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Design and Assessment of a Membrane and Absorption Based Carbon Dioxide Removal Process for Oxidative Coupling of Methane Alberto Teixeira Penteado, Erik Esche, Daniel Salerno, Hamid Reza Godini, and Günter Wozny Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b04910 • Publication Date (Web): 14 Jun 2016 Downloaded from http://pubs.acs.org on June 20, 2016

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Industrial & Engineering Chemistry Research

Design and Assessment of a Membrane and Absorption Based Carbon Dioxide Removal Process for Oxidative Coupling of Methane Alberto Penteado,∗ Erik Esche, Daniel Salerno, Hamid R. Godini, and Günter Wozny Process Dynamics and Operations Group, Technische Universtität Berlin, Sekr. KWT-9, Straße des 17. Juni 135, D-10623 Berlin, Germany E-mail: [email protected],[email protected] Phone: +49 30 314 29515 Abstract The oxidative coupling of methane (OCM) is a direct path for converting methane into ethene (ethylene), which is one of the most important building blocks for the chemical industry. Carbon dioxide is generated as a byproduct in the reactor and must be separated in order to produce the pure olefin. This step is commonly achieved by amine scrubbing, in which a significant amount of energy is consumed for regenerating the amine. In this contribution, a hybrid system employing gas separation membranes and absorption is modeled and simulated in order to carry out preliminary engineering design and economic evaluations to assess the feasibility of applying this process on industrial scale. It is demonstrated the hybrid process offers economic advantage when compared to the standalone absorption process when carbon dioxide is used as diluent in the OCM reactor.

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Introduction The oxidative coupling of methane (OCM) is the catalytic conversion of methane into ethene (ethylene), which is one of the most valuable building blocks for the chemical industry with a global production of over 141 million tones in 2011. 1 Thus, OCM allows for a direct path from feedstock such as natural gas, shale gas, or biogas into value added chemicals, avoiding the intermediate production of syngas. 2 Since ethylene is normally produced through naphtha or ethane cracking, which are very energy demanding and generate a significant amount of carbon dioxide emissions, OCM also presents an opportunity for emission reduction in the chemical industry. 3 In 1982 Keller and Bhasin published the first determination of active catalysts for OCM. 4 Up to date, several catalysts have been tested for OCM such as La2 O3 /CaO, 5 Li/MgO, 6 and Mn-Na2 WO4 /SiO2 . 7 Also, different reactor concepts, such as membrane reactors 8 and fludized bed reactors, 9 as well as different feeding policies have been studied experimentally and through simulations 10 and optimizations. 11 The high temperature required for the OCM reaction, which is typically above 973 K, also leads to the oxidation of the educt methane (CH4 ) and the products ethane (C2 H6 ) and ethylene (C2 H4 ) into carbon monoxide (CO) and carbon dioxide (CO2 ). Besides that, the reaction network is highly exothermic and higher temperatures tend to favor the combustion reactions, reducing selectivity towards products. This highlights the importance of novel active and selective catalyst development combined with efficient reactor design. Besides that, Su et al. suggests an upper bound of 28 % for the C2 yield (combined C2 H4 and C2 H6 yield) considering a continuous single-pass process with CH4 and O2 co-feed. 12 This means that even with enhanced reaction performance, the downstream separation task is far from trivial. Hence, an integrated approach considering both the reactor and the separation steps is needed in order to synthesize an efficient OCM process. The OCM process structure commonly consists of a reaction section; a compression section, in which the reactor outlet is cooled, compressed and the condensed water is removed; a carbon 2


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and Piperazine (PZ). Using a mixture of 37 wt%MDEA and 3 wt%PZ, the energy consumption is reduced to 3.7 MJ/kgCO2 . 14 The synthesis of new energy efficient fluids for CO2 absorption is currently an active research field, however industrial applicability is still limited due to the fact that these fluids are not yet produced in large scale and at commercial prices. Gas separation membranes can alternatively be employed for CO2 separation. Glassy polymeric membranes, such as cellulose acetate and polyimides, can be applied for natural gas sweetening, specially for small to medium size capacities (6000-50000 Nm3 /h) or even for larger capacities in remote or offshore applications; 15 and also for biogas upgrading. 16,17 Applicability for flue gas treatment is limited, because separation in membranes is driven by the pressure differential. This is not a drawback for the OCM process, since compression is already required to carry out the cryogenic distillation step. Gas separation membranes offer many advantages such as straightforward installation, small spatial footprint, quick start-up, simple operation and reduced environmental impact. 15 Scale-up is also relatively easy since it happens mainly in terms of adding more modules operating in parallel. This is however of economic disadvantage if compared to traditional separation techniques, such as distillation and absorption, which tend to scale up very well. A further concern for the OCM process is the ethylene loss as it is a high value product. Current membrane technology does not offer a significant CO2 selectivity towards C2 H4 , so that a standalone membrane-based separation would hardly make any sense. The combination of membrane and absorption in a hybrid process could, however exploit the advantages of both techniques and provide a cost efficient solution for OCM. An exemplary process flow diagram for a hybrid CO2 separation system using a one-stage membrane separation and absorption/desorption (process configuration b) is presented in Figure 2. A hybrid process concept employing a single flat sheet envelope polyimide membrane module and absorption (such as in process configuration b) has been investigated on mini-plant scale at Technische Universität Berlin, showing a further reduction in the 4


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Modeling The overall model is divided into membrane section and absorption section. The model for the membrane module is initially implemented in the free modeling environment MOSAIC. 19 MOSAIC is a convenient equation oriented modeling, simulation, and optimization environment, providing a LATEX based equation input for creating a model. The model is, thus written together with its description at a documentation level and a problem specification can be created by selecting the input variables. Equation systems can be solved internally using the available BzzMath numerical library. 20 MOSAIC code generation tool can also automatically generate code for the solution in different programming languages and environments, such as Matlab, Scilab, gPROMs, Modelica, C++, Python, AMPL, GAMS or even a user-defined language. A code for the solution of the membrane model in Aspen Custom Modeler (ACM) is generated and the ACM model is subsequently exported as a unit operation into Aspen Plus for flowsheet simulations. The absorption section is modeled in Aspen Plus using the thermodynamic and unit operation models contained therein and using an example file provided in the software’s library as a starting point. The full absorption/desorption model is detailed and validated against experimental data elsewhere. 21

Absorption Section Thermodynamics An unsymmetric electrolyte model already implemented in Aspen Plus, namely ENRTLRK, is applied for the liquid phase. The set of electrolyte equilibrium reactions is introduced and the equilibrium constants are calculated by minimizing the system’s Gibbs free energy. A PC-SAFT equation of state is used for the vapor phase as it is the default in the example file used. Parameters for MEA and water are regressed or retrieved by Aspen Technology 21 from various sources, while the other parameters are taken from the software’s databank. 7



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The Redlich-Kwong equation of state is tested as a numerically cheaper alternative providing similar results. The binary parameters for Henry constants implemented in the example file are kept. Additionally, parameters for the Henry constant of ethylene in MEA are estimated based on data from Carroll and Mather 22 in order not to underestimate ethylene loss in the absorption process. Reactions The full set of liquid phase reactions can be seen below. Minimization of the Gibbs free energy is used for the equilibrium reactions (Equations 1-3) and kinetic parameters from Pinsent et al. 23 and Hikita et al. 24 are used for the kinetically controlled reactions (Equations 4-7).

− ⇀ MEAH+ + H2 O − ↽ − − MEA + H3 O+

(1)

− ⇀ 2 H2 O − ↽ − − H3 O+ + OH−

(2)

−− ⇀ HCO3− + H2 O ↽ − − CO32− + H3 O+ CO2 + OH− −−→ HCO3− HCO3− −−→ CO2 + OH− MEA + CO2 + H2 O −−→ MEACOO− + H3 O+ MEACOO− + H3 O+ −−→ MEA + CO2 + H2 O Mass and Heat Transfer The absorption and stripping columns are designed through rate-based simulations using Aspen Plus’ built-in correlations. Sulzer structured packings Mellapak 350X and 350Y are selected as the internals of the absorption and stripping columns respectively, as they are recommended by the vendor and commonly employed for CO2 capture. The study of the performance of different packing types is outside of the scope of this contribution. Plug flow is assumed for the vapor, while the liquid is assumed to be well mixed. Liquid and 8


(3) (4) (5) (6) (7)

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vapor films are non-ideal with reactions taking place in the liquid phase. Correlations by Bravo et al. (1985) 25 and Hanley and Chen (2012) 26 are applied for the mass transfer coefficients and interfacial area. The Chilton and Colburn method is used for the heat transfer coefficient. 27 Packing hold-ups are calculated as proposed by Bravo et al. (1992). 28 Pressure drop is calculated according to vendor’s correlation. These correlations have been previously assessed in literature showing good agreement with experimental data for the given application. 29,30 Both authors highlight the importance of fitting parameters such as the interfacial area factor, the heat transfer factor, and the liquid and vapor mass transfer coefficient factors to precisely mimic plant performance and behavior. The diameters of the absorber (4 cm) and the stripper (10 cm) available for experiments are rather small, which greatly increases wall effects that are negligible in industrial size equipment. Since the purpose of this work is to provide basic engineering design to allow for a preliminary cost estimation on industrial scale, no attempt to match experimental data from the mini-plant is made and these factors are left to unity.

Membrane Section A flat sheet envelope type membrane module is considered for this application. It consists of a cylindrical metal casting containing round membrane envelopes placed in series around an inner tube for the permeate outlet. The retentate flows through the shell side, being directed by the flow diverters. The module is schematically represented in Figure 4. The membrane envelopes are produced in two different versions. The first R one has a polyimide active layer (PI/Matrimid ), offering high CO2 selectivity towards

hydrocarbons, and the second one is a poly-(ethylene oxide)-poly(butylene terephtalate) thin film composite membrane (PEO/PolyActiveTM ), which offers lower CO2 selectivity but higher permeances. The membrane materials and modules are designed and produced by co-workers at the Helmholtz-Zentrum Geesthacht Centre for Materials and Coastal Research, Germany. 17 9




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dniP = Li · ( f iR − f iP ) · h dz dniR = − Li · ( f iR − f iP ) · h dz ! EaiPI 0,PI PI Li = Li · exp − R · TR LiPEO

(8) (9) (10)

NC EaiPEO 0,PEO 2 = Li · exp − + σi ∑ R · TR i =1

R P 1 0 fi + fi T R · mi · · exp mi · T 2 σi2

!!

(11)

The model has been validated using data from mini-plant experiments showing a good agreement in pressures up to 20 bar, but the accuracy decreases for higher pressures. 33 The model is later applied for the optimal design of the two-stage stripping cascades. 18 The model is herein further extended in MOSAIC to allow for integration into Aspen Plus simulations. For the model to be able to simulate industrial flow rates, a scale-up by numbering-up approach is applied. This consists of dividing the inlet flow by the design flow rate for the membrane module to obtain the necessary number of parallel modules. A single module is then calculated and the resulting flows for the permeate and retentate streams are multiplied to the number of modules to obtain the total outlet flows. An overall energy balance is formulated by introducing ACM enthalpy function calls in MOSAIC. Ports are introduced to allow for the model to be connected to streams in Aspen. MOSAIC’s code generation is used to automatically generate a code in ACM modeling language, which is compiled in Aspen Custom Modeler and exported as a custom unit operation into Aspen Plus. The compressors are modeled as isentropic with an efficiency of 0.72 and two stages with interstage and outlet coolers down to 308 K. The intermediate pressure is the geometric average between the suction and discharge pressures.

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Simulation and Optimization The models are solved using the available numerical methods inside Aspen Plus v8.6. The Sparse solver is applied for the membrane modules. The Broyden method is applied for converging tear streams, providing significantly better convergence than the default Wegstein method. The SQP method is applied for sequential optimization in the membrane section. No proper variable translation could be achieved when switching from sequential mode into equation oriented mode, so that simultaneous optimization could not be performed. Optimal design of the absorption section is thus carried out only in the form of extensive simulation studies and sensitivity analyses. The simulation and design of the CO2 capture processes are performed for an OCM plant producing 100 kt of ethylene per year in the reactor. The overall CO2 capture target of 97 % is adopted for both the standalone absorption and the hybrid process configurations. In order to reflect the different reactor design and operating policies currently under investigation, four feed scenarios with different compositions are considered in this study. It is assumed that the reactor outlet passes through a compression and cooling section, where all the water is removed and the gas enters the CO2 capture section compressed up to 10 bar and cooled down to 308 K. The different compositions are presented in Table 1. Composition I emulates the outlet of an OCM reactor using N2 dilution. Composition II mimics a reactor using a really high CH4 to O2 ratio instead. Finally compositions III and IV consider the use of CO2 dilution in the reactor. Table 1: Feed gas composition scenarios

Feed Scenario I II III IV

Composition (mol%) CO2 C2 H4 N2 CH4 11.0 6.0 33.0 50.0 9.0 9.0 16.7 65.3 24.5 4.5 8.0 63.0 38.0 5.0 10.0 47.0

12


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Utilities are introduced to the simulation in order to determine their consumption and cost rates. The default utility costs from Aspen Plus are used as they reflect typical chemical plant values. Additionally, a value is attributed to the ethylene loss based on current ethylene market prices. 34 These values are displayed in Table 2. Adopting the ethylene market price for the ethylene loss is a conservative approach. At this stage of the process, the ethylene price is certainly lower, since it still needs to be further processed to achieve chemical or polymer grade. Since the membranes are less selective to ethylene than the absorption, this assumption tends to favor the process configuration with standalone absorption. It ensures, however that the feasible scenarios are industrially applicable for a wide range of ethylene prices without the need for estimating the costs for the other process steps that are outside of the scope of this contribution. Table 2: Utility and ethylene loss costs Description Cost Steam 2.2 · 10−9 Cooling Water 2.12 · 10−10 Electricity 2.15 · 10−8 Ethylene Loss 1.00

Unit USD/J USD/J USD/J USD/kgC2 H4

Absorption Section An initial analysis is performed in order to determine the best operating pressure for the absorption column. The upstream water removal / compression section is typically operated at 10 bar and the downstream distillation section is operated at 32 bar. 35 These pressures are, therefore taken as the lower and upper bounds to be considered for the absorption section. The question is whether to perform the compression from 10 to 32 bar upstream or downstream of the absorber column. A higher pressure is beneficial, for it facilitates the CO2 absorption. This reduces the necessary liquid to gas ratio (L/G), which is the ratio between the total mass flow rate of amine solution and the total mass flow rate of gas, and also the reboiler duty. Therefore, the pressure of 32 bar has been previously 13




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cooling water are accounted for, but left out of the plot for sake of clarity. Despite the fact that the reboiler duty is lower at higher operating pressures, the ethylene loss and the total compression duty are higher, resulting in a higher operating cost. The ethylene loss cost rate is the most affected variable by the operating pressure. The main cost driver is the compression and performing the compression in the downstream of the absorber is preferred, for the total gas flow rate to be compressed is smaller given that the CO2 is removed. This is even more critical for the CO2 diluted scenarios, as the reduction in the total gas flow rate is more significant. Thus, the pressure of 10 bar is selected for the absorption column. By analyzing Figure 6, it also possible to picture that there may be a minimum total cost which lies below the lower bound of 10 bar imposed by the upstream water removal section. This shall be further investigated when synthesizing the full OCM process. The workflow developed to simulate the absorption section is very similar to the one previously described by Agbonghae et al., 36 however the applications described therein comprise CO2 capture for natural-gas-fired or coal-fired power plants as well as natural gas combined cycle power plants. In this contribution, further details are given on the specific tools inside Aspen Plus to provide a guideline on the simulation and optimal design of such systems using the software. This task is described by Agbonghae et al. as being "a combination of science and art based on experience". 36 The flowsheet starts with an equilibrium calculation of the absorption column. A proper starting value for the amine solution flow rate is specified based on previously reported L/G ratios to achieve convergence. The inlet lean amine solution is pre-loaded with approximately 0.1 molCO2 /molMEA since this will be a recycle stream from the stripping column once the flowsheet is complete. This strategy aids to close the recycle later. The absorber is operated at 10 bar and pressure drop is neglected at this point. With the column initialized in equilibrium mode, a packing type is selected to run a packing sizing, which calculates the column diameter for a given fractional capacity. The default fractional 15




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sensitivity analysis is then performed by changing the packed height and calculating the necessary amine solution flow rate in each case. Increasing the absorber packed height decreases the required amount of amine for a given CO2 capture target, thus reducing the L/G ratio. The sensitivity results for feed scenario II can be seen in Figure 7a. After arriving at 18 m of packed height, the reduction in the L/G ratio is below 0.1%, therefore this packed height is selected for the absorber. Once the absorber design is finished, the flash, the rich amine pump and the lean-rich exchanger are simulated. The pressure in the flash is nearly atmospheric to allow for a reasonable amount of CO2 to be removed from the rich amine stream. The rich pump increases the pressure to 2.5 bar. The lean-rich exchanger is initially simulated as a heater, given that no recycle stream is yet available. The procedure to design the stripping column is similar to that of the absorber, however specifications for the partial condenser and for the reboiler are necessary. The stripper is operated at 2.5 bar. The previously reported required duty of around 4 MJ/kgCO2 14 is employed to provide reasonable starting values for the reboiler duty while the distillate flow is set to be large enough compared to the inlet CO2 and H2 O flows. Once the stripper is converged in rate-based mode including the reaction kinetics, design specifications are introduced to fix a lean amine loading (αLEAN ) of 0.1 molCO2 /molMEA at the bottom by varying the reboiler duty, and the CO2 mass fraction at the top at 0.9 by varying the distillate flow rate. A sensitivity analysis is again performed by changing the stripper packed height and calculating the necessary reboiler duty on each case. The result of this sensitivity analysis for feed scenario II is presented in Figure 7b. The reduction in the reboiler duty by adding one meter of packing drops below 0.1% for a packed height of 12 m, therefore this packed height is selected for the stripping column. Figure 7 also presents a comparison between the two employed mass transfer correlations in both the absorption and stripping columns. While they differ in the lower range of packed heights, the results are identical in the actual operating range of both columns. In this case both correlations provide identical results and the one by Hanley and Chen 26 is 17




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preferred for convergence is achieved faster and easier when changing the packed height during the sensitivity analyses. The calculated pressure drop per unit height in the 2.7 m diameter absorber and 2.9 m diameter stripper are lower than the value of 2 mbar/m informed by the vendor and typically recommended in literature 37 as the maximum allowed pressure drop per unit height. The pressure drop are 1.2 mbar/m in the absorber and 0.4 mbar/m in the stripper. This indicates that the default fractional capacity of 0.62 adopted by the software is rather low for this type of packing, and a higher fractional capacity, between 0.7 and 0.8, could be applied resulting in smaller column diameters. The stripper’s bottom stream is pumped up to 10 bar in the lean pump and is now available for closing the energy recycle in the lean-rich-exchanger. The heater block is replaced by a heat exchanger block, which pre-heats the rich amine stream by cooling the lean amine stream. The difference between the hot stream outlet temperature and the cold stream inlet temperature is set to 10 K. The cooled amine solution enters a mixer/tank, in which a make-up stream is added to replace the water and amine losses in the gaseous streams. The make-up is calculated by a balance block. The outlet of the mixer is cooled down to 308 K and is recycled back to the absorber by the use of a transfer block, which copies this stream into the inlet lean amine stream, virtually closing the recycle. Since the design specification is fixing the αLEAN in the bottom of the stripper to the same value as the lean amine entering the absorber, these streams are nearly identical. Therefore, only few iterations are necessary to converge the flowsheet once the transfer block is activated. Attempts to close the recycle by actually connecting the streams are useless and would hardly permit any change in the operating point in order to perform any further analysis of the system. Thus, the presented strategy greatly improves the stability and usability of the model. A change in operating point can be easily achieved by the following steps. The transfer block is deactivated, essentially closing the recycle. The design variables are changed. The 19



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simulation is run once and once again after reactivating the transfer block. This procedure is used to conduct a sensitivity analysis for the lean amine loading. The rich amine loading at the bottom of the absorber (αRICH ) is bound by the amine absorption capacity to around 0.5 molCO2 /molMEA . It is possible to achieve the same CO2 capture duty by recirculating a higher amount of a more loaded amine (higher αLEAN ) or a lesser amount of a leaner amine (lower αLEAN ). Using a leaner amine solution is desired in the absorber, for its increased CO2 absorption capacity reduces the L/G ratio. However, more heat must be introduced in the stripping to regenerate it. Thus, using a higher αLEAN can decrease the reboiler duty up to a certain point. On the other hand, if αLEAN is too high, more amine is recirculated through the system, which causes additional heating, cooling and pumping costs and increases equipment size. In order to find the optimal operating point, the specification for the lean amine solution is varied from 0.1 to 0.2 molCO2 /molMEA and the utility costs for each case are reported. The sensitivity results for feed scenario II are presented in Figure 8. The minimum utility consumption can be achieved at around 0.14 molCO2 /molMEA , and the respective L/G ratio in the absorber is 2.12 kg/kg. The specific reboiler duty in this case is 3.85 MJ/kgCO2 . This methodology is repeated to all four feed scenarios, since the standalone absorption is to be used as reference to be compared with the hybrid processes. The total utility cost divided by the amount of captured CO2 is calculated as the specific absorption cost and assumed to be constant for each scenario during the optimization of the upstream membrane section.

Membrane Section The design of the membrane section is achieved through sequential optimization. This section is simulated in a separate flowsheet and different process configurations are considered. The standalone absorption (process configuration a) is taken as the base case for the comparison. For the hybrid processes, the following configurations are considered: (b) single membrane unit and absorption as in Figure 2, (c) single membrane unit with 20


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different manually introduced values of the permeate pressure. Splitters are added to allow for the membrane modules to be by-passed and for streams to be purged. These are not shown in the process flow diagrams for sake of clarity. The split ratios are also selected as optimization variables. By increasing the membrane areas, more CO2 can be captured, thus reducing the necessary amine flowrate and the reboiler duty in the absorption section. On the other hand, by increasing the membrane areas, more ethylene is lost and utility costs, such as electricity for compression, are also added in process configurations c, d, and e. This optimization intends to size the membrane section for each process configuration by providing a simple balance (simplified operational cost) between the costs on each section. For the flowsheets presenting a simplified operational cost lower than the operational cost of the base case (a), a rigorous simulation and design of the downstream absorption section is carried out as previously described in order to obtain a detailed operational cost and in order to design the equipment to allow for the equipment cost estimation.

Economic Analysis The operational cost (OPEX) is calculated as the sum of the cost rates for the utilities and for the ethylene loss. This is enough to allow for a comparison between the different process configurations studied and to analyze the CO2 capture step of the OCM process individually. An yearly operational factor of 0.95 is assumed, resulting in 8,322 operating hours per year. The capital investment cost (CAPEX) is calculated as the sum of all the installed equipment cost. The cost of equipment is estimated using Aspen Economic Analyzer, which allows for the on-the-fly sizing and cost estimation of equipment based on an up to date data bank. The absorber and stripper are structured packed columns as previously described. The centrifugal pumps are installed with a reserve in parallel. The leanrich-exchanger, the amine solution cooler, the condenser, and the gas coolers are classic 22


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TEMA BEM type shell and tubes exchangers, while the reboiler is a kettle type TEMA BKU exchanger. The mixer/tank and the condenser accumulator are horizontal pressure vessels. The compressor is a multi-stage centrifugal compressor. Costing for the membrane modules is carried out separately based on the pressure vessel, given that this is the major cost driver for the module. To keep calculations conservative, no degression factor for the serial production of the modules is applied and an installation cost factor of 1.2 is used. The base material for the equipment in the absorption section is stainless steel AISI 304, while carbon steel is adopted in the membrane section. The total cost of each process configuration is calculated as the sum of the operational cost (OPEX) with the annualized capital investment cost (CAPEX) as given in Equation 12. The IR is the interest rate of 15 % and N is the operational life of the plant, which is assumed to be 20 years.

TotalCost = OPEX + CAPEX ·

! IR · IR + 1) N (1 + IR) N − 1)

(12)

Results and Discussion All simulated flowsheets are able to achieve the CO2 removal target of 97 %. None of the hybrid process configurations present advantage over the standalone absorption for feed scenarios I and II. For these scenarios, the optimizer by-passes the membranes, or brings the membrane areas to the lower bound, or runs into local solutions worse than the base case. In these cases, the ethylene loss is costlier than the energy savings achieved by introducing an upstream membrane process. The low CO2 concentration in the gas streams does not allow for sufficient driving force for an effective separation through gas permeation. When considering gas flows diluted with N2 or CH4 in the OCM reaction step, standalone absorption offers a better performance for the downstream CO2 capture 23



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step. Most membranes for carbon dioxide separation, including the ones considered in this publication, have been developed and applied for natural gas sweetening and/or biogas upgrading, 17 to which product loss is not as critical given the lower value of the produced pipeline quality natural gas. The PEO membrane has a CO2 selectivity towards C2 H4 of around 3 at room temperature, which is found to be insufficient for the given industrial application. No economically feasible configuration applying the PEO membrane is found. For the hybrid process to become interesting with these feed scenarios, higher performance membranes, in terms of permeances and CO2 selectivity towards C2 H4 , are required. For the CO2 diluted gas streams (feed scenarios III and IV) the higher inlet CO2 concentration leads to a more efficient separation in the membranes. Furthermore, the higher CO2 amount to be removed requires a higher amine flowrate in the standalone absorption, which increases the ethylene loss in this section. By applying the PI membrane in both single and rectification cascade configurations (b and e respectively), the operational cost for the CO2 separation can be reduced. For configurations (c) and (d), the optimization led the membrane modules to be by-passed or brought the membrane areas to their lower bounds or the simplified operational cost is still higher than that of the base case (a), so that these configurations are left out of the detailed economic analysis. The single membrane configuration (b) offers the simplest alternative for the hybrid process, since only the installation of the membrane modules is necessary. For the rectification cascade (e), the permeate gas compressor and the coolers are also required. However, the rectification cascade allows for the recovery of plenty of the ethylene permeating through the first membrane module. Hence, it allows for increased membrane areas, drastically reducing the removal duty of the absorption section. For the case with the highest inlet CO2 concentration (feed scenario IV) the rectification cascade removes more than 60 % of the inlet CO2 while losing only 2 % of the inlet C2 H4 . Besides that, the permeate of the second membrane module is a nearly pure carbon dioxide stream that can be recycled directly as diluent to the reactor or employed in another process without any further processing. The main simulation results 24


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are presented in Table 3. The presented values are the membrane areas per module and the CO2 capture and C2 H4 loss in each section for the feasible combinations of feed scenario and process configuration. Table 3: Main simulation results for the feasible process configurations

Feed Scenario and Configuration IIIa IIIb IIIe IVa IVb IVe

Membrane Membrane Section Absorption Section Areas CO2 Capture C2 H4 Loss CO2 Capture C2 H4 Loss [m2 ] [%] [%] [%] [%] 97.0 1.5 1.0 17.9 1.3 79.1 1.3 3.6 & 6.2 38.9 0.8 58.1 1.0 97.0 2.3 2.5 39.5 3.7 57.5 1.6 5.5 & 7.2 63.5 2.0 33.5 1.1

From the capital investment point of view, the flowsheet for the hybrid process tends to be costlier as it has more equipment. However, since the membrane section takes up part of the CO2 removal duty, the equipment in the absorption section is smaller for the hybrid process. Notably, by adding the membrane section, the diameters of the columns can be reduced because the total gas flow rate is lower. The membrane modules are relatively cheap compared to the absorption and stripping columns. The construction materials that can be applied in the membrane section are also cheaper than those in the absorption section. Consequently, a trade-off between adding more equipment and reducing the size of current expensive equipment is observed. For feed scenario III, the equipment cost of both hybrid processes are slightly higher than that of the standalone absorption. For feed scenario IV, however the equipment cost for both hybrid processes are lower. As expected, the CAPEX for the rectification cascade is higher than the single membrane module, given that the compressor and the coolers are necessary. But, as this configuration reduces even further the removal duty of the absorption section, the actual increase in CAPEX is not that high. The calculated operational cost, the capital investment cost and the total cost are presented in Table 4. The process configuration with a single PI membrane (b) offers the 25



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lowest total cost when considering the feed scenario III. However, the difference in the total cost is relatively low and most likely way lower than the margin of error contained in the cost estimates. For feed scenario IV, the cheapest process configuration is the rectification cascade (e). The difference in this case is more significant, showing that the gas permeation is certainly a cost effective alternative when considering highly CO2 diluted gas streams for the OCM process. The annualized total cost for these CO2 removal steps are 104 USD/tC2 H4 and 128 USD/tC2 H4 . These costs are relatively low considering the current ethylene market price and that, for these feed scenarios, the CO2 capture step is to represent a large share in the final ethylene production cost. Table 4: Economic Analysis Results Scenario and Configuration IIIa IIIb IIIe IVa IVb IVe

OPEX CAPEX kUSD/year kUSD 8,142 15,653 7,696 17,024 7,518 18,706 12,122 23,922 11,262 18,542 9,726 19,764

Total Cost kUSD/year 10,643 10,416 10,507 15,944 14,224 12,884

Conclusion A model has been successfully developed in order to simulate an absorption and membrane based hybrid process for removing carbon dioxide from an outlet gas stream of the oxidative coupling of methane reactor. The workflow allows for the documentation-based modeling in MOSAIC with the subsequent integration of models into Aspen flowsheets. This is of particular interest for the implementation of customized unit operation models for novel process equipment. While MOSAIC is a convenient modeling platform, the custom unit operation model, in this case a gas separation membrane, can be later integrated into a process flowsheet in combination with well established unit operations for the process synthesis and cost estimation steps. 26


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The hybrid CO2 removal process for OCM is thoroughly analyzed and its economical feasibility is assessed taking into consideration the product (ethylene) loss. Different configurations of the hybrid process are compared with the state-of-the-art technology, which is the standalone absorption process with a 30 wt% Monoethanolamine (MEA) aqueous solution. It is demonstrated that for the CO2 capture step of the OCM process, the ethylene loss plays a major role in the economic feasibility, since ethylene is a high value product. For OCM reaction concepts employing N2 or CH4 dilution, the hybrid CO2 removal process is hardly of any economic advantage compared to the standalone absorption. While energy savings could be achieved, the product loss does not compensate it. The use of the PEO membrane for the given application is also found to be infeasible, given its low CO2 selectivity towards C2 H4 . On the other hand, when considering carbon dioxide diluted flows in the OCM reactor, the hybrid process is economically attractive. The concept of combining OCM with other CO2 intensive and specially endothermic processes, such as the dry reforming of methane, has a lot of potential and shall be further investigated. Gas permeation can bring a significant advantage for the CO2 removal step in this case. For feed scenario III, the best configuration is a single PI membrane, while for feed scenario IV a two-stage PI membrane rectification cascade is the most economical one.

Outlook It is important to model and to simulate in detail each step of the OCM process individually. This helps to enhance the understanding and to provide input for the design of the integrated flowsheet. However, for an optimal process synthesis, the whole process superstructure has to be considered simultaneously. To achieve that, surrogate models that are more suitable for optimization have to be applied. It is necessary to include economic terms, notably the operational cost, to the objective function to ensure that the solutions are 27



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economically feasible. Furthermore, the environmental impact of the OCM process must also be assessed within the process synthesis stage. Alternative separation techniques, such as pressure and temperature swing adsorption are still to be investigated. Finally, the OCM is also to be combined with other chemical processes, such as ethane cracking, considering that many shale gas compositions are also rich in ethane; dry reforming of methane, since it could use the heat and the carbon dioxide produced in the OCM reactor; and methanol production, given that it could make use of the H2 from the reformer and the available CO2 .

Acknowledgement The corresponding author gratefully acknowledges the support from CAPES, Coordination for the Improvement of Higher Education Personnel - Brazil (grant number: BEX 11946/130). Financial support from the Cluster of Excellence Unifying Concepts in Catalysis by the German Research Foundation (DFG EXC 314) is gratefully acknowledged.

Nomenclature Base Names σ

Lenard-Jones molecular diameter [Å]

CAPEX Capital cost/expenditure [USD] Ea

Activation Energy [kJ/kmol]

f

Fugacity [Pa]

h

Memnbrane height [m]

IR

Yearly interest rate [%] 28


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L

Permeance [kmol/(m2 s Pa)]

m0

Parameter in free-volume theory [Pa−1 ]

mT

Parameter in free-volume theory [K−1 ]

N

Number of operating years [years]

n

Molar flow [kmol/s]

OPEX Operational cost/expenditure [USD/year] R

Universal gas constant [kJ/(kmol K)]

z

Membrane length [m]

Subscripts Chemical component i

i

Superscripts 0

Reference condition

P

Permeate side

PEO Poly-(ethylene oxide) membrane PI

Polyimide membrane

R

Retentate side

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Graphical TOC Entry

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Design and Assessment of a Membrane and ... - ACS Publications

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