Table Of ContentPreface
This volume contains peer-reviewed manuscripts describing the scientific and
technological advances presented at the 6 ht Natural Gas Conversion Symposium held in
Girdwood, Alaska on June 17-22, 2001. This symposium continues the tradition of
excellence and the status as the premier technical meeting in this area established by
previous meetings in Auckland, New Zealand (1987), Oslo, Norway (1990), Sydney,
Australia (1993), Kruger National Park, South Africa (1995), and Taormina, Italy (1998).
ehT 6 ht Natural Gas Conversion Symposium is conducted under the overall direction of
the Organizing Committee chaired by Theo Fleisch (BP) and Ertrique Iglesia (University
of California at Berkeley). The Program Committee was responsible for the review,
selection, editing of most of the manuscripts included in this volume. A standing
International Advisory Board has ensured the effective long-term planning and the
continuity and technical excellence of these meetings. ehT members of each of these
groups are listed below. Their essential individual contributions to the success of this
symposium are acknowledged with thanks.
ehT Editors gratefully acknowledge the generous financial support given by the
Sponsors, who are listed below.
ehT Editors acknowledge the talents and dedication of .sM Sarah Key, who patiently
handled the details of the review, acceptance, and submission processes for the
manuscripts in this volume. The Editors also thank the authors of these manuscripts for
their help and patience during the arduous task of assembling this volume.
ehT Editors
Enrique Iglesia, University of California at Berkeley
James J. Spivey, North Carolina State University
oehT H. Fleisch, BP
V1
Program eettimmoC
Enrique Iglesia, Program Chair, University of California at Berkeley
James J. Spivey, Program Co-Chair, North Carolina State University
Carlos Apesteguia, INCAPE (Argentina)
Manfred Baems, Institut fuer Angewandte Chemie Berlin-Adler (Germany)
Calvin Bartholomew, Brigham Young University (USA)
Alexis Bell, University of California ta Berkeley (USA)
Steven Chuang, University of Akron )ASU(
Burtron Davis, University of Kentucky (USA)
Theo Fleisch, PB (USA)
Krijn ed Jong, University of Utrecht (The Netherlands)
Daniel Driscoll, U.S. Department of Energy (USA)
Rocco Fiato, ExxonMobil (USA)
Joep Font Freide (USA)
Karl Gerdes, Chevron (USA)
Anders Holmen, Norwegian University of Science dna Technology (Norway)
Graham Hutchings, University of Wales (United Kingdom)
Ben Jager, SASOL (South Africa)
Eiichi Kikuchi, Waseda University )napaJ(
yaJ Labinger, California Institute of Technology (USA)
Johannes Lercher, Technische Universitat Munchen (Germany)
Wenzhao ,iL Dalian Institute of Chemical Physics )anihC(
Jack Lunsford, Texas A&M University (USA)
Mario Marchionna (Italy)
David Marler, ExxonMobil (USA)
Claude Mirodatos, CNRS-Villeurbane (France)
Julian Ross, University of Limerick )dnalerI(
Jens Rostrup-Nielsen, Haldor Topsoe (Denmark)
Domenico Sanfilippo, Snamprogetti )ylatI(
Lanny Schmidt, University of Minnesota (USA)
Stuart Soled, ExxonMobil (USA)
Samuel Tam (USA)
David Trimm, University of New South Wales (Australia)
Wayne Goodman, Texas A & M University (USA)
Charles Mims, University of Toronto )adanaC(
vii
Organizing eettimmoC
Theo Fleisch, Co-Chair, BP
Enrique Iglesia, Co-Chair, University of California-Berkeley
Dennis Banasiak, Phillips Petroleum
Mark Bendersky, Alaska Science & Technology Foundation
Jeff Bigger, Syntroleum
Rocco Fiato, ExxonMobil
Safaa Fouda, CANMET Energy Technology Centre
Mike Koleda, Energy Frontiers International
Rob Motal, Chevron
Ron Sills, PB
James .J Spivey, North Carolina State University
Venkat Venkataraman, U.S. Department of Energy
International Advisory Board
Carlos Apesteguia (Argentina)
Manfred Baems (Germany)
Theo Fleisch (USA)
Anders Holmen (Norway)
Graham Hutchings (United Kingdom)
Enrique Iglesia (USA)
Ben Jager (South Africa)
Eiichi Kikuchi (Japan)
Wenzhao Li (China)
Jack Lunsford (USA)
Ian Maxwell (The Netherlands)
Claude Mirodatos (France)
Julian Ross (Ireland)
Jens Rostrup-Nielsen (Denmark)
Domenico Sanfilippo (Italy)
Lanny Schmidt (USA)
David Trimm (Australia)
vii
Organizing eettimmoC
Theo Fleisch, Co-Chair, BP
Enrique Iglesia, Co-Chair, University of California-Berkeley
Dennis Banasiak, Phillips Petroleum
Mark Bendersky, Alaska Science & Technology Foundation
Jeff Bigger, Syntroleum
Rocco Fiato, ExxonMobil
Safaa Fouda, CANMET Energy Technology Centre
Mike Koleda, Energy Frontiers International
Rob Motal, Chevron
Ron Sills, PB
James .J Spivey, North Carolina State University
Venkat Venkataraman, U.S. Department of Energy
International Advisory Board
Carlos Apesteguia (Argentina)
Manfred Baems (Germany)
Theo Fleisch (USA)
Anders Holmen (Norway)
Graham Hutchings (United Kingdom)
Enrique Iglesia (USA)
Ben Jager (South Africa)
Eiichi Kikuchi (Japan)
Wenzhao Li (China)
Jack Lunsford (USA)
Ian Maxwell (The Netherlands)
Claude Mirodatos (France)
Julian Ross (Ireland)
Jens Rostrup-Nielsen (Denmark)
Domenico Sanfilippo (Italy)
Lanny Schmidt (USA)
David Trimm (Australia)
iiiv
uo$11opS
Gold
Alaska Science & Technology Foundation
BP
ExxonMobil
Sasol-Chevron
Shell Global Solutions
United States Department of Energy
Silver
Conoco
Phillips Petroleum
Syntroleum
Texaco
Bronze
Air Liquide
ARCO
ITM Syngas Alliance
Natural Resources Canada
Nexant
Studies in Surface Science and Catalysis
J.J. Spivey, E. Iglesia and T.H. Fleisch (Editors)
(cid:14)9 2001 Elsevier Science B.V. All rights reserved.
MODELING MILLISECOND REACTORS
Lanny D. Schmidt
Department of Chemical Engineering and Materials Science
University of Minnesota
Minneapolis MN 55455
ABSTRACT
Catalytic partial oxidation processes at very short contact times have great promise
for new routes to chemical synthesis from alkanes because they are capable of producing
highly nonequilibrium products with no carbon formation using reactors that are much
smaller and simpler than with conventional technology. We summarize some of the
considerations which may be important in modeling and in interpreting partial oxidation
processes. The gradients in these monolith reactors are typically 01 6 K/sec and 10 5 K/cm,
and reactions are fastest in the regions of highest gradients. Therefore a conventional one-
dimensional model may be highly inaccurate to describe these processes, particularly when
used to attempt to decide between different reaction mechanisms. We argue that detailed
modeling which includes detailed descriptions of reactor geometry, gas and solid properties,
and surface and homogeneous reaction kinetics will be necessary to develop reliable
descriptions of these processes. Even with detailed modeling, it may be necessary to
consider the validity of these parameters under extreme reaction conditions.
INTRODUCTION
Oxidation processes in monolithic catalysts exhibit features not observed in
conventional packed bed reactors because they operate with gas flow velocities of--1 rn/sec
with open channel catalyst structures for effective contact times of the gases on the catalyst
of typically 1 millisecond and produce kilograms of product per day with less than a gram
of catalyst. These processes are autothermal and nearly adiabatic because the exothermic
oxidation reactions heat the gases and the catalyst from room temperature to typical
operating temperatures of ~ 1000oc and the rate of heat generation is too large for effective
wall cooling. After lightoff, the reactions usually run to completion of the limiting reactant,
so conversion and selectivities are independent of flow rates over typically an order of
magnitude of residence time.
A recently explored example is methane oxidation to synthesis gas
CH4 + 1/202 ~ CO + 2H2,
which occurs with 100% 02 conversion, >90% CH4 conversion, and >90% selectivity to
CO and H2 (based on C and H respectively) on Rh catalyst coated on it-alumina foam
monolithI. Another example is alkane oxidation to olefins2, for example,
C2H6 + 1/202 >--- C2H4 + H20,
which occurs with 100% 02 conversion, >70% C2H6 conversion, and -85% selectivity to
ethylene on Pt- Sn catalyst coated on cz-alumina foam monolith. As a final example, the
total oxidation of alkanes to CO2 and H20
CH4 + 202 >-- CO2 + 2H20
can be attained with >99% fuel conversion3. The oldest examples of monolith reactors
with millisecond contact times are the Ostwald process to prepare nitric acid by ammonia
oxidation
NH3 + 5/402 ~-- NO + 3/2H20,
and the Andrussow process to prepare HCN4,
CH4 + NH3 + 02 ~-- HCN + 2H20 + H .2
Both of these processes take place on multiple layers of woven Pt-10%Rh gauze catalysts
operating at 800 and 1100~
All of these processes occur with approximately millisecond contact times with the
exothermicity of the reactions providing the energy to heat the gases and catalyst from room
temperature to operating temperatures from 800 to 12~ in <10 3- .s
We3-6 and many others7 have attempted to model these processes in detail to try
to determine the mechanisms by which these product distributions are formed and to find
conditions to optimize a particular product. We argue that these apparently simple
processes are in fact far more complicated than the usual packed bed catalytic reactor
assumptions used for typical modeling. First, the temperatures are sufficiently high that
some homogeneous reaction may be expected to occur, even at very short reaction times.
Second, the gradients in all properties are so large that all conventional assumptions may be
inaccurate. It is the purpose of this manuscript to address these issues.
ONE-DIMENSIONAL MODELS
We first consider the geometry of millisecond reactors. These typically occur in
open monolith catalyst structures which may consist of extruded, foam, or fiber ceramics or
woven or sintered metal structures, as sketched in the left panel of figure .1 All of these
structures can be approximated as a collection of tube wall reactors of length L and channel
diameter d.
One Dimensional Models
Most simple models of these processes have assumed one-dimensional
approximations3 to the geometries of figure .1 Radial mass and heat transfer are included
through effective mass and heat transfer coefficients to the walls. Either plug flow (no axial
mass or heat transfer in the gas) or models including axial diffusion (a boundary value
problem) are assumed, and the resultant model is easy to solve even for many equations.
The simplest approximation to the temperature is to assume a step change from the
feed temperature (25~ to the final catalyst and gas temperature (~1000~ so that the
energy balance can be ignored. Monolith temperatures typically vary by less than 100 ~
)a
I,f 4, / /.//I _L
i,':////i . . + I rT/I"//'///~5
0 L
Z.---
2 r/~,",rl/- ~'/" //~"- J
1ooo
c) "-"
I
I
i
25- , ,
__O|162 0
Q 0 0 )c I rzx / /////~' 0 i
OQ@ i
~ '=~
d) t-.f/z/~///~--'-~
25~ -" i -
z
erugiF .I Left :slenap tnereffiD monolith seirtemoeg approximated by a thgiarts .ebut
Center :slenap elbissoP erutarepmet seliforp down eht .rotcaer Right :slenap elbissoP
snoitacol of homogeneous noitcac~ and .semalf
from front to back, so the assumption of constant wall temperature is reasonable. Gases
should attain the wall temperature within a few channel diameters, so the temperatures
should be constant within less than 1 mm of the entrance. The expected temperature
profiles for gas and catalyst are sketched in curve a of the center panel of figure .1
Wall heat conduction is an important mechanism for backflow of heat which
maintains the monolith and the gas isothermal, and the temperatures can be calculated by
solving simultaneously for the gas and solid temperatures Tg and "Is, still in a one
dimensional approximation.
Upon heating from 25oc to 1000~ the kinematic viscosity increases by more than
a factor of ,01 as do the thermal diffusivity and mass diffusivity. Since reaction occurs very
quickly upon entering the monolith, these variations in properties near the entrance must be
included in any calculation of reactor performance.
Although Reynolds numbers are sufficiently small that laminar flow is a good
approximation and heat transfer coefficients and solid thermal conductivities are sufficiently
large that nearly isothermal gas and solid may be assumed, there may be serious problems
in the one dimensional, plug flow assumption in approximating reactor behavior. First,
most reaction appears to occur within a few tube diameters where the temperatures are
varying strongly, so the large gradients in this region may be significant. Second, the gases
strongly accelerate in the entrance region (typically by factors of 4 to 10), and the
decoupling of the fluid flow from the reaction and temperature equations may lead to
significant errors.
TWO-DIMENSIONAL SIMULATION
We have simulated quantitatively the temperature and velocity profiles for cold gases
entering a hot tube for the reactions and conditions of the methane to syngas[6] and for
ethane to ethylene[7]. These calculations were done using FLUENT to calculate fluid
properties. All fluid properties are properly accounted for including temperature and
mixture variations of diffusivities, thermal conductivity, and viscosity. We used the
Hickman model[5,6] of syngas generation for the surface reaction mechanism.
The CH ,4 temperature, and axial velocity profiles predicted by this model are shown
in figure 2. The calculations shown are for three tube diameters: 0.025, 0.05, and 0.1 cm.
These correspond to 80, 40, and 20 pores per linear inch which are typical foam ceramic
sizes. The region shown is only the first millimeter near the entrance to the hot catalyst
section. This section is preceded by an inert tube which produces a fully developed laminar
flow profile before entering the catalytic section.
The predicted velocity profile is especially interesting. Even though Red<30
throughout the entrance region and the velocity profiles before and within the catalyst
section are parabolic, the axial velocity is not parabolic in the entrance region, and the
velocity has a minimum in the center. A very thin boundary layer is established near the
entrance to the catalytic section as the temperture and all properties vary strongly in very
short distances.
Mass noitcarf HC 4 Dnularo~vroT IK] Axial volooty Ira/s]
.o~ ,.L lSa~ \~'~" -~:"~---- - ~ " "~"-~' ~~ I ~ "~~176176 ---''' I
,o,. o o o, ,o
z )mnr( (cid:12)9 )mar( (cid:12)9 )mar(
.OS20 o+~ -- ~ , ,
r 01251
(mm) 0000 i
0- )121
0- n052 .a2,s~ . .o.2,,~, o'o '" o!5- ~ :~!.o
oo z (nvn) 0 s o.s o o (mm)o S
z z (ram)
o,500 ~ o.soo
o- f
)n~rf( oooo
ojr~.o. o52.o-
o.o 05 o.! )4 oo5 o o 5.0 !o oos.o- 0.0 0 5
z )mar( z )mm( z )mar(
Figure 2. Calculated methane, temperature, and velocity profiles in a single cell of the
reactor for syngas from methane.
The concentrations predicted by one dimensional calculations must therefore be very
different from those calculated using these "exact" calculations. The gradients in
temperature, velocity, and composition are so large in the entrance region that only a
"complete" simulation should be expected to yield more than qualitative conclusions.
REACTIONS
Mechanisms and heat generation
Many modeling studies of these reaction systems have attempted to explain the
reaction mechanism qualitatively by fitting proposed mechanisms to experimental data. In
order to eliminate the large temperature gradients which exist in the millisecond reactor,
these studies have frequently used either dilution with an inert such sa He or 2N or lower
temperatures to slow down the rate of reaction so that intrinsic kinetics can be measured
without interference from temperature or concentration gradients. It is our opinion that the
changes caused by dilution or cooling can effectively disguise the kinetics to change the
overall mechanism significantly.
Two zone model
Syngas formation has been suggested to proceed by either "direct" or "indirect"
mechanisms. In the direct mechanism, reaction is initiated by 4I-IC decomposition followed
by H atom dimerization and C oxidation to produce syngas. The "indirect" mechanism
assumes that the first reaction step is total oxidation to CO2 and H20, and that CO and 2H
are formed by reaction of the remaining CH4 with H20 and 2OC in methane reforming.
These reactions would then occur in two distinct zones: an exothermic combustion zone
where 02 is present and an endothermic zone without 02.
.... It is important to note that, by the indirect mechanism, the primary reaction is
extremely exothermic, while 2OC and H20 reforming are extremely endothermic. These
sequential processes would tend to first pull the temperature up strongly and then pull it
down strongly, so one should expect a large variation in temperature from front to back in
the catalyst, even larger than would be predicted by heat losses through conduction.
Olefins must be made by the initial formation of alkyl species, followed by H
elimination to form the olefin,
C2H6g ~ C2H5 ~ ~---4H2C C2H4
for either a surface reaction or a gas phase (steam cracking) reaction.
Therefore, any indirect mechanism of partial oxidation suggests that there should be
two different zones in the catalyst where different reactions dominate. These predictions are
in strong contrast to a single mildly exothermic process, which is predicted by the direct
processes.
