Tapping Singular Middle Eastern Ultrasour Gas Resources Combining

Article Cite This: Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Tapping Singular Middle Eastern Ultrasour Gas Resources Combining Membrane and Absorption Systems: Potential for Energy Intensity Reduction Mohammed Alkatheri,† Ricardo Grandas,†,‡ Alberto Betancourt-Torcat,† and Ali Almansoori*,† †

Department of Chemical Engineering, Khalifa University of Science and Technology, The Petroleum Institute, Sas Al Nakhl Campus, P.O. Box 2533, Abu Dhabi, United Arab Emirates ‡ Department of Chemical Engineering, Universidad Industrial de Santander, Carrera 27 calle 9, Bucaramanga, Santander, Post Code 678, Colombia S Supporting Information *

ABSTRACT: A process design and techno-economic analysis is proposed for sweetening ultrasour natural gas containing over 20% H2S and 30% total acid gases using a hybrid scheme approach. This type of gas resource is unique and can only be found in the Middle East. The hybrid scheme combines membrane and amine gas absorption systems. The study was made on the basis of process simulations and sensitivity analyses to find the most suitable process design and operating parameters using the software ProMax. A Pebax-based membrane module(s) is used as the primary sweetening method, whereas gas absorption is applied to meet the final gas product specifications. The hybrid scheme is benchmarked against the current stand-alone absorption system used to process this rare type of gas. It was found that the gas sweetening energy intensity can be substantially reduced using the hybrid scheme and be more cost-effective than conventional stand-alone absorption units for treating Middle Eastern ultrasour natural gas.

1. INTRODUCTION Natural gas is a primary energy source for domestic, commercial, and industrial applications worldwide. Meeting the increasing global gas requirements involves improving the existing processes for production, treatment, energy harvesting, and pollution control. However, one of the most serious problems facing natural gas processing technologies is the presence of high contents of acid gases, typically, CO2 and H2S.1 These acid gases are susceptible to cause environmental pollution after burning. Moreover, H2S in the gas stream is toxic and flammable and can cause corrosion in the pipelines. However, CO2 removal is necessary to increase the energy yield in gas combustion. Recently, an increasing number of countries have expressed their concerns in meeting their future gas demands and export commitments. This trend is particularly palpable in the Middle East countries as a result of high population growth, increasing power and desalinated water demand, higher living standards, economic diversification into energy intensive industries (e.g., petrochemical, metallurgic), and gas injection to increase the lifetime of mature oil fields.2 To meet the increasing natural gas requirements, producers have been forced to develop highly sour gas fields formerly considered economically unattractive. The technical difficulties © XXXX American Chemical Society

of processing sulfur-rich (sour) gas once posed a great impediment for the development of unconventional gas resources. However, technology advances in sour gas processing and demand pressures have accelerated the development of these unconventional resources. The Middle East region is particularly interested in tapping these challenging reserves, because they represent approximately 60% of their gas resources.3 The most conventional sour gas processing technology is chemical absorption by solvents such as alkanolamines for natural gas cleanup, and Claus technology for H2S conversion to elemental sulfur. Though these technologies are both wellknown and proven, they can be challenging and uneconomical to process ultrasour gas resources, this given the high operating costs associated with ultrasour gas sweetening. The gas sweetening process costs, in general, depend on the feed gas sulfur content. Also, this type of operation generates large Special Issue: PSE Advances in Natural Gas Value Chain Received: Revised: Accepted: Published: A

May 28, 2017 October 16, 2017 November 3, 2017 November 3, 2017 DOI: 10.1021/acs.iecr.7b02200 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. One-stage membrane hybrid scheme process flow diagram.

simulating the operation of the absorber and distillation towers in a gas sweetening unit. Also, the interaction between the two towers at varying operating parameter values can be analyzed. For instance, during plant commissioning, simulations software help engineering companies meet the plant’s performance requirements according to the contractor specifications.6 Previous studies on sour gas sweetening have been performed using computer-aided methods such as simulation software (e.g., ProMax and Aspen Hysys) and mathematical modeling systems (e.g., Visual Basic, Matlab). For example, a wide range of research has been done on sour gas sweetening using amine gas absorption applying computer-aided approaches. Those investigations have found that gas absorption is an efficient method for removing acid gases from natural gas. The process energy costs were found to be directly related to the amount of acid gases in the feed.1,7−10 Also, studies on the use of mixed amine solvents for sour gas sweetening can be found in the literature. Those studies have found opportunities for energy savings by mixing amines at different concentrations.11−15 Other studies have been conducted to assess the economics of using membranes for gas sweetening. Those studies have also found that certain membrane system schemes can compete with amine-based gas sweetening methods.16−20 Additionally, only a few studies have focused on the application of hybrid methods for sour gas sweetening considering a combination of membrane and gas absorption systems. According to those studies, hybrid systems can yield lower costs than stand-alone gas absorption systems under certain ranges of feed flow rates, pressures, and acid gas concentrations.21,22 Although the use of hybrid systems for gas sweetening has been previously addressed in the literature, those studies have only focused on analyzing systems treating natural gas feeds with high concentrations of CO2 and low concentrations of H2S. As a consequence, they are neither a

volumes of elemental sulfur that cannot be absorbed by an already oversupplied market. Currently, sour gas production facilities are amassing ever-larger sulfur mountains. Thus, there is a need for less energy intensive and more cost-effective and sustainable gas processing alternatives.4 Membranes have proven to be a reliable technology for CO2 removal since its first application in the early 1980s. More recently, new polymeric membranes have been applied for bulk H2S removal from natural gas, including at very high H2S concentrations and operating pressures. This technology may represent a sustainable option for the development of new highly sour gas resources or retrofitting existing gas plants (e.g., absorption-based plants). In a primary step, the membrane system can be used to reduce a great portion of the acid gases (e.g., 70−90%) content from the feedstock gas. Successively, the final sweet gas product specifications can be met using a conventional absorption system. The permeate gas from the membrane system can preferably be reinjected, instead of being converted into elemental sulfur and stored onsite. As a result, possible reductions in capital and operating costs as well as sulfur production costs can be attained using this new hybrid approach.4 This hybrid scheme denotes the approach considered in the present study. Computer-based simulation techniques are valuable tools for studying chemical and physical processes such as those involved in sour gas processing.5 Process simulation allows one to examine alternative process configuration designs, incorporate changes in plant design, perturb critical operational variables, and study their response, troubleshoot, debottleneck, conduct techno-economic analysis, and estimate optimum operating conditions with minimum time, effort, and cost. The representation of chemical absorption systems completely relies on comprehensive thermodynamic and kinetic models. The effective representation of these systems involves B

DOI: 10.1021/acs.iecr.7b02200 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research Table 1. Pebax Membrane Properties23

good representation of the Middle East ultrasour gas resource characteristics nor the gas processing systems and conditions. To the authors’ knowledge, no previous study has been conducted to assess the treatment of ultrasour gas with such high concentration of H2S in the feed using a hybrid system. The present study aims to expand the limited literature on hybrid systems consisting of membrane and gas absorption systems. Specifically, the present study will examine the possibility of retrofitting an existing stand-alone absorption plant with a permeation stage(s) for sweetening ultrasour gas (i.e., over 20% of CO2/H2S or combined). The resulting hybrid systems are analyzed by applying a simulation approach. Currently, there is a lack of analyses on ultrasour gas processing given its incipient development stage, particularly when the H2S concentration in the natural gas is much higher than that of CO2. These sour gas resources are uncommon on a global scale, but relatively recurrent in the Middle East. Although the concept itself is old, these gas resources were not exploited until recently given the associated technical challenges, relatively small quantities of marketable natural gas yields, and high investment needs. The hybrid process was simulated using ProMax v3.2. The remainder of this work is organized as follows: Section 2 presents the hybrid scheme description. Section 3 describes the methodology followed in the present analysis. Section 4 explains the economic analysis model applied in this study. Section 5 discusses the performance of the proposed hybrid process for sweetening ultrasour gas. Conclusions are presented in section 6.

component

permeability [(std cm3)/(cm2·s·cm Hg)]

selectivity with respect to CH4

H2S CO2 CH4

140 31 1.8

77.78 17.22 1

The membrane considered for this analysis (i.e., Pebax) is more permeable to CO2 and H2S than to CH4 and higher hydrocarbons. Thus, the CO2 and H2S are concentrated in the permeate, whereas the CH4 and higher hydrocarbons will be concentrated in the retentate. 2.1.2. Two-Stage Membrane with Recycle Streams. In the one-stage membrane system, the permeate stream’s hydrocarbon loss rate is high. Thus, they significantly contribute to increase the membrane’s operation costs (as discussed later). However, the loss hydrocarbons can be recovered utilizing a two-stage membrane configuration connected in series and provided with recycle streams. Figure 2 shows the two-stage membrane hybrid scheme. The flow-sheet differences between the one-stage and multistage membrane systems are given as follows: (1) The permeate stream from the first membrane stage is recompressed to a pressure equivalent to the feed. This helps to keep a high/efficient separation driving force in the second permeation stage. The compressors’ maximum discharge temperature should not exceed 149 °C (300 °F).24 Accordingly, it is necessary to multistage the compression process in a series of operations.24 Cooling the gas (after partial compression) to the intake temperature reduces the power required in the following stage. Theoretically, minimum power is required by fixing the compression ratio in all stages to an equivalent value. Therefore, three compressions with intercooling stages were found to be the most suitable configuration for saving power, while keeping the compressors discharge temperature within limits (i.e., maximum 300 °F). (2) The compressed permeate stream is sent to the second stage membrane. (3) The permeate stream from the second stage (i.e., rich in acid gases components like CO2 and H2S) is discharged from the system. (4) The second membrane stage’s retentate stream mainly contains methane; thus, it is recycled and mixed with the feed gas in a mixer. 2.2. Gas Absorption Step. The second step of the hybrid scheme is a conventional gas absorption process. The retentate or “sour gas stream” from the previous step is sent to the bottom of an 8 ideal stage absorber with 33% efficiency.25 Typical absorbers for this process consist of approximately 24 real stages; this is based on years of industrial experience with gas sweetening and chemical solvents (this implies a real/ideal stage ratio of 3). A 40−50 wt % aqueous solution of DGA and MDEA were considered as solvent for the absorption of CO2 and H2S from the second step feed, i.e., from the retentate product of the membrane system. Moreover, a system factor is used as an estimate of the reduction in performance due to foaming. This is determined empirically and for chemical solvents is usually around 0.8. The column pressure drop was assumed to be 34.5 kPa (5 psi). The rich amine stream leaves the absorber through the bottom, while the sweet gas rises to the top of the absorber. A parametric analysis was performed on the rich amine stream to calculate the acid load (mole of acid gases/mol of amine). This is to guarantee that the maximum rich loading (which depends on the type of amine) was not overpassed, thus, helping to

2. HYBRID PROCESS DESCRIPTION A typical stand-alone absorption plant used for sweetening ultrasour natural gas is proposed to be retrofitted with a permeation stage(s) to potentially reduce the plant’s utility costs. Accordingly, the resulting hybrid scheme under study consists of two steps: First, a permeation step involving a membrane system. Second, a conventional gas absorption step/ unit. Additionally, two permeation configurations were investigated: (1) one-stage membrane and (2) two-stage membrane with recycle streams. Each of these steps are described in detailed next. 2.1. Permeation step (Membrane/Permeation System). The permeation step or membrane/permeation system consists of a new polymeric membrane applied for bulk H2S removal from natural gas, including at very high H2S concentrations and operating pressures. The membrane system can include one or more stages/modules, depending on the desired acid gas removal levels and methane loss rate. 2.1.1. One-Stage Membrane Configuration. Figure 1 denotes the one-stage membrane hybrid scheme analyzed in the present study. This scheme consists of two steps: a first step (permeation step) involving a one-stage membrane system, and a second step considering a gas absorption unit. First, the ultrasour gas stream enters a separation vessel to extract the condensed water content. The rich-methane top stream goes into the membrane module, where a significant portion of the acid gases are removed in the permeate (low-pressure) stream. The permeate gas can be reinjected rather than converted into elemental sulfur and stored onsite. However, most of the methane and higher hydrocarbons exit in the retentate (highpressure) stream. The membrane data used for this study such as permeability of CH4, CO2, and H2S as well as selectivity of CO2 and H2S, relative to CH4, can be found listed in Table 1.23 C


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Figure 2. Two-stage membrane with recycle stream hybrid scheme process flow diagram.

3. METHODOLOGY The simulation software ProMax was used to model and find the optimal operating parameters for the proposed retrofitted hybrid sweetening plant treating ultrasour gas. The model consists of a hybrid scheme composed of a new membrane separation step (permeation), followed by an already existing absorption step. Some commercial membranes have been reported to separate 70−90% of sour gases, which would significantly help reducing the operating cost of the downstream amine unit (absorption) used to reach the product specifications. To study the effect of the membrane permeation step, the commercial Pebax26−28 membrane module was considered for the present simulation analysis. According to George et al.,28 polyether block amides (i.e., diverse Pebax grades) are considered to be one of the few materials that simultaneously show high H2S/CH4 selectivity and good permeability. Thus, they are classified in the useful polymer category for acid gas removal. The membrane module’s retentate stream is generally rich in methane and hydrocarbons, whereas the permeate is rich in acid gases. The membrane is selective toward H2S and CO2 (with much lower permeability rate for CO2 than for H2S). In the present section, the focus is posed over the membrane module details because the amine unit (absorption) has been thoroughly analyzed in several previous studies.6,29−33 The ProMax membrane tool enables performing the separation process calculations. Accordingly, the description of the operating fundamentals of the Pebax membrane module in ProMax is presented next.25 3.1. Membrane Step. In the membrane module, the transport of gases through the dense (nonporous) polymer membranes occurs by a solution−diffusion mechanism. The gas dissolves in the polymer at the membrane high-pressure side, diffuses through the polymer phase, and desorbs or evaporates at the low-pressure side. More details on membrane separation

avoid corrosion problems in carbon steel equipment. Next, the rich amine stream passes through a flash drum to remove most of the hydrocarbons (because they are more volatile than the acid gases) before entering a rich-lean amine heat exchanger. In the heat exchanger, the rich amine stream exchanges heat with the lean amine stream, rising the former’s temperature to 210 °F. The flashing process minimizes erosion in the lean/rich amine exchanger, whereas its vapor may be used as fuel gas. Later, the hot rich amine stream goes to an 11 ideal stages amine regenerator. The regenerator is coupled to a condenser and a reboiler. The reboiler is heated using high-pressure steam, whose flow rate is correlated to that of the solvent through a specific steam-to-solvent ratio (kg/L or lb/gal). The condenser temperature is maintained at 49 °C (120 °F). The lean amine leaving the reboiler passes through the rich-lean amine heat exchanger and goes to an amine makeup unit to compensate for the loss amine. The lean amine composition leaving the makeup unit is adjusted depending on the solvent type (e.g., 20, 35, 50, and 70% mass fraction for MEA, DEA, MDEA, and DGA, respectively). Subsequently, the lean solvent enters a recirculation pump to get the pressure back up to that of the absorber. This is required due to the solvent’s pressure loss in the regenerator. After the lean solvent is pressurized, it is cooled down to within 5 °C (9 °F) of the feed gas temperature using a cooling water heat exchanger. This approach ensures that no hydrocarbon condensation occurs in the absorption process. Moreover, the rich loading is calculated on the basis of the solvent flow and the specified amine concentration. Therefore, for the recirculation flow rate a ProMax solver was set to maintain the appropriate level for each type of alkanolamine studied. D


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package indicates the liquid phase model,; whereas the equation of state designates the model for the vapor phase. In the present work, the Electrolytic ELR package was selected because it is specially tailored for chemical solvent applications. For the vapor phase, either the Peng−Robinson or Soave− Redlich−Kwong equation of state are recommended; the difference in results between them is generally minimal. The operation of the absorber is modeled using the “TSWEET Kinetics” model. This model accounts for the varying absorption rates of the different acid gases; particularly the relative slow absorption of CO2 that is kinetically limited. Using this kinetic model requires the residence time to be known for each stage. Accordingly, the residence time was estimated on the basis of the following column’s information: fraction flooding, tray spacing, weir height, and system factor. In general, the recommended fractional flooding is around 70%, the tray spacing is 2 ft. (0.6 m), the weir heights are 2−3 in. (5−8 cm), and the system factor is around 0.8. The convergence algorithms for the absorber and regenerator were both set to account for enthalpy changes considering wide or narrow boiling point differences between the components (i.e., enthalpy model = “Composition-Dependent”, inner loop model = “Boston-Sullivan”).

theory and calculation are clarified in the Supporting Information file (see section 1 in Supporting Information). In this study, for the one-stage membrane the area is determined as a function of the H2S removal fraction, whereas for the twostage membrane the H2S removal fraction is estimated as a function of the membranes’ areas. Some hydrocarbon losses occur as the acid gases selectivity is not as optimum as sought. The hydrocarbon losses were calculated for each simulation as they represent a loss opportunity cost in the sweetening process. Some assumptions were required to simulate the membrane unit. First, the retentate side pressure is similar to that of the inlet gas; this assumption is acceptable and has been used in other studies.20,22,34 Second, the pressure in the permeate side was assumed to be 2.2 bar(g); this value is within the optimum range of permeate pressures (i.e., minimum permeate pressure needed for condensates not to form due to the Joule− Thomson effect) determined by other researchers.22,34 Third, the membrane module was assumed to work under isothermal conditions. This assumption is valid for the retentate side (i.e., high-pressure side) because there is no pressure drop. However, in reality there is a temperature difference between the membrane unit’s inlet and permeate streams. This value varies as a function of the pressure difference between the inlet and permeate streams, due to the Joule−Thompson effect. Unfortunately, the ProMax v3.2 membrane tool does not allow users to specify different temperatures for the retentate and permeate streams. However, it does include an adiabatic option to run membrane modules where the permeate and retentate temperatures drop to the same level (value). Accordingly, some simulations were performed under adiabatic membrane conditions, and it was found that the maximum temperature drop was around 20 °C. In fact, no significant effect on the overall separation cost was noticed as a result of the aforementioned temperature drop due to the Joule− Thomson effect. Therefore, the isothermal assumption can be considered acceptable for evaluating the separation process of the proposed hybrid system. Moreover, the 68 bar natural gas feed pressure assumed in this study falls within the range 10− 100 bar. According to Marić35 under this pressure range, the effect of temperature drop for natural gas characterized by Joule−Thomson effect is negligible. Fourthly, the membrane element life was assumed to be 1 year (i.e., a little shorter than the range assumed by other studies17,20−22,34 given the extreme sour feed conditions in this study) to overcome fouling caused by liquid depositing in the membrane surface that might form due to the Joule−Thomson effect.22 Lastly, the selectivity and permeability data of the Pebax membrane are assumed not to change with respect to temperature and pressure.20,22,34 3.2. Absorption Step. For the gas absorption step, the first phase consists of selecting the applicable equation of state that better represents the process. Amines are chemical solvents that generate electrolytes, or solutions containing ions. Traditional equations of state do not describe electrolytes very accurately. Only a few thermodynamic packages are available in ProMax for modeling electrolytes. All these packages are Gibbs Excess Energy (Activity Coefficient) models, specifically designed to model the behavior of electrolytes. For nonreactive systems, the concentration of ionic species in the vapor phase is negligible; as a result, the vapor phase can be modeled using a traditional equation of state. The built-in electrolytic packages in ProMax are coupled with an equation of state (e.g., Peng−Robinson, Soave−Redlich−Kwong). Accordingly, the ProMax electrolytic

4. ECONOMIC ANALYSIS MODEL The economics of the separation process were assessed using the hybrid scheme’s operating costs as basis. This is because the permeation step (i.e., one-membrane and two-membrane module with recycle stream) is analyzed as a potential utilities-saving strategy in addition to the current stand-alone absorption process used in the Middle East, for example, in the United Arab Emirates to process ultrasour gas extracted from the Shah field.36 Accordingly, comparisons between the standalone and hybrid schemes are made to assess the technoeconomic feasibility of adding the permeation step into the gas sweetening process. As a result, for the present analysis only the permeation step flow-sheet costs were considered as capital investments. As shown in Figure 1, the one-stage membrane configuration consists of a single membrane module. Conversely, the two-stage membrane configuration includes two membrane modules, three heat exchangers, and three compressors (Figure 2). This suggests that the operating costs of both permeation configurations differ from one another. The membrane’s module cost is a one-time expense; however, the membrane elements need to be replaced periodically, which is why the latter was considered to be an operating cost for the purposes of this study. Additionally, the opportunity cost due to CH4 losses in the permeate is accounted for as part of the hybrid system operating cost.21 In other words, the permeate gas stream leaving the membrane module was considered to be a waste stream. This is the reason the permeate stream’s methane and other hydrocarbons losses are penalized in terms of opportunity cost in the economic analysis. A potential alternative is using the waste permeate stream as feedstock fuel for steam generation in the solvent regenerator’s reboilers. For instance, Niu et al. found that permeate gas reuse as fuel can lead to 10% to 28% reductions in the overall separation cost.22 Nevertheless, it is worth pointing out that reusing the permeate gas as fuel involves a series of challenges. First, the waste permeate stream cannot be directly fed to the reboiler because it would not meet the minimum required methane concentration (i.e., 40%) for boiler fuels.37 Second, one aim of this study is to retrofit existing amine E


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Industrial & Engineering Chemistry Research Table 2. Economic Parameters and Assumptions

absorption units where the new design can accommodate different levels of acid gas concentration in the feed. Reusing the waste permeate stream may remove this flexibility because the stream differs from case to case. Conventionally, the cost of processing natural gas is defined per thousand standard cubic feet (MSCF) of feed. However, this definition is typically used when sour gas containing small amounts of acid gases is sweetened. For natural gas with high acid gas content, it is more appropriate to define the cost in terms of MSCF of product. This because the production of ultrasour gas yields relatively small quantities of marketable natural gas (e.g., just 504 cubic feet of product per 1000 cubic feet of processed feed at Abu Dhabi’s Shah field)36,38 even after extensive treatment. Moreover, sweetened gas is sold on a product volume basis rather than processed feed. As a result, the cost definition used in this study is given in terms of MSCF of product.18,19

The economic evaluation of a separation process mainly depends on the chosen method of analysis, and the values assigned to the economic parameters. As a result, the economic evaluation may differ among sources. Nevertheless, these differences can be instructive if the applied methodology is clearly described;18,19,21 this is the approach applied in the present analysis. The main objective of the present study consists of identifying the optimum values of key process variables. For example, the H2 S concentration in the membrane’s retentate stream (which enters the gas absorber), and the permeation step configuration (one or two-stage membrane and second membrane stage’s size) are key independent variables that significantly impact the hybrid scheme’s overall cost. A summary of the economic analysis procedure, cost equations, and techno-economic assumptions is shown in Tables 2 and 3.21,24,39 F


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Industrial & Engineering Chemistry Research Table 3. Permeation Step Operation Cost Calculations: Two-Stage Membrane with Recycle Streams

approximated to be equal to the membrane capital cost divided by the membrane lifetime. The operating labor costs were neglected in the present study. This because all process configurations analyzed are similar, and labor costs depend on the number of process equipment. Thus, this particular cost factor can be neglected as a distinctive benchmark for comparison purposes. Additionally, the capital cost of the absorption unit was not considered in this study for several reasons. First, the lifetime of the amine sweetening unit is much larger than that of the membrane module (e.g., the membrane lifetime is approximately 1−3 years whereas the amine sweetening unit can reach up to 35 years). Second, the present analysis aims to compare processes with similar equipment and almost the same sizes. Therefore, the no consideration of the amine sweetening unit will not significantly impact the final results. Third, one of the main objectives of the present study is to assess the viability of adding membrane module(s) to existing amine-based gas sweeting plants. As a result, only the operating costs of the amine sweeting unit were calculated by summing up its corresponding utilities costs (e.g., process steam, cooling water, and electricity requirements), amine losses, and hydrocarbon losses from the flash drum and regenerator. For more details on the economic analysis refer to Tables 2 and 3.

The cost of the membrane elements was assumed to be $5/ ft.2, whereas the membrane module (consisting of a pressure vessel, membrane element, associated piping, valving, etc.) was assumed to be $10/ft.2. The elements cost is an estimate for a spiral-wound membrane. Also, because the membrane elements must be periodically replaced, this cost is considered as an operating cost. Note that the unit price of the membrane element was assumed as the average from the values reported by Niu et al.22 and Peters et al.34 In both studies the membrane module cost is assumed to be 2 times that of the membrane element cost. Additionally, to see the effect of this parameter, a sensitivity analysis evaluating the effect of the membrane element cost on the overall performance of the process was performed (see section 5.1.2 for details). As shown in Table 2 (eq 10a), the total cost of the sweetening process is equal to the summation of the membrane module(s) and absorption unit operating costs. For the one-stage membrane, the operating cost (eq 2) is a function of the membrane’s capital recovery cost (eq 1), hydrocarbon losses (eq 7), and replacement cost (Table 2). However, for the two-stage membrane with recycle stream the operating cost (eq 16) additionally includes the equipment capital recovery cost (e.g., heat exchangers and compressors) (eq 15) and the permeation step utility cost (Table 3). The equipment capital recovery cost is calculated by dividing the capital costs (bare module costs) of the permeation step’s heat exchangers and compressors by their corresponding payback period (i.e., 15 years). Moreover, the heat exchangers’ capital costs were calculated as a function of their areas. The areas were calculated by the heat exchanger design shortcut method.24 More details on the membrane system operating cost for the two-stage with recycle stream are presented in Table 3. For both configurations, as the membrane lifetime is assumed to be short, the membrane capital recovery cost was

5. HYBRID SCHEMES ANALYSES The present case study compares the costs of stand-alone absorption system with those of hybrid schemes. DGA and MDEA were selected for the hybrid system analysis. The selection is supported by a preliminary simulation-based study done to evaluate the performance of different alkanolamines (i.e., MEA, DEA, DGA, and MDEA) for treating ultrasour natural gas (see section 2 in the Supporting Information). The proposed hybrid schemes consist of two steps: (1) membrane-based process followed by (2) solvent-based G


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construction material.41 Experience has indicated that keeping the rich loading within certain limits is critical to preventing corrosion in equipment constructed in carbon steel.41 Furthermore, CO2 is more corrosive toward carbon steel than H2S.41 In this study, the H2S concentration is much greater than that of CO2, which further supports the use of carbon steel as material of construction. Additionally, under the considered feed pressure and temperature (i.e., P = 68 bar, T = 38 °C) carbon steel can be acceptably used as the primary material of equipment construction. This is, the ultrasour natural gas at the current study pressure is significantly below carbon steel’s maximum allowable stress pressure (i.e., 944 bar).42 Thus, carbon steel maximum allowable stress is high enough to withstand the design pressure. Moreover, carbon steel’s maximum allowable stress only drops rapidly after 350 °C. This value is far from the maximum operating temperature considered in the study (Figures 1 and 2).42 Accordingly, this allows avoiding the need for stainless steel, which is approximately 4 times more expensive. 5.1. Sensitivity Analysis on One-Stage Membrane Hybrid Scheme. The present case study considers a one-stage membrane system for the permeation step of the hybrid scheme. Various sensitivity analyses were done to assess the proposed hybrid scheme techno-economic performance. The applied analysis approach consists of the case study method.16,43 The corresponding sensitivity analyses are presented next. 5.1.1. Effect of H2S Removal in the Permeation Step over the Hybrid Scheme Costs. For the present analysis a parametric study was performed to determine the cost sensitivity of hybrid schemes to variations in the membrane retentate acid gas concentration. Figure 3 shows the cost variation of the one-stage membrane hybrid scheme’s components (i.e., permeation and absorption steps) as a function of the H2S fraction removed by the membrane module. From the figure, it can be observed that the higher the fraction of H2S removed by the membrane module, the lower the total hybrid scheme’s cost. Nonetheless, higher H2S removal rates result in larger operating costs for the membrane module, this because greater H2S removal rates demand larger membrane areas and are associated with higher methane losses. A comparison between the stand-alone and hybrid schemes shows that the later involves lower total operating costs (Figure 3). According to the simulation results, when MDEA (e.g., a tertiary and very aggressive alkanolamine to both acid gases) is employed as a solvent instead of DGA, the required solvent and utilities demand are lower (Figure 3). The difference is more noticeable when lower H2S percentages are removed by the membrane (permeation) step. However, at higher H2S removal levels by the membrane module the performance of both amines tends to be closer. This trend occurs due to the variations in the H2S/CO2 ratio determined by the membrane module’s H2S removal percentage. For instance, when the absorber’s feed H2S concentration is 22% the H2S/CO2 ratio is 2.10, whereas for a H2S concentration of 4% the H2S/CO2 ratio falls to 0.52. As MDEA is a tertiary amine, it is more reactive toward H2S. As a result, it performs better at higher H2S concentrations. Figure 4 illustrates the cost components of the membrane module. As shown in the figure, the hydrocarbon (i.e., methane) losses constitute a large fraction of the membrane module cost. In fact, methane losses are strongly correlated to

process. For the solvent-based step DGA and MDEA were selected as alkanolamines. These alkanolamines were found to be the most suitable solvents for highly/ultrasour natural gas processing according to a preliminary simulation-based study and information available in the literature.40 Combining membranes and absorption systems (e.g., solvent-based process) could potentially help to take advantage of each technology benefits while reducing their corresponding drawbacks. For instance, membranes operate best at high acid gas partial pressures whereas solvents are better for reducing acid gas content to very low concentrations (e.g., pipeline specifications H2S < 4 ppm and CO2 < 1 mol %). Therefore, for ultrasour gas feeds at high pressure and acid gas concentrations, which must be conditioned to meet pipeline specifications, the use of a hybrid system consisting of a membrane unit followed by an amine unit is a good option. The gas sweetening process’ feed composition, flow rate, pressure, and temperature may vary considerably depending on the natural gas source. The rich loading values and solvents concentrations used in the simulations are reported in Table 4.41 For the present analysis the ultrasour gas feed conditions Table 4. Rich Loading and Solvent Concentration Limits for Different Alkanolamines41 amine

rich loading (mole of acid gases/mol of amine)

amine concn (wt %)

MEA DEA DGA MDEA

0.35 0.40 0.40 0.50

20 35 50 50

were obtained from data available in the literature (details in Table 5).36,40 Moreover, the rich loading and solvent concentrations were set in the range of values that would allow carbon steel to be used as the plant’s primary equipment Table 5. Ultrasour Gas Feed Operating Conditions36,40 feed pressure feed temperature feed standard flow sweet gas specification main components nitrogen water CO2 H2S CH4 ethane propane isobutane n-butane isopentane n-hexane trace components benzene ethylbenzene toluene m-xylene o-xylene methyl mercaptan ethyl mercaptan

68 bar 38 °C 250 MMSCFD 4 ppm of H2S & CO2% < 1% mol % 0.02 3.976 10 25 54 3 1.5 0.5 0.8 0.4 0.4 ppm 300 100 1500 1500 500 100 50 H


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Figure 3. One-stage membrane hybrid scheme costs as a function of % H2S removed by the membrane module.

Figure 4. Total and itemized costs for the permeation step of the one-stage membrane hybrid scheme as a function of the % H2S removed by the membrane module.

utility cost. Over 90% of the absorber’s operating costs are utility expenses. 5.1.2. Effect of the Membrane Element Price over the Hybrid Scheme Cost. A sensitivity analysis was carried out to assess the effect of the membrane elements price over the total hybrid scheme cost. It was assumed the annual membrane module capital cost value to be 2 times that of the membrane element cost similar to that in previous studies.22,34 Figure 5 shows the total cost of the hybrid scheme as a function of the H2S removed by the membrane module for different membrane element prices. It is worth noticing that the separation cost is very sensitive to the membrane element price. Thus, the membrane element price plays an important role on determining the overall performance of the process. Furthermore, as shown in the figure, there is an optimum separation cost at a certain level of H2S removed by the membrane. Also, as the membrane price increases, the optimum value is more tangible. For instance, for hybrid systems using MDEA and DGA as solvents, and considering a membrane element price of $50/ft.2, the optimum separation

the membrane type. The lower the selectivity of the acid gases relative to methane, the higher the methane losses. One of the greatest concerns when membranes are used is the amount of valuable gas components that can be loss in the permeate, which can substantially decrease the system’s overall methane recovery rate. Additionally, higher membrane module’s H2S removal levels involves larger areas and hydrocarbon losses (see subsection 3.1, Figure S2 in the Supporting Information). As a result, it is worth pointing out that relying exclusively on membranes to reduce acid gas content into ppm levels is not feasible. This will require infinite membrane area combined with very large hydrocarbon losses. More details on the effect of H2S in the absorber’s feed over utility consumption and costs are presented in the Supporting Information file (see subsection “3.2. Effect of H2S concentration in the absorber’s feed over utility costs” in the Supporting Information). It is worth pointing out that the utilities demand are strongly correlated to the concentration of H2S in the absorber’s feed. Process steam is by far the leading I


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Figure 5. Effect of the membrane element price over the one-stage membrane hybrid scheme total cost (MDEA and DGA).

Figure 6. Membrane area (continuous line) and hydrocarbon losses (dashed line) for the one- and two-stage membrane hybrid schemes as a function of the H2S fraction removed by the membrane module(s).

10% of the first module. scenario 2 assumes the second membrane module area to be 40% of the first module. Scenario 3 considers the second membrane module area to be 80% of the first module. 5.2.1. Effect of H2S Removal in the Permeation Step over the Membrane Area and CH4 Losses. Figure 6 depicts the total membrane area and hydrocarbon losses for the three scenarios and one-stage membrane configuration as a function of the H2S removed in the permeation step. As shown in the figure, the one-stage membrane configuration not only requires the smallest area by H2S fraction removed but also includes the largest hydrocarbon losses. However, scenario 1 not only considers the lowest hydrocarbon loss rates but also includes the largest area by fraction of H2S removed. In general, it is worth pointing out that the two-stage membrane configurations with recycle streams consider lower hydrocarbon loss rates than the one-stage membrane.

cost can be found in the 65−72% H2S removal level. Accordingly, the minimum gas sweetening cost can be obtained by setting the membrane module’s area value to that corresponding to the optimum H2S removal level. The optimum H2S removal level is a function of feed composition, operating conditions, solvent type, process configuration, and cost estimation methodology. However, it is worth noticing that any hybrid process features an optimum H2S removal level/membrane area (i.e., optimum scheme’s cost). 5.2. Comparisons between the One-Stage and TwoStage Membrane with Recycle Streams. For the two-stage membrane with recycle stream three scenarios were studied. In all scenarios the first membrane area (of the two-stage configuration) was set to be equivalent to the corresponding one-stage membrane configuration for each H2S removal percentage. Moreover, the three studied scenarios consider variations in the second stage membrane area. For instance, scenario 1 considers the second membrane module area to be J


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Figure 7. Membrane operation cost for the hybrid schemes as a function of the fraction of H2S removed by the permeation module(s).

Figure 8. Total separation cost for the one- and two-stage membrane hybrid schemes as a function of the fraction of H2S removed by the membrane module(s).

5.2.2. Effect of H2S Removal in the Permeation Step over the Hybrid Scheme Costs. Figure 7 shows the membrane module(s) operating cost for the three scenarios and one-stage membrane configuration as a function of the H2S fraction removed by the permeation step. As shown in the figure, scenario 1 features the highest membrane modules’ operating costs throughout most of the H2S removal rates. This is the result of high methane recovery rates, which translate into large recycle stream flow rates requiring significant recompression power and intercooling water. Nonetheless, beyond 67% of H2S removal rate scenario 3 features the highest membrane modules’ operating costs mainly due to very large hydrocarbon losses (Figure 7). Although the one-stage membrane configuration includes the highest hydrocarbon loss rates of all studied options, it also features the lowest permeation costs given its simplicity. This translates into no recompression power requirements due to the lack of any recycle stream.

Figure 8 illustrates the total separation cost for the different configurations (i.e., using MDEA) as a function of the H2S removed by the permeation step. The figure shows that scenario 1 considers the highest total separation cost, whereas the one-stage membrane and scenario 2 include the lowest separation costs. Accordingly, under current studied conditions and feed compositions the one-stage membrane and scenario 2 were found to be the most suitable ultrasour gas processing configurations. However, one could expect that under high gas price levels scenario 2 would prevail as the leading processing configuration because it involves significantly lower hydrocarbon losses in comparison with the one-stage membrane scheme (Figure 7). As displayed in Table 6, comparisons between the standalone absorption system and proposed hybrid schemes (underestimated optimum conditions) show that the latter outperform the former in terms of energy intensity, utility savings, overall costs, and methane recovery rates. Moreover, K


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5.3. Effect of Absorption Stage’s Capital Cost over the Total Separation Cost. First, in this section the effect of the capital cost for a stand-alone absorption unit as a function of the acid gas concentration in the natural gas feed is investigated. MDEA was selected as the primary solvent for the absorption unit because it was found to be the most suitable amine for treating ultrasour gas. ProMax was run for different acid gas concentrations in the feed gas. Total equipment capital costs were calculated as a function of size. The sizes (i.e., diameters, length, and areas) of main unit operations such as absorber, regenerator, process vessel, and heat exchangers were estimated on the basis of the heuristics design shortcut method.24 The summation of the bare module costs of all equipment was denoted as the total capital investment (TCI). The TCI can be converted into annual amortized capital cost (capital recovery cost (CRC)) using the α factor presented in (eq 25) (Table 7).

Table 6. Comparison between the Optimum Annual Costs for the Different Gas Processing Schemes and Scenarios processing scheme/scenario stand-alone absorption unit one-stage membrane hybrid scheme two-stage membrane with recycle streams (scenario 1) two-stage membrane with recycle streams (scenario 2) two-stage membrane with recycle streams (scenario 3)

MDEA (MM $/yr)

DGA (MM $/yr)

64 25 37

84 33 50

25

33

32

39

the main differences between the hybrid schemes are process flowsheet complexity, hydrocarbon loss rates in permeation step, and the permeation step effect on following absorption step performance.

Table 7. Absorber Unit Capital Cost and Total Separation Cost Calculations

L


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Figure 9. Total separation cost of stand-alone absorption and its amortized capital cost (capital recovery cost (CRC)) as a function of entering acid gases.

Figure 10. Total capital investment and total utility cost of the stand-alone absorption unit as a function of acid gas concentration in the feed.

Figure 11. Total and single separation costs for a two-stage membrane with recycle stream (scenario 2) hybrid scheme as a function of % H2S removed by the membrane module(s).

numbers are in agreement with the ultrasour natural gas feed commonly found in the Middle East region. As shown in Figure 9, at high acid gas concentration (i.e., high H2S concentration) the contribution of the absorption unit’s amortized capital cost to the total absorption separation cost is rather small. However, at very low acid gas concentration the absorption unit’s amortized capital cost contributes significantly to the absorption unit separation cost. Because

More details on the absorption unit capital are presented in Table 7. Figure 9 displays the total separation cost of a stand-alone absorption unit and the amortized capital cost (capital recovery cost (CRC)) for a unit’s feeds at different acid gas concentrations. For all cases an acid gas concentration of 70% in H2S was considered and the balance in CO2. This M


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Currently there are many hybrid systems operating worldwide.44 However, the great majority of these applications are mainly directed to process natural gas from reservoirs with a high concentration of CO2 and low H2S content. For example, early applications of membranes were done in the area of enhanced oil recovery (EOR) with CO2 flood. Several EOR projects in West Texas use a combination of membranes and amines to separate CO2 from valuable gaseous hydrocarbons. The CO2 is returned to the subsurface while valuable gas components are recovered to be further processed in downstream facilities. In general, a high concentration of CO2 in the natural gas is a good indicator for the use of membranes and/or hybrid systems. Hybrid systems have proven to be very flexible and cost savings. Economic analysis performed for CO2 removal from natural gas in West Texas has shown that the total separation cost of a hybrid system consisting of a membrane unit followed by an amine unit can be up to 27% lower compared with both a stand-alone amine and hot sodium system.44 The approach proposed in this study showcases the potentiality of applying hybrid systems to process natural gas with a very high (ultrasour) concentration of H2S. The present approach is proposed given the latest developments on membranes that are more permeable to CO2 and H2S than to CH4 and higher hydrocarbons, also highly selective toward CH4.28,45 Although most of these membranes are currently being tested at a labscale and pilot plants, there is a lot of potentiality in this area given the substantial reserves of natural gas with these characteristics in the Middle East. These gas resources are very rare on a global scale, but recurrent in the Middle East region. Thus, new methods need to be developed so that Middle East gas producers can tap their ultrasour gas resources, which have been left unexploited for decades given the involved technical challenges, significantly high-energy requirements, and low final gas product yields.38

this study focuses only on ultrasour natural gas (very high acid gas concentration) processing, the absorption unit’s amortized capital cost can be considered small compared with the remainder cost components. Moreover, the study’s main target is retrofitting an existing amine unit by adding membrane separation unit(s); therefore, the capital cost of the absorption stage has no significant effects on the performance of the hybrid system. Figure 10 shows a comparison between the total capital investment and the total utility cost for a MDEA absorption unit at different acid gas feed concentrations. An absorption plant with a feed of 250 MMSCF of sour natural gas was considered for this analysis. From this figure, it is important to note that even the total capital cost is lower than the utility cost for acid gas concentrations higher than 8%, let alone the amortized capital cost. This trend also supports the notion that the operating cost (i.e., mainly the utility costs) is the most important component of the absorber unit’s total cost when ultrasour natural gas feeds are treated. Second, in this section the effect of the absorption stage capital cost on a new hybrid plant design was also investigated. The capital costs for an absorption unit similar to that of scenario 2 (i.e., two-stage membrane with recycle, the second membrane module area to be 40% of the first module) hybrid scheme were calculated for different H2S removal rates at the permeation step. Scenario 2 was selected for this further analysis because it was found to be the most suitable ultrasour gas processing configurations (see section 5.2.2). Figure 11 shows the variation in both the total and single separation costs for the hybrid system. The analysis was performed by considering the capital cost of the absorption stage to be a function of the H2S concentration removed by the permeation stage. Similarly to what was previously observed in Figure 9, Figure 11 also shows that the absorption stage separation cost is much more significant than the corresponding annual amortization capital cost, whereas the latter is the lowest among the observed costs. Consequently, one can consider the capital cost of the absorption stage in a hybrid scheme to play a minor role in evaluating the total cost of a separation system for treating ultrasour natural gas feeds compared with other costs. Finally, for existing absorption units the composition and operating conditions of the gas feed vary on time. Thus, overcoming such changes might require major adjustments to adapt the existing units. Nonetheless, the addition of a membrane separation step before an existing amine absorption unit could enable us to operate under higher feed range capacities and acid gas content. In other words, hybrid schemes are more flexible in terms of operability. However, the economics of such retrofitted hybrid plants might differ from those presented in this study. This is due to additional factors that have not been considered in the present analysis. The most important factors are the presence of higher (i.e., heavier) hydrocarbons and corresponding effects on costs, and crude natural gas streams content. For instance, higher hydrocarbons may condense to various extents during the membrane process. The condensation is undesirable and can be prevented by heating the feed stream. Nonetheless, this would translate into additional equipment increasing capital and operating costs. However, hybrid schemes also offer investment flexibility because the membrane capacity can be easily expanded by incorporating additional elements to existing membrane modules or by installing incremental membrane modules.

6. CONCLUSION This study presents a hybrid processing scheme for sweetening ultrasour natural gas. The proposed gas sweetening configuration consists of a membrane permeation step followed by a conventional amine-based absorber unit. Two configurations for the permeation step are proposed: one-stage membrane, and two-stage membrane with recycle streams. On the basis of a preliminary simulation study and information available in the literature, stand-alone absorption units using DGA and MDEA as solvents were found to be the least energy intensive (i.e., minimum electricity and steam demands) processes for sweetening highly/ultrasour gas. Accordingly, the stand-alone absorption units, using DGA and MDEA as solvents, were considered as base case scenarios for comparison purposes. Furthermore, both solvents (DGA and MDEA) were selected for the hybrid processing schemes. An economic analysis was performed on the one-stage membrane hybrid scheme for treating ultrasour natural gas. It was found that for hybrid schemes the higher the portion of H2S removed by the permeation step, the lower the total system’s cost. However, there is a limit on the H2S removal level that can be accomplished with the membrane module(s) alone, and relying exclusively on membranes to reduce acid gas content to ppm levels is infeasible. Additionally, methane losses constitute a large fraction of the membrane module cost, which represents one of the major concerns for employing membrane modules due to reductions in methane recovery rates. N


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Industrial & Engineering Chemistry Research Cph

However, hydrocarbon losses can be reduced using a twostage membrane with a recycle streams scheme. But, this modification entails additional expenses such as extra membrane modules, recompression equipment capital costs (heat exchangers and compressors), and recycle stream recompression energy requirements (e.g., power). Overall, both hybrid processing schemes can potentially be more costeffective than a conventional stand-alone absorption system. Furthermore, retrofitting existing gas absorption plants with membrane modules for bulk H2S removal may potentially reduce gas sweetening costs, while providing a higher operating flexibility for handling feed flow rate and composition changes. Also, for a hybrid gas sweetening process, an optimum point (cost) can be obtained at certain membrane module’s H2S removal level.

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Cpp Cpt Cpv CRC D E EC ECRC ECRC FM FP HCUC HF

ASSOCIATED CONTENT

S Supporting Information *

HLC HV

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.7b02200. Additional information on operating fundamentals of the membrane module in ProMax, a stand-alone absorption case study including a table of simulation results and operating costs, results of single membrane stage including figures of membrane area and hydrocarbon losses and utility costs and a table of simulation results, the amine sweetening package in ProMax, and the system factor (PDF)

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HXCC J L MHLC MMC MOC MRC Pi1 Pi2

AUTHOR INFORMATION

Corresponding Author

q Q

*A. Almansoori. Fax: +971 2 607 5200. E-mail: aalmansoori@ pi.ac.ae. ORCID

S

Ali Almansoori: 0000-0002-0789-5105

SC TUC V

Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors gratefully acknowledge the support of the Khalifa University of Science and Technology, The Petroleum Institute of Abu Dhabi, and the Abu Dhabi National Oil Company (ADNOC).

VCC X Y

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purchase cost of centrifugal compressor at ambient temperature and pressure [$] purchase cost of a centrifugal pump at ambient temperature and pressure [$] purchase cost of one tray [$] purchase cost of a vertical vessel at ambient temperature and pressure [$] membrane capital recovery cost [$/yr] permeate gas diffusion coefficient [m2/s] energy required by compressor [kW] electricity cost [$/yr] equipment capital recovery cost [$/yr] equipment capital recovery cost [$/yr] material of construction factor pressure factor heat exchanger and compressor utility cost [$/yr] standard volumetric flow rate of hydrocarbon losses [scf/h] hydrocarbon losses cost [$/yr] conventional sweet natural gas high heating value [MMBTU/scf] heat exchanger capital cost bare module cost [$] gas flux [mol/(m2·s)] molar flow rate at the membrane feed side (highpressure side) [mol/s] membrane hydrocarbon losses cost [$/yr] membrane capital cost [$] membrane operating cost [$/yr] membrane replacement cost [$/yr] partial pressure of gas i at feed side [Pa] partial pressure of gas i at permeate side (i.e., lowpressure side) [Pa] permeability coefficient [mol/(m·s·Pa)] permeability coefficient per membrane thickness [mol/(m2·s·Pa)] Henry’s solubility coefficient of the permeated gas [mol/(m3·Pa)] steam cost [$/yr] total utility cost [$/yr] molar flow rate at the membrane permeate side (low pressure side) [mol/s] vertical vessels or towers capital cost (bare module cost) [$] gas composition at feed side [%] gas composition at permeate side [%]

Subscripts

NOMENCLATURE αAB selectivity coefficient (i.e., the ratio between the permeability coefficient of gas A and B) [%] ALC amine losses cost [$/yr] AOC absorption stage operating cost [$/yr] C1 equilibrium concentration of diffused gas at feed side (i.e., high-pressure side) of membrane module [mol/ m3] C2 equilibrium concentration of the permeated gas at permeate side (i.e., low-pressure side) of membrane module [mol/m3] CA corrosion allowance [m] CC cooling water cost [$/yr] CCC compressor capital cost (bare module cost) [$] Cph purchase cost of floating head heat exchanger at ambient temperature and pressure [$]

0 i N nT

feed/inlet membrane point component i membrane stage final membrane stage (total number of membrane stages)

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REFERENCES

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Tapping Singular Middle Eastern Ultrasour Gas Resources Combining

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