Ind. Eng. Chem. Res. 1995,34, 3777-3783
3777
Elimination of Ammonia from Coal Gasification Streams by Using a Catalytic Membrane Reactor Edward N. Gobina, Jaafar S.Oklany, and Ronald Hughes*?+ Department of Chemical Engineering, University of Salford, Salford M5 4WT,U.K.
The application of a catalytic membrane process for the removal of dilute concentrations of ammonia resulting from the gasification of coal has been investigated. Mathematical simulations have been conducted for a n experimental reactor comprising a thin but continuous layer of a Pd-Ag alloy deposited on a porous substrate. Data for the permeation of hydrogen through the composite membrane were determined experimentally. A rate equation for ammonia decomposition over Ni/A1203catalysts was coupled with the experimentally determined hydrogen permeation rate data to obtain results for various reactor configurations. The results presented here have demonstrated the superiority of such a composite membrane reactor system over other methods. Furthermore, complete destruction of the ammonia was achieved for the relatively short space times considered i n this study and can also be applied in operations carried out under gasification conditions.
Introduction Environmental issues are set to play an increasingly prominent role especially in the chemical and petrochemical industries in the immediate future. Of major concern is the so-called greenhouse effect, considered by many t o be due to the high concentration of carbon dioxide, methane, and nitrogen oxides in the atmosphere. The increasing levels of these gases have been traced to exhaust gases from industrial plants, fossil fuel utilization for energy production, and, last but not least, traffic exhaust gases. The chemical process industry has the task of developing the expertise to respond to the challenge by devising new processes and finding solutions which should not produce anomalies such that the target pollutant is reduced at the expense of increasing or creating another. In the development of an industrialized modern society, provision of energy in a reliable, economic, and “clean” form should be considered an important aspect. However, at present some energy supply technologies have a significant environmental impact. The use of fossil fuels in advanced power generation systems is a typical example and is the subject of recent studies (Bozza et al., 1994;Smoot et al., 1993;Backer, 1993). Coal gasification may result in the production of small but potentially harmful amounts of “3. The concentration of NH3 in the synthesis gas stream will however depend on the gasifier design and operating conditions but may be as high as 0.38 mol % (Robson et al., 1976). In order to utilize this synthesis gas for power generation in the gas turbine, the NH3 must be removed to prevent NO, formation at conditions encountered in the adiabatic gas generator. Water based NH3 removal systems involve cooling the synthesis gas to around 400 K, and although a substantial amount of the NH3 is removed, residual levels of up to 400 ppm have been reported (Harris and Shah, 1976). Although the sensible heat recovered in this cooling process may be effectively utilized elsewhere, the process results in reduced cycle efficiency. High temperature clean up systems presently being developed remove very little or no NH3 at all. Therefore, a strong economic incentive
* Author to whom correspondence should be addressed. +
E-mail address:
[email protected].
Table 1. Typical Composition of Coal Gasification Stream component concn level component concn level species (mol %) species (mol %) Hz 20.0 Nz 48.0 “3
HzO HzS
0.3 1.0 -0.01-0.5
co + coz CHI
26.0 4.2
exists for the development of a process capable of effectively controlling NO, emissions resulting from “3.
The application of a high temperature catalytic membrane reactor process integrally linked with the gasification process appears to offer a unique route toward an efficient and cost effective method of removing the NH3 from synthesis gas streams. This is because membrane based separation systems possess low energy consumption, reduced environmental impact, low costs of maintenance, and space and weight efficiency and are relatively easier to install and operate. Poku and Plunkett (1989)have assessed the potential applications of membranes to coal gas clean up processes. Their study included various coal conversion processes such as the direct coal-fueled turbine (DCFT), hydrogen production, and the integrated gasification combined cycle (IGCC),among others. It was concluded that inorganic membranes that could separate gases by mechanisms other than Knudsen diffusion would be preferable. This conclusion was reinforced by a recent study by Collins et al. (1993)in their study of ammonia decomposition assisted by a membrane reactor. They concluded that only membranes with hydrogen selectivities greater than 50, which also have the capability of operating at large total pressure differences between the tube and shell sides, would achieve a high NH3 conversion. Our study aims to verify this claim by using permeability data obtained for a membrane having exclusive hydrogen permeability. The major difference between the present study and the earlier ones discussed above is that the membrane now considered possesses 100% permeability to hydrogen alone. Furthermore, simulations have been carried out under a wide range of operating conditions, and therefore the results can be applied to operations carried out under gasification conditions.
Q888-5885/95/2634-3777$09.Q~JQ 0 1995 American Chemical Society
3778 Ind. Eng. Chem. Res., Vol. 34,No. 11,1995 Table 2. Kinetic Param&m for N& Decomposition on NVAhOs Catalyst (Collins et al., 1993) varameter
Value 0.614 5.744 x 1019
2.304 x los
We have demonstrated in a n earlier study that enhancement of the final conversion in a reversible dehydrogenation system is possible even at a relatively low temperature by utilization of a membrane reactor which allows the exclusive and continuous removal of hydrogen from the reaction mixture (Gobina and Hughes, 1994a.b). The reaction can therefore he enhanced, resulting in a higher conversion far beyond that obtainable in conventional fmed bed reactors (Itoh, 1987). In the present investigation, a composite membrane consisting of a very thin (6pm), but continuous, layer of palladium-silver alloy deposited on the outer surface a of thermostable mesoporous support has been used to establish the hydrogen Permeability data used for the mathematical simulation of NH3 decomposition. The principles explored in this study can also be used for the recovery of hydrogen from waste gas streams such as those encountered in ammonia synthesis.
Analysis (a) Reaction Rate and Kinetics. It was assumed that the kinetics of NHs followed the Temkin-Pyzhev mechanism which describes the rate Of NH3 decomposition on a solid catalyst. The rate equation based on this mechanism is given by eq 1.
The preexponential factor, k., the activation energy, E, and the exponent @ have been determined experimentally for NH3 decomposition over NVAl203 catalysts in the pressure range 9-36 atm and a temperature range of 673-823 K (Collins et al., 1993) and are presented in Table 2. Here, Keqis the thermodynamic equilibrium constant defined by eq 2 (Harrison and Kobe, 1953;Neilsen et al., 1964).
2250'322 - 0.85340 - 1.51049 log T 25.8987 x lo-?'+
14.8961 x 1O-'P
Figure 1. SEM photngraph of Pd-Ag composite membrane on Vycor glass.
pm. A typical SEM of a Pd-Ag sputtered Vycor substrate is shown in Figure 1 and indicates a coherent film with good adherence as a result of which the surface morphology of the original substrate is translated on the film itself. The present simulation study has been based on a film thickness of 6 pm. The permeability of hydrogen through the composite membranes has been determined experimentally, the details of which have been described elsewhere (Gobina et al., 1994; Gobina and Hughes, 1993). Briefly measurements were made between 473 and 653 K and pressure differentials of up to 2.2atm. The permeation rate of hydrogen, Q H ~(molls) was found to obey the half power pressure relationship given by eq 3 (Sieverts and
Kumbhaar, 1910;Ackerman and Koskinas, 1972).where 8. is the permeability constant of hydrogen gas through the membrane (molvm thicknessl(cm%atmm)); A is the membrane area available for flow (cm2),6 is the PdAg alloy film thickness (em), and P I and Pz are the partial pressures (atm) of hydrogen gas in the upstream and downstream sides, respectively. The permeability constant, Q., was found to obey the Arrhenius law as given by eq 4, where R is the universal gas constant
Qo = 3.2027 x lo-' e-(*)
(2)
(b) Hydrogen Permeation Rate. The substrate material used for the preparation of the composite metallized membrane is a porous Vycor glass tube (4nm nominal pore size, lOmm o.d., and 1.1-mmthickness) supplied by Coming Ltd. The tube lengths used were 200 mm. Prior to magnetron sputtering, the substrates were cleaned using deionized water dried in a vacuum and then temperature stabilized. After cleaning by sputtering in the deposition chamber, the magnetron deposition technique was used to prepare thin coherent films on the tubular substrates. Film thickness (and hence deposition rates) was estimated using additional samples in the deposition chamber placed adjacent to the actual substrate of interest. Using various values of magnetron power, the deposition rates were varied and the minimum thickness required to prepare pinhole free dense, continuous films was ascertained. This was found to be not less than 4
RT
(4)
(kJ/(mol.K)) and T i s the absolute temperature (K).The permeation constant, activation energy, and permeation rates derived from eqs 3 and 4 are in good agreement with results reported in the literature (Govind and Atnoor, 1991)for hydrogen permeation through a composite membrane consisting of a thin but continuous Pd-Alloy layer on a porous substrate. (c) Mathematical Model. In order to derive the mathematical formulations used to represent this system, the following assumptions were made: (1)There is a negligible pressure drop. Experimental reactors such as those considered in this study have a typical pressure drop of roughly 0.004 a t d c m . In the case of our experimental system with a relatively small tube length, this would result in a negligible pressure drop when compared with the base operating pressure of 36 atm. Therefore, this assumption is acceptable. (2)The membrane has exclusive selectivity for Hz. This assumption was confirmed by the experimental
Ind. Eng. Chem. Res., Vol. 34, No. 11,1995 3779 NIIALP, CATALYST
THIN Pd-Ag LAYER
POROUS CERAMIC SUPPORT
AMMONIA-RICH SYNTHESIS GAS
MIXTURE
HYDROGEN-RICH STR€AM SWEEP GAS
STAINLESS STEEL SHELL
\
HYDROGEN PERMEATION
Figure 2. Catalytic membrane reactor concept in countercurrent flow mode.
work already carried out on hydrogen permeation (Gobina et al., 1994). (3) The operation is isothermal. This assumption is acceptable due to the low ammonia concentration in the feed. (4) Plug flow prevails in the shell and tube sides. This assumption will be shown t o be valid under the range of operating conditions that these simulation runs were carried out. (5) Radial tube-side and shell-side concentration gradients are negligible. (6) Axial diffusion in the membrane is negligible compared with radial diffusion. The ratio of the thickness of the membrane to its length is of the order of thus validating the assumption. (7) There is steady state operation. (d) Mathematical Formulations. (i) Tube-Side Material Balance. A schematic of the proposed catalytic membrane reactor is presented in Figure 2. Four components are present in the tube side, namely, ammonia, hydrogen, nitrogen, and inerts (this includes all other components in the feed stream). A component material balance on the tube side yields
Table 3. Summary of Input Base Data for the Mathematical Simulation volume of the catalyst bed (cm3) palladium-silver alloy film thickness (cm) diameter of catalyst pellets (cm) pressure of synthesis gas (atm) pressure of sweep gas stream (atm) flow rate of synthesis gas (cm3(STP)/min) flow rate of sweep gas stream (cm3(STP)/min)
(6) where the positive sign is for the cocurrent model and the negative one is for the counter current model.
1.0
350.0 300.0
(iii) Numerical Solution. The system of six ordinary differential equations (ODES) that comprise the cocurrent model, eqs 5 and 6, is solved using the fourth order Runge-Kutta technique with the following initial conditions:
4;
= [4;1°
(8)
For the countercurrent model, the following initial conditions apply:
4; = [4fI" L = 0 4;
where WH, is the ammonia decomposition reaction rate, vi is the reaction stoichiometry of each component (negative for reactants), &i is the permeation rate which is calculated from eq 3 for Hz and is zero for the other components, 4; is the component molar flow rate divided by the flow rate of ammonia in the initial feed lo, and L and LO are the dimensionless reactor l e n d h and reactor length, respectively. (ii) Shell-SideMaterial Balance. The model consists of two components in the shell side, namely, hydrogen and the sweep gas used. A material balance for each component on the shell side yields
6.689 6.0 10-4 0.15-0.2 36.0
= [4;'10
L = 1.0
(9) (10)
The Shooting method is used to solve the resulting boundary value problem in the countercurrent case. The cocurrent results are used to provide the initial guess for the countercurrent problem (Press et al. (1986)). No stability problems have been encountered for all the runs carried out in this study.
Results and Discussion Various options of the membrane reactor were examined and used as the basis for comparison. These include fixed bed operation (permeation = 01, no sweep flow but permeation, and operation as a membrane reactor (permeation and reaction). Each of these configurations will be examined and the results discussed below. The base data for reactor simulation are presented in Table 3. (a) Fixed-Bed Operation (Permeation = 0). In this operation, the permeation rate of hydrogen through the composite membrane was zero, and therefore operation corresponded to a conventional fured-bed reactor. The reaction of ammonia decomposition can be written
3780 Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 100
.-"
-7
u 0 '5 31 .25 f
58.75
N
80 70
'f
60
>
50
45.00
E E
U
90
80
40
'F
30
17.50
20 10 0
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9 1 .O
0.0
L/LO
Figure 3. Comparison of ammonia conversion in fixed-bed (permeation = 0) and membrane reactor operation.
0.1
0.2
0.3
0.4
0.6
0.5 L/LO
0.7
0.8
0.9
1.0
Figure 4. Effect of temperature on ammonia conversion (countercurrent mode).
simply as k,
2NH3 k,N,
+ 3H,
where kl/kZ = Kes,and can be calculated as given from eq 2 for a given temperature. Depending on the concentration of hydrogen in the feed gas stream, therefore, the equilibrium can either shift t o the right (resulting in NH3 decomposition) or to the left (resulting in NH3 production). At the hydrogen concentration of 20 vol %, which is typical of most synthesis gaseous streams, our simulation results show that the equilibrium shifts toward NH3 production. This is depicted in Figure 3, confirming the unsuitability of utilizing a fmed-bed reactor on its own for NH3 destruction. (b) Membrane Reactor Operation without Sweep Gas. During operations in this mode, hydrogen was allowed to permeate through the membrane, but the sweep flow rate was set to zero. The effect of this type of operation on the conversion of NH3 is also shown in Figure 3. Here, due to the permeation of hydrogen, the equilibrium shifts to the right and ammonia decomposition occurs. However, although high NH3 conversion is attainable, complete destruction of the ammonia cannot be achieved due t o the fact that the shell-side and tube-side hydrogen partial pressures become equal, resulting in no net driving force for hydrogen permeation. Complete NH3 destruction could be achieved by a reduction of the shell-side hydrogen partial pressure. In this study a sweep gas consisting of pure nitrogen was employed. (c) Membrane Reactor Operation with Sweep Gas. There is a clear advantage to be gained by employing a sweep gas. The driving force for hydrogen permeation is the difference in the square roots of the hydrogen partial pressure in the shell and tube sides, respectively. Using a sweep gas in the shell side increases this driving force by reducing the hydrogen partial pressure in the shell side. More hydrogen is therefore removed from the reaction zone, thus increasing the decomposition rate of NH3 to completion. The profile for such an operation is also included in Figure 3. Various other parameters were examined,the results of which are presented below. (i)Effect of Temperature. The effect of operating temperature on the decomposition rates of NH3 along the reactor tube length is depicted in Figure 4 (countercurrent) and Figure 5 (co-current),respectively. For the countercurrent operation, complete decomposition is achieved even at a temperature as low as 673 K. Our base temperature for this study was selected at 873 K.
90
80
h
N 7 70 'f 60
9
I /
40
/
/
, /
10
0.0
/
0.1
0.2
0.3
0.4
0.5 L/Lo
-
C0-current k d e O.OOBJmin Pmwn Ritie = 0.027 Sweep Rdie,.._.." o.a?7.....-.'-
Sposa rime
-
_... ............. 0.6
0.7
0.8
0.9
1.0
Figure 5. Effect of temperature on ammonia conversion (cocurrent mode). 100
-7 w
'z
90
80 70 60 50
s 40 .-5 30 : 20 l o0t " " 0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
L/LO
Figure 6. Effect of sweep ratio on the ammonia conversion.
For operations at this temperature, although complete decomposition is attained for both modes of operation, a smaller membrane area is required in the countercurrent mode of operation. However, the difference in these modes of operation is very marginal at this temperature, and only below 723 K does any significant difference in these modes become apparent. The results that follow are reported for the countercurrent mode of operation. (ii) Effect of Sweep Ratio. Figure 6 shows the effect of varying the sweep flow rate on the decomposition of NH3 at 873 K. The total feed flow rate was maintained at 350 cm3(STP)/min,and the NH3 mole fraction was 0.3%. The results indicate that although the variation of the sweep gas ratio (which is defined as the ratio of the sweep rate t o that of the feed rate) has an effect on the NH3 decomposition rate, sweep ratios above 0.143 (sweep gas flow rates above 50 cm3(STP)/min)did not achieve spectacular results. This is because the difference in the square roots of the
Ind. Eng. Chem. Res., Vol. 34,No. 11,1995 3781 - 0.03%
0.0223
0,0167 _ - - 0.0131 _- _
-40 -60
-80
1I - . . 00.01 .ws - 00.m71 . m - - _0.W67 - o.we3 - 0,0055
1
-loor.
0.0
Carntersurnnt Wade P-um ratio = 0.027 hrnp.rotun 871K
-
"
0.1
' 0.2
"
0.3
"
0.4
"
"
0.5 0.6
"
"
"
0.7 0.8 0.9
'
I
1.0
L/Lo
c 0.0
-.. .
--. -
. 4
0.1
0.2
0.3
0.4 0.5
0.6
0.7
0.8
0.9 1.0
L/LO
Figure 7. Effect of inlet hydrogen concentration on ammonia conversion.
Figure 8. Effect of space-time on ammonia conversion.
Table 4. Effect of Shell-Side Operating Pressure without Sweep Gas on Ammonia Conversion (Tube-SidePressure = 36 bar) ammonia conversion (%) condition of shell side 95.0 no sweep, pressure at 1.0 bar 100.0 no sweep, pressure at 0.36 bar
of mass transfer from the bulk gas phase to the catalyst pellet is small. However, under conditions of varying feed flow rate at constant catalyst weight, the external mass transfer rate changes with the velocity of the gas stream. The effects of such changes are considered here. The mass transfer flux, N,, of the hydrogen to the inner surface of the membrane is given by
hydrogen partial pressure which is the driving force for Hz permeation does not change significantly beyond a sweep gas flow rate of 50 cm3(STP)/min, and hence equilibrium shift is not effected significantly. In these simulations, we have employed pure nitrogen as a sweep gas so as t o enable higher hydrogen removal rates for a particular upstream pressure. "his hydrogen would be of more value if it could be easily recovered. One possibility is to employ a vacuum in the shell side to extract the hydrogen and so produce pure hydrogen or operate without purge gas at all. These two alternatives are shown in Table 4 using the base case data (Table 3). A subatmospheric pressure of 0.36 bar is sufficient for eliminating the ammonia. Another possibility is to use steam as the sweep gas. This is preferable since the hydrogen can easily be recovered by condensation. However, the possible effects of steam on the permeability of hydrogen through Pd-Ag membranes must first be determined. (iii) Effect of the Inlet Concentration of H2 on the Ammonia Decomposition. The major advantages of employing a membrane reactor with exclusive hydrogen permeability is depicted in Figure 7. The decomposition of NH3 to give Nz and H2 is equilibrium limited and therefore depends on the composition of the reactants in the initial mixture. This would either result in more ammonia formation or decomposition as indicated earlier. In Figure 7, for feed mixtures containing greater than 20 mol % hydrogen, the profiles but show an initial increase in the production of "3, as hydrogen permeation starts to occur, the equilibrium is driven toward more NH3 decomposition. For all the hydrogen concentrations considered (5-40 mol %), complete decomposition of NH3 was eventually achieved for the reactor length of 14 cm considered in this study. (iv) Effect of Space Time on the3" Decomposition. The effect of variation of the space time (r&, defined as the volume of the catalytic bed divided by the volumetric flow rate of the feed gas stream, on the decomposition of NH3 along the reactor axis is depicted in Figure 8. As the space time is increased, the membrane area required to achieve complete NH3 decomposition decreases. This is due to the fact that the contact time of the NH3 is increased. Since this system involves dilute concentrations of "3, the rate
N, = (PI - pJ(Sh)D,B/d,
(11) where p l and p2 are the molar densities (moVcm3) of hydrogen on the reaction and permeate sides, respectively, D m is the diffusivity of hydrogen in the reaction mixture (cm2/s),dt is the reactor diameter (cm), and Sh is the Sherwood number (k&D). The permeate side concentration of hydrogen was very small due to the presence of the sweep gas. Under the conditions of the present study, the flow in the reaction (tube) side was laminar and the Sherwood number is given by eq 12 (Bennett and Myers, 1982),where Re and Sc are the
(yy-J4
Sh = 1.86(Re)1'3(Sc)1'3
(12)
Reynolds and Schmidt numbers, p and p, are fluid viscosities at average bulk temperature and the wall, respectively, and Lo the length of reactor across which permeation of hydrogen occurred. Using this correlation for Sh and the flux eq 11, the hydrogen mass transfer flux from the bulk gas to the inner membrane surface was calculated and found t o be always greater than the permeation flux, and therefore mass transfer in the tube side was not limiting despite the fact that external rates are changing with feed velocity. Calculations were also made for the mass transfer of hydrogen when there was no chemical reaction. These calculations showed that the mass transfer flux was always greater than the permeation flux. This is further confirmed by the fact that all the hydrogen contained in the original synthesis feed gas stream could be removed by the membrane during such operations. The size of the catalyst particles used in this study was in the range 1.5 Id, I 2 mm. Particles in this size range are sufficiently small for the impact of intraparticle mass resistance t o be neglected. Estimates of the external mass transfer resistance were also made for this size of particle and the flow rates considered. Surface concentrations of ammonia were between 90 and 95% of the bulk gas values, and therefore external mass transfer was not a significant factor. In the development of the mathematical model to describe the gaseous flow behavior in the catalytic membrane reactor system, we assumed perfect plugflow. This was confirmed by making estimates of the
3782 Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 100
-7 N
'' 5
90
80
Tubeside Press.(Preoo. Ratio)
.
70 60 50
0
u 40 0
z
'E
30
E
2o
--
Countwcurnnt Mod. SPOC. Tim. o.owmin T i m p r o t u m 873K Sweep Ratio 0.057 '
0.0 0.1
L
0.2
,
j
"
,
0.3 0.4
j
,
8
'
8
0.5 0.6 L/Lo
,
8
0.7
'
8
-
-
'
0.8
0.9
1.0
Figure 9. Effect of operating pressure (pressure ratio) on ammonia conversion. Table 5. Membrane Reactor Performance at a Higher Shell-Side Pressure (Tube-Side Pressure = 36.0 bar; Shell-Side Pressure = 30.0 bar) ammonia conversion (%) sweep gas flow rate (cm3(STP)/min) cocurrent countercurrent 300
64
600
81
1000
89
90 99.8 100
Peclet number for mass transfer, Pe, in the empty shell side and packed tube, respectively. In the empty shell side, the Pe, value was greater than 40 for the range of flow rates considered, and therefore the perfect plugflow assumption is valid. For the packed tube side, eq 13 (Young and Finlayson, 1973) was used to check any deviation from perfect plug-flow behavior.
(13) For the range of ammonia concentrations and feed flow rate studied in this investigation, the left hand side was always far smaller than the Peclet number (assumed to have a value of 1-21, validating the plugflow assumption. (v)Effect of Operating Pressure. Figure 9 shows the effect of the operating pressure or more appropriately the pressure ratio; defined as the ratio of the shellside pressure to the tube-side pressure, on the ammonia decomposition. For the range of pressure ratios considered for this study (0.027-0.21, there was complete The present membrane condestruction of the "3. sisted of a 6 pm Pd-Ag film on a mesoporous glass tube having a 40-Apore size. This reactor has been tested experimentally for upstream pressures of up to 5 atm (Gobina, 1992). Such a system is therefore a feasible proposal for use, in, for example, an IGCC process that operates in the higher pressure ratio range. For lower pressure ratios (high upstream pressures) more robust substrates such as alumina ( y or a),which are capable of sustaining very high pressure differentials, would be preferred. We are currently experimenting on the application of such substrates for membrane reactor applications. For simulations carried out at a pressure drop of 35 bar (Le., 36 to 1bar) as indicated earlier, the pressure energy of the gas is being "wasted". The permeate hydrogen would be more valuable if it was recovered at pressure. Simulations have been carried out for pressure drops from 36 to 30 bar. These are presented in Table 5 using the base case data (Table 3). It is seen that any gain in hydrogen pressure is achieved at the
expense of a reduction in the ammonia conversion. Furthermore, higher decomposition rates can only be achieved for very high sweep rates. The effectiveness of the countercurrent mode of operation is more evident here due to the higher shell-side partial pressure. (d)Removal of H2S. The potential of Pd-Aglporous support composite membranes as viable candidates for the removal of NH3 from synthetic coal gas streams is unique due to its ability to withstand high temperatures. However, even very small amounts of H2S (-10 ppm) have the capability of poisoning the Pd-Ag film (Hurlbert and Konecny, 1961; Yoshida et al., 1983; Chabot et al., 1988). Therefore in order to apply such a system, an H2S removal stage must precede the composite membrane (Jarr et al., 1989). This might involve for example a molten salt liquid membrane prior to the Pd-Ag membrane reactor. Alternatively, Edlund and Pledger (1993) have proposed a method of utilizing Pd-based membranes in the presence of hydrogen sulfide. They utilized a vanadium foil (30 pm thick) which was sandwiched between palladium and platinum foils, both of which were 25 pm thick. The Pd foil is placed on the permeation side so as to prevent the vanadium from oxidation during start-up and shutdown operations. The platinum layer was on the feed side and served to provide chemical resistance to the hydrogen sulfide and functions also as a catalyst for hydrogen sulfide thermolysis. Complete conversion of the hydrogen sulfide was reported at test conditions as follows: residence time, 15 min; temperature, 700 "C; and pressure, 7.8 atm for a 1.5 vol % H2S system. In a more recent investigation, Robinson and Winnick (1993) have presented results of their study on the use of a self-fabricated zirconia membrane for the electrochemical removal of hydrogen sulfide from coal gasification streams. Such a system can clearly be incorporated into the gasification process. Possible modifications can also be accommodated in the present membrane system to handle H2S. Work in this area is proceeding.
Conclusions The potential application of a composite metalceramic membrane to the elimination of the dilute concentrations of NH3 contained in coal derived synthesis gas streams has been explored. A well established rate equation for ammonia decomposition on solid catalysts was applied in conjunction with experimentally determined permeation data for hydrogen gas through a Pd-Aglporous ceramic composite membrane. The results of this analysis have clearly shown that NH3 decomposition rates far higher than those attainable in conventional fured-bed operations and other gas cleanup systems are feasible at gasifier conditions. The countercurrent mode of operation was shown t o be superior at relatively lower temperature (i723 K) compared t o the cocurrent mode, but there was very little difference between these modes when no sweep gas was used. Also the sensitivity of the NH3 decomposition rate on the composition of the feed gas mixture was found to be an important factor in the operation of such a system, and since hydrogen was the only permeable species, such a membrane system is versatile and complete destruction of the NH3 could be achieved even at a relatively low upstream pressure of 5 atm. The use of such catalyst-membrane systems would be important for the prevention of global warming and acid rain, which at present have a significant environmental impact.
Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 3783
Acknowledgment We thank the EPSRC for their continued financial support for this work. We are also grateful to Dr. Dermot Monaghan for carrying out the magnetron deposition on the substrate and SEM photographs. Palladium film stoichiometry was carried out by Johnson Matthey Research Centre, Sonning Common, to whom we express our gratitude.
List of Symbols A = membrane area (cm2) co = ammonia concentration in the feed (moycm3) d, = pellet diameter (cm) dt = tube diameter (cm) E = activation energy (kJ/mol) Kes = equilibrium constant (atm) ko = preexponential factor (m~l/(cm~.s*atm-~.~~*) LO= reactor length (cm) PI, Pz = hydrogen partial pressure in the upstream and downstream side of the tube (atm) Pe, = Peclet number for mass transfer P, = component partial pressure (i = "3, Hz,Nz, and inerts) (atm) Qi = component permeation rate (mol/s) QO= permeation constant (mol*cd(cm2*s.atm1'2)) R = universal gas constant (kJ/(mol.K)) Q H ~ = reaction rate (moy(cm3*s)) T = temperature (K) us = feed-side velocity (cdmin) Greek Symbols
p
= exponential constant in rate equation 6 = membrane thickness (cm)
= reaction stoichiometry = component flow rate divided by ammonia flow rate at the tube inlet (i = "3, Hz, and Nz) p = fluid viscosity (cP) vi
@i
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