The Study of an Adiabatic Parallel Plate Methanation Reactor W. P. Haynes, R. R. Schehl,' J. K. Weber, and A. J. Forney Pittsburgh Energy Research Center, U.S. Energy Research and Development Administration, Pittsburgh, Pennsylvania 152 13
Pilot plant data are presented from a 1291-h test on an adiabatic catalytic methanation reactor employing recycled product gas for cooling. Topics discussed include conversion, product yields, catalyst properties, and deactivation rates. A mathematical model postulating an irreversible Langmuir-Hinshelwood mode of deactivation is presented and applied.
Introduction Virtually all of the processes currently under development for making synthetic natural gas from coal require a catalytic methanation step for upgrading the raw gas to a high Btu pipeline quality gas. The Energy Research and Development Administration, Pittsburgh Energy Research Center, is conducting bench-scale and pilot-plant scale studies of methanation with a variety of types of catalytic methanation reactors. This paper describes a 1291-h pilot plant test of an adiabatic parallel plate methanator. This type of reactor is frequently referred to as a hot-gas-recycle reactor, as large quantities of partially cooled product gas are circulated through the catalyst bed in order to remove the heat of reaction. Previous experiments with the hot-gas-recycle reactor have been reported (Forney e t al., 1965; Haynes et al., 1974). One of the purposes of this experiment (HGR-15) was to conipare the performance of a longer catalyst bed (8 ft) with previous experiments using shorter beds (2 and 5 ft). Reactor Description Experiment HGR-15 was performed in a pilot plant reactor constructed of type 304 stainless steel 3-in. schedule 40 pipe, 10 f t long, and flanged at each end. The catalyst bed consisted of grid assemblies of parallel type 304 stainless steel plates flame-sprayed with Raney nickel. Each grid assembly was 6 in. long, contained 15 plates, and was shaped to conform to the inside diameter of the reactor. Pertinent catalyst bed data are presented in Table I. As indicated in Figure 1,adjacent grid assemblies were rotated 90" with respect to each other. Prior to assembly, individual plates were sand-blasted on both sides with an iron-free grit and then flame-sprayed with a 0.007-in. thickness coat of bonding material. Subsequent to the bond coat, Raney nickel alloy powder (80-200 mesh) was flamesprayed onto the surface to a desired thickness of 0.022 in. The grids were assembled, placed in the reactor to form an 8-ft bed, and then activated by passing a 2 wt % solution of NaOH through the reactor until 70% of the aluminum in the Raney alloy was reacted. The extent of reaction was determined by measuring the quantity of hydrogen evolved according to the amount of 3 mol of H:! for every 2 mol of A1 reacted. After activation, the catalyst was washed with distilled water until the effluent water attained a pH of 7.2. The reactor was maintained under hydrogen until the system was brought to synthesis conditions (300 "C and 300 psig) before synthesis feed gas was introduced. A simplified flowsheet of the hot-gas-recycle pilot plant is shown in Figure 2. Only the primary reactor, consisting of the parallel plate grid assemblies, is considered in this report. The second stage reactor, an adiabatic reactor charged with a precipitated nickel catalyst, was operated only during the later periods of' the run. Additional heat exchangers, not illustrated, were used to compensate for system heat losses, to achieve a
measure of heat recuperation, and to control the gas temperature into both the hot gas compressor and primary reactor. Cooling of the catalyst bed is achieved by direct transfer of the heat of reaction to the slightly cooler gas stream flowing through the bed. The hot recycle stream is cooled by heat exchange with the reactor synthesis gas, and a portion of the product gas is sufficiently cooled to condense out the water vapor for recycle to the synthesis gas stream. The mixture of the cooled recycled gas and the fresh feed gas constitute a feed to the hot gas recycle reactor a t a controlled temperature that may be 50-150 "C lower than the reactor outlet temperature, depending upon the total amount of gas recycled and the extent of heat exchange. Operating Procedures a n d Results The operating conditions for run HGR-15 were similar to those for run HGR-14 (Haynes et al., 1974) (HGR-14 consisted of a 2-ft bed of catalytic grids) with the exceptions that the fresh feed space velocity remained constant a t 2000 SCFH/ft3 (in the final periods of the experiment the space velocity was decreased to 1500 and then to 1000 SCFH/ft:' and the cold gas recycle (CGR) to fresh feed ratio was maintained consLant a t 3. Space velocity calculations were based on an empty reactor volume. The operating parameters studied and the product gas characteristics are presented in Figures 3 and 4, respectively. Note: the hot-gas-recycle stream flow rate was calculated. Small errors in gas analyses produced relatively large fluctuations in total recycle flow rates. Performance of run HGR-15 compared unfavorably with run HGR-14 as the catalyst life was shorter (1291 h compared to 2307 h for HGR-14) and the initial product CO concentration was higher (0.09%compared to 0.01%).Five shutdowns occurred during the run a t 22, 109, 207, 494, and 811 h on stream due to compressor and heat exchanger repair. The spent catalyst was kept in a hydrogen atmosphere during cooling, depressurizing, and while remaining in stand-by conditions. A t 955 h the temperature rise across the reactor was increased to 125 "C. The CO conversion temporarily increased during the first 48 h, but then continued to decrease a t the same rate as the A T = 100 "C operating condition. At 1150 h AT = 100 "C was reinstated, resulting in a further decrease in conversion. Since the overall catalyst activity was already declining rapidly, the effect of increased temperature difference is difficult to assess. The product CO concentration a t the end of operation at 2000 SCFH/ft' fresh feed gas space velocity was 1.10 ~ 0 1 %The . total CH3 production per pound of catalyst was 15.0 MSCF/lb. Selected tabulated periods in run HGR-15 are presented in Table 11. Period 2 represented performance a t 2000 SCFH/ft:' fresh feed gas space velocity and 3:l cold recycle ratio when the catalyst is fresh. Periods 33 and 34 provide a comparison of prior to and immediately following a change Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
113
5)-
450
'Exp HGR - I5 Moxlmum temverotures
v)
400
k-g
W
250
I d - 1
- 3'
--I
I--
\I
Fresh feed qos rote
p i p e sch 40
Pressure : 300 psiq Fresh feed spoce veloc~tyfor 821 scfh : Z O O 0 hr.' Bed helghl : 80 ft
304 S S
Toto1 recycle rotio
Catalyst
Thermocouples
-
grids
+ - - -
~I
Cold recyrle rotio
J -
&
200
I
600
800
IO00
I200
Figure 3. Reactor conditions.
1 3 3
g
12
w
20-
s
-
Exp HGR -15
-
08
zl-5
Gas outlet
400
TIME ON STREAM, hours
Carbon monoxide in product gas
:?
no gg:
04
Ea n x5 0
Figure 1. Hot-gas-recycle methanation reactor.
,
0
Main
reoctor
L.
I U compressor
I
TIME ON STREAM, hours
k l
Figure 4. Product gas characteristics.
Hot recycle
1
(Cold recycle
Condenser
J p
l
Table I. Catalvst Bed Data for HGR-15
1
1
Condenser
Gas
Type catalyst
High
b J
-Btu gas Second stage reoctor
Woter
L - J
Figure 2. Flowsheet of hot-gas-recycle process.
in catalyst temperature spread to 125 "C AT. Periods 41 and 42 represent the conclusion of 125 "C AT and the beginning of 100 " C AT operating conditions. Period 44 represents operation a t 1500 SCFH/ft" fresh feed gas space velocity a t 3:l CGR and 7:l HGR ratios. Period 45 represents operation a t 114
Methane produced
J
I-----I
l
9001
Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
Wt oh nickel oh Activated Plate thickness, in. Space between bare plates, in. Average bond coat thickness, in. Average catalyst thickness, in. Bed diameter X length, in. Bed volume, f t j Weight of unactivated catalyst, lb. Superficial area of catalyst, ft' Void fraction
Flame-spray Raney nickel 42" 70 0.048 0.135
0.007 0.022
3.07 X 96 0.41 15.47 45.64 0.500
Before leaching. Leaching of aluminum stopped when 70% of the theoretical amount of hydrogen had evolved. (1
Table 11. Experiment HGR -15-Selected
Test Data Period no.
Fresh gas: Rate, scfh HZ, VOI O h co, vol% CO?, vol % Nj, VOI O h CH~,vol% Hn/CO Exposure velocity, scfh/ftz Space velocity, hk' Mixed feed gas (wet): Rate, scfh Hy, VOI % co,VOlO?O COZ, vol % Nn,VOI % CHi,vol O h H20, VOI O h HnICO Inlet superficial vel., ft/s Inlet Reynolds no. Exposure velocity, scfh/ft2 Space velocity, h-I Volume total recycle/volume fresh gas Volume cold recycle/volume fresh gas Temperatures: Gas inlet, "C Maximum catalyst, "C Pressure, psig Product gas: Rate, scfh Hs, VOI % co, vol% cos, vol 0% N?,VOI O h CHI, VOI 96 H20, ~ o%l H?/CO Conversion: H2,OO/ fresh feed CO, % fresh feed (HZ + CO), % fresh feed H2, % mixed feed CO, YO mixed feed (H?+ CO), % mixed feed Usage ratio Heating value, Btu/scf Carbon recovery, %
2
33
34
41 42 Hours on stream
91
955
979
810 75.1 24.5 0.1 0.3
821 75.3 23.8
820 74.6 24.5 0.1 0.7 0.1 3.0 18.0 2013
44
45
1147
1171
1219
1243
822 75.1 24.5 0.2 0.2 3.1 18.0 2015
823 75.3 24.2 0.1 0.3 0.1 3.1 18.0 2017
617 74.7 24.2 0.2 0.8 0.1 3.1 13.5 1512
411 75.5 23.7 0.2 0.5 0.1 3.2 9.0 1008
3.0
7600 14.0 3.0 0.5 0.8 77.7 4.0 4.7 4.0 3440 166 18600 8.5 3.0
8700 13.4 3.O 0.7 1.1 77.6 4.2 4.5 4.6 3920 190 21250 9.8 3.0
10150 13.7 2.7 0.7 0.8 77.5 4.6 4.9 5.4 4580 222 24900 11.6 3.0
6780 13.0 2.8 0.6 0.9 78.4 4.3 4.6 3.6 3060 149 16620 10.2 3.0
4470 12.5 2.4 0.5 0.3 79.9 4.4 5.1 2.4 2040 98 10950 10.1 3.1
302 395 300
301 400 300
302 423 300
301 425 300
300 400 300
300 399 300
302 400 300
206.0 4.0 .03 1.17 3.9 90.7 0.2 133.0
212.3 8.2 0.6 0.5 89.5 0.2 13.7
208.6 6.9 0.4 0.6 0.9 91.0 0.2 16.9
210.9 7.2 0.7 0.8 1.2 89.9 0.2 9.6
211.7 8.6 0.9 0.8 0.9 88.6 0.2 9.3
156.4 7.1 0.7 0.6 1.0 90.4 0.2 10.4
103.8 6.4 0.3 0.5 0.3 92.3 0.2 21.4
98.7
97.6 99.5 98.1 45.6 78.8 50.8 3.1 937 100
98.0 99.7 98.4 56.3 88.0 61.8 3.0 947 96.7
98.0 99.4 98.3 52.0 77.6 56.7 3.0 937 97.0
91.6 99.2 97.9 43.4 70.0 47.9 3.0 931 97.3
98.3 99.5 98.6 51.2 78.3 56.0 3.0 943 96.0
98.8 99.8 99.0 54.4 89.1 60.1 3.1 959 96.5
0
3.1 17.8 1990 12000 8.5 1.7 1.0 3.5 79.2 6.1 5.1 6.4 5660 263 29400 14.0 2.5
100
99.0 58.4 98.4 65.0 3.0 934 95.0
1000 SCFH/ft? fresh feed gas space velocity a t 3:l CGR and 7:l HGR ratios. The analyses of catalyst plate scrapings are shown in Table 111. It is apparent from these analyses that sulfur poisoning and fouling due to carbon deposition during run HGR-15 increased over that during HGR-14. Several items concerning these data should be emphasized. (1) Increased sulfur poisoning in HGR-15 existed; 0.26% sulfur compared to 0.16% sulfur in HGR-14 on the reactor inlet plates. (2) Increased carbon deposition in HGR-15; 4.5%compared to 3.25% on same inlet plates and 0.9% to 0.61% on the exit plates. Carbon was measured by the standard ASTM combustion technique and, hence, it was not possible to distinguish carbidic carbon from amorphous carbon. X-ray analysis, of course, indicated the presence of appreciable carbide.
0.1
0.7 0.1 3.2 18.0 1986 9750 13.5 2.5 0.5
0.9 78.0 4.6 5.4 5.2 4400 214 23900 11.1
1.0
0
(3) A consistent decreased amount of nickel in the HGR-15 samples, Le., an average of 57.33% on the HGR-15 plates compared to an average of 78.05% in HGR-14. This would indicate incomplete leaching of aluminum. (4) X-ray analysis of the spent catalyst revealed metallic nickel and nickel carbide, Ni&, in the catalyst near the gas inlet and only metallic nickel near the gas outlet, indicative of carbiding occurring during both runs. (5) An increased percentage of nickel on the bottom plate of HGR-15; 64.4% compared to 52.2% on top plate, indicating nonuniform leaching. (6) Plate thicknesses of the top and bottom assemblies were measured after removal of the activated portion of the catalyst. The average plate thickness of the bottom plates was 0.078 in. and the top plates 0.091 in. This is further evidence of nonuniform leaching occurring more extensively a t the Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
115
Table 111. Properties of Spent Raney Nickel Catalyst Run 14"
Table V. Surface Areas, Pore Volume, and Pore Radii of Spent Raney Nickel Catalyst
Run 15 HGR-14
Reactor Reactor Reactor outlet Reactor outlet inlet (botinlet Reactor (bot(top) tom) (top) middle tom) X-ray analysis Chemical analysis,
Ni, Ni:jC
Ni
Ni, Ni:IC Ni, Nil$
Ni A1
78.45 8.41
77.65 9.43
52.60
C Fe Na
3.25 0.88
0.61 0.96
-
-
S
0.16
0.07
4.50 0.24 0.16 0.26
-
54.70 -
Ni 64.70 -
Oh
0.90 0.27 0.12
3.80 0.25 0.12 0.16
0.003
Average of two samples Table IV. Average Plate Thickness Data of HGR-15 after Removal of Spent Catalyst (in.) Top plate assemblies
Bottom plate assemblies
End
Middle
End
End
Middle
End
0.091
0.085
0.098
0.081
0.073
0.081
reactor bottom. Average plate thickness data are represented in Table IV. Experiments subsequent to HGR-15 have utilized recirculation of the caustic solution in order to achieve more uniform leaching of the Raney nickel catalyst. Results of this technique are not yet available. The spent catalyst was examined for surface area and pore volume characteristics. These data are tabulated in Table V. The higher surface area a t the inlet of run HGR-15 is very likely a result of the larger percentage of carbon deposits on these plates. Metal surface area measurements were not performed on the spent catalyst samples. Measurements made on samples from previous runs indicated that the nickel metal sites were fewer a t the reactor inlet even though the BET surface area was greater a t the inlet than a t the exit. Table VI lists superficial deactivation rates for the catalyst bed in run HGR-15. The superficial deactivation is given by the change in product CO percent per mscf of CH4 produced per pound of catalyst in the bed. It will be noticed that there is a general increase in superficial deactivation rate with time on stream. This is precisely the behavior a "zonal burn-out'' type catalyst bed deactivation would exhibit. Application of a Mathematical Model t o R u n HGR-15 A recent publication (Ralston et al., 1974) presented a one-dimensional mathematical model along with some applications of a tube-wall methanation reactor to be utilized in the ERDA Synthane process. This model made no provisions for deactivation of the catalyst. The model has been improved by taking into account catalyst deactivation with increasing time on stream and to extend the model to applications in hot-gas-recycle methanation. The mechanism responsible for deactivation of flamesprayed Raney nickel catalyst is not thoroughly understood a t the present time. I t is possible that deactivation is caused by several processes acting simultaneously, such as sintering, poisoning by trace amounts of sulfur in the feed stream, and trace amounts of iron carbonyl depositing on the active sites which subsequently act as a center for carbon deposition. The model presumes the following. (1) The catalyst is poisoned via some mechanism independent of that of the main reaction, that is, the feed contains 116
Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
Inlet
Plates Surface area, m,'/g Pore vol. cm:'/g Average pore radius, 8, Percent pore volume with the following radii 60 Percent total surface area with the following radii (A) 60
Outlet
HGR-15 Inlet
Outlet
30.95 29.63 45.10 29.04 0.077 0.094 0.128 0.094 49.52 63.35 56.67 65.06
(A) 16.09 15.59 13.16 45.11
5.59 13.43 11.65 8.70 60.63
13.46 10.16 9.60 9.78 57.00
9.59 8.84 5.53 9.12 66.92
32.16 22.18 14.59 9.09 21.98
13.11 24.30 16.53 10.08 35.98
30.38 16.66 12.09 10.10 30.77
24.13 16.63 7.88 10.80 40.56
10.05
Table VI. SuDerficial Catalvst Deactivation Rate Data Nominal fresh feed space velocity, h-' From 2000 2000 2000 2000
2000
Stream period, h To
207 (to 3rd shutdown) 494 (to 4th shutdown) 931 (to AT = 125 "C operating conditions 1003 1150 (AT = 125 "C operating conditions) 1150 1196 (return to AT = 0
207 494
100 " c
Catalyst deactivation rate, %I MSCF/lb of catalyst 3.7 x lo-' 21.9 X lo-' 87.7 x 10-j 147.5 X lo-' 569.1 X lo-'
operation) a trace amount of some unspecified poison. The reactions may be written as k
CO + 3H2 e CH4 + H20 (main reaction, fast) k,
P --+W (poisoning reaction, slow) I t is assumed that the poison reaction is much slower than the main reaction. Consequently, the main reaction may be assumed to approach steady-statebehavior over a short period of time. This assumption has the effect of decoupling the equations associated with the main reaction model from those associated with the poisoning model. (2) Since there is little experimental evidence with regard to the poisoning mechanism, an irreversible Langmuir--Hinshelwood type of expression for the rate of formation of absorbed poison is assumed. The kinetic rate expression for the conversion of CO to CH4 was taken to be that proposed by Lee et al. (1970,1973) which provides a reasonable fit to data reported by IGI'. Lee's rate expression developed from supported nickel catalysts is of the form
360 b / h15
350
43O-420 -
---
---1 - -
-1
--
410
/!
400 390 -
380 370 360 350 -
J.
443 h r s
340 330 310 310
-
300c
I
---'I
I
1005 h r s
1
0
02
04
06
F R A C T I O N A L DISTANCE 'HRU
Figure 5 . (;as temperature profiles for different times on stream; model (-);
In the present analysis, this rate equation was expressed in terms of the molar concentrations of the respective species.
c
The symbols signify that these are concentrations a t the surface of the catalyst as opposed to concentrations in the bulk stream of the gas. Note that the concentration of methane at the catalyst surface is assumed to be the same as the bulk methane concentration. Due to the relatively large quantity of product gas that is recycled through the catalyst bed, methane is the major constituent in the gas a t every point in the reactor. I t is assumed that negligible error is introduced in the kinetic rate by ignoring the methane concentration gradient between the bulk gas and the catalyst surface. A t steady-state conditions there is no net change in reactant concentrations a t the catalyst surface; thus, the rate of reaction must equal the rate a t which CO and Hz diffuse to the surface. The mass transfer rates are given by -r = kc(Cro -
e,,)
/
t
(34
where kc and kc^ are the film mass transfer coefficients. The mass transfer coefficients are calculated, for a given Reynold's number, from the standard j-factor correlation (Treybal, 1955; Welty, et al., 1969). In order to calculate the global rate of reaction at any point i n the reactor, eq 2,3a, and 3h must be solved simultaneously. The method of finite differences is utilized to calculate local bulk ( oncentrations and temperature through the reactor.
08
I
0
0.2
0.4
0.6
0.8
I
REACTOR
experimental ( 0 ) .
Since the reactor is operating adiabatically with a temperature rise through the catalyst bed on the order of 100 "C, there could be an appreciable temperature dependence in the global rate expression which is accounted for through the Arrhenius expression. Knowledge of the catalyst temperature is required to evaluate this expression; hence we write I'
= h ( T - T')
(4)
and
dT' r = C p M -aA h is the film heat transfer coefficient and is evaluated in a manner similar to the mass transfer coefficient. Equation 5 is used to calculate the gas temperature as a function of distance through the reactor and, given T', the catalyst temperature may be estimated from eq 4. The poisoning mechanism assumed in this paper is that of the irreversible adsorption of compounds other than the principle reactants and products on the catalytically active sites. This is commonly referred to as Type I, or independent poisoning. In the absence of any experimental evidence to the contrary, it is reasonable to assume an irreversible Langmuir--Hinshelwoodtype of expression for the rate of formation of adsorbed poison rp
= kpCn(1 - u/q)
(6)
In dimensionless form, eq 6 and a differential mass balance of the poison species across an element of the catalyst surface become, respectively Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
117
i7*
-\
95 h r s 4
837hrs
443hrs
\
\
0.2
0.4
C6
00
1
0
I 0.2
\
1
;..-I‘ I
0.4
0.6
,-0.0
o
I
0.2
0.8
0.6
0.4
I
F R A C T I O N A L DISTANCE T H R U REACTOR
Figure 6. CO concentration profiles for different times on stream: model (-); experimental ( 0 ) .
410
m
I
I
I
I
I
I
. .. 2
340
4
330 320
2
300 -1ME
Figure 7. CO concentration in the product stream as a function of time on stream.
and
Equations 7 and 8 must be solved simultaneously for $ and
4 as functions of 0 and a. If the catalyst bed is initially free of adsorbed poison and the concentration of poison precursor in the feed remains constant with time, then the following boundary conditions are applicable
$40, a ) = 0
d e ,0)= 1
(9)
The analytical solution of this system is of the form
+(e,
a ) = (1 - e-”)/(l
+ epH(eNDn- 1))
and the function $(e, a ) may be directly related to the catalyst activity a t any point (0, CY) as follows
k = kinitial(1 - $1 provided that the initial catalyst activity is uniform over the entire catalyst bed. $(e, a ) contains two parameters which must be determined from experiment as there are insufficient independent data to predict them a priori. Let these parameters be called p1 and p2 and be defined by pi = %/t= kpCDO/d
p~ = N D d A = kp/V 118
Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
0 0
ON S T R E A M . h o u r s
02
04
06
08
IO
FRACTIONAL DISTANCE THROUGH REACTOR
Figure 8. Comparison of calculated catalyst temperature (- -) with experimental ( 0 )and calculated (-) gas temperatures at 347 h on stream. Relative catalyst activity ( - - - - -).
The ratio p I / p 2 is a measure of the “velocity” of the deactivation front as it traverses the catalyst bed, while p2 is proportional to the slope of the activity profile a t its point of inflection. The model outlined above has essentially three parameters which must be obtained from experimental data: kinitial, p1, and p2. Other parameters such as physical properties of the gas, transfer coefficients, activation energy, and heat of reaction may be estimated from data published in the literature. Ideally, kinitial can be determined from data obtained early in the run when deactivation of the catalyst bed is negligible. Similarly, the ratio p J p 2 may be estimated by measuring the “velocity” of the global reaction zone as it moves down the reactor. This leaves only p 1 or p2 to be fitted to experimental temperature profile data. Unfortunately, there were numerous unscheduled shutdowns during run HGR-15. During a shutdown, the catalyst is maintained in a 50 to 100 psig hydrogen environment which induces a partial recovery of the catalyst activity. This adds another dimension, not easily accountable, to the poisoning model. Another complication was the apparent greater deactivation rate during the first 200 h of the run. At 207 h on stream the hot-gas-recycle compressor was overhauled. Prior to this time it is possible that the catalyst was exposed to higher concentrations of oil vapor from this compressor. A fit of the model to experimental data, using an activation energy of 1.25 X 104 Btullb-mol (Lee, 1973) yielded the values k N 1.1 X lo6 ft5I2/h-lb p1 = 0.0114 (0-207 h on stream), p 1 N 0.00625 h-l (times succeeding 207 h), and P2
N 0.2 ft-'. Typical comparisons of the model predictions with experimental data are shown in Figures 5 and 6. Figure 5 illustrates the fit of calculated gas temperature profiles to experimentally measured temperatures for several different times on stream. Figure 6 compares the calculated and observed CO concentrations as functions of distance through the reactor and time on stream. Figure 7 demonstrates the simulation of product CO concentrations for the first 1150 h. Agreement between model and experiment is seen to be remarkably good. A t 955 h the temperature rise through the catalyst bed was increased to 125 "C. Figure 8 illustrates a comparison of calculated catalyst temperature with experimental and calculated gas temperatures a t 347 h on stream. The maximum calculated temperature difference between catalyst and gas is approximately 20 "C for this particular time on stream. The relative catalyst activity is also plotted t o indicate the extent of catalyst deactivation a t this time.
Conclusions Because of the increased number of shutdowns and unsatisfactory leaching, an equitable comparison of runs HGR-15 and HGR-14 cannot be made. Another significant difference between the runs was the flow condition of the gas stream. HGR-14 was operated in laminar flow ( N R =~ 1200-1800) whereas HGR-15 was turbulent ( N K=~ 4000-5000). A simple one dimensional mathematical model with an empirical catalyst deactivation mechanism is found capable of describing reactor performance. Additional experiments are scheduled to be performed in a 2 in. X 14-ft reactor capable of going to pressures as high as 1000 psig. Such experiments shall test the ability of the model to predict reactor performance under dimensional scale-up and different operating conditions such as total system pressuure and space velocity. Acknowledgments We acknowledge with thanks J. J. Elliott and H. Pennline for their work in supervising the operation of the pilot plant methanation reactors. Nomenclature A = superficial catalyst surface area, ft2 A , , = total superficial catalyst surface area, ft2
C = concentration in bulk gas, lb-mol/ft3 (3 =
concentration a t catalyst surface, lb-mol/ft3
Cr, = poison concentration in gas, lb-mol/ft? CI," = poison concentration in feed, lb-mol/ft? Cp = specific heat of gas, Btu/lb-mol O F E = activation energy, Btu/lb-mol h = heat transfer coefficient, Btu/h ft2 O F k , k l = preexponential factor kn, h,i = equilibrium constants k p = poisoning rate constant, ft/h k c = mass transfer coefficient, lb-mol/h ft2 concentration difference
M = gas molar flow rate, lb-mol/h NI,= Aohp/V, dimensionless constant P = partial pressure, psi R = gas constant, Btu/lb-mol OR r = rate of CO conversion, lb-mol/h ft2 of catalyst r 1 = rate of methane formation, lb-mol/h g of catalyst r p = rate of formation of adsorbed poison, lb-mol/h ft2 of catalyst t = real time, h T = catalyst temperature, OR T' = gas temperature, OR V = volumetric flow rate of gas through reactor, ft"h N = A/Ao, dimensionless superficial catalyst surface area 0 = kpCDOt/uT, dimensionless time u = poison concentration on catalyst surface, lb-mol/ft2 cr'r = value of u corresponding to complete deactivation, lbmol/ft2 4 = C~JCDO,dimensionless concentration $ = U / U T , dimensionless concentration L i t e r a t u r e Cited Forney, A. J., Demski, R. J., Bienstock, D.,Field, J. H.. U.S. Bur. Mines Rept. Invest., 6609,(1965). Haynes, W. P., Forney, A. J., Elliott, J. J., Pennline, H. W., Am. Chem. SOC.Div. Fuel Chem. Prepr., 19 (3), 10 (1974). Lee, A. L., "Methanation for Coal Gasification", Clean Fuel fw Coal Symposium, Chicago, Ill., Sept 1973. Lee, A . L., Feldkirchner, H. L.,Tajbl, D. J., Am. Chem. SOC.Div. FueIChem. Preor.. 14 (4). 126 (19701. Ralston, T. D.: Haynes: W. P., Forney, A . J., Schehl, R. R., Bur. Mines Rept. lnvest., 7941 (1974). Treybal, R. E., "Mass Transfer Operations", p 55, McGraw-Hill Book Co., New York, N.Y., 1955. Welty, J. R., Wicks, C. E., Wilson, R. E., "Fundamentals of Momentum, Heat, and Mass Transfer", p 580, Wiley, New York, N.Y., 1969.
Received f o r reuieu: February 23,1976 Accepted J u l y 29,1976
Ind. Eng. Chem., Process Des. Dev., Vol. 16, No. 1, 1977
119
