and $9.95, respectively, per tonne of CO2. Although the total production cost of alternative process 2 was reduced by about 8.2% compared to the base case, overall, the CO2 liquefaction cost decreased by only 5.5% and 5.1% for alternatives 1 and 2 because of the increase in capital investments. As both CAPEX and OPEX reached to their maximum values, the energy savings of the alternative processes gradually decreased. The maximum CO2 liquefaction costs were calculated as $23.83/t of CO2 for the base case and $23.39/t of CO2 and $23.49/t of CO2 for alternatives 1 and 2, respectively. The total cost savings for the alternatives were less than 2% for both cases. However, the results showed that the total CO2 liquefaction costs for the alternatives were less than that of the base case in all cases.

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NOMENCLATURE F = molar flow rate of the CO2 gas stream FT = log-mean temperature difference correction factor q = head r = compression ratio T = temperature W = compression work x1 = fraction of equipment cost to FCI x2 = fraction of installation cost to FCI x3 = fraction of instrumentation cost to FCI x4 = fraction of piping cost to FCI x5 = fraction of electrical cost to FCI η = isentropic efficiency of the compressor REFERENCES

(1) Hegerland, G.; Jørgensen, T.; Pande, J. O. Liquefaction and Handling of Large Amounts of CO2 for EOR. In International Conference on Greenhouse Gas Control Technologies (GHGT-7); Elsevier: Vancouver, Canada, 2004. (2) Decarre, S.; Berthiaud, J.; Butin, N.; Guillaume-Combecave, J. L. CO2 maritime transportation. Int. J. Greenhouse Gas Control 2010, 4 (5), 857−864. (3) Aspelund, A.; Mølnvik, M.; De Koeijer, G. Ship Transport of CO2: Technical Solutions and Analysis of Costs, Energy Utilization, Exergy Efficiency and CO2 Emissions. Chem. Eng. Res. Des. 2006, 84 (9), 847−855. (4) Aspelund, A.; Jordal, K. Gas conditioningThe interface between CO2 capture and transport. Int. J. Greenhouse Gas Control 2007, 1 (3), 343−354. (5) Desideri, U.; Paolucci, A. Performance modelling of a carbon dioxide removal system for power plants. Energy Convers. Manage. 1999, 40 (18), 1899−1915. (6) Soave, G. Equilibrium constants from a modified Redlich− Kwong equation of state. Chem. Eng. Sci. 1972, 27 (6), 1197−1203. (7) Coutinho, J. A. P.; Kontogeorgis, G. M.; Stenby, E. H. Binary interaction parameters for nonpolar systems with cubic equations of state: A theoretical approach 1. CO2/hydrocarbons using SRK equation of state. Fluid Phase Equilib. 1994, 102 (1), 31−60. (8) Li, H.; Yan, J. Evaluating cubic equations of state for calculation of vapor−liquid equilibrium of CO2 and CO2 mixtures for CO2 capture and storage processes. Appl. Energy 2009, 86 (6), 826−836. (9) Austegard, A.; Solbraa, E.; De Koeijer, G.; Mølnvik, M. Thermodynamic Models for Calculating Mutual Solubilities in H2O−CO2−CH4 Mixtures. Chem. Eng. Res. Des. 2006, 84 (9), 781− 794.

4. CONCLUSIONS Alternative CO2 liquefaction processes for the ship transportation were proposed in this study. The product pressure of the liquid CO2 stream was 6.5 bar, which is the operating pressure of storage terminals and CO2 carriers. The alternative processes used multistage expansion and multistream heat exchangers to lower input stream temperatures for the compressor using the cold vapor stream generated from the expansion. Optimized operating conditions for each process were also obtained by solving nonlinear programming problems. Economic feasibility was also evaluated in this study. Despite having the highest CAPEX value, alternative process 2 produced liquid CO2 with the lowest cost because of its small OPEX. With alternative processes 1 and 2, the operating energy was reduced by 7.5% and 8.2%, respectively. As a result, the costs of CO2 liquefaction for the alternative processes were reduced by 5.1% and 5.5% compared to the base case.

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ACKNOWLEDGMENTS

The authors gratefully acknowledge the Korea Science and Engineering Foundation for support provided through the Advanced Environmental Biotechnology Research Center (No. 2011-0001114); the Brain Korea 21 Project initiated by the Ministry of Education of Korea (ME); the Energy Efficiency & Resources and Human Resources Development of the Korea Insitute of Energy Technology Evaluation and Planning (KETEP) grant funded by the Ministry of Knowledge Economy (MKE); the Industrial Strategic Technology Development Program “Design of Topside LNG Regasfication Plant of LNG FSRU (10031883)” by the MKE; and the LNG Plant R&D Center funded by the Ministry of Land, Transportation and Maritime Affairs (MLTM) of the Korean government.

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Carbon Dioxide Liquefaction Process for Ship Transportation

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