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maintains that fact throughout the calculation. Actually, after each column solution, the distribution of all components can be checked by infinite section equations utilizing the conditions a t the top and bottom of the calculated column, and the distribution of each component can be easily found. This serves either to verify that a component is not distributing or to calculate the amount in each product if it does distribute. Also, it is apparent that other specifications of the light and heavy keys could be used, for example, the concentration of a key in a product, or even a total product amount or a product boiling point. However, setting the amounts of the two key components in the products is the most common way of specifying the separation. The number of stages to be inserted between the feed stage and the corresponding infinite section in each expansion of the column is arbitrary. If the nondistributing diluents are considerably different from the corresponding key components, only a few plates are necessary. However, since many distillation systems are fairly close-boiling, there are often components which differ in volatility from a key component by only a small amount, perhaps l o % , and which still do not distribute. To reduce the amount of these components at the ends of the calculated column to amounts which do not affect the reboiler and condenser loads usually requires 10 to 15 stages. If the number of stages added in each column expansion is small, many columns are calculated unnecessarily. The authors have found it reasonable to add five stages at a time. Results The calculational procedure has been tested on many examples and has been found to be extremely stable,
yielding a solution to each column in a few iterations. The method has been used on systems in which the equilibrium constants were independent of liquid composition, but it could be readily extended to include such effects. One interesting result of the test calculations, which were done on hydrocarbon systems, was the finding that Underwood's method was usually reasonably correct. If the minimum sectional flows given by Underwood are assumed to be flows at the respective pinch points, and energy balances are used to obtain flows and loads a t the top and bottom of the column, the error was normally only a few percent. However, the calculation by Newton-Raphson is sufficiently rapid that there is no reason not to use the correct calculation. A typical close-boiling four-component system which required solution of a 14-stage column, a 24-stage column, and a 34-stage column was solved in 12.5 s on a CDC 6400 computer. A six-component system which required a 14-stage calculation and a 24-stage calculation required 16.9 s. The authors have written a Fortran program for single-feed two-product columns for systems up to ten components. The program is available a t no charge on request. Literature Cited Bachelor, J. B., Pet. Refiner, 36 (6), 161 (1957). Erbar, R. C., Maddox, R. N., Can. J . Chem. Eng., 40, 25 (1962). Naphtali, L. M., Sandholrn, D. P., AIChE J . , 17, 148 (1971). Newman, J., Ind. Eng. Chem. Fundam., 7, 514 (1968). Ricker, N. L., Grens, E. A. 11, AIChE J . , 20, 238 (1974). Shiras, R. N., Hanson, D. N., Gibson, C. H., Ind. Eng. Chem., 42, 871 (1950). Underwood, A . J. V., J . Inst. Pet., 32, 614 (1946). Underwood, A . J. V., Chem. Eng. Prog., 44, 603 (1948).
Received for reuiew February 9, 1978 Accepted August 7 : 1978
Operation of a Tube Wall Methanation Reactor Henry W. Pennline," Richard R. Schehl, and William
P. Haynes
Process Engineering Division, Pittsburgh Energy Technology Center, U.S.Department of Energy, Pittsburgh, Pennsylvania 152 13
The design and operation of a PDU-scale tube wall methanation system are discussed. The tube wall reactor was constructed from a stainless steel pipe with a surrounding jacket. The inside surface of the tube was flame sprayed with Raney nickel catalyst for 14 ft of length. The catalyst coating was activated by leaching with a caustic solution. A liquid coolant in the outer shell of the reactor removed the heat of methanation and maintained the system at near isothermal conditions. Exposure velocity and recycle ratio were varied during the run. Concentration and temperature profiles were periodically measured along the length of the catalyst bed. Catalyst performance, methane production, deactivation, and other results are presented for the test which was terminated after 1179 h on stream.
Introduction An ever increasing demand for natural gas coupled with growing supply shortages has renewed interest in producing synthetic natural gas from coal. The Pittsburgh Energy Technology Center of the 17,s.Department of Energy has developed the SYNTHANE Process. This process, as well as most gasification schemes, uses a final methanation step to convert a synthesis gas from the gasification of coal into a high-Btu pipeline quality gas. It is the objective of the process development unit (PDU) methanation program to develop efficient catalytic reactor systems with emphasis on the development of improved This article
methanation reactor systems and on the enhancement of catalyst performance. There are three basic reactor systems currently being developed at PETC (Haynes et al., 1972,1974,1977; Schehl et al., 1976). All employ a Raney nickel catalyst because of its high activity and selectivity toward methane, and all have very small pressure drops across the catalyst bed. However, the major difference in the designs is in the novel method of heat removal. In the tube wall reactor system, Raney nickel is thermally sprayed on the inside of a pipe. This reactor is surrounded by a jacket containing a coolant which removes the exothermic heat of reaction and
not subject to U.S. Copyright. Published 1978 by the American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 18,No. 1, 1979
157
condenser
T C wells
-
Reactant pas
Dowlhcrm reservoir
I
4 ” s C h 40 plpe dowlherm jacket
cootea internally with Raney nickel
L - 7
Product par
Figure 1. Tube wall methanation reactor.
maintains the catalyst a t near isothermal conditions. In the hot gas recycle system, sufficient quantities of product gas are recycled over a fixed bed of catalyst-sprayed plates to remove the heat. In the hybrid reactor, catalyst inserts are charged into an uncoated tube wall reactor. The heat of reaction is removed from the bed by both the recycle of product gas, as in the hot gas recycle system, and by boiling Dowtherm in the reactor jacket, as in the tube wall reactor. The purpose of this report is to present the investigation of one type of process development unit-the tube wall methanator. Reactor Description The reactor, shown in Figure 1, consisted of a 2-in. schedule 40 pipe flanged on the ends and surrounded by a 4-in. schedule 40 jacket. All reactor material was 304 stainless steel. Raney nickel catalyst (58% aluminum, 42% nickel) was flame sprayed onto the inner 2-in. pipe wall for 14 ft of length. Dowtherm, a liquid coolant with high boiling points a t low pressures, was added to the annular shell to remove the exothermic heat of reaction during the run. A reservoir was adjacent to the vertical reactor and was used as part of the Dowtherm vapor-liquid recycling system. The temperature of the saturated coolant was regulated by controlling the electrical resistance heating and the pressure of the cooling system. Nucleate boiling took place on the outer surface of the reactor pipe and thereby provided a natural convective circulation of Dowtherm. Both the reactor and reservoir were insulated, The Dowtherm vapor was condensed by cooling water or air convection and returned to the reservoir. In a commercial plant the heat removed by the Dowtherm would be used to produce process steam. Probes could be inserted into the reactor through the top and bottom flanged reactor heads. Four thermowells with thermocouples were inserted to get representative catalyst temperatures; the sample ports were machined so that the thermowells were against the catalyst surface. Thermocouples were positioned to obtain temperatures at 12 different locations along the 14 ft of catalyst length. Gas concentration profiles were taken by using four stationary stainless steel tube probes which sampled gases a t the 2, 6,9, and 12 f t levels down from the catalyst top (reactant gas inlet).
t
Reactor
I lHeliym
Used t u ~ s t i c tokeoff
LL
U
Figure 2. Catalyst activation assembly.
The reactor had been used twice before. The inner tube surface was prepared for catalyst coating by grit-blasting with compressed air and iron-free virgin aluminum oxide pellets until any trace of the previous spent catalyst coating was removed. In a new reactor, a bond coat of nickel aluminide (80% nickel, 20% aluminum) is first deposited to about a 6-mil thickness to prevent the catalyst coat from spalling from the roughened stainless steel surface. Since this dense bond coat was applied to this particular reactor before and the grit-blasting did not remove it, there was no need for reapplication. The internal flame spraying of the Raney nickel catalyst was done with a wire-fed metallizing gun using a hydrogen-oxygen flame. The Raney nickel “wire” is specially made for the gun and is a 1 to 4 mixture of a polypropylene/polyethylene binder to a 80-200 mesh size of powdered catalyst. In the past, the reactor jacket was cut before the 2-in. pipe was internally flame sprayed because of expansion stresses. However, for this particular time, water near its boiling point was circulated through the jacket while the internal flame spraying was in progress. No evidence of reactor distortion was detected. Respraying the tubes in the SYNTHANE prototype plant would also be a simplified process using the same technique. For run 16, a reactor length of 14 f t was sprayed to approximately a 25-mil catalyst thickness which covered 7.4 ft2 of reactor surface area. The spray rate was 1.04 linear feet of reactor per hour. From previous sprayed pieces of pipe, a laydown figure of 100 g/ft2 pipe area was found, and the total catalyst weight for this run was 740 g. Battelle Columbus Labs has produced a more efficient spray gun which can deposit catalyst at a rate of 20 linear feet of reactor per hour. After spraying, the reactor was vertically placed in the carbon steel system and leached in situ, Since catalytic activity is directly related to surface area, a very porous layer of nickel is desired. This is accomplished by reacting the aluminum in the Raney nickel alloy with a 2 wt 5% solution of ACS-pure sodium hydroxide. The gravity-fed system is shown in Figure 2. Before the caustic flow was started, the reactor was filled with deionized water. The extent of activation was determined by metering the amount of hydrogen evolved according to three moles of hydrogen for every two moles of aluminum reacted. The reaction was stopped after 4.25 h when 70.9% of the theoretical amount of aluminum in the Raney alloy was
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Table I. Catalyst Bed Data for TWR-Run 16 catalyst type weight percent percent activated estimated catalyst thickness, in, reactor diameter x length, in. void reactor volume, ft3 weight of unactivated catalyst, lb geometric area of catalyst, ft2 effective diameter, ft a
Before leaching.
flame sprayed Raney nickel 42a 70.gb 0.025 2.067 x 168
0.311
1.63
7.39
0.168
Theoretical. GOI
prehealer
Fresh 90s
n
Reoclor
comprersar
Condenser
Woler
w
Product
90s
Figure 3. Flow scheme of the tube wall reactor system.
reacted. Temperature of the leaching reaction never exceeded 50 "C. Previous experience indicated that the unleached Raney nickel substrate acted as an adhering agent between the leached catalyst and the bond coat. When leaching was finished, the reactor was drained under a helium atmosphere, and a continuous stream of deionized water was flowed over the catalyst. Many times during the rinsing procedure the catalyst was batch flushed with this water and always under a helium atmosphere. No traces of spalled catalyst were found in the effluent water. Washing was stopped after 93.5 h when the pH of the rinse water (6.7) was approximately the same as the demineralized water (6.2). The leaching apparatus was disassembled and the reactor was incorporated into the system, all steps performed under a helium flow. The reactor system was then placed under hydrogen until the temperature and pressure were brought to methanation conditions, at which time the synthesis feed gas was gradually fed into the system to start the run. Pertinent catalyst bed data are shown in Table I. Reactor System The synthesis gas in the experiment was made by steam reforming natural gas in a Girdler plant. The hydrogen to carbon monoxide ratio ranged from 3.1:l to 3.3:1, always with a slight excess of hydrogen to prevent possible carbon deposition. Two activated carbon traps were located in the fresh gas inlet line and one in the product recycle line. This precaution was taken to eliminate the possibility of sulfur contamination of the nickel catalyst. Sulfur concentrations entering the system were maintained around 66 ppb and were detected by the methylene blue method. Figure 3 illustrates a simplified flow diagram of the tube wall methanation system. Fresh synthesis gas enters the system and mixes with a dry product gas recycle stream or "cold recycle" stream. In TWR-run 16 this mixed gas stream was preheated through several heat exchangers and entered the reactor a t 353 "C. The boiling Dowtherm coolant system was maintained a t 375 "C and controlled the maximum temperature variance of the catalyst to
within 20 "C. The product gas exits the reactor and splits; one stream leaves the system and is flared and the other is recycled. The recycle stream is cooled and dried by a double pipe heat exchanger and then compressed. Flow streams are metered and checked by orifices. During the test volumetric gas samples which were representative of a 24-h period were analyzed by both mass spectrometry and gas chromatography. An on-line gas chromatograph checked sample results. The gas analyses along with condensed product water weights and metered flows were used to calculate the mass balances for each 24-h period. A sample stream from the mixed gas before the reactor was sent to an iron deposit unit consisting of a transparent Vycor tube heated to 300-350 "C in a furnace through which the sample stream flows. As shown in the past, any iron or nickel carbonyl in the sample gas will decompose and form a metallic coating on the tube.
Results and Discussion TWR-Run 16 was terminated after 1179 h on stream. The following conditions were investigated during the experiment: a fresh feed exposure velocity of 10, 20, 30, and 40 (1 unit exposure velocity = 1 SCFH of fresh synthesis gas/ft2 of geometric catalyst surface area) and a cold-recycle-to-fresh-gasratio ranging from 3:l to 0. Inlet gas temperature (av 353 " C ) , Dowtherm coolant temperature (av 375 "C), and system pressure (300 psig) were held reasonably constant throughout the test. Three brief unscheduled power outages occurred but did not affect catalyst life. No shutdowns occurred during the run. Operating parameters and product gas characteristics are shown in Figure 4. The following select representative periods are tabulated in Table I1 to provide more detailed information in run 16. Period 3 represents performance of the fresh catalyst at a 10-exposure velocity and a 3.06:l recycle ratio. Period 8 indicates the effect of decreasing the recycle ratio to 1.50:l at a constant exposure velocity of 10. Periods 11 and 18 show the results of decreasing the recycle ratio from 1.54:l to 0.50:l at an exposure velocity of 20. The effect of increasing fresh feed can be compared between period 8 and period 11. Periods 21 and 35 provide a comparison of identical flow conditions a t a 30-exposure velocity and a 0.50:l recycle ratio but at different times. Period 26 represents performance a t an exposure velocity of 30 with no recycle. Periods 41 and 44 represent performance at the highest exposure velocity attained (40) and the change of recycle from 0.25:l to 0 a t the end of the run. Since carbon monoxide was the limiting reactant, a good indication of the Raney nickel catalyst performance was the percent carbon monoxide in the product gas. In general, an increase in exposure velocity a t a constant recycle ratio brought about an increase in product CO concentration. At 263 h on stream, the fresh gas was increased from an exposure velocity of 10 to 20 a t 1.50:l recycle ratio and the product gas CO concentration increased from 0.13 to 0.52%. At 503 h the exposure velocity was boosted from 20 to 30 a t 0.50:l recycle ratio with an increase in product gas CO concentration. At 959 h the 30-exposure velocity condition was raised to 40 and again the greater the fresh feed the lower the CO conversion. Although the fresh feed was decreased at 1127 h from a 40-exposure velocity to 30, the percent CO in the product did not decrease as would be expected. The catalyst was appreciably deactivated a t this point.
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 1, 1979 2 51
1 400 .-
Maximum catalvst
I
I
I
I
159
nrl
1
Exp TWR-16
1
350
300
t
05 1000
c
Product qos heating value
W
U
1 vi
800
a W
1
o
t
I
I
LCL1200
Cold
400
iI
recycle ratio
600
eo0
io00
1200
0
200
600
400 TIME
TIME ON STREAM h o u r s
ON
IO00
800
I200
STRELM.hours
Figure 4. Reactor conditions and product gas characteristics.
It should also be mentioned that, at a constant fresh feed rate and cold recycle ratio, the percent CO in the product gas gradually increased with time. An example of this observation can be seen at the exposure velocity of 30 and 0.51 recycle ratio in the period from 697 to 963 h when the percent CO increased from 0.66 to 0.84. Results indicate a time-dependent poisoning which deactivates the catalyst throughout the run. The recycle ratio also directly affected the catalytic reactor performance. When the recycle ratio was initially increased from 1-51to 3:l a t an exposure velocity of 10, the CO percent in the product increased. Later the CO concentration decreased when the recycle was brought back to 1.5:l. A t a 30-exposure velocity the recycle ratio was changed from 0.51 to 0 back to 0.51. Here the best reactor performance was at the no recycle level. A decrease in recycle at an exposure velocity of 20 or 40 also brought about a lowered product CO percent and thus a better CO conversion. From the previous tlndings a simple plug flow model can be developed. For a catalyst sprayed tubular reactor, mass transfer of reactants and products to and from the tube wall becomes an important factor. In the reactor the bulk gas diffusion resistance is assumed much greater than that of the surface reaction, and thus the global rate of reaction is limited by the rate of flow of reactant to the surface. Since bulk mass transfer is limiting in this laminar flow region, the reactant concentration at the catalyst surface will be negligible, and the global reaction rate for the disappearance of CO is directly proportional to the bulk CO concentration. For this isothermal isobaric system the fractional change in volume of the system can be considered by making the restriction that the volume of the reacting system varies linearly with conversion. This contraction factor, e , accounts for both the reaction stoichiometry and the presence of recycled inerts. Solving the design equation for a plug flow reactor in terms of CO conversion yields
An equation for the mass transfer coefficient can be found
ol--L---
30
2 0
IO
l/"t
Figure 5. Least-squares fit through plotted representative data.
by dimensionlesss group analogies between heat and mass transfer. Specifically, the following correlation of Sieder and Tate for laminar flow was used (Bennett and Myers, 1962).
(;y3(
Nu = 1.86Re1/3Sc1/3
$14
(2)
Finding the mass transfer coefficient and substituting into the design equation yields
SD0.667 2*37D2.333L0.333
(:)"'I
- 1-
-
u0.667
If the model represents the reactor, a plot of the right side of eq 3 vs. l/v0.667 should yield a straight line. In Figure 5 data from selected periods were plotted as indicated above. A linear least-squares fit through the points has a slope of 1.63 ( f t / ~ ) O . ~If~ the . bracketed theoretical slope ~~ of eq 3 is calculated, a value of 2.43 ( f t / ~ ) Ois. ~found. Agreement between the slopes is good considering the various assumptions and approximations in the calculation. The design equation does not take into consideration catalyst deactivation and, as previously noted, deactivation did occur throughtout the run. At the end of the test extensive deactivation had occurred and a point plotted a t that time does not fall near the straight line in Figure 5.
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Ind. Eng. Chern. Process Des. Dev., Vol. 18, No. 1, 1979
0 - 5
2 ,TIPI
4
6 8 IO REACTOR LENGTH, f e e t
12
e1;4
Figure 6. Actual gas concentration profiles for two different times on stream but same exposure velocity and recycle ratio.
The design equation can further be simplified in terms of the fresh synthesis gas rate and the recycle flow. It can be seen that the reactor CO conversion decreases with an increase in fresh feed or an increase in cold recycle. These results have been verified as shown before in the data. Besides the predominant methanation reaction a secondary reaction is noticeable due to the CO, concentration in the product gas. The "water gas shift" reaction combines CO and HzO to form COPand Hz CO + HzO + COZ + Hz (4) Apparent equilibrium constants were calculated for the water gas shift reaction and methanation reaction and are tabulated for select periods in Table 11. True equilibrium values for the reactions at 380 "C are 14.4 and 19600 atm-*, respectively. The small values of the apparent equilibrium constant for the water gas shift reaction indicate that the reaction does not proceed to any appreciable extent, unlike the results in methanation bench-scale testing units (Savinell et al., 1976). The magnitudes of the methanation reaction constants are much greater and indicate that the reaction is dominant over the Raney nickel catalyst. At the end of the test the methanation constant decreases considerably due to catalyst deactivation. The stationary gas probes were used to take samples along the catalyst length at various steady-state conditions a t a specific time. However, the CO in the product gas increased with time a t constant flows and thus the CO in the mixed gas to the reactor increased, which made profile comparisons difficult. The concentration profiles at 625 and 1177 h can be compared at a 30-exposure velocity due to the no recycle situation. Figure 6 shows the CO concentration profiles along the catalyst length. In the earlier profile most of the reaction took place along the top 2-6 ft of bed. In the later profile extensive deactivation had occurred and the 9-14 ft of catalyst length was more reactive. Deactivation of the inlet portion of the bed was so extensive that the exit percent CO never went below 3.3 a t the 14-ft level. The above findings would suggest a deactivation which proceeds zonally rather than uniformly along the catalyst length. Catalyst temperature profile data indicate temperature fluctuations along the catalyst length. Figure 7 represents the temperature profiles a t several different conditions. The maximum temperature or "hot spot" is the position along the reactor length a t which a major portion of the exothermic methanation reaction takes place. As catalyst performance decreases, the hot spot moves down the length of the reactor as seen in Figure 8, again suggesting zonal deactivation. Similar results are found in the bench-scale methanation studies of Savinell et al. (1976). The hot spot location is affected only by the fresh feed rate and time and not by recycle, for although a decrease in recycle will increase CO conversion, the hot spot does not move up the
Figure 7. Actual temperature profiles for various times on stream.
0
240
480
720
960
I200
7 ME ON STWEAM hour5
Figure 8. Hot spot location at various exposure velocities.
reactor length. If a particular concentration profile is compared with the temperature profile, it is found that the hot spot falls within the catalyst length where most of the conversion takes place. The movement of the hot spot along the catalyst length combined with the concentration profile findings strengthens the theory of a zonal deactivation. Spent Catalyst Analysis At 1179 h the synthesis gas feed to the reactor was halted. The electrical resistance heaters and the compressor were turned off, and the system was bled, pressure purged with hydrogen, and left under a hydrogen atmosphere until cool. Later the reactor was removed from the system and scrapings of the brittle spent catalyst were obtained from the top 2 ft of catalyst (reactor inlet), middle 2 ft, and bottom 2 ft (product exit). Bulk samples were subjected to chemical analysis, X-ray diffraction analysis, and BET surface and pore size distribution measurements. Results of the chemical analysis and X-ray diffraction are shown in Table 111. The sulfur concentrations found by wet chemical analyses are roughly the same in all three spent catalyst samples. These are within the limits of 0.2-0.3% sulfur in unactivated Raney nickel. Gas samples were analyzed for sulfur by the methylene blue method both before and after the final carbon trap in the fresh gas system. The average sulfur concentration entering the reactor was 66 ppb. The iron concentrations were greater than the 0.1-0.3% found in unactivated Raney nickel. The mixed gas into the reactor was sampled for iron carbonyl by using the iron deposit units as previously described, Two trials, one a t the beginning and one at the end of the run, yielded after 453 and 573 h 15.5 ppb and 4.2 ppb, repectively. The iron carbonyl concentration is probably greater than above because the units are not 100% efficient. Since most of the reactor system was carbon steel and under high partial pressures of carbon monoxide a t room temperature, it is conceivable that Fe(C0)5 was formed, entered the hot reactor in the feed stream, decomposed, and then deposited iron on the Raney nickel catalyst (Schehl et al., 1977).
Ind. Eng. Chem. Process Des. Dev., Vol.
18, No. 1, 1979
161
Table 11. Selected Test Data TWR-Run 16 period number 3
11
8
18
21
26
35
41
45
hours o n stream fresh gas: rate, scfh H, , vol % co, vol % co,, vol % N,, vol % CH,, vol % H,O, vol % H,/CO exposure velocity, scfhlft' space velocity, h" mixed feed gas (wet): rate, scfh H,, vol % co, vol % co,, vol % N,, vol % CH,, vol % H,O, vol % H,/CO inlet superficial velocity, f/s inlet Reynolds no., exposure velocity, scfh/ft2 space velocity, h-l vol. cold recycle/vol. fresh gas temperatures: gas inlet, " C maximum catalyst, " C average catalyst, C pressure, psig product gas (wet): H,, vol % c o , vol % co,,vol % N,, vol % CH,, vol % H,O, vol % H,KO conversion: H,, % mixed feed CO, % mixed feed (H, + CO), % mixed feed usage ratio heating value, Btu/scf system carbon recovery, % scf CH,/lb of catalyst K,, apparent: methanation, atm-l water gas shift
119
239
311
479
551
67 1
887
1031
1127
75.8 75.5 23.6 0.1 0.7 0.1 0 3.20 10.26 244
73.9 75.8 23.7 0.1 0.3 0.1 0 3.20 10.00 237
148.0 75.2 24.2 0.2 0.2 0.2 0 3.11 20.03 476
147.3 7 5.4 23.6 0.1 0.8 0.1 0 3.20 19.93 414
223.8 75.2 24.3 0.1 0.3 0.1 0 3.10 30.28 720
225.2 74.3 24.3 0.2 1.2 0 0 3.06 30.47 724
223.6 75.7 23.8 0.2 0.3 0 0 3.18 30.26 719
291.5 74.1 24.0 0.2 1.6 0.1 0 3.09 40.26 957
299.7 76.1 23.1 0.3 0.5 0 0 3.29 40.56 964
308.0 21.53 6.06 0.33 0.97 70.99 0.12 3.55 0.43 1702 41.68 990 3.06
183.5 33.92 9.57 0.20 0.99 55.22 0.10 3.54 0.25 876 24.83 590 1.50
375.0 33.20 10.01 0.37 0.65 55.67 0.10 3.32 0.52 1797 50.75 1206 1.54
220.4 52.47 15.78 0.24 0.91 30.56 0.05 3.33 0.31 802 29.82 709 0.50
336.1 52.64 16.36 0.38 0.52 30.04 0.06 3.22 0.46 1248 45.48 1081 0.51
223.5 74.30 24.30 0.20 1.20 0 0 3.06 0.31 557 30.24 719 0
331.0 54.10 16.25 0.49 0.69 28.42 .06 3.33 0.44 1221 44.79 1064 0.49
371.0 62.41 19.47 0.46 1.48 16.11 0.04 3.21 0.50 1159 50.20 1193 0.25
298.0 76.10 23.10 0.30 0.50 0 0 3.29 0.39 7 54 40.33 958 0
366 392 385 300
359 394 386 300
373 396 386 300
371 395 388 300
361 396 389 300
373 392 383 300
341 395 385 300
350 397 385 300
335 396 383 300
3.76 0.33 0.38 1.02 90.45 4.06 11.39
5.31 0.12 0.24 1.30 82.45 10.58 44.25
4.97 0.67 0.40 0.79 76.67 16.50 7.42
5.23 0.16 0.39 0.86 69.99 23.37 32.69
6.14 0.50 0.70 0.73 68.13 23.79 12.28
5.08 0.11 1.04 0.53 46.62 46.57 46.18
7.45 0.58 0.81 1.13 65.97 24.06 12.84
11.44 1.10 1.06 0.71 56.21 29.37 10.40
14.42 1.45 1.60 0.60 42.01 39.76 9.94
85.25 95.46 87.50 3.17 967.9 93.04 10.94
87.92 99.06 90.37 3.15 952.9 88.44 10.96
87.51 94.43 89.11 3.07 951.1 88.25 21.79
93.56 99.34 94.89 3.13 947.1 93.69 21.16
92.30 97.97 93.65 3.03 933.0 92.88 33.10
96.66 99.78 97.43 2.96 916.5 95.44 32.70
90.98 97.67 92.53 3.10 913.4 93.66 32.09
89.00 96.60 90.80 2.95 865.8 95.22 42.17
90.58 96.89 92.05 3.08 796.3 96.51 39.57
457 1.07
1059 1.00
336 0.18
1559 0.55
306 0.36
3285 1.03
144 0.43
21.9 0.38
8.38 0.40
Table 111. Properties of Spent Raney Nickel Catalyst reactor t o p (gas inlet)
reactor middle
reactor bottom (gas exit)
38.7 31.3 1.3 0.47 0.14 0.19
39.2 32.9 0.8 0.38 0.14 0.3
51.7 28.3 0.6 0.44 0.12 0.34
Ni,C, Ni,Al, NiAI,
Ni,C, Ni,AI, NiAl,
Ni
chemical analysis, %
Ni A1 C Fe Na S X-ray diffraction
Carbon content was determined by the standard ASTM combustion technique. The concentration is greater at the entrance and all samples were greater than the 0.1-0.270 inherently in unactivated Raney nickel. Carbon may be present on the catalyst as the result of several reactions:
(1) the decomposition of iron carbonyl, (2) the decomposition of carbon monoxide, and (3) the formation of nickel carbide. As shown, there is evidence of iron being carried into the reactor and some carbon associated with carbonyl deposition is possible. The iron probably further promotes carbon deposition through the decomposition of carbon monoxide. Furthermore, X-ray diffraction studies have verified the presence of a nickel carbide a t the reactor inlet. As shown in the chemical analyses, the nickel concentration a t the reactor exit is greater than a t the reactor entrance. Two possibilities can be speculated: first, provided no catalyst eroded during the run and the spent catalyst samples were scraped to the same depth, then unequal leaching along the 14 f t of catalyst occurred. In the leaching diagram, Figure 2 , the gravity-fed caustic solution enters the reactor bottom first. Since the activation time was short (4.25 h) and a forced circulation of the caustic solution was not used, the bottom leached the most as seen by the highest Ni content (51.7%) and the
162
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 1, 1979
Table IV. BET Surface Areas, Pore Volumes, and Pore Radii of Spent Raney Nickel Catalyst reactor top
surface area. m2/e pore volume, cA'/g average pore radius, A percent pore volume
reactor
bottom
2 ft (inlet)
middle 2 ft
(outlet)
43.17 0.118 54.88
39.40 0.121 61.22
28.80 0.132 91.87
12.69 14.14 11.45 9.91 51.81
6.04 13.74 12.14 10.69 57.39
0 0 4.09 11.69 84.22
26.12 22.46 14.04 9.93 27.45
13.46 24.20 16.59 11.95 33.80
0 0 8.25 19.55 72.20
2 ft
with the following
radii: 60 A
percent total surface area with the following radii: 60 A
lowest A1 content (28.5%). Also, X-ray diffraction results indicate A1-Ni compounds, which are found in unleached Raney nickel, in the top section of the catalyst. The other possibility is replacement of the active nickel a t the reactor inlet by migration of the aluminum from pockets of unleached catalyst or from aluminum complexes formed by leaching. Most of the reaction took place along the catalyst top length during the experiment and the local heat of reaction could provide a driving force for the migration. Although metal surface area measurements were not performed on the spent catalyst samples, BET surface measurements were obtained and shown in Table IV. The surface areas of the samples are not as great as freshly leached Raney nickel (64 m2/g). However, there is a decreasing trend from reactor inlet to outlet. The high amount of carbon at the inlet would account for the large surface area. Amorphous carbon would also indicate the smaller pore radii associated with carbon a t the reactor inlet as compared to the outlet. Summary The tube wall methanator provides an efficient method of heat removal and a relative simple design which is an
advantage in future scale-up work. Even without the use of a recycle flow, catalyst temperatures can be controlled with relative ease. The reactor is capable of converting fresh feeds with a CO concentration as high as 25%. With different flows there is negligible pressure drop across the reactor. Recent improvements in the catalyst flame spraying techniques and equipment have decreased spraying laydown time by 20-fold. Although this particular run lasted only 1179 h with a wide range of fresh feed and recycle flows, the catalytic reactor did show relative stability during the test with little difficulty in operation. Since run 16, the process development system has been modified. Vessels and piping of carbon steel have been replaced with stainless steel to retard or eliminate the formation of the iron carbonyl poison. A future test will utilize this new system. Nomenclature D, inside diameter of reactor, ft D, diffusivity, ft2/s e, contraction factor h , mass transfer coefficient, ft/s L , catalyst length, ft Nu, Nusselt number Sc, Schmidt number Re, Reynold number Q, reactor inlet volumetric flow, ft3/s S, geometric catalyst surface area, ft2 wo, catalyst surface gas viscosity, lb/ft-s W , bulk gas viscosity, lb/ft-s u , reactor inlet gas velocity, ft/s Xco, CO conversion L i t e r a t u r e Cited Bennett, C. O., Myers, J. E., "Momentum, Heat, and Mass Transfer", Mc(3awHIl1, New York, N.Y.. 1962. Haynes, W. P., Elliott, J. J., Forney, A. J., Am. Cbem. Soc. Div. FuelChem. Prepr., 16 (2),47 (1972). Haynes, W. P., Forney, A . J. Elliott, J. J.. Pennline, H. W., Am. Chern. Soc. Div. Fuel Chern. Prepr.. 19(3), 10 (1974). Haynes, W. P., Schehl, R. R., Weber, J. K.,Forney, A. J., Ind. Eng. Chem. Process Des. Dev., 16, 113 (1977). Savinell, R. F., Youngblood, A. J.. Haynes, W. P., ERDA PERU RI-7614,(1976). Schehl, R. R., Pennllne, H. W., Strakey, J. P., Haynes, W. P., Am. Chern. Soc. Div. Fuel Cbem. Prepr., 21 (4),2 (1976). Schehl, R. R., Pennline, H. W., Youngblood, A. J., Baird, M. J., Strakey, J. P., Haynes, W. P., ERDA PERC/RI-77/10, (1977).
Received f o r review April 7, 1978 Accepted July 20, 1978
Presented at the Division of Colloid and Surface Chemistry, 2nd Joint Conference Chemical Institute of Canada/American Chemical Society, Montreal, Canada, May 1977.
