Conceptual Process Design of CO2 Recovery Plants for Enhanced Oil

Article pubs.acs.org/IECR

Conceptual Process Design of CO2 Recovery Plants for Enhanced Oil Recovery Applications Dong-Hun Kwak,† Donghyun Yun,‡ Michael Binns,† Yeong-Koo Yeo,† and Jin-Kuk Kim*,† †

Department of Chemical Engineering, Hanyang University, Wangsimni-ro 222, Seonggdong-gu, Seoul, 133-791, Republic of Korea Plant Basic Engineering Team, GS Engineering and Construction, Gran Seoul 33, Jongno-gu, Seoul, 110-121, Republic of Korea

‡

S Supporting Information *

ABSTRACT: Processes for recovering CO2 (carbon dioxide) from CO2-rich gas are important in the CO2 EOR (enhanced oil recovery) field. From an environmental point of view EOR through the injection of CO2 is very beneficial because it allows for the storage of part of the CO2 injected while increasing oil recovery. To make this process even more environmentally friendly, the fraction of CO2 which exits (is produced from) the oil well can be captured and recycled for reinjection giving a more efficient CO2 storage strategy. The mixture of gases produced from an oil well contains light hydrocarbons, heavy hydrocarbons, water, and CO2. Dehydration units and numerous other separation units for separating CO2 and hydrocarbons can be used to recover CO2, so various potential configurations should be investigated to find the one which is most appropriate. In this study, the TEG (triethylene glycol) and adsorption dehydration processes are used for gas dehydration and a combination of amine, Selexol and distillation processes are used for CO2 separation. Unisim is used to simulate the processes and they are evaluated economically in terms of plant installation costs and energy consumption. A case study is presented to demonstrate the feasibility of various design configurations.

1. INTRODUCTION Enhanced oil recovery (EOR) has been widely used in oil production to increase the production rate from oil wells, and with the aid of EOR approximately 30 to 60% more oil can be extracted.1 One method of EOR is to inject CO2 into the oil reservoir. In a high pressure reservoir, CO2 can act as an effective fluid for EOR because it is miscible with the oil, and injected CO2 makes oil swell up, resulting in reduced viscosity and surface tension.2 EOR through the injection of CO2 is also applied for the sequestration of captured CO2, because some fraction of the injected CO2 remains in the reservoir during the EOR operation. Hence, in addition to the economic benefits resulting from increased oil recovery the process also has environmental benefits related to the removal and storage of CO2 which may have been created by other processes such as those in the industrial energy sector. As part of the CO2 EOR operation, typically more than 50% of injected CO2 comes back to the surface together with the oil and gas products.2 Hence, a ground facility for the recovery of CO2 is necessary to reduce CO2 emissions and CO2 purchasing costs. The environmental benefits of this process are also enhanced as a result of the recovery and reinjection of CO2 which then gives a more efficient CO2 storage strategy with minimal CO2 released to the atmosphere. The operational strategy for CO2 recovery process applied in CO2 EOR considered in this paper is divided in two phases: phases I and II (Figure 1). In the early years of operation at a CO2 EOR plant, the amount of gas produced from the oil well and its composition will change significantly. In this situation and under these conditions the production of hydrocarbons is likely to be economically infeasible. Thus, in phase I, dehydration units are operated to remove water from the gas mixture and compression processes are introduced to inject low © 2014 American Chemical Society

Figure 1. Two-phase arrangement for CO2 EOR ground facility.

purity CO2 back into the oil well. Phase II is defined as the period starting from the year when gas compositions stabilize and do not vary significantly. Hydrocarbon products and high purity CO2 product are produced during phase II and so additional processing facility for CO2 separation in phase II operate in parallel with the dehydration and compression processes used in phase I, as shown in Figure 1. The gas mixture which is produced by oil wells using CO2 EOR contains considerable amounts of CO2, its composition rises quickly to above 30% and often stays in the range of 70− 90%.3 Received: Revised: Accepted: Published: 14385

May 22, 2014 August 10, 2014 August 26, 2014 September 5, 2014 dx.doi.org/10.1021/ie502110q | Ind. Eng. Chem. Res. 2014, 53, 14385−14396

Industrial & Engineering Chemistry Research

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to provide conceptual design guidelines. CO2 recovery processes using distillative separation are considered for separating CO2 due to their capability to deal with a wide range of CO2 compositions in a feed mixture.3 Also, hybrid processes which use distillation followed by solvent processes are considered, in which bulk CO2 separation is carried out using distillation, reducing the CO2 composition and concentration levels so that they are suitable for separation using a solvent-based process. In the case study considered, different configurations of CO2 recovery using combinations of distillation and solvent-based processes are introduced and their economics are investigated. To identify the energy saving associated with hybrid processes, the simulation program Unisim is used to model the CO2 recovery processes in a given situation. To calculate the economics of each configuration and to compare the different options, the estimation of capital costs is based on information given by Peters et al.,12 and utility costs are estimated using information from Smith and Varbanov.13

Therefore, CO2 recovery processes often require a series of separation units consuming considerable amounts of energy. As a result, the design of a CO2 recovery process is not straightforward due to the nature of the complex design interactions involved. However, it is important to improve energy efficiency and to improve the economics of the CO2 EOR ground facilities. There are many studies conducted for separating CO2 using a wide range of chemical and physical solvents. Amine-based chemical solvents, for example, monoethanolamine (MEA), diethanolamine (DEA), and methyldiethanolamine (MDEA) separate CO2 using chemical reactions, which require large amounts of heat for recovering and regenerating the solvent. Physical solvents, for example Selexol and Rectisol, remove CO2 by dissolving it in to the liquid phase.4 They require less energy for solvent recovery but use higher operating pressures, and hence they are more suitable for gases with a high concentration of CO2. However, physical solvents are also not appropriate for sweetening acid gas which contains CO2 concentration above 70%. For EOR applications, CO2 recovery processes should deal with high inlet CO2 compositions, and in these cases distillation columns are mainly used for CO2 separation.3,5−8 The separation of CO2 through distillation has two problems: solidification of CO2 and the azeotrope between ethane and CO2. Solidification of CO2 due to freezing can be prevented by the injection of additives5 or using a specially modified column with a CO2 freezing zone.6 For example, to overcome the azeotrope between ethane and CO2, additives can be introduced to change the composition in the distillation column.7 The Ryan/Holmes process is a typical CO2 recovery process based on distillation,3,8 which autonomously overcomes both problems using produced C4+ components as an additive. CO2 separation using only the Ryan/Holmes process can be implemented without the addition of other solvent processes, however the distillation operation requires considerable energy consumption. According to Jung et al.,9 the operating energy for distillation is relatively high compared to solvent processes. Thus, hybrid processing combining distillation and solvent processes are expected to have lower energy consumption depending on the particular hybrid configuration of the two separation technologies. For CO2 EOR applications at the industrial scale, it is often necessary to combine different separation technologies together. For example a distillation process combined with a solvent-based process might be used rather than employing a single unit operation or process. This might be necessary because of a number of factors including the high concentration of CO2 in the feed or requirements placed on the recovered gas products, or due to economic factors affecting plant efficiency. Although there have been a number of studies considering the combination of different separate CO2 methods,10,11 most of these studies consider the use of a membrane combined with another separation technology, and they mainly focus on CO2 recovery from inlet feed gases with lower fractions of CO2 compared with the feed gas streams used for CO2 recovery as part of the CO2-injection EOR case. For example Peters et al.10 combine membranes with absorption systems for CO2 removal from natural gas and Belaissaoui et al.11 combine membranes with cryogenic processes for CO2 capture. This paper aims to screen different technologies and configurations for CO2 recovery in CO2 injection EOR applications, to evaluate their techno-economic impacts, and

2. DEHYDRATION AND CO2 RECOVERY PROCESSES 2.1. Gas Dehydration. When a gas mixture is produced from an oil well it usually contains water, which is either at saturated conditions or at the water dew point. This water can cause freezing problems and form hydrates in the cold sections of the process.14 Also, when the water is liquefied during processing, it can contain acid gases such as H2S (hydrogen sulfide) and CO2 which cause corrosion. Thus, water removal is necessary to prevent equipment breakdown. TEG (triethylene glycol) is one of the most common solvents used to remove water from wet gas. Figure 2 shows a

Figure 2. TEG dehydration process.

typical TEG dehydration process scheme. In the absorber, TEG selectively absorbs water and is withdrawn from the bottom of the absorber while dry gas exits the absorber at the top. Waterrich TEG is subsequently regenerated in a distillation column (desorber) and the temperature of reboiler in this column is typically limited to 204 °C to avoid TEG thermal degradation.14 As an alternative, adsorption dehydration processes (e.g., molecular sieve, silica gel) can be used when the target dew point of the dry gas is lower than that achievable through TEG dehydration, and they can also be used when the wet gas has a low water content.15 For example, a molecular sieve facilitates separation between water and other gas molecules based on their size differences. Water molecules with smaller size pass through the pores in the molecular sieve while larger molecules cannot penetrate. To model the thermodynamic properties of TEG and the dehydration process, the Peng−Robinson equation fluid package is utilized. This simulation model is compared against 14386

dx.doi.org/10.1021/ie502110q | Ind. Eng. Chem. Res. 2014, 53, 14385−14396


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Table 1. Feed Compositions for Model Validation composition (mol %) C1 C2 C3 nC4 iC4 nC5 iC5 nC6 nC7 nC8 nC9 nC10 nC11 nC12 N2 CO2 H2S H2O flow rate

TEG dehydration16 92.28 3.54 1.94 0.58 0.34 0.16 0.18 0.14 0.07 0.02 0.003 0.0005 4.4 × 10−05 1.1 × 10−05 0.29 0.39 0.06 2068.0 kmol/h

MDEA process17

Selexol process4

85.05 5.44 2.01 0.57 0.37 0.15 0.17 0.19

70.18 0.83 0.22 0.08 0.05 0.03 0.03 0.11

25.7 7.9 6.6 3.1

3.47 1.83 0.69 0.06 26698.0 kmol/h

0.45 27.78 0.001 0.02 15831.1 kmol/h

4.9 50 0.18

a case with results in the literature16 in which the temperature, pressure, and flow rate of feed in this case are 25 °C, 71 bar (g) and 37 t/h, respectively. The composition of this inlet and the specifications for the columns are shown in Tables 1 and 2.

absorber

desorber

4 1 70.75

7 2 1.05

1.3 0.32

907.2 kmol/h

temperature and composition profiles inside the absorber (Figure 3) and the desorber (Figure 4). 2.2. CO2 Separation. High concentrations of CO2 in the gas mixture extracted from CO2 EOR oil wells make the separation more difficult and hence considerable amounts of energy are needed for the separation. For this reason it is necessary to understand the features of each CO2 separation technology and to select sustainable and economic configurations. In this paper, three typical processes are considered: amine (chemical solvent), Selexol (physical solvent), and distillation processes are considered for CO2 separation, which are all proven processes widely used in industry. Each technology has its own separate characteristics in terms of their handling of different CO2 contents, their operating ranges, and energy consumption. In addition to CO2 recovery, separation of H2S should also be carried out simultaneously to meet the hydrocarbon product specifications. In many cases, CO2 and H2S walk the same path inside acid gas separation processes.

Table 2. Column Specification of TEG Dehydration Process16 number of tray feed tray outlet pressure [bar(g)]

Ryan/Holmes process8

Additionally the temperatures inside the desorber are set to 100 °C in tray 1 and 204 °C in tray 4, and the inlet temperature of desorber is set to 175 °C. This model is shown to produce results very similar to the reference data considering the

Figure 3. Comparison between ref 16 and simulation data inside the absorber of the TEG dehydration process: (a) temperature profile, (b) liquid mole fraction profiles. 14387



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Figure 4. Comparison between ref 16 and simulation data of the desorber in TEG dehydration process model: (a) temperature profile, (b) liquid mole fraction profile.

2.2.1. Amine Process. Amine processes use amine solvents to separate CO2 from an acid gas mixture. Chemical reactions occur between the amine and CO2 which allow these processes to catch CO2 even when the gas has a low CO2 content. However, large amounts of energy are required to regenerate the amine solvents due to the strong bonds formed between amine and CO2. Therefore, amine-based processes are appropriate when the gas contains small amounts of CO2, but are not recommended for bulk separation of CO2. Amine solvents which absorb CO2 from the mixture are recovered in a desorber (distillation column), using typically low pressure steam (Figure 5). Additionally in this process, H2S is also absorbed into the amine solvents and it is stripped from the amine solvents together with CO2.

Table 3. Comparison between Reference Data and Simulation Results for the MDEA Amine Solvent Process17 amine type lean/rich amine concn (wt %) circulation rate (m3/h) gas flow (Sm3/h) H2S out (ppm vol) CO2 out (ppm mol) total rich loading (mol/mol) reboiler duty (MW) absorber top/bottom pressure (bar) absorber top/bottom temperature (°C) desorber top/bottom pressure (bar) desorber top/bottom temperature (°C)

ref

Unisim

MDEA 47.5/43.1 305.1 629382 1 8350 0.378 22.85 65.9/66.3 43.8/36.4 2.35/2.6 109.7/131.5

MDEA 47.49/44.95 304.69 629382 0.85 8287 0.361 23.97 65.9/66.3 43.54/36.13 2.35/2.6 114.43/131.6

chemical reaction to absorb CO2, and so they can be regenerated using only flashing which requires less energy compared with amine (chemical solvent) processes (Figure 6).

Figure 5. Schematic diagram of the amine process.

The MDEA solvent process model for removal of CO2 and H2S is validated here through a comparison between the model and experimental data from the literature.17 The absorber feed composition from the reference is shown in Table 1 and the specified number of trays in the absorber and desorber are 34 and 25, respectively. For simulation the amine package is used to predict the thermodynamic properties of MDEA. Table 3 shows the comparison between the reference data and the simulation study results for MDEA process using the feed conditions from Table 1. It can be seen that the model reproduces the reference data with reasonable accuracy. 2.2.2. Selexol Process. The Selexol process is a CO2 recovery process using physical solvent DEPG (dimethyl ethers of polyethylene glycol). Physical solvents do not rely on any

Figure 6. Schematic diagram of the Selexol process.

As with the amine solvent, H2S is removed from the gas together with CO2 simultaneously. Thus, the Selexol process can be implemented to separate CO2 from mixtures which have a relatively high CO2 content compared to those using an amine solvent. However, this process has some difficulties when the specifications require a product with very low CO2 composition.18 The Selexol process is also known to be unable to remove CO2 from a natural gas liquid (NGL) stream 14388



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conditions. These equations are used to check for the existence of solid CO2 after distillation column simulations and to avoid column operating conditions under which CO2 freezing occurs. When high purity methane is separated from CO2-rich gas, the top of the distillation column can reach very low temperatures (−90 °C ≈ −70 °C), causing CO2 solidification. One way to prevent CO2 solidification is with the addition of extra additive5 and another way is to use a distillation column modified to overcome the CO2 freezing problem.6 Also, azeotrope issues can occur when CO2 is separated from NGL components, in particular ethane. In this case, additional solvent is required to break the azeotropic conditions. One method to achieve this is to use a slip stream of C4+ products as the solvent. This method is called extractive distillation.7 The Ryan/Holmes process is a typical distillation process, in which extractive distillation is used to overcome the CO2 freezing issue and the azeotrope existing between ethane and CO2 while producing methane, CO2 and NGL.3 Figures 7 and

because the solubilities of C3+ components are higher than that of CO2, causing difficulties in the separation between NGL and CO2.4 Validation of the Selexol process model is carried out using a comparison with reference data using a feed gas specified in the literature (Table 1).4 The temperature and pressure of feed gas are set to 48.9 °C and 66.6 bar(g) and in the absorber the number of trays and the pressure are set to 10 and 60.4 bar(g). Furthermore, in this case (based on the reference) four flash tanks are used, operating at 33.5 bar(g), 17.1 bar(g), 0.71 bar(g), and −0.67 bar(g). The result of modeling the Selexol process is shown in Table 4. For simulation of the Selexol Table 4. Comparison between Reference and Simulation Data in Selexol Process Modeling4 Residue Gas CO2 C1 Atmospheric Flash Gas CO2 C1 Vacuum Flash Gas CO2 C1

ref (kmol/h)

Unisim (kmol/h)

395.7 10721.4

393.2 10713

3490.5 388.1

3497.5 385.4

513.0 0.7

515.1 7.8

process, the Glycol fluid package is used to describe the behavior of Selexol (DEPG). The model results are compared against results from the literature in Table 4 where the flow rates of CO2 and methane contained in streams are shown to be similar to the reference values for the Selexol process with the exception of the methane flow rate in vacuum flash gas which is predicted to be slightly higher than the reference. This difference is due to an accumulation of errors occurring in the front end of the vacuum flash gas production. However, as this difference is only small we can say that this model is suitable to represent the Selexol process. 2.2.3. Low Temperature Distillation. The main advantage of the distillative approach for recovery of CO2 is its ability to handle acid gas with any CO2 content.5 In particular it is useful in the case of CO2 EOR where the CO2 composition in the gas feed can be very high and changes significantly during the EOR period of operation. However, it requires a refrigeration cycle to separate CO2 from the hydrocarbons. Also, this method has two potential obstacles which may cause difficulties during the operation: (1) CO2 solidification, (2) azeotrope between ethane and CO2. The equations for calculating the freezing temperature of CO2 are given by19

Figure 7. Three-column Ryan/Holmes scheme.

Figure 8. Four-column Ryan/Holmes scheme.

8 show the Ryan/Holmes process using 3 and 4 column arrangements, respectively. The differences between these two configurations include different separation sequences and the existence of a CO2 recovery column in the four-column arrangement. In the three-column process, methane is separated in the first column (demethanizer) and CO2 is removed in the second column (ethane recovery column), while NGL/H2S and additives are extracted in the third column (additive recovery column). However, in the four-column process, the first column (ethane recovery column) acts to separate ethane and CO2, and as a result the top product consists of methane and CO2, while the remaining components are removed at the bottom of this column. CO2 is partially removed in the CO2 recovery column but the top products still contain approximately 15−30% of CO2. This CO2 is fully separated in the demethanizer, and the separated CO2 is

⎤ ⎡ VCO2solid L sat sat sat exp ( P P ) xCO2 ΦCO P = P Φ − ⎥ ⎢ CO solid CO CO solid 2 2 2 2 ⎦ ⎣ RT (1)

⎤ ⎡ VCO2solid V sat sat sat exp⎢ (P − PCO )⎥ yCO ΦCO P = PCO ΦCO 2 2 solid 2 2 solid 2 ⎦ ⎣ RT (2)

In these equations (referring to CO2) x and y are the liquid and gas mole fractions, φL and φV are liquid and gas fugacity coefficients, P refers to the pressure, VCO2solid refers to molar volume of solid CO2, R is the universal gas constant, T is the temperature, and sat refers to saturation (or sublimation) 14389



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dehydration process. However, the CO2 recovery process at a CO2 EOR ground facility treats gas mixtures containing large quantities of CO2 and consumes considerable amounts of energy for the CO2 separation. Hence, there is a demand for the design of different configurations which can reduce the energy requirements in the whole process. In this study, amine, Selexol, and distillation processes are considered for the design of CO2 recovery processes. The operational plan for the CO2 recovery process is specified with a planned 14 year period of operation for CO2 recovery which is divided into two phases. In phase I, the early years of CO2 EOR, only dehydration and compression processes are used to reinject the gas mixture produced from the oil well. This is accomplished using a TEG dehydration process. In phase II, the CO2 recovery processes are operated in parallel with the processes operated in phase I and therefore high purity CO2 and low purity CO2 are both reinjected simultaneously. For this case study, four different configurations of CO2 recovery equipment are simulated with Unisim using different combinations of the models validated in section 5. Subsequently comparisons are made to identify which options are the most beneficial. An adsorption dehydration process is additionally introduced in order to satisfy water content specification for products in phase II. To simplify the modeling in phase II, the adsorption process is modeled as a splitter in Unisim. The composition and flow rate of the gas mixture produced from the oil well change every year in the CO2 EOR period. The CO2 content of the gas mixture increases drastically during the first 5 years and then stabilizes but still increases continuously. It is shown that the flow rate of the gas mixture increases every year until year 11 and then decreases after this peak (Tables 7 and 8). The temperature and pressure of feed

recycled back to the ethane recovery column. Finally, the additive recovery column produces NGL/H2S and additive, the same as the additive recovery column in the three-column configuration. Although one additional column is required in a four-column process, it has advantages including a higher temperature in the condenser of the demethanizer column and the fact that it produces liquid CO2 (where the three-column arrangement produces gas-phase CO2). Because of this higher condenser temperature, the four-column configuration requires less energy compared to the three-column configuration, and it has a lower risk of CO2 freezing in the column. Also, producing liquid-phase CO2 is beneficial for the CO2 EOR process because the separated CO2 is pressurized before being injected into the oil well, and in this situation using CO2 in the liquid phase consumes less power compared with using the vapor phase CO2. For validation of the Ryan/Holmes process model a comparison is made between US patent data8 concerning a three-column Ryan/Holmes process and the simulation model constructed using Unisim. The feed gas composition and the column specifications for the reference case are given in Tables 1 and 5. In the simulation model the Peng−Robinson equation Table 5. Column Specifications for the 3-Column Ryan/ Holmes Process8

number of tray inlet tray operation pressure (bar(g))

demethanizer

ethane recovery column

20 12 (feed) 1 (additive) 41.4

50 42 (feed) 5 (additive) 20.7

additive recovery column 30 15 17.2

is used to describe the component properties affecting the process. A comparison between the simulation results and the reference data is given in Table 6 where it can be seen that there is a good agreement between the model results and the literature values.

Table 7. Yearly Flow Rate and CO2 Content of Gas Produced

3. CASE STUDY The two main components of the CO2 EOR ground facility are the dehydration process and the CO2 recovery process. In the case of dehydration, there are no combinatorial complexities associated with the choice of using either an adsorption or TEG Table 6. Comparison between Reference Data8 and Simulation Results for the Ryan/Holmes Three-Column Process

Methane Product C1 CO2 C2 CO2 Product C1 CO2 C2 NGL Product C1 CO2 C2

ref (mol %)

Unisim (mol %)

81.64 1.7 0

81.96 1.66 0

1.8 94.3 3.8

2.03 93.39 3.79

1.1 0.8 33.6

0.3 0.88 34.59

year

gas flow rate (MMSCMD)

CO2 (mol %)

phase

1 2 3 4 5 6 7 8 9 10 11 12 13 14

0.4 0.8 2.0 4.1 6.1 8.1 10.0 11.3 12.1 12.6 12.7 12.6 12.1 11.8

8 21 56 72 77 80 82 82 82.5 84 85 87 88 90

I I I I I II II II II II II II II II

gas are set to 40 °C and 20 bar(g). The target product specifications for hydrocarbon (HC) gas are below 3 mol % of CO2, 4 ppm of H2S, and 112.1 kg/mmscm of water. Also the pressure and heating value of HC gas should be above 69 bar(g) and 37.3 MW/scm, respectively. The target specifications for NGL product are 60−90% of C2 recovery, with a minimum 95% recovery of C3 and H2S content below 50 ppm. The high purity CO2 product which is separated in the CO2 recovery process should be pressurized up to 172 bar(g) and should contain less than 100 ppm of H2S. To perform the 14390



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Table 8. Detailed Composition of the Gas Produced in Years 3, 5, and 11 C1 C2 C3 iC4 nC4 iC5 nC5 C6+ H2S CO2 H2O

3 year (mol %)

5 year (mol %)

11 year (mol %)

19.7 10.6 7.9 1.4 1.4 0.8 0.8 0.8 0.6 56 saturated

10 5.4 4 0.7 0.7 0.4 0.4 0.4 1 77 saturated

5.3 2.9 2.6 0.6 0.7 0.26 0.27 0.57 0.6 86.2 saturated

simulation, the ambient temperature is assumed to be 30 °C with a minimum temperature difference in heat exchangers set to be 10 °C for temperatures above ambient and 5 °C for refrigeration utility, respectively. Three different fluid packages are used to model the three different CO2 recovery methods (amine, glycol, and Peng−Robinson fluid packages) as mentioned in sections 2.2.1−2.2.3. Additionally the refrigeration duty is supplied only using a simple propane cycle in all cases. 3.1. Phase I. Before simulating and comparing different CO2 recovery processes, a TEG dehydration and compression model operated in phase I is simulated to calculate the overall energy consumption for a CO2 EOR ground facility during this phase. The feed gas from year 5 is used for phase I simulation because this has the maximum flow treated in phase I, and it is assumed that this process continuously treats this same amount of gas mixture in phase II. The number of trays and the pressure in the absorber are set to 4 and 70 bar(g), and after simulation this process is found to consume 0.73 MW of steam at 204 °C, 19.46 MW of shaft power, and 46.11 MW of cooling water. In this process large quantities of shaft power and cooling water are required to be used for the compression of the dry gas for reinjection. 3.2. Phase II: Design 1. Among the CO2 recovery methods mentioned in section 2.2, the Ryan/Holmes process is the only option which allows both the separation of CO2 and the generation of hydrocarbon products. For this reason the Ryan/ Holmes process will be considered as the first design option and the base case against which other configurations will be evaluated and compared. Here the four-column process is operated in moderate temperature conditions that allow a lower energy consumption than the three-column arrangement, and so it is considered to be a more appropriate base case for comparison with other designs.3 This first design is completed with the addition of an amine solvent process to remove H2S contained in the NGL product. The amine solvent MDEA is chosen here, which requires relatively lower regeneration energy compared with other amine solvents such as MEA or DEA. In all designs considered in this work the CO2 recovery processes are modeled using the feed gas from year 11 of operation (as this year has the maximum flow within phase II). So the flow rate of feed gas passed to the CO2 recovery equipment in phase II is 6.7 MMSCMD. The process diagram showing the configuration of design 1 is presented in Figure 9 and column operating conditions are shown in Table 9 (detailed stream data is given in Table S1 in the Supporting

Figure 9. Design 1 process diagram.

Information). The operating pressures of each distillation column are determined on the basis of the values from the literature8 plus those calculated through simulation. In particular, the reference pressures are in the ranges: 27.58 to 44.82 bar(g) for the demethanizer, 20.68 to 34.47 bar(g) for the ethane recovery column, and 6.895 to 20.68 bar(g) for the additive recovery column. In this study the demethanizer and CO2 recovery column pressures are set equal to 44.82 bar(g). These high values are used because they allow lower condenser temperatures which enhances the separation efficiency between methane and CO2. Hence, in this design the condenser temperatures are set to −33 °C, which is the lowest feasible temperature when using a propane refrigeration cycle. The additive recovery column is operated at 30 bar(g) which is higher than the reference values, and this is chosen to allow the use of cooling water in the condenser. This pressure is higher than the range given in ref 8, but pressures higher than this range have also been considered in the literature.7 The additive recovery column consumes large amounts of cooling duty for refrigeration when operated at lower pressures, giving very large operating costs. However, at higher pressures cooling water can be used for cooling which greatly reduces costs. Also, the required extra pump duty is relatively small in this case making a higher pressure more economical, reducing the overall energy costs. In the ethane recovery column the pressure is set to that of the feed gas, which is only slightly lower than the range of values in the reference. The MDEA process is used here to remove acid gas, and the separation efficiency here increases as the temperature of the absorber inlet stream decreases. So the temperature of inlet gas is fixed to 40 °C which is the lowest possible temperature when cooling water is the only source of cooling. Also, the desorber in the MDEA process operated at 1 bar(g). The flow rate of additive stream is calculated to satisfy the H2S content in the HC gas product. At the ethane recovery column C2+ components are separated from the inlet gas. Also, in this column the azeotrope between ethane and CO2 is overcome through the addition of additives. The stream A-1, containing methane and CO2 is passed to the CO2 recovery column to produce CO2. Highpurity CO2 is produced in the column and pressurized for reinjection into the oil well. Stream A-3 which comes from CO2 recovery column contains both methane and CO2. In the 14391



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Table 9. Design 1 Column Specifications ethane recovery column number of tray feed tray operation pressure [bar(g)]

50 25 (feed) 5 (additive) 20

CO2 recovery column 30 15 44.82

demethanizer 10 5 (feed) 1 (additive) 44.82

additive recovery column

absorber in MDEA process

30 15

20 1 (lean MDEA) 20 (feed) 2.92

30

desorber in MDEA process 20 1 0.98

specifications for the HC gas. Additionally the Selexol process contains three flash tanks for Selexol regeneration, which are operating at 35 bar(g), 20 bar(g), 1 atm, respectively. The outlet gas from the HP flash tank is recycled to the absorber in the Selexol process, and the gas outlets from the lower pressure flash tanks are recycled to the ethane recovery column. The other column specifications and operation parameters are the same as those used in Design 1 (detailed stream data for design 2 is given in Table S2 in the Supporting Information). Because of the CO2 separation in the CO2 recovery column, the stream B-3 can be sent to a Selexol process in this design. The Selexol process removes CO2 and H2S which is recycled back to the ethane recovery column (stream B-4) while HC gas is produced meeting the product specifications. In comparison with design 1, the amount of additive inserted in the ethane recovery column is now relatively small. This is because in design 1 a large amount of additive is recycled and injected into the ethane recovery column in order to satisfy the H2S content specification for the HC gas. However, in design 2, because of the high H2S removal performance of the Selexol process, less additive is required resulting in decreased energy consumption in the ethane and additive recovery columns. 3.4. Phase II: Design 3. Designs 1 and 2 are based on the four-column Ryan/Holmes process (with modifications in design 2). Alternately, design 3 is a modification of the threecolumn Ryan/Holmes process. The original three-column Ryan/Holmes process is not considered in this case study because it is not as efficient as the four-column process.3 However, another possibility considered in design configuration 3 is for the cooperative combination of a Selexol process with the demethanizer in the three-column Ryan/Holmes process. In the three-column process, very low temperature is required in the condenser because of the sharp separation between methane and CO2 in the demethanizer. Also, a higher condenser temperature is required as CO2 composition in the top product increases. This top product contains methane, ethane, and CO2 so it is suitable to use the Selexol process to remove CO2 in this gas. As a result, combining the Selexol process with the three-column process can reduce the separation load in the demethanizer, and this option is expected to increase condenser temperature in the demethanizer and reduce the shaft power requirement in the associated refrigeration cycle. As with designs 1 and 2 an MDEA solvent process is used after the additive recovery column to remove H2S from the NGL product (Figure 11). The calculated stream data for design 3 is given in Table S3 in the Supporting Information. Here the Selexol process for design 3 requires an additional heater in order to satisfy the H2S content specification in the HC gas product. In design 2, the H2S is removed in two distillation columns prior to the Selexol process. However, in design 3, the Selexol process is used to remove both CO2 and H2S, and because the specification for the HC gas requires a very small H2S content, the additional heater is needed to

demethanizer, methane gas specification in the top stream is satisfied, and the bottom stream which contains CO2 is recycled to the ethane recovery column. Additive is also inserted into the demethanizer which increases the column temperature. So the temperature of the demethanizer can be controlled by changing the amount of inserted additive. The additive recovery column acts to separate C2, C3, and C4+ components. Some fraction of the C4+ products is recycled to other columns, where it is used as the mentioned additive. Finally, the amine process removes H2S and gives an NGL product. 3.3. Phase II: Design 2. The main idea behind the design of the next configuration is to replace some part of the Ryan/ Holmes process with a Selexol process. The Selexol process is preferentially considered here due to its lower energy consumption compared with that of amine solvent processes. Selexol can be applied for the separation of methane or ethane from CO2. The distillation columns which separate methane or ethane from CO2 are the demethanizer and CO2 recovery columns in the four-column Ryan/Holmes process and the demethanizer in a three-column process. The inlet gas of the CO2 recovery column in a four-column process contains almost all CO2 from an oil well, so treating this gas using Selexol would require a large amount of solvent and shaft power. However, the inlet gas of the demethanizer in a four-column process contains much less CO2 due to the separation in the CO2 recovery column. Hence, in design 2 a Selexol process is used to replace the demethanizer. This option is expected to reduce refrigeration duties. As with design 1 an MDEA solvent process is used to remove H2S to give the NGL product (Figure 10). Also, the pressure of the absorber in the Selexol process is set to 69 bar(g) and the number of trays is set to 10 in order to meet the product

Figure 10. Design 2 Process Diagram. 14392



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The CO2 recovery units after the MDEA solvent process are arranged in the same configuration as in design 3 (the presence of H2S has a greater influence on design 3 compared with design 2). For this design the process scheme is presented in Figure 12, and stream data is given in Table S4 in the


increase the separation efficiency of the Selexol process. This additional heater is inserted in front of the flash tank operating at atmosphere pressure, and increases the stream temperature to 150 °C in order to vaporize more of the H2S and to satisfy the 4 ppm specified limit in the HC gas. Except for these specified changes the remaining structure has the same design parameters as those used in design 2. The demethanizer in the original 3-column performs a sharp separation between methane and CO2 to produce HC gas directly. However, in this case a sloppy separation occurs in the demethanizer, and the HC gas condition is satisfied using the Selexol process. This approach results in decreased energy consumption and an increased condenser temperature in the demethanzier. Also, in this case the stream C-1 which is a top product of the demethanizer contains some ethane. However, most of this ethane is absorbed in the Selexol process together with the CO2 and recycled back to the demethanizer (in stream C-3). The bottom products of the demethanizer (stream C-2), enter the ethane recovery column to separate CO2 and other components. In comparison with design 2, design 3 has a lower energy consumption in the ethane recovery column. This is due to the methane separation prior to ethane recovery column which means methane is not present in stream C-2, which results in a reduced refrigeration duty in the ethane recovery column condenser. However, because of this methane recovery prior to ethane recovery, the recovery of ethane is more difficult in this design compared with design 2. Additionally design 3 has the disadvantage that CO2 is produced in the vapor phase meaning more energy is required for compression of the CO2 prior to injection into the oil well. 3.5. Phase II: Design 4. After the various ways in which a Selexol process could be integrated in designs 2 and 3 were considered, a configuration in which an amine process is inserted at an early stage was considered for design 4. The feed gas contains large quantities of CO2 (above 70 mol %) so an amine process is not considered here for the bulk removal of CO2. Instead, the amine process is used primarily for removal of H2S. In the previous design this was performed as a final step. However, the presence of H2S in the CO2 recovery equipment may cause increased corrosion. So considered here is the removal of H2S in front of the CO2 recovery process, using an MDEA amine solvent process which can effectively remove H2S in the presence of CO2.


Supporting Information. Except for the MDEA solvent process, all design parameters are the same as those specified in design 3. The MDEA solvent process selectively absorbs H2S in the presence of CO2 and because the H2S is absorbed at a higher rate than CO2 a lower residence time is preferred in order to remove a smaller fraction of the CO2. Hence, in this design the number of trays in the absorber is set to 8 and column pressure is set to 10 bar(g) which are lower values than those used in the MDEA solvent process in the other three designs considered. Also, since the H2S is removed by the MDEA solvent process, the Selexol process does not need the additional heater which is required in design 3. Except for these differences the design parameters used here are same as those in design 3. While H2S is separated using the MDEA solvent process, some fraction of the CO2 is also absorbed in MDEA, removed together with the H2S (acid gas stream). As a result the CO2 recovery of the whole process is reduced. Because of the absence of H2S and the decreased flow rate of CO2, the energy consumption of CO2 recovery units after the MDEA process is also reduced compared to that in design 3. However, a design this large requires a large amount of heat to regenerate MDEA in the desorber. 3.6. Comparison of CO2 Recovery Configurations. The performance of the CO2 recovery processes and their energy consumption are calculated for each configuration, and they are presented in Table 10. The temperatures of HP and LP steam are set to 240 and 180 °C, respectively. Cooling water is assumed to be available at the ambient temperature of 30 °C, and the refrigeration temperatures are different in each design case. Design 1 uses the largest amount of HP steam and refrigeration duty. The main reason for this high energy consumption is due to the relatively large quantities of additive used in order to satisfy the H2S content specifications in the HC gas product. One effect of this large additive flow rate is that the ethane recovery in the NGL product is above 95%. In the other design cases, ethane recovery is set to 90% which is the target ethane recovery in the NGL product specifications. If high ethane recovery is not required, this specification can be decreased and the energy consumption can be reduced in most 14393



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Table 10. Comparison of the Performance of Different Design Configurations HC gas product NGL product

high purity CO2

utility consumption

economic evaluation

flow rate [kmol/h] C1 composition (mol %) flow rate [kmol/h] C2 recovery (%) C3 recovery (%) flow rate [kmol/h] CO2 composition (mol %) CO2 recovery (%) HP steam (240 °C) [MW] LP steam (180 °C) [MW] refrigeration [MW] cooling water [MW] Shaft power [MW] capital cost [M$] energy cost [M$/yr] annual cost [M$/yr]

design 1

design 2

design 3

design 4

645.7 96.03 901.7 97.20 98.46 10084.4 99.77 99.70 117.7 199.8 123.9 283.3 31.3 247.8 143.8 209.2

641.8 96.66 866.9 90.02 95.50 10119.1 99.46 99.74 68.1 144.5 113.7 203.3 33.1 214.3 102.2 158.7

638.7 95.99 887.3 90.23 95.54 10113.4 99.53 99.74 76.7 193.7 98.1 275.6 52.9 227.8 118.5 178.6

647.3 96.27 871.1 90.17 95.71 6080.5 99.13 59.73 0.8 292.7 61.9 258.4 44.4 210.2 98.7 154.1

The capital costs in all cases are calculated using the results of simulation, combined with correlations from the literature describing the purchase cost of equipment12 and modified using an installation factor of 1.5 and the CEPCI (chemical engineering plant cost index) for 2011.20 This gives the costs of a CO2 EOR CO2 recovery plant. The particular correlations used for calculating the purchase costs of equipment are given below in eqs 3−7.21 The detailed process information used for estimating capital costs of design 2 are presented in Table S5 as an example. The cost of heat exchanger is given as

cases. The lowest refrigerant stream temperatures required are −33 °C in design 1, −21 °C in designs 2 and 3, and −10 °C in design 4. The lower refrigeration temperature required in design 1 results in an increased shaft power required in the propane cycle (used for refrigeration). In comparison with design 1, design 2 requires less utilities to satisfy product specifications (with the exception of shaft power which is slightly higher in design 2 as a result of the added compressor and pump in the Selexol process). Because of the presence of the Selexol process used in design 2, it demonstrates a better H2S removal performance using lower quantities of additive than design 1. Design 3 requires less refrigeration duty than design 2 because of the absence of methane in the ethane recovery column (removed beforehand using the demethanizer and the Selexol process) which reduces the energy consumption in its condenser. However, all the other utilities including steam, shaft power, and cooling duty are higher in design 3 than those required in design 2. For example, the Selexol process in design 3 requires additional LP steam to meet H2S content specification in HC gas. Increased energy consumption in terms of cooling water and shaft power are also required as a result of the phase of the high purity CO2. In design 2, CO2 is produced in the liquid phase and so a pump is used for increased pressure. However, in design 3, CO2 is produced in the vapor and so it must be pressurized using a compressor, and additional cooling duty is required because of the resulting high temperatures. Design 4 has the lowest requirements for refrigeration duty and HP steam, but requires very large amounts of LP steam which is consumed during solvent regeneration in the MDEA solvent process. Additionally, in design 4 the fraction of CO2 that is absorbed into MDEA while removing H2S in the absorber is not recovered. For this reason the CO2 recovery achieved in design 4 is only around 60%.

C hx = 94093 + 1127A0.98

(3) 2

where A is the area of heat exchanger (m ). All heat exchanger, condenser, and reboiler costs are calculated using eq 3. The compressor costs in these processes are calculated using the following equation: log Cc = 0.9518 log W + 2.9184

(4)

where W is the shaft work done in compressor (kW). The cost of pump is evaluated as a function of the flow rate (F, m3 s−1) into the pump: log Cp = 0.279(log F )2 + 1.7642 log F + 5.9583

(5)

The column costs are calculated based on the following equations: log C tc = 0.5633(log D)2 + 1.0566 log D + 3.8087

(6)

log Cpc = 1.3402 log D + 4.2677

(7)

where Ctc is the cost of tray column per each tray and Cpc is the cost of packing column per meter. Equations 6 and 7 are functions of D which is the diameter of column (m). For the purpose of comparison the cost of an adsorption dehydration unit prior to the CO2 recovery process is not considered as part of the economic evaluation as it is operated in exactly the same manner with the same equipment and energy required in designs 1−3. Although the costs of the dehydration process for design 4 are expected to smaller because of the smaller flow rate, this is not considered important because that design is already the least expensive option and because design 4 shows a significant deficit in the CO2 recovery which makes it undesirable.

4. ECONOMIC EVALUATION Economic evaluation for each design case is performed considering the costs of energy and the capital costs of the equipment required. The unit costs of HP steam (240 °C) and LP steam (180 °C) are $7.63/ton and $7.03/ton, respectively. Also, cooling water and electricity costs are $0.005/(kW h) and $0.05/(kW h). 14394



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Notes

The overall economic evaluation of each of the four designs considered in the case study is presented in Table 10. The annualized capital costs are calculated based on an interest rate and payback time of 10% and 5 years, respectively. In Table 10, design 4 appears to be the least expensive design configuration. However, CO2 recovery in design 4 is only around 60%, so this is only appropriate if the process specifications do not require a higher percentage of recovered CO2. After design 4, design 2 shows the lowest annual costs and although the refrigeration duty required for design 2 is higher than that of design 3, the total energy and capital costs are lower in design 2. This result indicates that placing the ethane recovery column before the demethanizer in the separation sequence is more beneficial in terms of reduced energy requirements and reduced costs. Design 2 is also attractive because it produces CO2 in the liquid phase, reducing the shaft power required to compress the CO2 prior to injection.

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work was supported by “The development of CO2 geological storage technologies through 1,000 ton CO2-EOR pilot test” of the Korea Institute of Energy Technology Evaluation and Planning (KETEP) grant funded by the Ministry of Trade, Industry and Energy (MOTIE) (No. 2012T100201728).

■

5. CONCLUSIONS Four different configurations for CO2 recovery have been simulated and investigated for the purpose of CO2 EOR. These four options have been evaluated economically base on simulation results and costing data from the literature. The process schemes of the four configurations are introduced in Figures 9−12, and the performance in terms of energy consumption and economic evaluation of each option are presented in Table 10. In terms of economics and overall performance, design 2 appears to be the least expensive and the most feasible option. Design 4 gives lower costs but only recovers around 60% of the CO2, and so may be considered infeasible because of this low recovery. Design 3 requires less refrigeration duty compared to design 2, but requires more hot utility and shaft power when pressurizing CO2 and removing H2S in the Selexol process. Design 1 gives the highest capital and energy costs. This is due to the large amount of additive required to aid the removal of H2S from the HC gas product. The comparison of four different options in this study is for a fixed feed gas flow rate and composition. However, this feed flow rate and composition will change throughout the life of the CO2 EOR operation (see Table 10) and so in future work the effectiveness of each design should be evaluated for a range of different conditions. In addition, the fraction of gas sent to dehydration and compression (low purity reinjected using phase I equipment) versus the fraction sent to CO2 recovery equipment (high purity in-injection using phase II equipment) can also be varied and would be an important design parameter to consider during the operation. Also, in the ideal case a superstructure containing all possible combinations could be considered and evaluated to consider all possibilities simultaneously through optimization.

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Superscripts

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V = vapor L = liquid sat = saturation

REFERENCES

(1) U.S. Department of Energy. http://energy.gov/fe/scienceinnovation/oil-gas/enhanced-oil-recovery (accessed April 28, 2014). (2) Global CCS Institute. http://www.globalccsinstitute.com/ publications/accelerating-uptake-ccs-industrial-use-captured-carbondioxide/online/28496 (accessed April 28, 2014). (3) Ryan, J. M.; Schaffert, F. W. CO2 recovery by the Ryan/Holmes process. Chem. Eng. Prog. 1984, October. (4) Bucklin, R. W.; Schendel, R. L. Comparison of flour solvent and Selexol processes. Energy Prog. 1984, 4, 3. (5) Holmes, A. S.; Ryan, J. M. Cryogenic distillative separation of acid gases from methane. United States Patent No. 4318723, 1982. (6) Valencia, J. A.; Mentzer, B. K.; Denton, R. D.; Mart, C. J. The controlled freeze zone technology for the development of sour gas resources. Soc. Pet. Eng. 2012, 156944. (7) ZareNezhad, B.; Hosseinpour, N. An extractive distillation technique for producing CO2 enriched injection gas in enhanced oil recovery (EOR) fields. Energy Convers. and Manag. 2009, 50, 1491− 1496. (8) Holmes, A. S.; Ryan, J. M. Distillative separations of gas mixtures containing methane, carbon dioxide and other components. United States Patent No. 4462814, 1984. (9) Jung, J.; Lim, Y.; Jeong, Y. S.; Lee, U.; Yang, S.; Han, C. CO2 capture process using aqueous monoethanolamine (MEA): Reduction of solvent regeneration energy by flue gas splitting. Korean Chem. Eng. Res. 2011, 49, 764−768. (10) Peters, L.; Hussain, A.; Follmann, M.; Melin, T.; Hagg, M.-B. CO2 removal from natural gas by employing amine absorption and membrane technologyA technical and economical analysis. Chem. Eng. J. 2011, 172, 952−960. (11) Belaissaoui, B.; Moullec, Y. L.; Willson, D.; Favre, E. Hybrid membrane cryogenic process for post-combustion CO2 capture. J. Membr. Sci. 2012, 415−416, 424−434.

ASSOCIATED CONTENT

S Supporting Information *

Detailed stream data for the 4 design configurations considered and detailed equipment specifications for design 2. This material is available free of charge via the Internet at http:// pubs.acs.org.

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NOMENCLATURE P = pressure (Pa) Φ = fugacity coefficient (dimensionless) VCO2solid = molar volume of solid CO2 (m3/(g mol)) R = universal gas constant (J/(g mol K)) T = temperature (K) A = area of heat exchanger (m2) Cc = compressor cost ($) Chx = heat exchanger cost ($) Cp = pump cost ($) Cpc = packing column cost per tray ($/tray number) Ctc = tray column cost per meter ($/m) D = column diameter (m) F = liquid volume flow rate (m3/s) W = shaft work (kW)

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: +82 (0)2 2220 2331. 14395



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(12) Peters, M. S.; Timmerhaus, K. D.; West, R. E. Plant Design and Economics for Chemical Engineers; McGraw-Hill: New York, U.S., 2004. (13) Smith, R.; Varbanov, P. What’s the price of steam?. Chem. Eng. Prog. 2005, July. (14) Piemonte, V.; Maschietti, M.; Gironi, F. A triethylene glycol− water system: A study of the TEG regeneration processes in natural gas dehydration plants. Energy Sour. 2012, 34, 456−464. (15) Netušil, M.; Ditl, P. Natural Gas Dehydration, Natural Gas Extraction to End Use; Gupta, S., Ed.; InTech: Rijeka, Croatia, 2012; DOI: 10.5772/45802; http://www.intechopen.com/books/naturalgas-extraction-to-end-use/natural-gas-dehydration (accessed May 18, 2014). (16) ProSimPlus Application Example: Natural Gas Dehydration Unit with Triethylene Glycol; ProSim: Labege, France, 2010. (17) Mohamadirad, R.; Hamledhdar, O.; Boor, H.; Monnavar, A. F.; Rostami, S. Mixed amine application in gas sweetening plants. Chem. Eng. Trans. 2011, 24. (18) Stewart, M.; Arnold, K. Gas Sweetening and Processing Field Manual; Elsevier: Oxford, U.K., 2008. (19) Eggeman, T.; Chafin, S. Beware the pitfalls of CO2 freezing prediction. Chem. Eng. Prog. 2005, March. (20) Chemical Engineering Plant Cost Index (averaged over year). http://www.nt.ntnu.no/users//magnehi/cepci_2011_py.pdf (accessed April 28, 2014). (21) Peters, M. S.; Timmerhaus, K. D.; West, R. E. Plant Design and Economics for Chemical Engineers; McGraw-Hill: New York, 2004.

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Conceptual Process Design of CO2 Recovery Plants for Enhanced Oil

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