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Status report on the operation of the bench-scale plant for hydrogen production by the mark-13 process
0360-3199/82/080629-08 $03.00/0 Pergamon Press Ltd. 1982 International Association for Hydrogen Energy.
Int. 1. Hydrogen Energy, Vol. 7, No. 8, pp. 629-636, 1982. Printed in Great Britain.
STATUS REPORT ON THE OPERATION OF THE BENCH-SCALE PLANT FOR H Y D R O G E N P R O D U C T I O N BY THE MARK-13 PROCESS D. VANVELZEN and H.
LANGENKAMP
Chemistry Division, Commission of the European Communities, Joint Research Centre, Ispra Establishment, 1-21020 Ispra, Italy
(Receivedfor publication 2 October 1981) Abstrad--The bench-scale plant for thermochemical hydrogen production by the Mark-13 process has been successfully operated since May 1978. The nominal hydrogen production rate of the plant working at atmospheric pressure is 100 STP 1. h -1. A concise description of the process flow scheme is given. Operation of the plant demonstrated that thermochemical water decomposition is a viable process and a valid alternative to other processes for hydrogen production from water. The availability of the plant for experimental work was very high and no special start-up procedure was required. All reactions and product separation units were working within their specifications and no by-product formation was observed. INTRODUCTION
- - to ascertain whether the specified reactant conversions and concentrations could be reached and maintained; - - to check the possible formation of by-products; - - to investigate process control problems connected with the continuous operation of a closed cycle process; -to obtain data for the preparation of specifications for possible future larger size plants. It is felt that these goals were fully reached during the successful operation of the plant for a year and a half. This statement will be discussed in more detail in the following sections. Additionally, another objective emerged unexpectedly, i.e. the study of the behaviour of various cell designs for the electrolysis of HBr on the scale required in the plant. Indeed, this point turned out to be one of the main items of our work and, consequently, covers one of the main subjects of this paper.
The application of thermochemical cycles for hydrogen production from water should circumvent the conversion of large amounts of thermal energy into electricity, followed by electrolysis as in the direct electrolysis of water. Research on thermochemical cycles started in about 1970 in several places of the world. In 1974, a variant of the purely thermochemical cycles was introduced : the hybrid cycle. Here, electrical energy is used in one reaction of the process as a means to overcome the positive free energy barrier inevitably occurring in any purely thermochemical cycle. Obviously, the electrolysis step in hybrid cycles is performed at very low voltages in comparison with direct water electrolysis. Examples of potentially promising hybrid cycles are the Westinghouse sulfur cycle (also known as the Ispra Mark-11 cycle) and the Ispra Mark-13 sulfur-bromine cycle. The latter cycle is the subject of this paper. It was invented by Schiitz and Fiebelmann [1] and consists of the following three reactions: 2
HBr --->H2 + Br2 (the electrochemical step)
(1)
SO2 + Br2 + 2H20 --->2 HBr + H2SO4
(2)
H2SO4---> H20 + SO2 + 1/2 02
(3)
The development of this cycle is well past the laboratory stage and a bench-scale continuous model plant was designed and constructed. The plant has been in full operation since May 1978 [2]. At the time of its inauguration, it represented the first demonstration of a complete hydrogen production process by the thermochemical route. The nominal hydrogen production rate of the plant is 100 1. h -1. The plant is constructed mainly of commercial glass and quartz equipment connected with PTFE tubing and is operated at atmospheric pressure. The original objectives of the operation of the circuit were:
PROCESS DESCRIPTION A detailed description of the process has already been published on various occasions [2, 3]. The process description is therefore confined to a schematic drawing of the model (Fig. 1) and a concise lay-out of the reactor arrangement. The plant consists of three main sections: (A) HBr-electrolysis comprising the bromine distillation/HBr absorption column; (B) The acid formation section, consisting of two reactors for the acid formation reaction, where gaseous HBr and sulfuric acid of about 75 wt% are formed, coupled with a sulfuric acid concentration column for the production of 95 wt% sulfuric acid; (C) The sulfuric acid decomposition reactor and a separation unit for SO2 and O2. (A) HBr-electrolysis
629
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Fig. 1. Mark-13 flow diagram. the electrolysis cell. The liquid leaving the cell contains about 40 wt% HBr and 4 - 6 wt% Br2. The formed bromine is completely dissolved in the electrolyte. The exit liquid is pumped into a distillation column. In this column, bromine is distilled out of the top and is fed to the acid formation section to be reduced to hydrogen bromide. This product from the acid formation section enters the distillation column as a vapour. Two processes take place simultaneously in this column, i.e. the stripping of bromine as well as the reconcentration of the electrolyte with HBr from 40 to 45 wt%. The hydrogen formed in the electrolytic cell is measured by a wet gas meter and leaves the plant as the main product. (B) The acid formation section The reaction between bromine, sulfur dioxide and water to gaseous hydrogen bromide and liquid sulfuric acid is carried out in two packed columns in series. In one column, Br2 and SOz are contacted with a sulfuric acid solution of about 75 wt%. Bromine is slightly in stoichiometric excess to SO2. The solution is steadily recycled over the column to warrant a satisfactory gas/liquid contact. During the passage of the gas through this column, partial conversion of Br2 and SO2 is obtained. The formed H2SO4 dissolves in the liquid and leaves the column by an overflow system. This solution is fed to the H2SO4 concentration column, where it is further concentrated to 95 wt% H2SO4. The gas stream leaving this first reactor consists of
HBr, Br2 and SO2. This mixture is fed to a second reactor where a solution of about 55 wt% H2SO4 is circulating. Here, the reaction proceeds to completion and gaseous HBr with a small percentage of bromine, but free from SO2, leaves the reactor and is fed to the bromine distillation column (section A). The formed sulphuric acid dissolves in the circulating solution and is fed by overflow to the reactor containing the 75 wt % H2SO4 solution. The water required for the reaction is introduced into the second reactor. (C) The sulfuric acid decomposition section The concentrated sulfuric acid produced in the acid formation section is pumped to the sulfuric acid evaporator, where it is heated to about 900 K. The vapour is then further heated to 1000--1100 K and passed over a catalyst. A residence time of about 0.3 s is sufficient to decompose 70% of the inlet sulfuric acid into SO2 and 02. The water formed in this reaction and the unreacted H2SO4 are cooled and condensed and passed to the sulfuric acid concentration column (section B). The separation of SO2 and 02 is performed primarily by cooling down the gas mixture to 225 K. About 95% of the sulfur dioxide is condensed in this step. This liquid SO2 is subsequently evaporated and recycled to the acid formation section. The gas mixture, leaving the cooling trap, consists of approximately 91 vol % 02 and 9 vol % SO2. The remaining SO2 is removed by contacting the gas with
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS bromine and water, i.e. reaction (2). The purified oxygen leaves the process as a secondary product. A photograph of the plant is given in Fig. 2. O P E R A T I O N OF THE PLANT As mentioned above, the plant was put into operation in May 1978 and has always run very smoothly. During the first months, a number of minor modifications had to be carried out, mainly in the bromine distillation column and the H2SO4 decomposition section. The only other major difficulty encountered was the failure of the direct current generator for the HBr-electrolysis in the summer of 1979. This caused the plant to be shut down for 6 weeks. However, the fact that the plant has been unavailable for experimental work due to the failure of an ancillary piece of equipment is much more the exception than the rule. Under normal conditions, the availability of the plant was high. The circuit can be started up in a short time without any special problems. Apart from a preparatory period of about 2 h when the two distillation columns and the sulfuric acid decomposition reactor are pre-heated, and
631
the refrigerating unit for the SO2-condensation is cooled, no special start-up procedure is required. Once the required temperature levels are reached, the plant can be put on stream immediately at its nominal hydrogen and oxygen production flow-rate. It is then very easy to maintain the circuit in near steady-state operation, due to the fair reliability of the metering pumps for flow-rate control. There are neither major instabilities nor formation of by-products or side-reactions. In conclusion, it is technically feasible to put the plant into operation at any desired time and to maintain it in continuous operation for very long periods. There are no technical reasons requiring any interruption of operation. RESULTS (A)
HBr-electrolysis
As mentioned above, the study of the performance of various cells for the electrolysis of HBr turned out to be one of the main subjects of research during the operation of the bench-scale plant. Three different cells
Fig. 2. Laboratory scale demonstration plant for the Mark-13 process.
632
D. VAN VELZEN AND H. LANGENKAMP
have been tested. Two of the three also underwent some minor modifications, mainly to improve their hydrodynamic behaviour. All three cells were so-called horizontal controlled convective diffusion cells [4]. This cell type works without a diaphragm, its principle being that the flow-rate of the electrolyte through the cell is kept well within the laminar regime, avoiding as much as possible convective currents from the anode to the cathode. In this way, bromine, generated at the anode, is kept away from the cathode region, and high (near to 100%) current efficiencies can be obtained. Schlitz et al. [4, 5] reported promising results in small scale experiments, so that testing of the controlled convective diffusion cells on a somewhat larger scale in the bench-scale plant was a logical follow-up. The tested cells are monopolar and graphite is the electrode material for both the anode and the cathode. In all experiments we used platinum as a homogeneous electrocatalyst, the application of this noble metal considerably reducing the overvoltage at the cathode. The amount of Pt applied corresponds generally to the formation of a Pt-layer of 3-4 mg cm -2 on the electrode surface. The first cell is a rectangular cell, whereas the second cell is circular. In these two cells, the evolved hydrogen flows through the cell co-currently with the electrolyte flow. The third experimental cell is also a circular one with slightly inclined electrodes. Here, the hydrogen flow is counter-current to the electrolyte flow. Rectangular cell. The rectangular cell was the first cell used in the plant. A cell lay-out is given in Fig. 3. The electrode surface is 352 cm 2 and the electrode distance can be varied between 3 and 5 mm. A typical cell voltage curve is given in Fig. 4. Cell voltages are 0.85 V at 200 m A c m -2 and 1.03 V at 500 mA cm -2 at 350 K. The electrolyte is 45 wt% HBr. In the early experiments carried out with this cell, the current efficiency could not be raised above 91%. Moreover, this value did not remain constant with time, but decreased slowly during the course of each experiment [2]. For example, after 30 h of continuous operation, the current efficiency dropped from 91% to values around 75%. The run was interrupted at this point and 47%
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300
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the cell was flushed with bromine-free electrolyte for several hours. In a subsequent experiment, the original value of 91% was again found at the start of the run. It was anticipated that an insufficiently regular liquid flow pattern through the cell was the cause of this phenomenon. As a remedy, a modification of the cell was therefore made, consisting of the replacement of the liquid feed and outlet system. Originally there was only one inlet and one outlet tube, whereas the new arrangement comprises three parallel inlets and a liquid overflow with three separate outlets. The inlets and outlets are at a mutual distance of 60 mm. The modified cell performs very well. The observed cell voltages are slightly higher than those for the original cell, i.e. 0.90V at 200mA cm -2 and 1.10V at 500 mA cm -2 (previously 0.85 and 1.03 V, respectively). The current efficienc),, however, improved considerablyi at 200 mA c m - ' i t is now 99%, and at 500 mA cm- 94%. It follows that the loss in cell voltage is amply compensated by the gain in current efficiency and the electric energy consumption is practically constant. The most important improvement, however, is that the modified cell can now be operated for long periods without any concomitant decrease of current efficiency. The cell operated in a 30 hr experiment at 400 mA cm -2 at 1.03 V, at an average current efficiency of 95%. The electrolyte flow rate was maintained in this experiment at 8 kg h -1, and the temperature was 350 K. The favourable effect of the above modification is illustrated in Fig. 5, where the change of the current efficiency with time is shown before and after the three tube inlet system was mounted. Limiting current density. It must be noted that the current density of a given horizontal cell (and thus the hydrogen production rate) cannot be increased indefinitely. In general, running a cell above a limiting condition gives rise to the following course of events. Initially the cell operates well, cell voltage and current efficiency are in the anticipated range. However, after
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS
633
Current
Efficiency (%)
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a certain running time (usually about 3 h), a rapid decrease of the current efficiency occurs. In a few minutes it becomes virtually impossible to operate the cell, The hydrogen production rate drops dramatically and bromine vapours appear in the hydrogen outlet. At this point, bromine is evaporating from the electrolyte, severe mixing of the liquid phase takes place and, obviously, the flow regime is far from laminar. At this stage, the current must be switched off and the electrolyte flow rate maintained. After approximately I h the cell can again be put into operation. This cell failure has two main origins: (a) an excessive H2 production rate which causes mixing in the liquid phase due to the presence of gas bubbles; and (b) high bromine concentrations which favour the local formation of bromine vapour bubbles and, consequently, mixing of the liquid phase. Both (a) and (b) increase with increasing hydrogen production rate. A badly developed liquid flow pattern through the cell leads to the formation of preferential flow regions and more or less stagnant zones. This causes very high local bromine concentrations in the "stagnant" zones, and a largely increased probability of (b). It is therefore of paramount importance that the liquid flow pattern through a horizontal cell be well controlled. From the foregoing it follows that the limiting current density of a cell depends on the cell design, dimensions, electrode distance and configuration and can only be found empirically for each cell. The limiting current density for the rectangular cell is 500 mA c m -2. At 450 mA crn -2 the cell could be operated for long periods (20 h) under stable conditions, but operation at 540 and 600 mA cm -2 led to cell failure after 1-3 h of operation. Values of 80-85% of the limiting current density are recommended as safe values for cell operation in long duration experiments. This implies that the rectangular cell can be safely operated at current densities of
401)--440 m A c m -2. Under these conditions the cell produces 60-65 I. h -1 of hydrogen. Circular co-current cell. A schematic diagram of the cell is shown in Fig. 6. It essentially consists of two parallel circular graphite electrodes, fixed in a Teflon housing. The HBr solution is fed at the centre and flows radially towards the periphery of the electrodes, where it is collected and leaves the cell. Hydrogen is collected in the annular ring and leaves the cell by an exit in the upper part of this ring. The total electrode surface is 283 cm 2 and the electrode distance of the cell is 5 mm. The initial results obtained with this cell were rather disappointing. The typical cell voltage was very near to the one obtained for the modified rectangular cell, as shown in Fig. 4, but the current efficiency was very low: 85% maximum. An estimation of the pressure head due to the friction of the electrolyte flow through the cell indicated that this pressure drop was extremely low (about 10 -3 cm liquid head). In this case there was a strong preferential ~2
HBr
HBr ~-Brz ELectrode material grophi'l'e Electrode disl"once 5ram Electrode surfoce 283 cm2 Fig. 6. Circular co-current electrolytic cell.
634
D. VAN VELZEN AND H. LANGENKAMP
flow from the inlet to the only outlet tube and very bad flow distribution resulted, with detrimental effects on the current efficiency. The installation of an overflow system with four outlets at 90°, with a common flow resistance downstream, was introduced to improve the flow pattern. Indeed, this modification considerably improved the cell performance. Current efficiencies of nearly 100% could now be reached, with cell voltages of 0.90 V at 250 mA cm -2 and 0.99 V at 450 mA cm -2, at 46 wt% HBr at 350 K. These results were obtained in short-time tests whilst the testing of the cell during longer periods combined with regular operation of the total plant was the next stage of the programme. In the first test of this series, the cell was run for a period of 2 h at a relatively low current density: 245 mA cm -2. During this period, the current efficiency was excellent: 99%. The aforesaid conditions correspond to a hydrogen production ~'ate of approximately 30 STP 1. h -1, which is too low for regular operation of our bench-scale unit. We were therefore obliged to run the cell at much higher current densities, i.e. about 450 mA cm -2. In these experiments initially high values for the current efficiency were obtained (95%), but after a running time of 2.5-3.0h, the current efficiency decreased dramatically and finally it became impossible to operate the cell. After shutting down and rinsing the cell for about 1 h, with the recycle electrolyte solution, the original high initial current efficiencies were found again, but after 2.5 h, the cell failed once more. Obviously, for this cell the limiting current density is lower than 450 mA cm -2 and will be about 300 mA cm -2. This limits the hydrogen production capacity to approximately 30 1. h -1. Circular counter-current cell. This last cell type has to be considered as a drastic modification of the circular co-current cell. Its dimensions are nearly identical and it is mounted in the same PTFE housing as the old, co-current cell. A schematic diagram of the new cell is given in Fig. 7. H2
I
The electrodes are no longer horizontal, but slightly inclined towards the centre. The inclination angle is 8 °. The main difference, however, is the hydrogen outlet. In this new cell the hydrogen leaves at a single outlet at the centre of the cell, whereas the bromine-loaded electrolyte is still flowing radially towards the annular ring. The flow of hydrogen through the cell is thus counter-current to the flow of the electrolyte. This arrangement gives the new cell a number of advantages: (i) The old, co-current cell is very susceptible to the control of the liquid level in the cell. A low level causes the cathode to run dry, whereas an excessively high level leads to increased liquid mixing and thus to low current efficiencies. The range between high and low critical level is usually not more than 2-3 mm. The configuration of the new cell eliminates this disadvantage completely. (ii) It is relatively simple to build a bipolar cell on the principle of the new cell, which is much more difficult for the old, co-current cell. Among other things, this is due to difficulties in level control for each single cell compartment for the old cell. (iii) In the new cell the hydrogen leaves the cell in contact with the fresh, bromine-free electrolyte, instead of being in contact with the electrolyte containing bromine as in the old cell. The evolved hydrogen is therefore essentially free of bromine. We measured bromine contents of about 40 mg m -3 H2 for cell operation at 350 K. For co-current cells this value is about 1000 mg m -3 H2. The electrode surface of the new cell is 283 cm 2, like the old cell, and the electrode distance is 5 mm. The performance of the cell is similar to that of the cocurrent cell. Typical cell voltages are 0.89 V at 250 mA cm -2 and 1.01 V at 450 mA cm -2 at 350 K and in 47 wt% HBr solution. Under these circumstances the current efficiency is between 93 and 96%. As in the old, co-current cell, high current densities cannot be applied. Above 300 mA cm -2, difficulties in the operation of the cell occur. On the other hand, operation at 220 mA cm -2 is completely stable. Over long periods (> 10 h) the cell voltage could be maintained at 0.89 V, and the current efficiency at 93-94%. Under these conditions the cell produces 26 STP 1. h -1 of hydrogen.
(B) Acid formation
t 4 o * HBr+ 5% Br2 4?'% HBr Electrode mol-erial graphil"e Electrode disl"once 5rnrn Elecl"rode surfece :>83 cmz
Fig. 7. Circular counter-current electrolytic cell.
The required performance of the acid formation section is the simultaneous production of a gas stream of 8.0 gmol HBr h -1 (which must be as pure as possible) and a liquid phase of high H2SO4 concentration. van Velzen and Langenkamp [6] showed by equilibrium measurements that high sulfuric acid concentrations are attainable, only, however, in equilibrium with considerable partial pressures of bromine and sulfur dioxide in the gaseous phase. On the other hand, the partial pressures of bromine and/or sulfur dioxide are
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS negligible when the liquid produced has a sulfuric acid concentration below 55%. These observations imply that the reaction must be carried out in some type of counter-current staged reactor as otherwise considerable amounts of SO2 and Br2 must be separated from HBr and recycled. The problem has been solved by performing the reaction stage-wise in two co-current packed columns in series, analogous to a gas absorption operation. The overall SO2 conversion in the acid formation section has to be higher than 99%. It is found that this target can be easily reached provided that the relative sulfuric acid concentration in the first column is maintained below 55%. Typical liquid compositions and operation conditions are given in Table 1. Under the conditions given in Table 1, the sulfur dioxide conversion in the concentrated column amounted to 45.3%, and the overall SO2 conversion was as high as 99.7%. The residual SO2 concentration in the HBr stream was only 1300 ppm. The operation of this reactor section proved to be very simple and reliable. The working conditions can be varied over a wide range and are easily adaptable to the requirements of the operational parameters for the rest of the plant. Further scale-up of this reactor section is relatively easy, because common laws for gas absorption with chemical reaction can be followed. The equipment can be operated for any length of time without instabilities and there is no indication of by-product formation. It can be concluded that the acid formation section is a versatile and reliable part of the plant. (C) H2SO4
decomposition
The required duty of this reaction section is the production of 2 gmol oxygen h -1 at a reactor temperature around 1073 K. For this purpose a reactor consisting of
Table 1. Typical operation conditions for the acid formation section Into app. From app. Exit app. No. 1 No. l t o a p p . No. 2 No. 2 Liquids Flow rate (g h -1) Composition (wt %) Br2 HBr H2SO4 H20
198.5
445.4
377.6
-5.3 -94.7
3.9 32.2 34.7 29.4
2.8 75.0 22.2
Vapours Flow rate (STP I. h -1) Composition (vol. %) Br2 HBr SO2
140.4
182.7
145.5
50.1 -49.9
22.5 56.5 21.0
3.9 96.1 1300 ppm
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Fig. 8. Temperature profile of the sulfuric acid decomposition furnace. a quartz tube (3.5 cm internal diameter) heated by a split tube furnace was used. A residence time of i s (calculated at STP) in the catalytic zone seemed a suitable target. This calls for a catalyst volume of 85 cm 3. This quantity was placed in two separate sections, of 40 and 45cm 3, respectively, the one 5 cm from the other. This gap serves to reheat the reaction mixture following its passage through the first catalyst bed. The catalyst used is a special catalyst developed by the firm P6chiney-Ugine Kuhlmann, based on ferric oxide on a porous alumina support. The sections of the reactor tube not occupied by the catalyst are packed with quartz chips to improve the flow pattern and the heat transfer. The performance of the sulfuric acid decomposition reactor is very good. The average reactor conditions obtained during a 30 h run are given in Table 2. This conversion has been obtained with the axial temperature profile in the converter displayed in Fig. 8. Here, the two catalyst zones are clearly recognizable. Obviously, it is not possible to assign one fixed value for the reactor temperature in the catalyst zones, due to the very pronounced axial temperature profile in these zones. These temperatures range from 998 K as the minimum to 1040 K as the maximum recorded at the exit of the second catalyst zone. The degree of conversion obtained in this experiment is near to the equilibrium value at 1070K, i.e. higher than the recorded maximum catalyst temperature. This apparent contradiction can be explained as follows: the temperature is recorded in the centre of the reactor tube. Table 2. Typical operation conditions for the HzSO4 decomposition section. Average reactor temperature 1015 K IN: 95 wt% H2SO4 flow rate
462 g h -1 4.50 gmol h -1
OUT: Oxygen SO2 condensed at 220 K SO2 not condensed (8.8 vol. % of SO7]O2
39.1 !. h -4 1.61 gmol h -1 200 g h-l 3.12 gmoi h-I 3.15 1. h -1 0.13 gmol h -1
mixmre) Total SO2 production SO3-conversion
3.25 gmol h -~ 72%
636
D. VAN VELZEN AND H. LANGENKAMP
Besides the axial temperature profile shown in Fig. 8, there also exists a pronounced radial temperature profile with a minimum temperature in the centre and the maximum at the tube wall. The bulk of the gases are moving in the vicinity of the wall, i.e. at the higher temperature. The average value for the liquid hourly space velocity for the present reactor conditions, calculated for a hypothetical gas flow at 298 K and atmospheric pressure, is 4400 h -1. As a general conclusion, it can be stated that the sulfuric acid decomposition section is also performing very well. There are no instabilities in the course of time and there is no formation of any by-products.
CONCLUSIONS The bench-scale plant for hydrogen production by the Mark-13 process has been successfully operated since May 1978. During this period practically all our pre-set objectives could be realized: - - I t was convincingly demonstrated that thermochemical water decomposition is a viable process and a valid alternative for hydrogen production from water. - - T h e specified reactant concentrations and conversions can be reached and stably maintained over long periods. - - I n some cases (e.g. the acids formation reaction), the performance of the plant is even better than the original specifications, so that improvements in the reference design flow sheet could be made.
- - U n t i l now, no by-products have been detected in any reactor section. - - T h e horizontal cells without a diaphragm for HBrelectrolysis are a promising alternative to conventional cells. Attractive voltages and high current efficiencies could be reached in the bench-scale plant. The liquid flow pattern in the cell is of paramount importance for the working of this cell type. In further development work this last point will be given further attention. REFERENCES 1. G. H. Schiitz and P. Fiebelmann, Patent Application No. 71037, filed 3 October 1974, Luxemburg. 2. D. van Velzen and H. Langenkamp, Thermochemical hydrogen production by the Mark-13 process: a status report, Proc. 14th Intersoc. Energy Conversion Eng. Conf., Boston, Vol. 1, pp. 783-789 (1979). 3. D. van Velzen, H. Langenkamp, G. H. Schiitz, D. Lalonde, J. Flamm and P. Fiebelmann, Development and design of a continuous laboratory-scaleplant for hydrogen production by the Mark-13 cycle, Int. J. HydrogenEnergy 5,131 (1980). 4. G. H. SchiRz and D. Lalonde, A new electrolytic cell type for hydrogen production in hybrid cycles, Int. Conf. on Alternative Energy Sources, Miami Beach, December 1977. 5. G. H. Schiitz, P. J. Fiebelmann and D. Lalonde, Electrolysis of hydrobromic acid, in Hydrogen Energy System (Proc. 2nd World Hydrogen Energy Conf., Ziirich, Switzerland) (T. N. Veziro~lu and W. Seifritz, eds.) Vol. 2, pp. 709730. Pergamon Press, Oxford (1978). 6. D. van Velzen and H. Langenkamp, The oxidation of sulfur dioxide by bromine and water, Int. J. Hydrogen Energy 5, 85 (1980).
Int. 1. Hydrogen Energy, Vol. 7, No. 8, pp. 629-636, 1982. Printed in Great Britain.
STATUS REPORT ON THE OPERATION OF THE BENCH-SCALE PLANT FOR H Y D R O G E N P R O D U C T I O N BY THE MARK-13 PROCESS D. VANVELZEN and H.
LANGENKAMP
Chemistry Division, Commission of the European Communities, Joint Research Centre, Ispra Establishment, 1-21020 Ispra, Italy
(Receivedfor publication 2 October 1981) Abstrad--The bench-scale plant for thermochemical hydrogen production by the Mark-13 process has been successfully operated since May 1978. The nominal hydrogen production rate of the plant working at atmospheric pressure is 100 STP 1. h -1. A concise description of the process flow scheme is given. Operation of the plant demonstrated that thermochemical water decomposition is a viable process and a valid alternative to other processes for hydrogen production from water. The availability of the plant for experimental work was very high and no special start-up procedure was required. All reactions and product separation units were working within their specifications and no by-product formation was observed. INTRODUCTION
- - to ascertain whether the specified reactant conversions and concentrations could be reached and maintained; - - to check the possible formation of by-products; - - to investigate process control problems connected with the continuous operation of a closed cycle process; -to obtain data for the preparation of specifications for possible future larger size plants. It is felt that these goals were fully reached during the successful operation of the plant for a year and a half. This statement will be discussed in more detail in the following sections. Additionally, another objective emerged unexpectedly, i.e. the study of the behaviour of various cell designs for the electrolysis of HBr on the scale required in the plant. Indeed, this point turned out to be one of the main items of our work and, consequently, covers one of the main subjects of this paper.
The application of thermochemical cycles for hydrogen production from water should circumvent the conversion of large amounts of thermal energy into electricity, followed by electrolysis as in the direct electrolysis of water. Research on thermochemical cycles started in about 1970 in several places of the world. In 1974, a variant of the purely thermochemical cycles was introduced : the hybrid cycle. Here, electrical energy is used in one reaction of the process as a means to overcome the positive free energy barrier inevitably occurring in any purely thermochemical cycle. Obviously, the electrolysis step in hybrid cycles is performed at very low voltages in comparison with direct water electrolysis. Examples of potentially promising hybrid cycles are the Westinghouse sulfur cycle (also known as the Ispra Mark-11 cycle) and the Ispra Mark-13 sulfur-bromine cycle. The latter cycle is the subject of this paper. It was invented by Schiitz and Fiebelmann [1] and consists of the following three reactions: 2
HBr --->H2 + Br2 (the electrochemical step)
(1)
SO2 + Br2 + 2H20 --->2 HBr + H2SO4
(2)
H2SO4---> H20 + SO2 + 1/2 02
(3)
The development of this cycle is well past the laboratory stage and a bench-scale continuous model plant was designed and constructed. The plant has been in full operation since May 1978 [2]. At the time of its inauguration, it represented the first demonstration of a complete hydrogen production process by the thermochemical route. The nominal hydrogen production rate of the plant is 100 1. h -1. The plant is constructed mainly of commercial glass and quartz equipment connected with PTFE tubing and is operated at atmospheric pressure. The original objectives of the operation of the circuit were:
PROCESS DESCRIPTION A detailed description of the process has already been published on various occasions [2, 3]. The process description is therefore confined to a schematic drawing of the model (Fig. 1) and a concise lay-out of the reactor arrangement. The plant consists of three main sections: (A) HBr-electrolysis comprising the bromine distillation/HBr absorption column; (B) The acid formation section, consisting of two reactors for the acid formation reaction, where gaseous HBr and sulfuric acid of about 75 wt% are formed, coupled with a sulfuric acid concentration column for the production of 95 wt% sulfuric acid; (C) The sulfuric acid decomposition reactor and a separation unit for SO2 and O2. (A) HBr-electrolysis
629
An aqueous solution of 45 wt% HBr is passed through
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Fig. 1. Mark-13 flow diagram. the electrolysis cell. The liquid leaving the cell contains about 40 wt% HBr and 4 - 6 wt% Br2. The formed bromine is completely dissolved in the electrolyte. The exit liquid is pumped into a distillation column. In this column, bromine is distilled out of the top and is fed to the acid formation section to be reduced to hydrogen bromide. This product from the acid formation section enters the distillation column as a vapour. Two processes take place simultaneously in this column, i.e. the stripping of bromine as well as the reconcentration of the electrolyte with HBr from 40 to 45 wt%. The hydrogen formed in the electrolytic cell is measured by a wet gas meter and leaves the plant as the main product. (B) The acid formation section The reaction between bromine, sulfur dioxide and water to gaseous hydrogen bromide and liquid sulfuric acid is carried out in two packed columns in series. In one column, Br2 and SOz are contacted with a sulfuric acid solution of about 75 wt%. Bromine is slightly in stoichiometric excess to SO2. The solution is steadily recycled over the column to warrant a satisfactory gas/liquid contact. During the passage of the gas through this column, partial conversion of Br2 and SO2 is obtained. The formed H2SO4 dissolves in the liquid and leaves the column by an overflow system. This solution is fed to the H2SO4 concentration column, where it is further concentrated to 95 wt% H2SO4. The gas stream leaving this first reactor consists of
HBr, Br2 and SO2. This mixture is fed to a second reactor where a solution of about 55 wt% H2SO4 is circulating. Here, the reaction proceeds to completion and gaseous HBr with a small percentage of bromine, but free from SO2, leaves the reactor and is fed to the bromine distillation column (section A). The formed sulphuric acid dissolves in the circulating solution and is fed by overflow to the reactor containing the 75 wt % H2SO4 solution. The water required for the reaction is introduced into the second reactor. (C) The sulfuric acid decomposition section The concentrated sulfuric acid produced in the acid formation section is pumped to the sulfuric acid evaporator, where it is heated to about 900 K. The vapour is then further heated to 1000--1100 K and passed over a catalyst. A residence time of about 0.3 s is sufficient to decompose 70% of the inlet sulfuric acid into SO2 and 02. The water formed in this reaction and the unreacted H2SO4 are cooled and condensed and passed to the sulfuric acid concentration column (section B). The separation of SO2 and 02 is performed primarily by cooling down the gas mixture to 225 K. About 95% of the sulfur dioxide is condensed in this step. This liquid SO2 is subsequently evaporated and recycled to the acid formation section. The gas mixture, leaving the cooling trap, consists of approximately 91 vol % 02 and 9 vol % SO2. The remaining SO2 is removed by contacting the gas with
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS bromine and water, i.e. reaction (2). The purified oxygen leaves the process as a secondary product. A photograph of the plant is given in Fig. 2. O P E R A T I O N OF THE PLANT As mentioned above, the plant was put into operation in May 1978 and has always run very smoothly. During the first months, a number of minor modifications had to be carried out, mainly in the bromine distillation column and the H2SO4 decomposition section. The only other major difficulty encountered was the failure of the direct current generator for the HBr-electrolysis in the summer of 1979. This caused the plant to be shut down for 6 weeks. However, the fact that the plant has been unavailable for experimental work due to the failure of an ancillary piece of equipment is much more the exception than the rule. Under normal conditions, the availability of the plant was high. The circuit can be started up in a short time without any special problems. Apart from a preparatory period of about 2 h when the two distillation columns and the sulfuric acid decomposition reactor are pre-heated, and
631
the refrigerating unit for the SO2-condensation is cooled, no special start-up procedure is required. Once the required temperature levels are reached, the plant can be put on stream immediately at its nominal hydrogen and oxygen production flow-rate. It is then very easy to maintain the circuit in near steady-state operation, due to the fair reliability of the metering pumps for flow-rate control. There are neither major instabilities nor formation of by-products or side-reactions. In conclusion, it is technically feasible to put the plant into operation at any desired time and to maintain it in continuous operation for very long periods. There are no technical reasons requiring any interruption of operation. RESULTS (A)
HBr-electrolysis
As mentioned above, the study of the performance of various cells for the electrolysis of HBr turned out to be one of the main subjects of research during the operation of the bench-scale plant. Three different cells
Fig. 2. Laboratory scale demonstration plant for the Mark-13 process.
632
D. VAN VELZEN AND H. LANGENKAMP
have been tested. Two of the three also underwent some minor modifications, mainly to improve their hydrodynamic behaviour. All three cells were so-called horizontal controlled convective diffusion cells [4]. This cell type works without a diaphragm, its principle being that the flow-rate of the electrolyte through the cell is kept well within the laminar regime, avoiding as much as possible convective currents from the anode to the cathode. In this way, bromine, generated at the anode, is kept away from the cathode region, and high (near to 100%) current efficiencies can be obtained. Schlitz et al. [4, 5] reported promising results in small scale experiments, so that testing of the controlled convective diffusion cells on a somewhat larger scale in the bench-scale plant was a logical follow-up. The tested cells are monopolar and graphite is the electrode material for both the anode and the cathode. In all experiments we used platinum as a homogeneous electrocatalyst, the application of this noble metal considerably reducing the overvoltage at the cathode. The amount of Pt applied corresponds generally to the formation of a Pt-layer of 3-4 mg cm -2 on the electrode surface. The first cell is a rectangular cell, whereas the second cell is circular. In these two cells, the evolved hydrogen flows through the cell co-currently with the electrolyte flow. The third experimental cell is also a circular one with slightly inclined electrodes. Here, the hydrogen flow is counter-current to the electrolyte flow. Rectangular cell. The rectangular cell was the first cell used in the plant. A cell lay-out is given in Fig. 3. The electrode surface is 352 cm 2 and the electrode distance can be varied between 3 and 5 mm. A typical cell voltage curve is given in Fig. 4. Cell voltages are 0.85 V at 200 m A c m -2 and 1.03 V at 500 mA cm -2 at 350 K. The electrolyte is 45 wt% HBr. In the early experiments carried out with this cell, the current efficiency could not be raised above 91%. Moreover, this value did not remain constant with time, but decreased slowly during the course of each experiment [2]. For example, after 30 h of continuous operation, the current efficiency dropped from 91% to values around 75%. The run was interrupted at this point and 47%
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the cell was flushed with bromine-free electrolyte for several hours. In a subsequent experiment, the original value of 91% was again found at the start of the run. It was anticipated that an insufficiently regular liquid flow pattern through the cell was the cause of this phenomenon. As a remedy, a modification of the cell was therefore made, consisting of the replacement of the liquid feed and outlet system. Originally there was only one inlet and one outlet tube, whereas the new arrangement comprises three parallel inlets and a liquid overflow with three separate outlets. The inlets and outlets are at a mutual distance of 60 mm. The modified cell performs very well. The observed cell voltages are slightly higher than those for the original cell, i.e. 0.90V at 200mA cm -2 and 1.10V at 500 mA cm -2 (previously 0.85 and 1.03 V, respectively). The current efficienc),, however, improved considerablyi at 200 mA c m - ' i t is now 99%, and at 500 mA cm- 94%. It follows that the loss in cell voltage is amply compensated by the gain in current efficiency and the electric energy consumption is practically constant. The most important improvement, however, is that the modified cell can now be operated for long periods without any concomitant decrease of current efficiency. The cell operated in a 30 hr experiment at 400 mA cm -2 at 1.03 V, at an average current efficiency of 95%. The electrolyte flow rate was maintained in this experiment at 8 kg h -1, and the temperature was 350 K. The favourable effect of the above modification is illustrated in Fig. 5, where the change of the current efficiency with time is shown before and after the three tube inlet system was mounted. Limiting current density. It must be noted that the current density of a given horizontal cell (and thus the hydrogen production rate) cannot be increased indefinitely. In general, running a cell above a limiting condition gives rise to the following course of events. Initially the cell operates well, cell voltage and current efficiency are in the anticipated range. However, after
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS
633
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a certain running time (usually about 3 h), a rapid decrease of the current efficiency occurs. In a few minutes it becomes virtually impossible to operate the cell, The hydrogen production rate drops dramatically and bromine vapours appear in the hydrogen outlet. At this point, bromine is evaporating from the electrolyte, severe mixing of the liquid phase takes place and, obviously, the flow regime is far from laminar. At this stage, the current must be switched off and the electrolyte flow rate maintained. After approximately I h the cell can again be put into operation. This cell failure has two main origins: (a) an excessive H2 production rate which causes mixing in the liquid phase due to the presence of gas bubbles; and (b) high bromine concentrations which favour the local formation of bromine vapour bubbles and, consequently, mixing of the liquid phase. Both (a) and (b) increase with increasing hydrogen production rate. A badly developed liquid flow pattern through the cell leads to the formation of preferential flow regions and more or less stagnant zones. This causes very high local bromine concentrations in the "stagnant" zones, and a largely increased probability of (b). It is therefore of paramount importance that the liquid flow pattern through a horizontal cell be well controlled. From the foregoing it follows that the limiting current density of a cell depends on the cell design, dimensions, electrode distance and configuration and can only be found empirically for each cell. The limiting current density for the rectangular cell is 500 mA c m -2. At 450 mA crn -2 the cell could be operated for long periods (20 h) under stable conditions, but operation at 540 and 600 mA cm -2 led to cell failure after 1-3 h of operation. Values of 80-85% of the limiting current density are recommended as safe values for cell operation in long duration experiments. This implies that the rectangular cell can be safely operated at current densities of
401)--440 m A c m -2. Under these conditions the cell produces 60-65 I. h -1 of hydrogen. Circular co-current cell. A schematic diagram of the cell is shown in Fig. 6. It essentially consists of two parallel circular graphite electrodes, fixed in a Teflon housing. The HBr solution is fed at the centre and flows radially towards the periphery of the electrodes, where it is collected and leaves the cell. Hydrogen is collected in the annular ring and leaves the cell by an exit in the upper part of this ring. The total electrode surface is 283 cm 2 and the electrode distance of the cell is 5 mm. The initial results obtained with this cell were rather disappointing. The typical cell voltage was very near to the one obtained for the modified rectangular cell, as shown in Fig. 4, but the current efficiency was very low: 85% maximum. An estimation of the pressure head due to the friction of the electrolyte flow through the cell indicated that this pressure drop was extremely low (about 10 -3 cm liquid head). In this case there was a strong preferential ~2
HBr
HBr ~-Brz ELectrode material grophi'l'e Electrode disl"once 5ram Electrode surfoce 283 cm2 Fig. 6. Circular co-current electrolytic cell.
634
D. VAN VELZEN AND H. LANGENKAMP
flow from the inlet to the only outlet tube and very bad flow distribution resulted, with detrimental effects on the current efficiency. The installation of an overflow system with four outlets at 90°, with a common flow resistance downstream, was introduced to improve the flow pattern. Indeed, this modification considerably improved the cell performance. Current efficiencies of nearly 100% could now be reached, with cell voltages of 0.90 V at 250 mA cm -2 and 0.99 V at 450 mA cm -2, at 46 wt% HBr at 350 K. These results were obtained in short-time tests whilst the testing of the cell during longer periods combined with regular operation of the total plant was the next stage of the programme. In the first test of this series, the cell was run for a period of 2 h at a relatively low current density: 245 mA cm -2. During this period, the current efficiency was excellent: 99%. The aforesaid conditions correspond to a hydrogen production ~'ate of approximately 30 STP 1. h -1, which is too low for regular operation of our bench-scale unit. We were therefore obliged to run the cell at much higher current densities, i.e. about 450 mA cm -2. In these experiments initially high values for the current efficiency were obtained (95%), but after a running time of 2.5-3.0h, the current efficiency decreased dramatically and finally it became impossible to operate the cell. After shutting down and rinsing the cell for about 1 h, with the recycle electrolyte solution, the original high initial current efficiencies were found again, but after 2.5 h, the cell failed once more. Obviously, for this cell the limiting current density is lower than 450 mA cm -2 and will be about 300 mA cm -2. This limits the hydrogen production capacity to approximately 30 1. h -1. Circular counter-current cell. This last cell type has to be considered as a drastic modification of the circular co-current cell. Its dimensions are nearly identical and it is mounted in the same PTFE housing as the old, co-current cell. A schematic diagram of the new cell is given in Fig. 7. H2
I
The electrodes are no longer horizontal, but slightly inclined towards the centre. The inclination angle is 8 °. The main difference, however, is the hydrogen outlet. In this new cell the hydrogen leaves at a single outlet at the centre of the cell, whereas the bromine-loaded electrolyte is still flowing radially towards the annular ring. The flow of hydrogen through the cell is thus counter-current to the flow of the electrolyte. This arrangement gives the new cell a number of advantages: (i) The old, co-current cell is very susceptible to the control of the liquid level in the cell. A low level causes the cathode to run dry, whereas an excessively high level leads to increased liquid mixing and thus to low current efficiencies. The range between high and low critical level is usually not more than 2-3 mm. The configuration of the new cell eliminates this disadvantage completely. (ii) It is relatively simple to build a bipolar cell on the principle of the new cell, which is much more difficult for the old, co-current cell. Among other things, this is due to difficulties in level control for each single cell compartment for the old cell. (iii) In the new cell the hydrogen leaves the cell in contact with the fresh, bromine-free electrolyte, instead of being in contact with the electrolyte containing bromine as in the old cell. The evolved hydrogen is therefore essentially free of bromine. We measured bromine contents of about 40 mg m -3 H2 for cell operation at 350 K. For co-current cells this value is about 1000 mg m -3 H2. The electrode surface of the new cell is 283 cm 2, like the old cell, and the electrode distance is 5 mm. The performance of the cell is similar to that of the cocurrent cell. Typical cell voltages are 0.89 V at 250 mA cm -2 and 1.01 V at 450 mA cm -2 at 350 K and in 47 wt% HBr solution. Under these circumstances the current efficiency is between 93 and 96%. As in the old, co-current cell, high current densities cannot be applied. Above 300 mA cm -2, difficulties in the operation of the cell occur. On the other hand, operation at 220 mA cm -2 is completely stable. Over long periods (> 10 h) the cell voltage could be maintained at 0.89 V, and the current efficiency at 93-94%. Under these conditions the cell produces 26 STP 1. h -1 of hydrogen.
(B) Acid formation
t 4 o * HBr+ 5% Br2 4?'% HBr Electrode mol-erial graphil"e Electrode disl"once 5rnrn Elecl"rode surfece :>83 cmz
Fig. 7. Circular counter-current electrolytic cell.
The required performance of the acid formation section is the simultaneous production of a gas stream of 8.0 gmol HBr h -1 (which must be as pure as possible) and a liquid phase of high H2SO4 concentration. van Velzen and Langenkamp [6] showed by equilibrium measurements that high sulfuric acid concentrations are attainable, only, however, in equilibrium with considerable partial pressures of bromine and sulfur dioxide in the gaseous phase. On the other hand, the partial pressures of bromine and/or sulfur dioxide are
HYDROGEN PRODUCTION BY THE MARK-13 PROCESS negligible when the liquid produced has a sulfuric acid concentration below 55%. These observations imply that the reaction must be carried out in some type of counter-current staged reactor as otherwise considerable amounts of SO2 and Br2 must be separated from HBr and recycled. The problem has been solved by performing the reaction stage-wise in two co-current packed columns in series, analogous to a gas absorption operation. The overall SO2 conversion in the acid formation section has to be higher than 99%. It is found that this target can be easily reached provided that the relative sulfuric acid concentration in the first column is maintained below 55%. Typical liquid compositions and operation conditions are given in Table 1. Under the conditions given in Table 1, the sulfur dioxide conversion in the concentrated column amounted to 45.3%, and the overall SO2 conversion was as high as 99.7%. The residual SO2 concentration in the HBr stream was only 1300 ppm. The operation of this reactor section proved to be very simple and reliable. The working conditions can be varied over a wide range and are easily adaptable to the requirements of the operational parameters for the rest of the plant. Further scale-up of this reactor section is relatively easy, because common laws for gas absorption with chemical reaction can be followed. The equipment can be operated for any length of time without instabilities and there is no indication of by-product formation. It can be concluded that the acid formation section is a versatile and reliable part of the plant. (C) H2SO4
decomposition
The required duty of this reaction section is the production of 2 gmol oxygen h -1 at a reactor temperature around 1073 K. For this purpose a reactor consisting of
Table 1. Typical operation conditions for the acid formation section Into app. From app. Exit app. No. 1 No. l t o a p p . No. 2 No. 2 Liquids Flow rate (g h -1) Composition (wt %) Br2 HBr H2SO4 H20
198.5
445.4
377.6
-5.3 -94.7
3.9 32.2 34.7 29.4
2.8 75.0 22.2
Vapours Flow rate (STP I. h -1) Composition (vol. %) Br2 HBr SO2
140.4
182.7
145.5
50.1 -49.9
22.5 56.5 21.0
3.9 96.1 1300 ppm
635
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Fig. 8. Temperature profile of the sulfuric acid decomposition furnace. a quartz tube (3.5 cm internal diameter) heated by a split tube furnace was used. A residence time of i s (calculated at STP) in the catalytic zone seemed a suitable target. This calls for a catalyst volume of 85 cm 3. This quantity was placed in two separate sections, of 40 and 45cm 3, respectively, the one 5 cm from the other. This gap serves to reheat the reaction mixture following its passage through the first catalyst bed. The catalyst used is a special catalyst developed by the firm P6chiney-Ugine Kuhlmann, based on ferric oxide on a porous alumina support. The sections of the reactor tube not occupied by the catalyst are packed with quartz chips to improve the flow pattern and the heat transfer. The performance of the sulfuric acid decomposition reactor is very good. The average reactor conditions obtained during a 30 h run are given in Table 2. This conversion has been obtained with the axial temperature profile in the converter displayed in Fig. 8. Here, the two catalyst zones are clearly recognizable. Obviously, it is not possible to assign one fixed value for the reactor temperature in the catalyst zones, due to the very pronounced axial temperature profile in these zones. These temperatures range from 998 K as the minimum to 1040 K as the maximum recorded at the exit of the second catalyst zone. The degree of conversion obtained in this experiment is near to the equilibrium value at 1070K, i.e. higher than the recorded maximum catalyst temperature. This apparent contradiction can be explained as follows: the temperature is recorded in the centre of the reactor tube. Table 2. Typical operation conditions for the HzSO4 decomposition section. Average reactor temperature 1015 K IN: 95 wt% H2SO4 flow rate
462 g h -1 4.50 gmol h -1
OUT: Oxygen SO2 condensed at 220 K SO2 not condensed (8.8 vol. % of SO7]O2
39.1 !. h -4 1.61 gmol h -1 200 g h-l 3.12 gmoi h-I 3.15 1. h -1 0.13 gmol h -1
mixmre) Total SO2 production SO3-conversion
3.25 gmol h -~ 72%
636
D. VAN VELZEN AND H. LANGENKAMP
Besides the axial temperature profile shown in Fig. 8, there also exists a pronounced radial temperature profile with a minimum temperature in the centre and the maximum at the tube wall. The bulk of the gases are moving in the vicinity of the wall, i.e. at the higher temperature. The average value for the liquid hourly space velocity for the present reactor conditions, calculated for a hypothetical gas flow at 298 K and atmospheric pressure, is 4400 h -1. As a general conclusion, it can be stated that the sulfuric acid decomposition section is also performing very well. There are no instabilities in the course of time and there is no formation of any by-products.
CONCLUSIONS The bench-scale plant for hydrogen production by the Mark-13 process has been successfully operated since May 1978. During this period practically all our pre-set objectives could be realized: - - I t was convincingly demonstrated that thermochemical water decomposition is a viable process and a valid alternative for hydrogen production from water. - - T h e specified reactant concentrations and conversions can be reached and stably maintained over long periods. - - I n some cases (e.g. the acids formation reaction), the performance of the plant is even better than the original specifications, so that improvements in the reference design flow sheet could be made.
- - U n t i l now, no by-products have been detected in any reactor section. - - T h e horizontal cells without a diaphragm for HBrelectrolysis are a promising alternative to conventional cells. Attractive voltages and high current efficiencies could be reached in the bench-scale plant. The liquid flow pattern in the cell is of paramount importance for the working of this cell type. In further development work this last point will be given further attention. REFERENCES 1. G. H. Schiitz and P. Fiebelmann, Patent Application No. 71037, filed 3 October 1974, Luxemburg. 2. D. van Velzen and H. Langenkamp, Thermochemical hydrogen production by the Mark-13 process: a status report, Proc. 14th Intersoc. Energy Conversion Eng. Conf., Boston, Vol. 1, pp. 783-789 (1979). 3. D. van Velzen, H. Langenkamp, G. H. Schiitz, D. Lalonde, J. Flamm and P. Fiebelmann, Development and design of a continuous laboratory-scaleplant for hydrogen production by the Mark-13 cycle, Int. J. HydrogenEnergy 5,131 (1980). 4. G. H. SchiRz and D. Lalonde, A new electrolytic cell type for hydrogen production in hybrid cycles, Int. Conf. on Alternative Energy Sources, Miami Beach, December 1977. 5. G. H. Schiitz, P. J. Fiebelmann and D. Lalonde, Electrolysis of hydrobromic acid, in Hydrogen Energy System (Proc. 2nd World Hydrogen Energy Conf., Ziirich, Switzerland) (T. N. Veziro~lu and W. Seifritz, eds.) Vol. 2, pp. 709730. Pergamon Press, Oxford (1978). 6. D. van Velzen and H. Langenkamp, The oxidation of sulfur dioxide by bromine and water, Int. J. Hydrogen Energy 5, 85 (1980).