688
Ind. Eng. Chem. Res. 1989,28, 688-693
*SF
=
2fAl - 4P,US2AL dt
(3) The proposed solids friction factor, f,, can be expressed by eq 15.
Literature Cited
Acknowledgment This study has been supported in part by the US.Department of Energy (Morgantown Energy Technology Center) under Contract DE-AC21-82MC19372.
Nomenclature Cd = drag coefficient d = diameter, m f = friction factor g = gravitational constant, 9.807 m/s2 G = mass flux, kg/m2/s AL = transport length, m n = Richardson-Zaki index N = cluster number P = pressure, kg/(m s2) AF' = pressure drop, kg/(m s2) Re = Reynolds number U = velocity, m/s Greek Symbols t
= voidage
p = p =
SPE = solid potential energy SSF = solid-solid friction SWF = solid-wall friction t = tube, terminal T = total
gas viscosity, kg/(m s) density, kg/m3
Subscripts
Breault, R. W. Hydrodynamic Characteristics and Coal Combustion Modeling of a High Velocity Fluidized Bed. Ph.D. Dissertation, The University of New Hampshire, Durham, 1985. Capes, C. E.; Nakamura, K. Vertical Pneumatic Conveying: An Experimental Study with Particles in the Intermediate Turbulent Flow Regimes. Can. J. Chem. Eng. 1973, 51, 31. Klinzing, G. E. Vertical Pneumatic Transport of Solids in the Minimum Pressure Drop Region. Ind. Eng. Chem. Process Des. Dev. 1979, 18(3),404. Mathur, V. K. Pressurized High Velocity Fluidized-Bed Combustion Modeling. Report DOE/MC-19372-4 submitted to Morgantown Energy Technology Center, May 1984; Morgantown, WV. Reddy, K. V. S.; Pei, D. C. T. Particle Dynamics in Solids-Gas Flow in a Vertical Pipe. Ind. Eng. Chem. 1969, 8(3), 490. Squires, A. M. Application of Fluidization Beds in Coal Technology. In Alternate Energy Sources; Hartnett, J. P., Ed; Hemisphere Publishing Corp.: Washington, DC, 1976. Stemerding, S. The Pneumatic Transport of Cracking Catalyst in Vertical Risers. Chem. Eng. Sci. 1962, 17, 599. Van Swaaij, W. P. M.; Buurman, C.; von Breusel, J. W. Shear Stresses on the Wall of a Dense Gas-Solid Riser. Chem. Eng. Sci. 1970,25, 181. Yang, W. C. A Correlation for Solid Friction Factor in Vertical Pneumatic Conveying Lines. AIChE J. 1978, 24(3), 548. Yerushalmi, J.; Cankurt, N. T. High - Velocity Fluid Beds. CHEMTECH 1978, Sept, 564. Yersuhalmi. J.: Cankurt. N. T. Further Studies of the Regimes of Fluidization: Pouder' Technol. 1979, 24, 187. Yerushalmi, J.; Cankurt, N. T.; Geldart, D.; Liss, B. Flow Regimes in Vertical Gas-Solid Contact Systems. AIChE Symp. Ser. 1978, 74(176), 1. Yerushalmi, J.; Gluckman, M. J.; Graff, R. A,; Dobner, S.; Squires, A. M. Production of Gaseous Fuels from Coal in the Fast Fluid Bed. In Fluidization Technology; Keairns, D. L., Ed.; Hemisphere Publishing Corp.: Washington, DC, 1975; Vol. 11. Yerushalmi, J.; Turner, D. H.; Squires, A. M. The Fast Fluid Bed. Ind. Eng. Chem. Process Des. Dev. 1976, 15(1),47. Y
c = cluster g = gas f = friction
GKE = gas kinetic energy GPE = gas potential energy GWF = gas-wall friction mf = minimum fluidization s = solid sl = slip velocity SF = solid friction SGF = solid-gas friction SKE = solid kinetic energy
Received for review J u n e 13, 1988 Revised manuscript received February 10, 1989 Accepted February 25, 1989
High-Velocity Fluidized Bed Hydrodynamic Modeling. 2. Circulating Bed Pressure Drop Modeling Ronald W. Breaultt and Virendra K. Mathur* Department of Chemical Engineering, University of New Hampshire, Durham, New Hampshire 03824
Circulating high-velocity fluidized beds (HVFB) have been proposed to eliminate some of the problems encountered in conventional fluidized beds. The loop fluidized bed is one such system being considered for pressurized combustion of coal in the presence of a sulfur sorbent such as dolomite. This study has been conducted to obtain fundamental knowledge of the hydrodynamics of the one particular HVFB, the loop fluidized bed (LFB), which operates in various flow regimes. Experimental data obtained from a bench-scale unit have been used for the development of a mathematical model to predict the pressure profile in the LFB. The model requires solid flow rate, gas flow rate, equipment geometry, and solid fraction as inputs to predict the pressure a t any point in the system. The predicted pressures have been compared with experimental data and show good agreement. Part 1 of this study (Breault and Mathur, 1989) has described a test fluidized bed which has several advantages
-
'Present address: Rilev Stoker CorDoration. Worcester. MA 01610.
0888-5885/89/2628-0688$01.50/0
over a conventional fluidized bed. A high-velocity fluidized bed (HVFB) can overate over a wide range of gas throughputs.' The ga$ rate may be reduced to such a ;egree that the bed becomes turbulent or even enters the bubbling regime without losing uniformity of the bed 0 1989 American Chemical Society
Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989 689 Sections A-B and C-D are regimes of aerated solid downflow and gas upflow. The pressure drop across these sections can also be represented by the sum of the potential energy and a frictional term: (4) UA-B (01 Wc-d = MSPE + flf
b+ Standpipe
Gar
,/h
A R 1
I
where A P s p E and APf can be obtained from eq 2 and 3, respectively. Standpipe Flow. Standpipe flow (section B-C) has been extensively studied by various workers (Yoon and Kunii, 1970; Kunni and Levenspiel, 1977; Leung et al., 197813). The pressure drop through this section of the LFB can be modeled by
E
0
inlet
Eductor
I PRESSURE
Figure 1. Pressure profile in high-velocity fluidized bed system.
temperature. It is also claimed that in a HVFB coal might be introduced at fewer points without excessive pressure drop. Furthermore, it may be possible to capture higher amounts of sulfur dioxide due to the use of fine dolomite or limestone particles in a HVFB. A special case of the high-velocity fluidized bed concept, recently developed at the Morgantown Energy Technology Center, Morganntown, WV, is the loop fluidized bed (LFB) shown in Figure 1of part 1. The loop fluidized bed is one such circulating fluidized bed and is being considered for pressurized combustion of coal in the presence of a sulfur sorbent such as dolomite. No hydrodynamic characteristics of such a fluidized bed system are available. This investigation has been conducted to study the hydrodynamics of the LFB with special reference to the riser section which operates in the high-velocity fluidization regime. Data obtained in a bench-scale experimental setup have been used for the development of a mathematical model to predict the pressure profile in a circulating high-velocity fluidized bed, the LFB.
Background for Pressure Drop Modeling in Circulating High-Velocity Fluidized Bed Systems The HVFB system which includes the solid recirculation section as well as the entrained flow section consists of several gas-solid flow regimes. These are aerated gas-solid downflow, aerated solids downflow, and gas upflow, horizontal pneumatic transport, vertical pneumatic transport, and standpipe flow. Research workers have modeled the pressure drop in each of these regimes as the sum of individual contributions due to acceleration, kinetic energy, potential energy, and frictional effects (Stemerding, 1962; Reddy and Pei, 1969; Capes and Nakamura, 1973; Van Swaaij et al., 1973; de Jong and Hoelon, 1975a,b; Kunni and Levenspiel, 1977; Singh, 1977; Yang, 1978; Klinzing, 1979). These flow regimes are discussed as per Figure 1. Aerated Solid Flow. This is the region (section H-A) of aerated gas-solid downflow that exists in the LFB. The pressure drop for this regime can be represented by the sum of the potential energy and a frictional term: (1) UHA = USPE - APf where U s p E = p , ( l - t)gAL sin 0 (2) and (3) for Re = dgUs < 1300 P
where the slip velocity is given by
Pneumatic Transport. The eductor is a region of horizontal pneumatic transport, section D-E. The pressure drop is modeled by summing contributions due to kinetic energy requirements and frictional resistance: (7) APE-F = USKE +flf The pressure drop due to the increase in the particle kinetic energy is where
u, = ug- u, The pressure drop due to the frictional resistances of the gas and solids is (9) = UGWF + MSWF The pressure drop due to the gas frictional resistance can be expressed by the Fanning equation: Af'f
where the friction factor can be estimated from the literature (Capes and Nakamura, 1973). The pressure drop due to the solid frictional resistance has been studied by a number of investigators. Two approaches are generally taken to express this effect. Rose and Barnacle (1957) obtained the solid frictional resistance term by modifying the gas frictional resistance:
The solid friction factor, f p , is presented in a graphical form (Capes and Nakamura, 1973). An alternative appwach has been to use a modified Fanning equation 2fsp,(1 P S W F
- ww dt
(12)
The solid friction factor, f,, has been obtained by several researchers and has been presented in Table I, part 1. Riser Section. The riser portion (section F-G) of the high-velocity fluidized bed has not been investigated in detail. However, most workers consider that the pressure drop can be expressed as the sum of the individual con-
690 Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989
tributions due to kinetic energy, potential energy, frictional resistance effects, and cluster formation: 2 72 5 [
PF-G = USKE + USPE + uf + p, (13) The pressure drops U s , and APf can be determined from eq 8 and 9, respectively. The solid velocity, U,, used in eq 8 and 9 is taken to be the time-average velocity:
Run No S I Run No S2
o
H
0
1.75-
.-..
-
p.
E
+- 1 2 5 I '3 W
I
where p , ( l - t) is the time-average apparent density in the riser. The potential energy term can be obtained from an equilibrium force balance and is expressed as USPE
= Ps(1 - e ) g u
0251
(15)
The pressure drop, AP,,due to solid cluster formation is actually caused by the continuous particle-gas frictional effects experienced by the solid particles as the clusters form and break apart. This pressure drop contribution is included in Us,. Breault (1985) found that this effect could be modeled by a Fanning-type equation with the solid friction factor expressed by f a = 12.2(1 - €)/Ua€3
075r
-0.25l
0
I
I
I
I
I
I
I
2
3
4
5
6
7
STATIC PRESSURE (kg/ms2 X
Figure 2. Pressure profile in LFB for sand particles. 2 75
2 25
(16)
Pressure Drop across Orifice Plates. Leung and Jones (1978a) have reviewed the data and models presented in the literature for gas-solid flow through orifice plates. They recommend the use of the following equation:
I
-E -
I75
+
I25
I
o
RUN L I
0
RUN L 2
0
RUN L 3
\
0 W
I
0 75
The pressure drop predictions of eq 17 agree with the results reported by the six investigators (Leung and Jones, 1978a). The values of CD range from 0.5 to 0.65. Pressure Drop in Bends Due to Gas-Solid Flow. Kunii and Levenspiel (1977) present an equation that predicts the pressure drop in bends due to gas-solid flow. This equation is used extensively in pneumatic transport. The pressure drop is given by
ap = fbpu:
0 25
-0 2 5
1
,
I
2000
I
4
I
I
6000
4000
I
8000
STATIC PRESSME (kg/ms2)
Figure 3. Pressure profile in LFB for limestone particles.
(18)
The bend friction factor, f b , is 0.375, 0.188, and 0.125 for ?-b/dt values equal to 2, 4 and 6+, respectively.
Experimental Setup and Procedure The experimental unit used in this study is the same as that described in part 1. Pressure Profile Modeling in a Circulating High-Velocity Fluidized Bed Pressure measurements for sand particles are taken a t pressure ports P,-P, as shown in Figure 2 of part 1. The pressure at the gas inlet to the LFB (pressure port PI) is the highest, while the pressure at the gas outlet (pressure port P5)is the least compared to any other point in the LFB. This means that the gas entering the LFB has two paths that it may take enroute to the exit. The gas flow up the standpipe creates a pressure drop equal to that produced by the gas-solid flow through the riser. Experiments for pressure measurements in the LFB have been conducted for sand, limestone, and gypsum particles a t various solid flux, air flux (including flow rate through nozzle No), solid fraction in the riser, and standpipe height. Typical experimental pressure profiles (static pressure versus loop height) for sand, limestone, and gypsum particles are presented in Figures 2, 3, and 4.
2 25
-
-E &
2
w I
t
175 125
t
I\
o
RUN G 1
e
RUN G2
n RUN G3
0751
T
-0 250
2000
4000
6000
STATIC PRESSURE (kg/ms2)
Figure 4. Pressure profile in LFB for gypsum particles.
These agree with the theoretical concept shown in Figure 1. The profiles for all three solid particles exhibit the same general behavior. Static pressure at point D (solid entrance to the horizontal bottom loop) varies from 5000 to 6700 kg/(m s2). This variation is considered due to the effect of operating parameters in general and solid fraction in the riser in particular. The static pressure variation at
Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989 691
0 0 A 8 Figure 5. Loop fluidized bed model.
PDAGUP
PDAGD
PDSTNP
PDSTNP
PDORFC
PDORFC
PDAGUP
PDAGD
I
RISER CALL PDAGP
11. CALL PDBND
11. CALL PDVPC
4 CALL PDBND
& CALL PDHPC
(*) Figure 6. RISER calculation flow chart.
point A is considered due to downstream pressure drop in the cyclone and bag filters. The vertical B position is the height of the solids in the standpipe which is controlled by the amount of solids charged to the system. A model is developed to construct the pressure-height profile based on the physical and operating parameters in the HVFB. The model is based on pressure drop correlations for the various flow regimes discussed earlier. Pressure drop for the riser in the high-velocity fluidized bed is estimated by the model developed elsewhere (Breault, 1985). The HVFB is simplified for the model developed and is shown in Figure 5. The simplification involves primarily the air nozzle system. The model developed utilizes only one nozzle through which all air is introduced to the HVFB. As shown later, this simplification greatly reduces model complexity without sacrificing the accuracy of the model. It consists of two parts. These are the riser section ABCDEFGH and standpipe section AJ'JIH as shown in Figure 5. The portion ABCDEFGH of the HVFB consists of a horizontal pneumatic transport section, a 90' bend, the riser, a 1 3 5 O bend, and an angled pneumatic transport section. The correlations for estimating pressure drops for each of these sections have been previously discussed.
L(ry Figure 7.
STNDPP
calculation flow chart.
v 11HVFBPP.FOR
j ] RISER
0 /I HlMlF
RISER
8 1 STNDPP
Figure 8. HvFBPP calculation flow chart.
These equations have been rearranged and incorporated in the RISER subroutine to predict the pressure at the points of interest. The RISER subroutine calls subroutines written for each section. The subroutines and the main program are presented elsewhere (Breault, 1985). A calculation flow chart of the RISER subroutine is presented in Figure 6. The portion AJ'JIH of the HVFB consists of the following sections: an aerated solid flow section, an orifice, a standpipe, and an aerated solid flow section. The equations used for estimating the pressure drop in each of these sections have been discussed previously. They are suitably arranged to predict the pressure at the points of
692 Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989
In this study, a bench-scale loop fluidized bed made of Pyrex glass has been designed, fabricated, and installed. The LFB has been operated using sand, limestone, and gypsum particles. Experimental data are obtained to study the effect of particle size, particle density, air flux, and solid flux on the pressure profile in a circulating high-velocity fluidized bed. Based on the data, a pressure profile model represented by the computer program HVFBPP.FOR has been developed which accurately predicts pressure drop at any point in the LFB system.
Table I. Absolute Average Percent Deviation between P r e s s u r e Profile Model a n d D a t a gas flux low medium high
runs s9 S10 SI1
sand AAPD
limestone runs AAPD L6 6.22 L7 L8 Ll L2 4.80 L3
21.33
s1 s2
6.25
S18 S19
6.99
7.62
L20
s20
gypsum runs AAPD GI G8 6.81 G9 G4 G5 13.21 G6 G1 5.59 G2 G3
Acknowledgment
75r
2 25
This study has been supported in part by the US.Department of Energy (Morgantown Energy Technology Center) under Contract DE-AC21-82MC19372.
-
c RUN S I o RUN
-
Nomenclature
52
l75C
E
k
I
125-
I
I 0 75c
- 0 25 1
0
I
I
I
2000
I
4000
I
I
6000
I
I
8000
STATIC PRESSURE ( k g / m s 2 )
F i g u r e 9. Pressure profile model prediction compared to the data for sand runs S1 and S2.
interest. The subroutines for estimating the pressure in the standpipe are assembled in the STNDPP subroutine (standpipe) and the flow chart is presented in Figure 7. The flow chart for the HVFB flow model HVFBPP (High Velocity Fluidized Bed Pressure Profile) is presented in Figure 8. The subroutines RISER and STNDPP develop the height versus static pressure plot as shown in Figures 2-4. The experimental data for sand, limestone, and gypsum particles at three values of gas mass flux (low, medium, and high) are compared with the predicted pressure profile using the proposed model. The absolute average percent deviation (AAPD) values are presented in Table I. The AAPD ranges from 4.8 for limestone runs L1, L2, and L3 to 21.33 for sand runs S9, S10, and S11. It should be noted that only two groups of data have AAPD values exceeding 10. This deviation is possibly due to plugging of the pressure ports, P, and P,. Elimination of poor data at these ports considerably reduces the error. The static pressures predicted by the model for sand runs S1 and S2 are plotted in Figure 9 along with the experimental data points. The largest error (horizontal deviation between points and line) is found to be in the standpipe with the experimental values being low. The riser section predictions agree well with the data. The model predictions for limestone and gypsum are presented elsewhere (Breault, 1985).
Conclusions During the past few years, considerable effort has been made on the research and development of fluidized bed combustion of coal. This technology holds a number of attractions, all stemming from the concept of maintaining low temperatures in the range 1100-1200 K in the combustion chamber.
C D = orifice discharge coefficient d = diameter, m f = friction factor g = gravitational constant, 9.807 m/sz G = mass flux, kg/m2/s AL. = transport length, m P = pressure, kg/(m s2) AF' = pressure drop, kg/(m sz) r = radius Re = Reynolds number U = velocity, m/s Greek Symbols t = voidage 0 = angle of inclination p = gas viscosity, kg/(m s) p = apparent viscosity, kg/(m s) CI# = sphericity p = density, kg/m3 /s = density of mixture, kg/m3
Subscripts
b = bend c = cluster f = friction g = gas GWF = gas-wall friction mf = minimum fluidization o = superficial p = particle s = solid sl = slip velocity SKE = solid kinetic energy SPE = solid potential energy SWF = solid-wall friction t = tube
Literature Cited Breault, R. W. Hydrodynamic Characteristics and Coal Combustion Modeling of a High Velocity Fluidized Bed. Ph.D. Dissertation, The University of New Hampshire, Durham 1985. Breault, R. W.; Mathur, V. K. High-Velocity Fluidized Bed Hydrodynamic Modeling. 1. Fundamental Studies of Pressure Drop. Ind. Eng. Chem. Res. 1989, preceding paper in this issue. Capes, C. E.; Nakamura, K. Vertical Pneumatic Conveying: An Experimental Study with Particles in the Intermediate and Turbulent Flow Regimes. Can. J. Chem. Eng. 1973,51, 31. de Jong, J. A. H.; Hoelon, Q. E. T. J. M. Aerated Solids Flow Through a Vertical Standpipe Below a Pneumatically Discharged Bunker. Powder Technol. 1975a, 7, 197. de Jong, J. A. H.; Hoelon, Q. E. T. J. M. Concurrent Gas and Particle Flow During Pneumatic Discharge from a Bunker Through an Orifice. Powder Technol. 1975b, 12, 201.
Ind. Eng. Chem. Res. 1989,28,693-697 Klinzing, G. E. Vertical Pneumatic Transport of Solids in the Minimum Pressure Drop Region. Znd. Eng. Chem. Process Des. Dev. 1979,18(3),404. Kunni, D.; Levenspiel, 0. Fluidization Engineering; Krieger Publishing Co.: Huntington, NY, 1977. Leung, L. S.;Jones, P. J. Coexistence of Fluidized Solids and Packed Bed Flow in Standpipes. In Fluidization; Davidson, J. F., Keairns, D. C., Eds.; Cambridge University Press: New York, 1978a; p 116. Leung, L. S.; Jones, P. J.; Knowlton, T. M. Analysis of Moving-Bed Flow of Solids Down Standpipes and Slide Values. Powder Technol. 1978b,7, 1. Reddy, K. V. S.; Pei, D. C. T. Particle Dynamics in Solids-Gas Flow in a Vertical Pipe. Ind. Eng. Chem. Fundam. 1969,8(3),490. Rose, H. E.; Barnacle, H. E. Flow of Suspensions of Non-Cohesive Spherical Particles in Pipes. Engineer 1957,203(5290),898,939.
693
Singh, B. Lean Phase Vertical Pneumatic Conveying of Particulate Material-An Analysis of the Pressure drop in the Non-Accelerating Zone. Chemeca 1977, 77, 315. Stemerding, S. The Pneumatic Transport of Cracking Catalyst in Vertical Risers. Chem. Eng. Sci. 1962,17, 599. Van Swaaij, W. P. M.; Buurman, C.; von Breusel, J. W. Shear Stresses on the Wall of a Dense Gas-Solid Riser. Chem. Eng. Sci. 1973,25,31. Yang, W. C. A Correlation for Solid Friction Factor in Vertical Pneumatic Conveying Lines. AZChE J. 1978,24(3),548. Yoon, S.M.; Kunii, D. Gas Flow and Pressure Drop Through Moving Bed. Znd. Eng. Chem. Process Des. Dev. 1970,9(4),559.
Received for review J u n e 13, 1988 Revised manuscript received February 10, 1989 Accepted February 25, 1989
Influence of Reaction Parameters on the Stereoselectivity of the Nickel-Catalyzed Gas-Phase Hydrogenation of o -Cresol. 1. Kinetics and Reaction Pathway Werner K. Schumann, Oemer M. Kut, and Alfons Baiker* Department of Industrial and Engineering Chemistry, Swiss Federal Institute of Technology (ETH), 8092 Zurich, Switzerland
The thermodynamics and kinetics of the nickel-catalyzed gas-phase hydrogenation of o-cresol to the two stereoisomers, cis- and trans-2-methylcyclohexanol,were studied. The kinetics were investigated in a continuous fixed-bed reactor over a commercial nickel on silica catalyst in the temperature range 150-240 "C and a t atmospheric pressure. The formation of the stereoisomer products was found to occur via 2-methylcyclohexanone as intermediate. The rate of formation of this intermediate is much smaller than the rate of the consecutive hydrogenation to the stereoisomer products. Consequently, the sorption equilibrium of the intermediate 2-methylcyclohexanone is not reached. The conversion of the o-cresol shows a maximum at about 200 "C, and the hydrogenation to the intermediate is first order in hydrogen. The isomer ratio of the 2-methylcyclohexanol production is only weakly influenced by the reaction conditions. The kinetic results of the o-cresol hydrogenation are compared to those obtained for the hydrogenation of o-tert-butylphenol. The cis selectivities reached with both reactants are higher than those predicted by equilibrium calculations. Catalytic hydrogenation of alkylphenols represents an economic way for the synthetic of alkylated cyclohexanols, some of which are important intermediates in the fragrance and perfume industry. In several applications, only one of the two stereoisomers of the alkylated cyclohexanolsis desired. The effects of different catalysts and operating parameters on the stereoselectivity of the hydrogenation of substituted phenols have been studied by various groups, and the present state of the art in liquid-phase hydrogenation has been reviewed recently by Bartok (1985). To our knowledge, kinetic studies of the stereoselective hydrogenation of alkylphenols are scarce and furthermore confined to the liquid phase only. Recently, Kut et al. (1988) reported the kinetics of the liquid-phase hydrogenation of o-tert-butylphenol over nickel, cobalt, and noble metal catalysts. Gas-phase hydrogenations of aromatics over metal catalysts have been studied extensively using benzene and phenol as reactants (Kiperman, 1986) but not with more complex 2-alkylphenols. With this in mind, we have investigated the kinetics of the nickel-catalyzed gas-phase hydrogenation of o-cresol to the two stereoisomers, cis- and trans-2-methylcyclohexanol. First, a thorough study of the thermodynamics of this reaction is presented, including a discussion of the 0888-5885/89/2628-0693$01.50/0
influence of the various reaction parameters onto the equilibrium concentrations of all reaction components. Subsequently, we shall discuss the influence of the different reaction parameters on the kinetics and the reaction pathway. Finally, in the concluding section, the kinetic results of the o-cresol hydrogenation are compared with corresponding results obtained for the hydrogenation of o-tert-butylphenol.
Experimental Section Materials. The kinetic experiments were performed with o-cresol (STIA, Pratteln, BL, Switzerland, >99% 1, 2-methylcyclohexanone (>99 % 1, 2-methylcyclohexanol (Fluka, Buchs, SG, Switzerland, purum, >98%, x , , - ~=~ 0.27), and o-tert-butylphenol (STIA, Pratteln, BL, Switzerland, >99%). The intermediate product, 2methylcyclohexanone, was produced by hydrogenating o-cresol over palladium (Engelhard, Iselin, NJ, Type 99812). o-Cresol, 2-methylcyclohexanone,and o-tert-butylphenol were purified by distillation. A commercial granular catalyst consisting of 8 wt % nickel supported on silica (Katalysator Werke Huls, Marl, W. Germany, Type H1207) was used for the kinetic measurements. A sieve fraction corresponding to a mean particle size of 2.4 mm was used. The specific surface area of the reduced 0 1989 American Chemical Society
