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Membrane module for pilot scale oxygen production
Journal of Membrane Science 574 (2019) 252–261
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Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci
Membrane module for pilot scale oxygen production a
Nicolas Nauels , Simone Herzog a b c
b,⁎
a,c
, Michael Modigell , Christoph Broeckmann
T
b
Aachener Verfahrenstechnik (AVT), Mechanical Process Engineering, RWTH Aachen University, Forckenbeckstraße 51, 52074 Aachen, Germany Institute for Materials Applications in Mechanical Engineering (IWM), RWTH Aachen University, Augustinerbach 4, 52062 Aachen, Germany German University of Technology in Oman (GUtech), Muscat, Oman
ARTICLE INFO
ABSTRACT
Keywords: BSCF OTM Membrane module Oxygen production MIEC
An oxygen transport membrane module capable of total 596 Ba0.5Sr0.5Co0.8Fe0.2O3-δ (BSCF) membranes was constructed and tested under various capacities. The module is designed to operate with up to 20 bar feed pressure and a vacuum on the permeate side. The separation performance was tested with of up to 96 membranes tubes. A maximum oxygen purity of 98.9% was obtained and a maximum oxygen flux of 2.8 ml cm−2 min−1 was measured under a feed pressure of 10 bar. The maximum permeate oxygen flow rate achieved was 22 l min−1. The performance is lower than expected, for which an inhomogeneous module temperature and suboptimal flow distribution can be accounted for, as CFD simulations show. A lot of membrane breakages occurred during operation. Some reasons for that could be identified and the module could be optimized throughout the 2 years of testing. Within that time, three long term tests were performed and a maximum continuous operational time of 1800 h was achieved.
1. Introduction The production of oxygen by means of membranes has attracted the interest of researchers for many years. Ceramic Oxygen Transport Membranes (OTMs) offer the potential to separate oxygen from air with an infinite selectivity. Numerous process simulations indicate, that by means of efficient process integration, the membrane based oxygen production can be less energy intensive than the state of the art oxygen production via cryogenic air distillation [1–5] for special classes of processes. Since OTMs provide high purity oxygen but also require high working temperatures, the integration into oxygen consuming high temperature processes is favourable. Special focus is on the oxyfuel combustion process for power plants as a carbon capture method, where a fuel is combusted in absence of N2 to obtain a highly CO2 enriched flue gas. Many process calculations indicate the benefit of the OTM technology in this field [6,7], although it is worth to mention, that still there is no general agreement about that [8]. The currently most investigated type of OTM is the mixed ion electron conducting membrane, which allows oxygen diffusion through lattice vacancies and a charge compensating flux of electronic charge carriers without external electrodes. The oxygen transport is usually described as a resistance in series model of surface exchange and bulk diffusion. For relatively thick membranes around 1 mm the flux is usually governed by the bulk diffusion, which can be described by the Wagner equation
⁎
jO2 =
RT · 16F 2
el· ion
ln pO2, f ln pO2, p
el
+
ion
dlnpO2
(1)
In Eq. (1) el and ion represent the electronic and ionic conductivities. F denotes the Faraday constant, R the gas constant and δ the membrane thickness. pO2 represents the oxygen partial pressure on feed- and permeate side, which is the driving force for the membrane process. The OTM material Ba0.5Sr0.5Co0.8Fe0.2O3-δ (BSCF) is one of the most investigated since it provides high fluxes. Baumann et al. reported fluxes above 10 ml cm−2 min−1 using supported thin layer membranes for separating air mixtures [9]. However, the material shows a chemical instability in contact with flue gas components like CO2 and SO2 which makes BSCF only suitable for 3-end integration [10]. Therefore, to provide a sufficient oxygen partial pressure gradient, a pressure difference between both sides of the membrane has to be applied. This can be achieved by pressurizing the feed air, by applying vacuum on the permeate side or a combination of both measures. All of these operating concepts bring their advantages and disadvantages. Applying only vacuum results in a high energy demand for the vacuum pump, which cannot be recovered in the process anymore. Furthermore one has to handle high volume flows which might cause high pressure drops especially in small channels e.g. in membrane hollow fibre systems. Reported oxygen fluxes of those asymmetric membranes in the 3-end mode under application oriented conditions with air on the feedside and a vacuum on the permeate side are about
Corresponding author. E-mail address: [email protected] (S. Herzog).
https://doi.org/10.1016/j.memsci.2018.12.061 Received 27 July 2018; Received in revised form 20 December 2018; Accepted 24 December 2018 Available online 25 December 2018 0376-7388/ © 2018 Elsevier B.V. All rights reserved.
Journal of Membrane Science 574 (2019) 252–261
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Fig. 1. Sketch and photo of the membrane module.
one magnitude of order below the idealised 4-end results [11,12]. Pressurizing the feed air and heating it afterwards to the required working temperature of the membranes offers the possibility to recover compression energy by expanding the retentate in a gas turbine, but highly compressed air at 850 °C sets high requirements to apparatus engineering and materials. Therefore, most concepts favour the combination of pressure and vacuum [7,13]. The optimal operating parameters depend on the integration into the process, meaning the level of external heat that is provided, the possibility of integrating waste heat into the process and the desired partial pressure ratio that determines the oxygen flux per membrane area. Although a lot of research is going on in the field of material synthesis and membrane manufacturing, only a small attention is paid to the actual module layout by now. First to mention is the OTM technology by Air Products, which is based on planar asymmetric membranes, assembled in multi wafer stacks. Air products has constructed a test unit with the capacity to produce up to 100 t per day [14]. Oxygen production of 16 t per day was achieved during a 300 h run at approximately 1/4 capacity. In a smaller scale high purity oxygen production under full process conditions for over 15,000 h is reported. Another concept with strong industrial participation is followed by Praxair, aiming on an application in syngas reforming with tubular membrane panels [9]. Since the pressurized reforming step takes place in the inner side of the tubes, there is no need of a complex pressure vessel. Tubular concepts with a focus on pure oxygen delivery are presented by [15] in form of a 19-tube BSCF demonstrator and in [16] with a 889-LaSrCoFe hollow fibre membrane module with a maximum yield 3.1 L min−1 oxygen with a purity of 99.9%. Both modules operate as only vacuum-driven concepts. This work presents the construction and operation of a membrane module that was constructed in the German federal research project Oxycoal-AC, which aims at investigating the integration of the OTM technology into an oxyfuel power plant process [7]. The module was developed in collaboration of the Institute for Materials Applications in Mechanical Engineering (IWM) and the Mechanical Process Engineering department (AVT) of the RWTH Aachen University. It can operate under elevated pressure up to 20 bar feed and a permeate vacuum < 50 mbar. The module is designed to provide oxygen with up to 596 dead-end BSCF membrane tubes. Furthermore, aspects like the fast replaceability of membranes were considered for the construction. Flow- and temperature distribution were determined via CFD simulations to identify optimization potential. This paper shows the results of the first testing phases from 2012 to 2014.
thickness of 0.8 mm and a length of approximately 500 mm are produced by the Institute for Materials Applications in Mechanical Engineering via cold isostatic pressing. The granulated powder (Treibacher) is filled into pressing moulds and is compacted with 180 MPa. The green tubes are cut to their desired length and sintered for 5 h at 1100 °C [17]. The membranes are joined to metal housings using Monoprox 1196 epoxy resin. Each membrane tube is subjected to quality assurance before being inserted into the module. This implies measuring the straightness of the tubes, the circularity and the wall thickness at selected points along the membrane length using a highend wall thickness measuring device MiniTest 7400FH. Finally, a burst test is conducted where the membranes are exposed to an outer pressure of 26 bar in a special designed testing facility. Besides the mechanical stability, also gas tightness of the membranes can be checked since small cracks in the membrane lead to upcoming air bubbles in a liquid column of the testing facility. 2.2. Module construction The module was constructed and built up in collaboration with the company Tisoma Anlagenbau. A schematic sketch is shown in Fig. 1. The module consists of a symmetric assembly of two pressure vessels with a water cooled flange in between. Each vessel is capable of 298 membrane tubes, which is equivalent to 14 m2 total membrane area when the module is fully equipped. In the upper half the membranes stand, while in the bottom vessel the membranes hang in the water cooled flange. The vessel is designed for a maximum operating temperature of 900 °C and a pressure of 20 bar. The incoming feed air is preheated electrically. Additionally, there is an electric filament inside the module to maintain a constant temperature and to compensate heat losses. The hot zone is insulated in order to keep the heat losses as low as possible and to keep inner wall temperatures of the vessel below 150 °C. The membranes are mounted in a water cooled flange located between the two pressure vessels. Gas tightness between the feed side and the permeate side is reached by rubber O-ring sealings for each membrane tube. This allows a replacement of a single membrane tube in case of damage. The flange was constructed and built by RWTH Aachen in collaboration with Peter Vossen GmbH. The water cooling assures on the one hand that the epoxy joint between the metal housing and membrane tube and the O-ring sealings do not exceed their critical temperatures. On the other hand, the permeating oxygen is cooled while passing the open end of the membrane tube so that no safety relevant hot oxygen has to be handled. In order to reduce the heat flux from the hot air side into the cooling water, the flange is covered with insulation material. Nevertheless, cooling the oxygen to ambient temperature leads to a loss of thermal energy which would reduce the energetic efficiency of the oxygen separation process.
2. Experimental 2.1. Membrane Fabrication and joining technique The dead-end BSCF tubes with an outer diameter of 15.5 mm, a wall 253
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Fig. 2. Process scheme of the pilot module, MFM = mass flow meter, MFC = mass flow Controller, PV = pressure Valve.
2.3. Process layout
membranes is calculated via this tool. Nowadays CFD simulations represent a well-established method to obtain detailed flow and concentration profiles in membrane separation processes. The commercial CFD software used in this work, ANSYS Fluent v. 14.5, solves the Navier Stokes Equations for finite volume elements to describe the viscous flow in the module. The diffusive species transport is considered by the Maxwell- Stefan diffusion model. For the simulation, a geometrical model of the fluid volume is built and divided into discrete cells, the mesh. Since highest resolution is required in the region of the membrane tubes, the mesh was built starting from sweeping the membranes with hexahedral layers. The rest of the fluid volume was filled with tetrahedral elements. The boundary conditions are listed in Table 1 and displayed in Fig. 3. The Navier Stokes Equations are solved for the feed compartment, while the membrane is treated as a boundary. The mass transport of oxygen across the membrane is implemented by a mathematical model for which a simplified approach of a method presented by Abdel-jawad [18] is used. A thin layer along the membranes' outer surface is defined via a user defined function as sink for mass, momentum and energy terms. The mass sink for each cell in this layer is calculated by a modified Wagner Eq. (2), where amb (T ) represents the ambipolar conductivity and follows an Arrhenius behaviour. Transport parameters used for the simulation were determined in laboratory experiments for single membrane tubes [19].
Fig. 2 shows the process scheme of the membrane module. The feed air flow (0–75 m3 h−1) is controlled by mass flow controllers (Bronkhorst). Unless specified differently, all flows and flow rates refer to Standard Temperature and Pressure (STP) conditions of 0 °C and 1 atm. Entering the module, the air is heated with LEISTER heating elements D8 up to 900 °C. The hot air passes the membrane tubes via an air distribution plate and the retentate is cooled at the exit of the module. Pressure (END-Armaturen), mass flow of the retentate (Bronkhorst) and the oxygen concentration (sensortechnics XYA) are measured and controlled in a cold state. On the permeate side of the membranes the oxygen is cooled and delivered to a vacuum pump (Pfeiffer Hena 200). The permeate pressure can be controlled from < 50 mbar to ambient pressure. The same components for measuring and controlling oxygen concentration, pressure and mass flow like in the retentate section are used in the permeate section, adapted to the measurement range. 20 thermocouples are used to measure and monitor the temperature at various points in the system, e.g. flange temperature, wall temperatures and the temperature in the inner part of the vessel. Also two thermocouples are inserted into selected membranes for measuring the inner membrane temperature. To guarantee a safe operation, wall temperatures of the vessel as well as the pressure in the vessels are constantly measured and monitored. An exceedance of critical values leads to an opening of a gas discharge safety valve which the piping is equipped with. Additionally, mechanical relief pressure valves are implemented redundant to the electronic monitoring. Furthermore, the flow of cooling water is monitored and connected to the shutdown measures. The vessels as well as the water cooled flange are pressure proved and certificated by the German technical control board TÜV. The expected surface temperatures and the resulting wall thicknesses of the pressure exposed parts were calculated by CFD and FEM simulations.
mO2,cell = A cell·
pO2,f RT amb (T) ·ln 16F 2 pO2,p
(2)
Both the temperature and the feed partial pressure can directly be accessed from each cell via the simulation. The adjacent partial pressure is calculated under the assumption of Dalton's law for an ideal gas mixture Table 1 Boundary conditions for the CFD simulation.
2.4. Simulation The efficiency of the oxygen transport via an OTM membrane is, according to Eq. (1), mainly determined by the process parameters temperature and oxygen partial pressure gradient across the membrane. A membrane module should therefore provide a sufficient and homogeneous temperature along the membranes and guarantee a sufficient feed oxygen transport to the membranes in terms of an optimized flow field distribution to avoid concentration polarization effects. Especially the last point is not or only poorly accessible via measurements in large scale membrane modules. In order to investigate the experimental results more deeply, selected experiments were analysed by using Computational Fluid Dynamics (CFD) simulations. Besides the fluid dynamic behaviour of the gas flow, also the oxygen transport across the 254
Parameter
Value
feed air inlet flow inlet temperature feed air pressure permeate pressure heating temperature upper flange temperature number of membranes effective membrane length effective active membrane area total built in membrane area heat conductivity flange (W m−1K −1) heat conductivity insulation (W m−1K −1) heat conductivity BSCF (W m−1K −1)
36 m³ h−1 850 °C 10 bar 50 mbar 900 °C 25 °C 298 370 mm 5.37 m2 7.25 m2 51 (20 °C) - 50.8 (100 °C) 0.19 (800 °C) - 0.27 (1000 °C) 2.26 (800 °C) - 3.1 (960 °C)
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Fig. 3. CAD Drawing of the geometry and mesh at the membrane tube.
pO 2, f = xO 2 ·ptot
(3)
The molar fraction x O2 and the local, static pressure ptot are easily accessible from the simulation. Since the module is operated in 3-end mode with a vacuum on the permeate side, the permeate oxygen partial pressure is assumed to be equal to the overall pressure on the inner side of the tubular membranes. Furthermore, pressure drop for the inner side of the membranes was neglected in awareness of the small volume flows and the relatively large diameter of the channels > 13 mm. According to the calculated mass flow of oxygen that is subtracted in the sink, a momentum and the enthalpy correction is performed in the simulation. 3. Results and discussion 3.1. Temperature distribution Fig. 4. Experimental temperature profile inside a hanging membrane, length 0 is the closed membrane end.
One essential point for the design of a membrane module is a homogeneous flow and temperature distribution, which turned out to be difficult to maintain constant within the vessel. While the vessel temperature was controlled to 850 °C, the membranes only reached a maximum temperature of about 830 °C at its tip and only 750 °C close to the insulation of the water cooled flange. Fig. 4 shows the axial temperature profile measured with a movable thermocouple which was placed inside a selected membrane during stationary operation. For a vessel temperature set point of 850 °C, only 30 cm of the membrane length are exposed to a temperature above 800 °C. The temperature drops rapidly within the flange insulation so that a surface temperature at the water cooled flange of about 150 °C is reached. At the joint between metal housing and ceramic membrane tube (50 cm) the temperature is uncritical for the used epoxy glue. With 37 cm of the 50 cm membrane protruding out of the insulation, only 74% of the membrane area take part in the air separation process, while the remaining 13 cm are required for cooling and sealing. That might seem like a waste of membrane material, but represents the only way to achieve a cold sealing between metal housing and ceramic membrane. Fig. 4 also points out the difficulty of determining the actual operation temperature. Unless otherwise noted, the temperature referred in the following diagrams is the vessel temperature, while the actual maximum membrane temperature is about 20 °C lower. A similar temperature distribution was obtained by CFD simulation, Fig. 5. The zones adjacent to the insulation of the water-cooled membrane fitting flange show a significant temperature difference to the zones close to the vessel wall. This may be interpreted as an indication that the heat loss over the flange cannot be compensated by the inner heating so that the set point of 850 °C cannot be kept constant in the entire feed regime. One reason for this is the spatial separation between
the inner flow chamber and the heating surfaces via the air distribution plate, whereby the convective heat transport from the heating surfaces is rather low. Higher temperatures were achieved in a later run by modifying the air distribution plate (compare long-term test C). 3.2. Permeation Measurements Various tests were performed with up to 96 membranes. In order to make the tests comparable to each other, results are expressed in terms of oxygen flux. All fluxes depicted in this paper were normalized to the active membrane area, meaning the part of the membranes that protrudes from the insulation was considered (370 mm length/membrane). Fig. 6 shows the oxygen flux and the purity of the membrane module with 47 membranes during start up of the module. Due to the low temperature gradient of 150 K/h a quasi-stationary behaviour of the module can be assumed. Additionally, the O2-flux of a single membrane tube part (10 cm), measured under lab-conditions, is depicted in that diagram for comparison. Main differences in that comparison are the well defined counter flow direction in the test rig compared to the cross flow in the actual module layout, but also the temperature distribution plays a main role, as the following diagrams show. With rising temperature, the permeation rate of the membranes increases according to the Wagner equation. At 850 °C and 12 m³ h−1 feed flow, a feed pressure of 5 bar and a vacuum pressure of 100 mbar, a flux of 1.68 ml cm−2 min−1 was obtained. This is significantly below the flux that can be obtained with a single tube in a laboratory 255
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Fig. 5. Calculated temperature distribution in the module.
Fig. 6. Oxygen flux, purity and leakage during start up with 47 membranes, feed air flow: 12 m³ h-1, pfeed :5 bar, ppermeate: 100 mbar. Fig. 7. Simulated velocity profile in the feed compartment.
experiment. Main reason for this deviation is the inhomogeneous temperature distribution in the module, but also maldistribution of flow accounts for a lower permeation rate, as CFD simulations show in terms of the velocity profile in Fig. 7. Highest fluid velocities appear at the inlet and outlet of the module as well as in the bottom part of the module below the hanging membranes. The high velocities indicate a short-cut flow below the membranes, whereby a big part of the feed air leaves the module without taking part in the membrane separation process. This is also indicated by the O2-mass fraction distribution in the module, Fig. 8. While the feed stream in the bottom part of the vessel contains an average oxygen mass fraction of about wO2 = 0.18, there is a depletion in the membranes zone with wO2 < 0.12. Thus the driving force for the oxygen transport is smaller than assumed when just taking the mean of the measured feed and retentate partial pressures as mean feedside partial pressure. An additional effect which decreases the driving force is the back swirl which can be identified in Fig. 8, which leads to a back mixing of already depleted air into the feed compartment. The CFD calculations reveal that the average O2-partial pressure of the cells adjacent to the membranes is about 14% below the average partial pressure of all cells in the module. During all start up procedures an increase of oxygen purity was
Fig. 8. Simulated oxygen mass fraction in the feed compartment. 256
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measured as shown in Fig. 6. The oxygen purity in that example increases from 79% at 640 °C to 98.3% at 850 °C. This results mainly from the fact that the amount of permeating oxygen increases while the amount of leaking nitrogen into the system remains constant or even drops with increasing temperature. At 640 °C, the leakage rate is around 0.1 ml cm−2min−1 and slightly decreases to 0.04 ml cm−2min−1 at 850 °C. It cannot be distinguished if the leakage is caused by material defects in the ceramic or from the system like leaky joints. The small decrease of the leakage with elevating temperatures could be a hint of resintering and closing of microcracks in the membrane, which might be assisted by the high feed pressures, but also the increasing gas viscosity with temperature, leading to a lower viscous flow inside small pores could be a more simple explanation. Nevertheless, the effect of the increasing oxygen purity was also seen during long time operation at constant temperature (compare Fig. 11). Fig. 9 shows the oxygen flux for a test with 96 membranes during a pressure increase after start up. Feed pressure was varied from 2.0 to 3.5 bar and permeate pressure from 50 to 100 mbar. With increasing feed pressure the temperature in the module drops due to the lower flow velocity at higher pressure, which reduces the convective heat transport from the heating elements to the inner fluid chamber. Despite of the temperature drop, there is a continuous increase of the permeation rate with increasing feed pressure. Maximum total permeate oxygen flow rate achieved in this experiment was 22 l min−1, resulting in a flux of 1.27 ml cm−2 min−1 at an oxygen purity of 98.9%. The theoretical O2-flux of the membrane tube, assuming a uniform membrane temperature under these pressure gradients, is by 25% higher than the experimental one for 100 mbar permeate pressure. This is mainly due to the inhomogeneous temperature distribution in the module. Lowering the permeate pressure to 50 mbar increases the difference up to 35% from the theoretical values. This indicates a significant feed side mass transport limitation caused by local short-circuit flow. The packing density with 96 membranes might intensify this effect, as it can be seen compared to an experiment equipped with 38 membranes depicted in Fig. 10. Some modifications in the air distribution were made to enable a better heat exchange between the heating elements and the feed air, so that a vessel temperature of 880 °C could be reached. Fig. 10 shows the oxygen flux of the experiment with 38 membranes (Aact = 0,7 m2) during a pressure increase from 5 to 10 bar. By elevating the
Fig. 10. Oxygen flux and temperature for a run with 38 membranes (Aact=0,7 m2), pperm = 50 mbar, feed air flow: = 12 m³h-1, modified air distribution.
temperature, a higher permeation rate could be achieved compared to the previously described experiments. The highest measured oxygen flux under these conditions was 2.8 ml cm−2 min−1, which is a total production of 19.6 l min−1. The deviation between the theoretical and the measured permeation rate is 23% at 5 bar feed air pressure. Increasing the feed pressure to more than 8 bar does not lead to a significant increase in permeation rate which is a further indication of a limitation of the feed side oxygen transport to the membrane surface, although the oxygen recovery rate is only between 0.3 and 0.4 under these conditions. 3.3. Long term tests In order to prove the feasibility of the module operation, long term tests were performed. Within the time of operation three tests, following named as (A), (B) and (C), were performed aiming on an operation time of 1000 h. Several changes were made in the module configuration between each long term test, which are described in the following:
• Test (A) was operated with 96 membranes at varying feed pressures •
•
between 1 and 5 bar abs. The vessel temperature was adjusted to 850 °C and a permeate pressure of 50 mbar was applied. Test (A) was stopped after 1000 h. Test (B) was operated with 47 membranes, a feed pressure of 5 bar and a permeate pressure of 100 mbar. Furthermore changes in the pressure control system were made and a new insulation material for the flange insulation was mounted. Test (B) only reached an operation time of 580 h due to an emergency shutdown resulting from a failure of the cooling water system. Test (C) was operated with 38 membranes at 5 bar and a permeate pressure of 50 mbar. Furthermore constructive changes in the air distribution plate were made that allowed a higher vessel temperature of 880 °C (Tmem=860 °C). The test reached an operation time of 1800 h, whereby after 1000 h an increase of the feed pressure from 5 to 10 bar was done. Test (C) stopped again with an emergency shutdown due to a failure of the mass flow controller.
Fig. 11 and Fig. 12 show the oxygen flux and purity for the different tests. Due to many stops and changing conditions during the test, the flux of Test (A) is not displayed continuously. For test (A) a maximum O2-flux of 1.27 ml cm−2 min−1 was measured. The oxygen purity at the beginning reached 95.4%, but varied strongly over the tested period
Fig. 9. Oxygen flux and temperature for a run with 96 membranes (Aact=1,73 m2) at varying feed pressure, pperm = 50 mbar, feed air flow: = 12 m³h-1. 257
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3.4. Membrane breakages and failures During the first test phases, more than 300 membranes have been set in. The reasons for membrane breakages were diverse, but certain trends were observed. The majority of the membranes broke during or after emergency shutdowns, which were mostly caused by external failures like malfunctions of temperature sensors, preheaters or cooling water. In that case, all pressure valves open and the heating shuts down. This results in mechanical and thermal stress for the membranes. The pressure instantaneously releases from operation pressure (10 bar) to ambient pressure within 2 min. Simultaneously the membrane temperature drops due to the air expansion and the shutdown of the heating and the preheaters by approximately 70 °C within that time, which means a high stress to the ceramics. During the first tests, also the pressure relief valves on the permeate side opened in case of an emergency shutdown, and in presence of a vacuum inside the membrane tubes cold air is sucked into the membrane tubes, which might have led to membrane breakage due to a thermo-shock. This problem has recently been solved by applying a back-pressure valve at the permeate duct. Besides the breakages that occurred during emergency shutdowns, most frequently membrane breakages were observed during the cool down phase between temperatures of 250 °C and 310 °C. This is reported to be the temperature region where BSCF shows a minimum in the characteristic fracture stress curve [20]. Even with a controlled shutdown of the module it was almost impossible to prevent membrane breakages, thus the cooling down phase represents a critical step in the module operation, although it is not yet understood what exactly causes the membrane to break. The vast majority of fractures occurred at the position where the membrane emerges the insulation and enters the hot zone, see Fig. 13. At this position, the membrane is exposed to a large local temperature gradient, which obviously leads to a significant degradation: while "new" membrane tubes were tested to stand an internal pressure of 1 bar (g), aged membranes after module operation showed a significant reduced stability to the inside pressure test and already broke at internal pressure tests at 200 mbar(g) at that position. Additionally, when breaking the vacuum on the permeate side, like in case of emergency shutdowns or membrane breakages, the incoming air is sucked via the watercooled flange into the membrane tubes and therefore the indicated position becomes the first spot to expire the thermo-shock as described before. Besides the described damages, secondary damage caused by falling membranes can be observed. In the worst case this can lead to a domino effect and damage the majority of the built in membranes. It is almost impossible to identify the origin of the membrane breakage in this case. In case of leakages or single membrane breakages the module could be shut down and the defect membrane could be replaced. Therefore the controlled shut down of the heating while maintaining the
Fig. 11. Oxygen purity during the 3 long term tests under varying conditions.
Fig. 12. O2 flux during long term tests (B) and (C).
due to membrane breakages and replacements. In total, the experiment had to be stopped 6 times during the 1000 accumulated hours due to breakages or major leaks of in total 20 membranes. The varying oxygen purity results from the replacement of single membranes during the stops. After the previously described constructive changes, test (B) showed a significantly improved stability. An oxygen purity after start-up of 97% was reached, slightly increasing to 98.5% after 580 h. A maximum flux of 1.69 ml cm−2 min−1 was obtained under the given conditions, which slightly decreased (-5%) during time to a value of 1.6 ml cm−2 min−1 after 580 h. Test (C) showed the highest flux, since through the module modifications a higher temperature could be applied. Small oscillations in the flux result from oscillations in the pressure control of the vacuum or the feed side. Highest measured oxygen flux within the first 1000 h was determined to 2.38 ml cm−2 min−1 after 560 h, while after 1000 h a flux of 2.3 ml cm−2 min−1 was obtained. In contrast to the improved permeation rate, low oxygen purity was obtained for this test. The purity at start-up only reached a value of 70%, caused by a high air leakage rate of ~ 1 ml cm−2 min−1. During the 1000 h long term test, a decrease of the air leakage rate to 0.21 ml cm−2 min−1 was observed leading to an oxygen purity of 93.3%. Since no changes in operation mode or module configuration were performed during the first 1000 h, the increase of the oxygen purity is an indication for resintering of microcracks, which was also investigated via grain size analysis in Section 3.4. After 1000 h, a change in feed pressure was made from 5 to 10 bars, resulting in a higher flux of 2.8 ml cm−2 min−1 but a decrease in oxygen purity, which partially recovered after longer operation time.
Fig. 13. Typical breakage after shutdown. 258
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Fig. 14. SEM (backscattered electron image) micrographs of the feed side of the membrane before and after 1800 h of testing,.
adjacent to the water cooled flange, so a formation of a hexagonal phase cannot be precluded, which could also be an explanation of the decreasing permeation rate for the long term tests. As it can clearly be seen in Fig. 15, this change in composition along the grain boundaries also occurs on the permeate side of the membrane. The superficial degradation of the permeate side appears to be lower than the one on the feed side. After 1800 h of testing, the grain boundaries on the surface are still visible with SEM. Also particles seem to have formed on the surface and especially along the grain boundaries, which again were identified as slightly changed BSCF compositions characterized by iron depletion. As expected in the 3-end mode, almost no other components like Cr or S occurred on the permeate side although some spots containing chloride were found during SEM. The source for that contamination is probably attributed to a lubricant paste used for a screw connection in the permeate chamber. It is remarkable, that no carbonate formations have been detected on the membranes, indicating that the CO2 content of the ambient air does not seem to influence the permeation behaviour in a significant way over the tested time range. Also, a growth of the grain size was observed via image analysis, as depicted in Fig. 16. The shift of the average grain size d50 from 18 to 23 µm shows a sintering activity even at the operation temperatures of 850 °C, what can be taken as an additional indication of the resintering of microcracks and the correlated increase of the oxygen purity.
operation pressure showed the best success since tensile stresses of the hanging membranes were avoided that way. An identification of the defect membrane tubes in the module could only be done by visual inspection or by localizing obvious flows through the single tubular membranes at a small pressure difference. 3.5. Aging and degradation Fig. 14 shows SEM images of the membranes from the long term test (C) for the membrane surface on the feed side. At the initial condition the grain boundaries are clearly visible. After 1800 h of testing, a clear surface degradation is obvious. The formerly smooth surface is covered with fine particles in the hot region of the membrane. Via EDX analysis of these particles, a significant amount of chromium was detected, resulting in with up to 15 at%, which has evaporated from the high temperature steels. Also spots containing about 2 at% sulfur were determined via EDX. Chromium and sulfur formations are known to hinder oxygen permeation and could therefore be an explanation for the slightly decreasing permeation rate in test (B). The source of the sulfur has not been clarified yet, but it is supposed to originate from oil residues from the compressor unit. The described contamination of the membrane could not be observed in cross-sections of the aged membranes, being a hint for a clear surface effect in the submicron range. However, this still has to be kept in mind especially when thinking about supported thin layer membranes with a lifetime demand of several thousand hours. A clear change after 1800 h compared to the initial membrane can be seen in the cross section image, where the grain boundaries are clearly visible by its bright colour. Along the grain boundaries a change in the material composition was determined via EDX, that can mainly be characterized by a significant depletion of iron and a minor depletion of strontium. This decomposition occurring predominantly at the grain boundaries has previously been reported by [11] and was there contributed to a formation of a hexagonal phase, but it is worth to mention that for our investigated membranes no clear evidence of a hexagonal phase was detected via XRD measurements and also the average temperature of membranes was with ~860 °C above that limit which is usually reported for hexagonal phase formation [11,21,22]. But as shown before, there exist zones of lower membrane temperature
4. Conclusions The design of modules is an essential point on the way to enable the OTM technology for high purity oxygen production. Since the oxygen transport membranes operate at elevated temperatures and high pressures, the large scale module construction holds some difficulties. A membrane module capable of 596 tubular membranes was constructed and tested with various capacities, at a maximum of 96 tubes. Under operating pressure of 10 bars and 880 °C, a flux of 2.8 ml cm−2 min−1 was obtained. The highest oxygen purity was determined to 98.9%. Despite of a quality control of every single membrane, a leakage could not be avoided completely and an identification of leaky membranes turned out to be very difficult. With oxygen purity below 99%, this technology loses one of its biggest advantages compared to the 259
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Fig. 15. SEM (backscattered) micrographs of the permeate side of the membrane as sintered and after testing.
testing period of 1800 h could not be exceeded. The testing phase showed a general feasibility of that concept, but there is a huge potential for improvement. Besides an improved membrane manufacturing and quality control, a hot membrane-metal joining/sealing and therefore an avoidance of the watercooled flange is mandatory to overcome the problems with the inhomogeneous temperature distribution in the module. This would also eliminate the temperature gradient along the membranes and therefore the spot where the vast majority of membrane breakages occurred. Acknowledgements This work was supported by BMWi and MIWFT as well as by the companies RWE Power AG, E.On AG, Hitachi Power Europe, Linde AG, MAN Turbo, WS-Wärmeprozesstechnik and Vattenfall [grant number 0326890]. References
Fig. 16. Grain size distribution of the membrane before and after 1800 h operation. Examination of approx. 4000 grains in thermally etched specimens by image analysis.
[1] P. Dyer, Ion transport membrane technology for oxygen separation and syngas production, Solid State Ion. 134 (1–2) (2000) 21–33, https://doi.org/10.1016/ S0167-2738(00)00710-4. [2] J. Nazarko, M. Weber, E. Riensche, D. Stolten, Oxygen Supply for Oxyfuel Power Plants by Oxy-Vac-Jül Process, in: D. Stolten, V. Scherer (Eds.), Efficient Carbon Capture for Coal Power Plants, Wiley, 2011, pp. 131–137. [3] R. Castillo, Thermodynamic evaluation of membrane based oxyfuel power plants with 700 ° C technology, Energy Procedia 4 (2011) 1026–1034, https://doi.org/10. 1016/j.egypro.2011.01.151. [4] J.M. Repasky, V.E. Stein, P.A. Armstrong, Ceramic and Coal: ITM Oxygen for Power Generation with Reduced CO2-Emissions, Detailed Engineering Study Results, in AIChE Spring National Meeting: San Antonio, TX, 2013. [5] A. Mak, E.P. Foster, P.A. Armstrong, J.M. Repasky, V.E. Stein, ITM Oxygen Technology for Clean Energy Applications, in: Proceedings of the 30th Annual International Pittsburgh Coal Conference 2013: a meeting held 15-18 September 2013, Beijing, China, pp. 270–300. [6] T. Ochs, D.B. Oryshchyn, C.A. Summers, Integrated Pollutant Removal: Modeling and Experimentation. 〈https://www.osti.gov/servlets/purl/899047〉 (accessed 26 July 2018), 2005. [7] H. Stadler, F. Beggel, M. Habermehl, B. Persigehl, R. Kneer, M. Modigell, P. Jeschke, Oxyfuel coal combustion by efficient integration of oxygen transport membranes, Int. J. Greenh. Gas. Con. 5 (1) (2011) 7–15, https://doi.org/10.1016/j.ijggc.2010. 03.004. [8] I. Pfaff, A. Kather, Comparative thermodynamic analysis and integration issues of CCS steam power plants based on oxy-combustion with cryogenic or membrane based air separation, Energy Procedia 1 (1) (2009) 495–502, https://doi.org/10. 1016/j.egypro.2009.01.066.
cryogenic ASU. An implementation of that the OTM technology therefore requires a higher purity which could be reached by an improved quality control that allows leak tests of single membranes also in the built in state. The permeation rate obtained was below that one measured on single membrane tubes, which was mostly due to an inhomogeneous temperature distribution that results from the water cooled parts of the module. Also the investigations showed some potential of improvement for the flow duct in the module, which also contributed to the lower permeation rate. The biggest problem during the experiments in the pilot scale were the breakages of membrane tubes, where one single breakage caused a shutdown of the module for several days and in the worst case a breakage of most built-in membranes. The fact that most breakages occurred due to externally caused shutdowns emphasizes the need for an optimized process control as well as for safety measures to avoid large membrane losses. During the test phase, significant improvement regarding the reliability of the module was reached, nevertheless a 260
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N. Nauels et al. [9] S. Baumann, J.M. Serra, M.P. Lobera, S. Escolástico, F. Schulze-Küppers, W.A. Meulenberg, Ultrahigh oxygen permeation flux through supported Ba0.5Sr0.5Co0.8Fe0.2O3−δ membranes, J. Membr. Sci. 377 (1−2) (2011) 198–205, https://doi.org/10.1016/j.memsci.2011.04.050. [10] S. Engels, T. Markus, M. Modigell, L. Singheiser, Oxygen permeation and stability investigations on MIEC membrane materials under operating conditions for power plant processes, J. Membr. Sci. 370 (1–2) (2011) 58–69, https://doi.org/10.1016/j. memsci.2010.12.021. [11] X. Li, T. Kerstiens, T. Markus, Oxygen permeability and phase stability of Ba0.5Sr0.5Co0.8Fe0.2O3−δ perovskite at intermediate temperatures, J. Membr. Sci. 438 (2013) 83–89, https://doi.org/10.1016/j.memsci.2013.03.017. [12] P. Niehoff, S. Baumann, F. Schulze-Küppers, R.S. Bradley, I. Shapiro, W.A. Meulenberg, P.J. Withers, R. Vaßen, Oxygen transport through supported Ba0.5Sr0.5Co0.8Fe0.2O3−δ membranes, Sep. Purif. Technol. 121 (2014) 60–67, https://doi.org/10.1016/j.seppur.2013.07.002. [13] J.F. Vente, W.G. Haije, R. IJpelaan, F.T. Rusting, On the full-scale module design of an air separation unit using mixed ionic electronic conducting membranes, J. Membr. Sci. 278 (1–2) (2006) 66–71, https://doi.org/10.1016/j.memsci.2005.10. 044. [14] Lori L. Anderson, Phillip A. Armstrong, Robert R. Broekhuis, Michael F. Carolan, Jack Chen, Mark D. Hutcheon, Charles A. Lewinsohn, Christopher F. Miller, John M. Repasky, Dale M. Taylor, Charles M. Woods, Advances in ion transport membrane technology for oxygen and syngas production, Solid State Ion. 288 (2016) 331–337, https://doi.org/10.1016/j.ssi.2015.11.010. [15] R. Kriegel, I. Voigt, Demonstration of oxygen separation with BSCF membranes, in:
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12. Jülicher Werkstoffsymposium: Proceedings of a meeting held 05-06 October, Jülich, Germany, 2011. X. Tan, Z. Wang, B. Meng, X. Meng, K. Li, Pilot-scale production of oxygen from air using perovskite hollow fibre membranes, J. Membr. Sci. 352 (1–2) (2010) 189–196, https://doi.org/10.1016/j.memsci.2010.02.015. E.M. Pfaff, A. Kaletsch, C. Broeckmann, Design of a Mixed Ionic/Electronic Conducting Oxygen Transport Membrane Pilot Module, Chem. Eng. Technol. 35 (3) (2012) 455–463, https://doi.org/10.1002/ceat.201100447. M.M. Abdel-jawad, S. Gopalakrishnan, M.C. Duke, M.N. Macrossan, P.S. Schneider, J.C. Diniz, da Costa, Flowfields on feed and permeate sides of tubular molecular sieving silica (MSS) membranes, J. Membr. Sci. 299 (1–2) (2007) 229–235. S. Engels, Mass transfer properties of mixed ion electron conducting membranes for the oxygen separation from air, [dissertation], 〈http://publications.rwth-aachen. de/record/82895/files/4148.pdf〉. E. Pfaff, M. Oezel, A. Eser, A. Bezold, Reliability of ceramic membranes of BSCF for oxygen separation in a pilot membrane reactor, in: Proceedings of the 10th Pacific Rim Conference on Ceramic and Glass Technology: a meeting held 2-7 June 2013 San Diego, CA, USA, 2013. Z. Shao, Investigation of the permeation behavior and stability of a Ba0.5Sr0.5Co0.8Fe0.2O3−δ oxygen membrane, J. Membr. Sci. 172 (1–2) (2000) 177–188, https://doi.org/10.1016/S0376-7388(00)00337-9. D.N. Müller, Structural and electronic investigations on the highly oxygen deficient mixed ionic electronic conductor Ba(x)Sr(1-x)Co(0,8)Fe(0,2)O3-delta (BSCF), [dissertation], 〈http://publications.rwth-aachen.de/record/63061/files/3746. pdf〉.
Contents lists available at ScienceDirect
Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci
Membrane module for pilot scale oxygen production a
Nicolas Nauels , Simone Herzog a b c
b,⁎
a,c
, Michael Modigell , Christoph Broeckmann
T
b
Aachener Verfahrenstechnik (AVT), Mechanical Process Engineering, RWTH Aachen University, Forckenbeckstraße 51, 52074 Aachen, Germany Institute for Materials Applications in Mechanical Engineering (IWM), RWTH Aachen University, Augustinerbach 4, 52062 Aachen, Germany German University of Technology in Oman (GUtech), Muscat, Oman
ARTICLE INFO
ABSTRACT
Keywords: BSCF OTM Membrane module Oxygen production MIEC
An oxygen transport membrane module capable of total 596 Ba0.5Sr0.5Co0.8Fe0.2O3-δ (BSCF) membranes was constructed and tested under various capacities. The module is designed to operate with up to 20 bar feed pressure and a vacuum on the permeate side. The separation performance was tested with of up to 96 membranes tubes. A maximum oxygen purity of 98.9% was obtained and a maximum oxygen flux of 2.8 ml cm−2 min−1 was measured under a feed pressure of 10 bar. The maximum permeate oxygen flow rate achieved was 22 l min−1. The performance is lower than expected, for which an inhomogeneous module temperature and suboptimal flow distribution can be accounted for, as CFD simulations show. A lot of membrane breakages occurred during operation. Some reasons for that could be identified and the module could be optimized throughout the 2 years of testing. Within that time, three long term tests were performed and a maximum continuous operational time of 1800 h was achieved.
1. Introduction The production of oxygen by means of membranes has attracted the interest of researchers for many years. Ceramic Oxygen Transport Membranes (OTMs) offer the potential to separate oxygen from air with an infinite selectivity. Numerous process simulations indicate, that by means of efficient process integration, the membrane based oxygen production can be less energy intensive than the state of the art oxygen production via cryogenic air distillation [1–5] for special classes of processes. Since OTMs provide high purity oxygen but also require high working temperatures, the integration into oxygen consuming high temperature processes is favourable. Special focus is on the oxyfuel combustion process for power plants as a carbon capture method, where a fuel is combusted in absence of N2 to obtain a highly CO2 enriched flue gas. Many process calculations indicate the benefit of the OTM technology in this field [6,7], although it is worth to mention, that still there is no general agreement about that [8]. The currently most investigated type of OTM is the mixed ion electron conducting membrane, which allows oxygen diffusion through lattice vacancies and a charge compensating flux of electronic charge carriers without external electrodes. The oxygen transport is usually described as a resistance in series model of surface exchange and bulk diffusion. For relatively thick membranes around 1 mm the flux is usually governed by the bulk diffusion, which can be described by the Wagner equation
⁎
jO2 =
RT · 16F 2
el· ion
ln pO2, f ln pO2, p
el
+
ion
dlnpO2
(1)
In Eq. (1) el and ion represent the electronic and ionic conductivities. F denotes the Faraday constant, R the gas constant and δ the membrane thickness. pO2 represents the oxygen partial pressure on feed- and permeate side, which is the driving force for the membrane process. The OTM material Ba0.5Sr0.5Co0.8Fe0.2O3-δ (BSCF) is one of the most investigated since it provides high fluxes. Baumann et al. reported fluxes above 10 ml cm−2 min−1 using supported thin layer membranes for separating air mixtures [9]. However, the material shows a chemical instability in contact with flue gas components like CO2 and SO2 which makes BSCF only suitable for 3-end integration [10]. Therefore, to provide a sufficient oxygen partial pressure gradient, a pressure difference between both sides of the membrane has to be applied. This can be achieved by pressurizing the feed air, by applying vacuum on the permeate side or a combination of both measures. All of these operating concepts bring their advantages and disadvantages. Applying only vacuum results in a high energy demand for the vacuum pump, which cannot be recovered in the process anymore. Furthermore one has to handle high volume flows which might cause high pressure drops especially in small channels e.g. in membrane hollow fibre systems. Reported oxygen fluxes of those asymmetric membranes in the 3-end mode under application oriented conditions with air on the feedside and a vacuum on the permeate side are about
Corresponding author. E-mail address: [email protected] (S. Herzog).
https://doi.org/10.1016/j.memsci.2018.12.061 Received 27 July 2018; Received in revised form 20 December 2018; Accepted 24 December 2018 Available online 25 December 2018 0376-7388/ © 2018 Elsevier B.V. All rights reserved.
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Fig. 1. Sketch and photo of the membrane module.
one magnitude of order below the idealised 4-end results [11,12]. Pressurizing the feed air and heating it afterwards to the required working temperature of the membranes offers the possibility to recover compression energy by expanding the retentate in a gas turbine, but highly compressed air at 850 °C sets high requirements to apparatus engineering and materials. Therefore, most concepts favour the combination of pressure and vacuum [7,13]. The optimal operating parameters depend on the integration into the process, meaning the level of external heat that is provided, the possibility of integrating waste heat into the process and the desired partial pressure ratio that determines the oxygen flux per membrane area. Although a lot of research is going on in the field of material synthesis and membrane manufacturing, only a small attention is paid to the actual module layout by now. First to mention is the OTM technology by Air Products, which is based on planar asymmetric membranes, assembled in multi wafer stacks. Air products has constructed a test unit with the capacity to produce up to 100 t per day [14]. Oxygen production of 16 t per day was achieved during a 300 h run at approximately 1/4 capacity. In a smaller scale high purity oxygen production under full process conditions for over 15,000 h is reported. Another concept with strong industrial participation is followed by Praxair, aiming on an application in syngas reforming with tubular membrane panels [9]. Since the pressurized reforming step takes place in the inner side of the tubes, there is no need of a complex pressure vessel. Tubular concepts with a focus on pure oxygen delivery are presented by [15] in form of a 19-tube BSCF demonstrator and in [16] with a 889-LaSrCoFe hollow fibre membrane module with a maximum yield 3.1 L min−1 oxygen with a purity of 99.9%. Both modules operate as only vacuum-driven concepts. This work presents the construction and operation of a membrane module that was constructed in the German federal research project Oxycoal-AC, which aims at investigating the integration of the OTM technology into an oxyfuel power plant process [7]. The module was developed in collaboration of the Institute for Materials Applications in Mechanical Engineering (IWM) and the Mechanical Process Engineering department (AVT) of the RWTH Aachen University. It can operate under elevated pressure up to 20 bar feed and a permeate vacuum < 50 mbar. The module is designed to provide oxygen with up to 596 dead-end BSCF membrane tubes. Furthermore, aspects like the fast replaceability of membranes were considered for the construction. Flow- and temperature distribution were determined via CFD simulations to identify optimization potential. This paper shows the results of the first testing phases from 2012 to 2014.
thickness of 0.8 mm and a length of approximately 500 mm are produced by the Institute for Materials Applications in Mechanical Engineering via cold isostatic pressing. The granulated powder (Treibacher) is filled into pressing moulds and is compacted with 180 MPa. The green tubes are cut to their desired length and sintered for 5 h at 1100 °C [17]. The membranes are joined to metal housings using Monoprox 1196 epoxy resin. Each membrane tube is subjected to quality assurance before being inserted into the module. This implies measuring the straightness of the tubes, the circularity and the wall thickness at selected points along the membrane length using a highend wall thickness measuring device MiniTest 7400FH. Finally, a burst test is conducted where the membranes are exposed to an outer pressure of 26 bar in a special designed testing facility. Besides the mechanical stability, also gas tightness of the membranes can be checked since small cracks in the membrane lead to upcoming air bubbles in a liquid column of the testing facility. 2.2. Module construction The module was constructed and built up in collaboration with the company Tisoma Anlagenbau. A schematic sketch is shown in Fig. 1. The module consists of a symmetric assembly of two pressure vessels with a water cooled flange in between. Each vessel is capable of 298 membrane tubes, which is equivalent to 14 m2 total membrane area when the module is fully equipped. In the upper half the membranes stand, while in the bottom vessel the membranes hang in the water cooled flange. The vessel is designed for a maximum operating temperature of 900 °C and a pressure of 20 bar. The incoming feed air is preheated electrically. Additionally, there is an electric filament inside the module to maintain a constant temperature and to compensate heat losses. The hot zone is insulated in order to keep the heat losses as low as possible and to keep inner wall temperatures of the vessel below 150 °C. The membranes are mounted in a water cooled flange located between the two pressure vessels. Gas tightness between the feed side and the permeate side is reached by rubber O-ring sealings for each membrane tube. This allows a replacement of a single membrane tube in case of damage. The flange was constructed and built by RWTH Aachen in collaboration with Peter Vossen GmbH. The water cooling assures on the one hand that the epoxy joint between the metal housing and membrane tube and the O-ring sealings do not exceed their critical temperatures. On the other hand, the permeating oxygen is cooled while passing the open end of the membrane tube so that no safety relevant hot oxygen has to be handled. In order to reduce the heat flux from the hot air side into the cooling water, the flange is covered with insulation material. Nevertheless, cooling the oxygen to ambient temperature leads to a loss of thermal energy which would reduce the energetic efficiency of the oxygen separation process.
2. Experimental 2.1. Membrane Fabrication and joining technique The dead-end BSCF tubes with an outer diameter of 15.5 mm, a wall 253
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Fig. 2. Process scheme of the pilot module, MFM = mass flow meter, MFC = mass flow Controller, PV = pressure Valve.
2.3. Process layout
membranes is calculated via this tool. Nowadays CFD simulations represent a well-established method to obtain detailed flow and concentration profiles in membrane separation processes. The commercial CFD software used in this work, ANSYS Fluent v. 14.5, solves the Navier Stokes Equations for finite volume elements to describe the viscous flow in the module. The diffusive species transport is considered by the Maxwell- Stefan diffusion model. For the simulation, a geometrical model of the fluid volume is built and divided into discrete cells, the mesh. Since highest resolution is required in the region of the membrane tubes, the mesh was built starting from sweeping the membranes with hexahedral layers. The rest of the fluid volume was filled with tetrahedral elements. The boundary conditions are listed in Table 1 and displayed in Fig. 3. The Navier Stokes Equations are solved for the feed compartment, while the membrane is treated as a boundary. The mass transport of oxygen across the membrane is implemented by a mathematical model for which a simplified approach of a method presented by Abdel-jawad [18] is used. A thin layer along the membranes' outer surface is defined via a user defined function as sink for mass, momentum and energy terms. The mass sink for each cell in this layer is calculated by a modified Wagner Eq. (2), where amb (T ) represents the ambipolar conductivity and follows an Arrhenius behaviour. Transport parameters used for the simulation were determined in laboratory experiments for single membrane tubes [19].
Fig. 2 shows the process scheme of the membrane module. The feed air flow (0–75 m3 h−1) is controlled by mass flow controllers (Bronkhorst). Unless specified differently, all flows and flow rates refer to Standard Temperature and Pressure (STP) conditions of 0 °C and 1 atm. Entering the module, the air is heated with LEISTER heating elements D8 up to 900 °C. The hot air passes the membrane tubes via an air distribution plate and the retentate is cooled at the exit of the module. Pressure (END-Armaturen), mass flow of the retentate (Bronkhorst) and the oxygen concentration (sensortechnics XYA) are measured and controlled in a cold state. On the permeate side of the membranes the oxygen is cooled and delivered to a vacuum pump (Pfeiffer Hena 200). The permeate pressure can be controlled from < 50 mbar to ambient pressure. The same components for measuring and controlling oxygen concentration, pressure and mass flow like in the retentate section are used in the permeate section, adapted to the measurement range. 20 thermocouples are used to measure and monitor the temperature at various points in the system, e.g. flange temperature, wall temperatures and the temperature in the inner part of the vessel. Also two thermocouples are inserted into selected membranes for measuring the inner membrane temperature. To guarantee a safe operation, wall temperatures of the vessel as well as the pressure in the vessels are constantly measured and monitored. An exceedance of critical values leads to an opening of a gas discharge safety valve which the piping is equipped with. Additionally, mechanical relief pressure valves are implemented redundant to the electronic monitoring. Furthermore, the flow of cooling water is monitored and connected to the shutdown measures. The vessels as well as the water cooled flange are pressure proved and certificated by the German technical control board TÜV. The expected surface temperatures and the resulting wall thicknesses of the pressure exposed parts were calculated by CFD and FEM simulations.
mO2,cell = A cell·
pO2,f RT amb (T) ·ln 16F 2 pO2,p
(2)
Both the temperature and the feed partial pressure can directly be accessed from each cell via the simulation. The adjacent partial pressure is calculated under the assumption of Dalton's law for an ideal gas mixture Table 1 Boundary conditions for the CFD simulation.
2.4. Simulation The efficiency of the oxygen transport via an OTM membrane is, according to Eq. (1), mainly determined by the process parameters temperature and oxygen partial pressure gradient across the membrane. A membrane module should therefore provide a sufficient and homogeneous temperature along the membranes and guarantee a sufficient feed oxygen transport to the membranes in terms of an optimized flow field distribution to avoid concentration polarization effects. Especially the last point is not or only poorly accessible via measurements in large scale membrane modules. In order to investigate the experimental results more deeply, selected experiments were analysed by using Computational Fluid Dynamics (CFD) simulations. Besides the fluid dynamic behaviour of the gas flow, also the oxygen transport across the 254
Parameter
Value
feed air inlet flow inlet temperature feed air pressure permeate pressure heating temperature upper flange temperature number of membranes effective membrane length effective active membrane area total built in membrane area heat conductivity flange (W m−1K −1) heat conductivity insulation (W m−1K −1) heat conductivity BSCF (W m−1K −1)
36 m³ h−1 850 °C 10 bar 50 mbar 900 °C 25 °C 298 370 mm 5.37 m2 7.25 m2 51 (20 °C) - 50.8 (100 °C) 0.19 (800 °C) - 0.27 (1000 °C) 2.26 (800 °C) - 3.1 (960 °C)
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Fig. 3. CAD Drawing of the geometry and mesh at the membrane tube.
pO 2, f = xO 2 ·ptot
(3)
The molar fraction x O2 and the local, static pressure ptot are easily accessible from the simulation. Since the module is operated in 3-end mode with a vacuum on the permeate side, the permeate oxygen partial pressure is assumed to be equal to the overall pressure on the inner side of the tubular membranes. Furthermore, pressure drop for the inner side of the membranes was neglected in awareness of the small volume flows and the relatively large diameter of the channels > 13 mm. According to the calculated mass flow of oxygen that is subtracted in the sink, a momentum and the enthalpy correction is performed in the simulation. 3. Results and discussion 3.1. Temperature distribution Fig. 4. Experimental temperature profile inside a hanging membrane, length 0 is the closed membrane end.
One essential point for the design of a membrane module is a homogeneous flow and temperature distribution, which turned out to be difficult to maintain constant within the vessel. While the vessel temperature was controlled to 850 °C, the membranes only reached a maximum temperature of about 830 °C at its tip and only 750 °C close to the insulation of the water cooled flange. Fig. 4 shows the axial temperature profile measured with a movable thermocouple which was placed inside a selected membrane during stationary operation. For a vessel temperature set point of 850 °C, only 30 cm of the membrane length are exposed to a temperature above 800 °C. The temperature drops rapidly within the flange insulation so that a surface temperature at the water cooled flange of about 150 °C is reached. At the joint between metal housing and ceramic membrane tube (50 cm) the temperature is uncritical for the used epoxy glue. With 37 cm of the 50 cm membrane protruding out of the insulation, only 74% of the membrane area take part in the air separation process, while the remaining 13 cm are required for cooling and sealing. That might seem like a waste of membrane material, but represents the only way to achieve a cold sealing between metal housing and ceramic membrane. Fig. 4 also points out the difficulty of determining the actual operation temperature. Unless otherwise noted, the temperature referred in the following diagrams is the vessel temperature, while the actual maximum membrane temperature is about 20 °C lower. A similar temperature distribution was obtained by CFD simulation, Fig. 5. The zones adjacent to the insulation of the water-cooled membrane fitting flange show a significant temperature difference to the zones close to the vessel wall. This may be interpreted as an indication that the heat loss over the flange cannot be compensated by the inner heating so that the set point of 850 °C cannot be kept constant in the entire feed regime. One reason for this is the spatial separation between
the inner flow chamber and the heating surfaces via the air distribution plate, whereby the convective heat transport from the heating surfaces is rather low. Higher temperatures were achieved in a later run by modifying the air distribution plate (compare long-term test C). 3.2. Permeation Measurements Various tests were performed with up to 96 membranes. In order to make the tests comparable to each other, results are expressed in terms of oxygen flux. All fluxes depicted in this paper were normalized to the active membrane area, meaning the part of the membranes that protrudes from the insulation was considered (370 mm length/membrane). Fig. 6 shows the oxygen flux and the purity of the membrane module with 47 membranes during start up of the module. Due to the low temperature gradient of 150 K/h a quasi-stationary behaviour of the module can be assumed. Additionally, the O2-flux of a single membrane tube part (10 cm), measured under lab-conditions, is depicted in that diagram for comparison. Main differences in that comparison are the well defined counter flow direction in the test rig compared to the cross flow in the actual module layout, but also the temperature distribution plays a main role, as the following diagrams show. With rising temperature, the permeation rate of the membranes increases according to the Wagner equation. At 850 °C and 12 m³ h−1 feed flow, a feed pressure of 5 bar and a vacuum pressure of 100 mbar, a flux of 1.68 ml cm−2 min−1 was obtained. This is significantly below the flux that can be obtained with a single tube in a laboratory 255
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Fig. 5. Calculated temperature distribution in the module.
Fig. 6. Oxygen flux, purity and leakage during start up with 47 membranes, feed air flow: 12 m³ h-1, pfeed :5 bar, ppermeate: 100 mbar. Fig. 7. Simulated velocity profile in the feed compartment.
experiment. Main reason for this deviation is the inhomogeneous temperature distribution in the module, but also maldistribution of flow accounts for a lower permeation rate, as CFD simulations show in terms of the velocity profile in Fig. 7. Highest fluid velocities appear at the inlet and outlet of the module as well as in the bottom part of the module below the hanging membranes. The high velocities indicate a short-cut flow below the membranes, whereby a big part of the feed air leaves the module without taking part in the membrane separation process. This is also indicated by the O2-mass fraction distribution in the module, Fig. 8. While the feed stream in the bottom part of the vessel contains an average oxygen mass fraction of about wO2 = 0.18, there is a depletion in the membranes zone with wO2 < 0.12. Thus the driving force for the oxygen transport is smaller than assumed when just taking the mean of the measured feed and retentate partial pressures as mean feedside partial pressure. An additional effect which decreases the driving force is the back swirl which can be identified in Fig. 8, which leads to a back mixing of already depleted air into the feed compartment. The CFD calculations reveal that the average O2-partial pressure of the cells adjacent to the membranes is about 14% below the average partial pressure of all cells in the module. During all start up procedures an increase of oxygen purity was
Fig. 8. Simulated oxygen mass fraction in the feed compartment. 256
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measured as shown in Fig. 6. The oxygen purity in that example increases from 79% at 640 °C to 98.3% at 850 °C. This results mainly from the fact that the amount of permeating oxygen increases while the amount of leaking nitrogen into the system remains constant or even drops with increasing temperature. At 640 °C, the leakage rate is around 0.1 ml cm−2min−1 and slightly decreases to 0.04 ml cm−2min−1 at 850 °C. It cannot be distinguished if the leakage is caused by material defects in the ceramic or from the system like leaky joints. The small decrease of the leakage with elevating temperatures could be a hint of resintering and closing of microcracks in the membrane, which might be assisted by the high feed pressures, but also the increasing gas viscosity with temperature, leading to a lower viscous flow inside small pores could be a more simple explanation. Nevertheless, the effect of the increasing oxygen purity was also seen during long time operation at constant temperature (compare Fig. 11). Fig. 9 shows the oxygen flux for a test with 96 membranes during a pressure increase after start up. Feed pressure was varied from 2.0 to 3.5 bar and permeate pressure from 50 to 100 mbar. With increasing feed pressure the temperature in the module drops due to the lower flow velocity at higher pressure, which reduces the convective heat transport from the heating elements to the inner fluid chamber. Despite of the temperature drop, there is a continuous increase of the permeation rate with increasing feed pressure. Maximum total permeate oxygen flow rate achieved in this experiment was 22 l min−1, resulting in a flux of 1.27 ml cm−2 min−1 at an oxygen purity of 98.9%. The theoretical O2-flux of the membrane tube, assuming a uniform membrane temperature under these pressure gradients, is by 25% higher than the experimental one for 100 mbar permeate pressure. This is mainly due to the inhomogeneous temperature distribution in the module. Lowering the permeate pressure to 50 mbar increases the difference up to 35% from the theoretical values. This indicates a significant feed side mass transport limitation caused by local short-circuit flow. The packing density with 96 membranes might intensify this effect, as it can be seen compared to an experiment equipped with 38 membranes depicted in Fig. 10. Some modifications in the air distribution were made to enable a better heat exchange between the heating elements and the feed air, so that a vessel temperature of 880 °C could be reached. Fig. 10 shows the oxygen flux of the experiment with 38 membranes (Aact = 0,7 m2) during a pressure increase from 5 to 10 bar. By elevating the
Fig. 10. Oxygen flux and temperature for a run with 38 membranes (Aact=0,7 m2), pperm = 50 mbar, feed air flow: = 12 m³h-1, modified air distribution.
temperature, a higher permeation rate could be achieved compared to the previously described experiments. The highest measured oxygen flux under these conditions was 2.8 ml cm−2 min−1, which is a total production of 19.6 l min−1. The deviation between the theoretical and the measured permeation rate is 23% at 5 bar feed air pressure. Increasing the feed pressure to more than 8 bar does not lead to a significant increase in permeation rate which is a further indication of a limitation of the feed side oxygen transport to the membrane surface, although the oxygen recovery rate is only between 0.3 and 0.4 under these conditions. 3.3. Long term tests In order to prove the feasibility of the module operation, long term tests were performed. Within the time of operation three tests, following named as (A), (B) and (C), were performed aiming on an operation time of 1000 h. Several changes were made in the module configuration between each long term test, which are described in the following:
• Test (A) was operated with 96 membranes at varying feed pressures •
•
between 1 and 5 bar abs. The vessel temperature was adjusted to 850 °C and a permeate pressure of 50 mbar was applied. Test (A) was stopped after 1000 h. Test (B) was operated with 47 membranes, a feed pressure of 5 bar and a permeate pressure of 100 mbar. Furthermore changes in the pressure control system were made and a new insulation material for the flange insulation was mounted. Test (B) only reached an operation time of 580 h due to an emergency shutdown resulting from a failure of the cooling water system. Test (C) was operated with 38 membranes at 5 bar and a permeate pressure of 50 mbar. Furthermore constructive changes in the air distribution plate were made that allowed a higher vessel temperature of 880 °C (Tmem=860 °C). The test reached an operation time of 1800 h, whereby after 1000 h an increase of the feed pressure from 5 to 10 bar was done. Test (C) stopped again with an emergency shutdown due to a failure of the mass flow controller.
Fig. 11 and Fig. 12 show the oxygen flux and purity for the different tests. Due to many stops and changing conditions during the test, the flux of Test (A) is not displayed continuously. For test (A) a maximum O2-flux of 1.27 ml cm−2 min−1 was measured. The oxygen purity at the beginning reached 95.4%, but varied strongly over the tested period
Fig. 9. Oxygen flux and temperature for a run with 96 membranes (Aact=1,73 m2) at varying feed pressure, pperm = 50 mbar, feed air flow: = 12 m³h-1. 257
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3.4. Membrane breakages and failures During the first test phases, more than 300 membranes have been set in. The reasons for membrane breakages were diverse, but certain trends were observed. The majority of the membranes broke during or after emergency shutdowns, which were mostly caused by external failures like malfunctions of temperature sensors, preheaters or cooling water. In that case, all pressure valves open and the heating shuts down. This results in mechanical and thermal stress for the membranes. The pressure instantaneously releases from operation pressure (10 bar) to ambient pressure within 2 min. Simultaneously the membrane temperature drops due to the air expansion and the shutdown of the heating and the preheaters by approximately 70 °C within that time, which means a high stress to the ceramics. During the first tests, also the pressure relief valves on the permeate side opened in case of an emergency shutdown, and in presence of a vacuum inside the membrane tubes cold air is sucked into the membrane tubes, which might have led to membrane breakage due to a thermo-shock. This problem has recently been solved by applying a back-pressure valve at the permeate duct. Besides the breakages that occurred during emergency shutdowns, most frequently membrane breakages were observed during the cool down phase between temperatures of 250 °C and 310 °C. This is reported to be the temperature region where BSCF shows a minimum in the characteristic fracture stress curve [20]. Even with a controlled shutdown of the module it was almost impossible to prevent membrane breakages, thus the cooling down phase represents a critical step in the module operation, although it is not yet understood what exactly causes the membrane to break. The vast majority of fractures occurred at the position where the membrane emerges the insulation and enters the hot zone, see Fig. 13. At this position, the membrane is exposed to a large local temperature gradient, which obviously leads to a significant degradation: while "new" membrane tubes were tested to stand an internal pressure of 1 bar (g), aged membranes after module operation showed a significant reduced stability to the inside pressure test and already broke at internal pressure tests at 200 mbar(g) at that position. Additionally, when breaking the vacuum on the permeate side, like in case of emergency shutdowns or membrane breakages, the incoming air is sucked via the watercooled flange into the membrane tubes and therefore the indicated position becomes the first spot to expire the thermo-shock as described before. Besides the described damages, secondary damage caused by falling membranes can be observed. In the worst case this can lead to a domino effect and damage the majority of the built in membranes. It is almost impossible to identify the origin of the membrane breakage in this case. In case of leakages or single membrane breakages the module could be shut down and the defect membrane could be replaced. Therefore the controlled shut down of the heating while maintaining the
Fig. 11. Oxygen purity during the 3 long term tests under varying conditions.
Fig. 12. O2 flux during long term tests (B) and (C).
due to membrane breakages and replacements. In total, the experiment had to be stopped 6 times during the 1000 accumulated hours due to breakages or major leaks of in total 20 membranes. The varying oxygen purity results from the replacement of single membranes during the stops. After the previously described constructive changes, test (B) showed a significantly improved stability. An oxygen purity after start-up of 97% was reached, slightly increasing to 98.5% after 580 h. A maximum flux of 1.69 ml cm−2 min−1 was obtained under the given conditions, which slightly decreased (-5%) during time to a value of 1.6 ml cm−2 min−1 after 580 h. Test (C) showed the highest flux, since through the module modifications a higher temperature could be applied. Small oscillations in the flux result from oscillations in the pressure control of the vacuum or the feed side. Highest measured oxygen flux within the first 1000 h was determined to 2.38 ml cm−2 min−1 after 560 h, while after 1000 h a flux of 2.3 ml cm−2 min−1 was obtained. In contrast to the improved permeation rate, low oxygen purity was obtained for this test. The purity at start-up only reached a value of 70%, caused by a high air leakage rate of ~ 1 ml cm−2 min−1. During the 1000 h long term test, a decrease of the air leakage rate to 0.21 ml cm−2 min−1 was observed leading to an oxygen purity of 93.3%. Since no changes in operation mode or module configuration were performed during the first 1000 h, the increase of the oxygen purity is an indication for resintering of microcracks, which was also investigated via grain size analysis in Section 3.4. After 1000 h, a change in feed pressure was made from 5 to 10 bars, resulting in a higher flux of 2.8 ml cm−2 min−1 but a decrease in oxygen purity, which partially recovered after longer operation time.
Fig. 13. Typical breakage after shutdown. 258
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Fig. 14. SEM (backscattered electron image) micrographs of the feed side of the membrane before and after 1800 h of testing,.
adjacent to the water cooled flange, so a formation of a hexagonal phase cannot be precluded, which could also be an explanation of the decreasing permeation rate for the long term tests. As it can clearly be seen in Fig. 15, this change in composition along the grain boundaries also occurs on the permeate side of the membrane. The superficial degradation of the permeate side appears to be lower than the one on the feed side. After 1800 h of testing, the grain boundaries on the surface are still visible with SEM. Also particles seem to have formed on the surface and especially along the grain boundaries, which again were identified as slightly changed BSCF compositions characterized by iron depletion. As expected in the 3-end mode, almost no other components like Cr or S occurred on the permeate side although some spots containing chloride were found during SEM. The source for that contamination is probably attributed to a lubricant paste used for a screw connection in the permeate chamber. It is remarkable, that no carbonate formations have been detected on the membranes, indicating that the CO2 content of the ambient air does not seem to influence the permeation behaviour in a significant way over the tested time range. Also, a growth of the grain size was observed via image analysis, as depicted in Fig. 16. The shift of the average grain size d50 from 18 to 23 µm shows a sintering activity even at the operation temperatures of 850 °C, what can be taken as an additional indication of the resintering of microcracks and the correlated increase of the oxygen purity.
operation pressure showed the best success since tensile stresses of the hanging membranes were avoided that way. An identification of the defect membrane tubes in the module could only be done by visual inspection or by localizing obvious flows through the single tubular membranes at a small pressure difference. 3.5. Aging and degradation Fig. 14 shows SEM images of the membranes from the long term test (C) for the membrane surface on the feed side. At the initial condition the grain boundaries are clearly visible. After 1800 h of testing, a clear surface degradation is obvious. The formerly smooth surface is covered with fine particles in the hot region of the membrane. Via EDX analysis of these particles, a significant amount of chromium was detected, resulting in with up to 15 at%, which has evaporated from the high temperature steels. Also spots containing about 2 at% sulfur were determined via EDX. Chromium and sulfur formations are known to hinder oxygen permeation and could therefore be an explanation for the slightly decreasing permeation rate in test (B). The source of the sulfur has not been clarified yet, but it is supposed to originate from oil residues from the compressor unit. The described contamination of the membrane could not be observed in cross-sections of the aged membranes, being a hint for a clear surface effect in the submicron range. However, this still has to be kept in mind especially when thinking about supported thin layer membranes with a lifetime demand of several thousand hours. A clear change after 1800 h compared to the initial membrane can be seen in the cross section image, where the grain boundaries are clearly visible by its bright colour. Along the grain boundaries a change in the material composition was determined via EDX, that can mainly be characterized by a significant depletion of iron and a minor depletion of strontium. This decomposition occurring predominantly at the grain boundaries has previously been reported by [11] and was there contributed to a formation of a hexagonal phase, but it is worth to mention that for our investigated membranes no clear evidence of a hexagonal phase was detected via XRD measurements and also the average temperature of membranes was with ~860 °C above that limit which is usually reported for hexagonal phase formation [11,21,22]. But as shown before, there exist zones of lower membrane temperature
4. Conclusions The design of modules is an essential point on the way to enable the OTM technology for high purity oxygen production. Since the oxygen transport membranes operate at elevated temperatures and high pressures, the large scale module construction holds some difficulties. A membrane module capable of 596 tubular membranes was constructed and tested with various capacities, at a maximum of 96 tubes. Under operating pressure of 10 bars and 880 °C, a flux of 2.8 ml cm−2 min−1 was obtained. The highest oxygen purity was determined to 98.9%. Despite of a quality control of every single membrane, a leakage could not be avoided completely and an identification of leaky membranes turned out to be very difficult. With oxygen purity below 99%, this technology loses one of its biggest advantages compared to the 259
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Fig. 15. SEM (backscattered) micrographs of the permeate side of the membrane as sintered and after testing.
testing period of 1800 h could not be exceeded. The testing phase showed a general feasibility of that concept, but there is a huge potential for improvement. Besides an improved membrane manufacturing and quality control, a hot membrane-metal joining/sealing and therefore an avoidance of the watercooled flange is mandatory to overcome the problems with the inhomogeneous temperature distribution in the module. This would also eliminate the temperature gradient along the membranes and therefore the spot where the vast majority of membrane breakages occurred. Acknowledgements This work was supported by BMWi and MIWFT as well as by the companies RWE Power AG, E.On AG, Hitachi Power Europe, Linde AG, MAN Turbo, WS-Wärmeprozesstechnik and Vattenfall [grant number 0326890]. References
Fig. 16. Grain size distribution of the membrane before and after 1800 h operation. Examination of approx. 4000 grains in thermally etched specimens by image analysis.
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cryogenic ASU. An implementation of that the OTM technology therefore requires a higher purity which could be reached by an improved quality control that allows leak tests of single membranes also in the built in state. The permeation rate obtained was below that one measured on single membrane tubes, which was mostly due to an inhomogeneous temperature distribution that results from the water cooled parts of the module. Also the investigations showed some potential of improvement for the flow duct in the module, which also contributed to the lower permeation rate. The biggest problem during the experiments in the pilot scale were the breakages of membrane tubes, where one single breakage caused a shutdown of the module for several days and in the worst case a breakage of most built-in membranes. The fact that most breakages occurred due to externally caused shutdowns emphasizes the need for an optimized process control as well as for safety measures to avoid large membrane losses. During the test phase, significant improvement regarding the reliability of the module was reached, nevertheless a 260
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