Kinetics of the water-gas shift reaction over Rh/Al2O3 catalysts

Full text

Applied
Catalysis
A:
General
470 (2014) 31–
44
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at
ScienceDirect
Applied
Catalysis
A:
General
j
ourna
l
h
om
epage:
www.elsevier.com/locate/apcata
Kinetics
of
the
water-gas
shift
reaction
over
Rh/Al2O3catalysts
Canan
Karakayaa,
Robin
Otterstättera,
Lubow
Maierb,
Olaf
Deutschmanna,b,∗
aInstitute
for
Chemical
Technology
and
Polymer
Chemistry,
Karlsruhe
Institute
of
Technology
(KIT),
D-76131
Karlsruhe,
Germany
bInstitute
for
Catalysis
Research
and
Technology,
Karlsruhe
Institute
of
Technology
(KIT),
D-76131
Karlsruhe,
Germany
a
r
t
i
c
l
e
i
n
f
o
Article
history:
Received
29
May
2013
Received
in
revised
form
14
October
2013
Accepted
16
October
2013
Available online 27 October 2013
Keywords:
Water-gas
shift
Kinetics
Rhodium
Carboxyl
Stagnation-flow
reactor
Modeling
a
b
s
t
r
a
c
t
The
kinetics
of
the
water-gas
shift
(WGS)
reaction
over
Rh/Al2O3catalyst
is
studied
experimentally
and
numerically.
Using
the
experimentally
determined
conversion
in
WGS,
reverse
WGS,
and
preferential
oxidation
of
CO
over
a
catalytically
coated
disk
and
over
a
honeycomb
monolith,
a
thermodynamically
consistent
multi-step
reaction
mechanism
with
the
associated
rate
expressions
was
developed.
Both
the
experimental
configurations
were
numerically
simulated
coupling
models
for
the
flow
field
with
this
heterogeneous
reaction
mechanism.
The
main
reaction
path
for
CO2formation
on
this
catalyst
is
concluded
to
be
the
direct
oxidation
of
CO
with
O
species
at
high
temperatures,
whereas
the
formation
of
the
carboxyl
(COOH)
group
is
significant
at
temperature
below
600 ◦C.
The
reaction
kinetics
reproduced
the
experimental
observations,
also
for
the
subsystems
of
hydrogen
and
CO
oxidation.
© 2013 Elsevier B.V. All rights reserved.
1.
Introduction
The
water-gas
shift
(WGS)
reaction
(Eq.
(1))
is
industrially
important
in
many
chemical
processes.
CO
+
H2O
↔
CO2+
H2H◦
298 =
−40.4
kJ/mol
(1)
In
conversion
of
hydrocarbons
by
total
and
partial
oxidation,
steam
reforming,
and
dry
reforming,
WGS
is
one
of
the
crucial
reactions
that
determine
the
overall
product
selectivity
[1–4].
WGS
technology
is
used
to
purify
reformates
from
CO
to
oper-
ated
low-temperature
fuel
cells
and
ammonia
synthesis
plants
with
hydrogen
[5–8].
In
exhaust-gas
after-treatment,
WGS
occurs
in
the
catalytic
converter
with
significant
effects
on
the
reduction
of
CO
emissions
and
thermal
stability
of
the
catalysts
[4,9,10].
1.1.
Low
temperature
applications:
synthesis
gas
purifications
In
commercial
applications,
in
which
the
removal
of
CO
from
the
synthesis
gas
stream
is
necessary,
the
WGS
reaction
takes
place
in
two
steps,
involving
high-temperature
shift
(HTS)
and
low-
temperature
shift
(LTS)
processes
[8].
Iron
oxide
and
chromium
oxide
catalysts
are
used
for
the
HTS
in
the
temperature
range
of
310–450 ◦C
[11].
For
the
LTS
reaction,
usually
carried
out
as
a
∗Corresponding
author
at:
Institute
for
Chemical
Technology
and
Polymer
Chem-
istry,
Karlsruhe
Institute
of
Technology
(KIT),
D-76131
Karlsruhe,
Germany.
Tel.:
+49
721
608
43064;
fax:
+49
721
608
44805.
E-mail
address:
deutschmann@kit.edu
(O.
Deutschmann).
second
step
after
HTS,
the
catalysts
are
based
on
zinc
and
copper
oxides
operated
between
200 ◦C
and
250 ◦C
[11,12].
There
is
an
interest
in
alternative
robust
WGS
catalysts,
because
LTS
catalysts
are
sensitive
to
air
and
steam
and
they
are
easily
poi-
soned
by
sulfur
[10].
Besides,
conventional
catalysts
imply
large
reactor
volumes,
in
which
mostly
a
packed-bed
reactor
configu-
ration
is
used.
The
large
size
limits
their
application
to
on-board
reforming
technologies
where
smaller
systems
are
required
[13].
Noble
metal-supported
catalysts
show
promising
activity
because
of
their
high
stability
in
low
and
high
temperature
regimes
and
high
tolerance
capacity
to
impurities
[4,10].
Pt,
Rh,
Pd,
and
Au-
promoted
catalysts
on
different
supports
(ceria,
La2O3)
have
been
investigated
for
WGS
reactions
[4,6].
Among
these
metals,
Rh
is
a
promising
catalyst
exhibiting
high
turnover
frequencies
(TOF)
and
stability.
In
understanding
the
reaction
mechanism
of
WGS
over
Rh,
the
interaction
of
the
catalyst
with
the
support
needs
to
be
considered
for
support
materials
like
ceria
as
well
[5,14,15].
Many
theoretical
and
experimental
studies
are
conducted
to
elucidate
the
reaction
kinetics
of
WGS
resulting
in
three
general
reaction
mechanisms.
The
first
one
relies
on
the
assumptions
that
the
redox
mechanism
is
dominant
and
that
CO2is
generated
by
a
reaction
of
CO,
which
is
adsorbed
on
the
metal
and
CeO2sup-
port,
and
that
H2is
formed
via
re-oxidation
of
Ce
with
H2O
[1].
In
the
second
one,
it
is
assumed
that
a
carboxyl
species
plays
a
decisive
role
[5].
For
supports
containing
CeO2,
regardless
of
the
metal
type
(Pt
or
Rh),
subtraction
of
H
from
water
leads
to
OH
for-
mation
on
the
support,
which
is
a
slow
step,
and
CO2is
primarily
formed
via
a
carboxyl
(COOH)
intermediate.
The
reaction
between
the
chemisorbed
CO
and
O
is
negligible
[5].
In
the
third
mechanism
0926-860X/$
–
see
front
matter ©
2013 Elsevier B.V. All rights reserved.
http://dx.doi.org/10.1016/j.apcata.2013.10.030

32 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
proposed,
a
formate
species
(HCOO)
is
the
decisive
intermediate
formed
by
adsorbed
CO
and
OH.
Olympiou
et
al.
studied
WGS
reactions
at
350–550 ◦C
on
Pt,
Pd,
and
Rh
supported
on
␥-Al2O3with
similar
noble
metal
dis-
persions
[8].
According
to
their
study
carried
out
at
steady
state
conditions,
adsorbed
HCOO
is
formed
upon
surface
diffusion
of
H
and
OH
present
on
the
alumina
support.
This
reaction
is
considered
as
a
slow
step
in
WGS
and
the
intermediate
HCOO
resides
on
the
alumina
support
and/or
the
metal.
A
variety
of
other
studies
pro-
posed
the
presence
of
a
formate
species
on
the
reduced
ceria,
e.g.,
on
Au/CeO2[6],
Rh/MgO,
Rh/ZnO,
and
Rh/CeO2[16],
Pt/MgO
and
Pt/ZrO
[17],
and
Pt/CeO2[18].
HCOO
and
the
carboxyl
species
(COOH)
are
isomers
and
both
the
structures
were
detected
experimentally
using
FTIR
and
DRIFTS
and
steady-state
isotropic
transient
kinetic
analysis
(SSITKA)
anal-
ysis
[19,20].
However,
there
is
a
disagreement
in
the
interpretation
of
the
experimental
data
as
to
whether
it
is
a
carboxyl
or
a
formate
species.
According
to
Hilaire
et
al.
[21],
the
band
in
the
spectra,
which
appears
in
the
range
of
1000–1700
cm−1originates
from
OCO
asymmetric
and
symmetric
stretching
vibrations.
Therefore,
it
is
difficult
to
distinguish
whether
the
bands
in
this
region
corre-
spond
to
carbonates
or
formates,
because
both
species
contain
OCO
vibrations.
Besides,
theoretical
calculations
favor
the
formation
of
the
carboxyl
species
[2,3].
1.2.
High
temperature
applications:
effect
of
WGS
reaction
on
synthesis
gas
production
For
high-temperature
applications,
Rh
supported
on
Al2O3is
a
well-known
catalyst
for
synthesis
gas
production
from
hydrocar-
bons
[22–24].
Two
routes
to
synthesis
gas
are
proposed
[25–28].
The
direct
route
postulates
the
formation
of
H2and
CO
via
partial
oxidation
of
CH4in
the
presence
of
gas-phase
O2[29],
whereas,
the
indirect
route
favors
a
two-zone
model,
in
which
first
total
oxida-
tion
takes
place
followed
by
steam
reforming
of
hydrocarbons
and
WGS
[28,30].
The
literature
addresses
different
types
of
mechanism
for
WGS
reaction
and
its
effect
on
catalytic
partial
oxidation
(CPOx)
and
reforming
of
hydrocarbons
depending
on
the
reactor
types,
in
which
the
kinetic
investigations
are
carried
out,
and
on
operat-
ing
conditions
and
analytics
used
[23,28,31–33].
Horn
et
al.
[22]
studied
the
effect
of
the
WGS
in
CPOx
over
a
Rh
catalyst
moni-
toring
concentration
and
temperature
profiles
along
the
catalyst
bed
under
transient
and
steady-state
conditions.
They
concluded
that
WGS
has
only
a
minor
effect
since
the
amount
of
CO2does
not
change
in
the
absence
of
O2and
that
the
contribution
of
the
WGS
varies
with
the
C/O
feed
ratio.
Michael
et
al.
[34]
have
studied
the
effects
of
H2O
and
CO2as
co-reactants
on
CPOx
by
also
resolv-
ing
spatial
concentration
profiles.
They
found
that
feeding
of
H2O
as
a
co-reactant
has
no
effect
on
conversion
of
CH4,
however,
the
selectivity
is
affected
due
to
WGS.
Maestri
et
al.
[23]
proposed
that
WGS
is
in
equilibrium
and
the
adsorbed
OH
species
is
the
main
oxygen
source
to
form
CO
via
adsorbed
carbon
generated
by
pyrol-
ysis
of
methane.
It
was
concluded
that
steam
reforming
(SR)
and
dry
reforming
(DR)
of
methane
always
occur
with
WGS
and
that
the
formation
of
CO2is
mainly
due
to
the
dissociation
of
the
car-
boxyl
(COOH)
species.
Wang
et
al.
[33]
studied
the
WGS
reaction
over
Rh/Al2O3at
600 ◦C
and
800 ◦C
under
vacuum
by
using
a
TAP
(Temporal
Analysis
of
Products)
reactor.
They
concluded
that
CO2
formation
mainly
occurs
via
fast
oxidation
of
CO
with
adsorbed
oxygen
or
via
a
nucleophilic
attack
of
adsorbed
OH
groups
on
the
alumina
support,
which
has
an
inverse
spill-over
effect
because
the
water
can
dissociatively
adsorb
on
alumina
by
producing
O(s)
and
OH(s)
species;
the
suffix
(s)
denotes
adsorbed
species
in
this
article.
The
isotopic
tracer
studies
of
Wei
et
al.
show
that
H2O
dissociation
Fig.
1.
Stagnation-flow
field
and
microprobe
sampling
technique.
is
quasi-equilibrated,
i.e.,
H(s)
and
OH(s)
species
recombine
rapidly
to
form
gas-phase
hydrogen
and
water
[32].
1.3.
Reactors
for
kinetic
studies
There
are
many
different
reactors
used
for
kinetic
studies;
only
few
remarks
will
be
made
here.
Despite
their
simplicity,
kinetics
derived
from
integral,
tubular
flow
reactors
containing
powders,
pellets,
foams
or
honeycomb
monoliths
as
catalysts
are
usually
difficult
to
be
interpreted
on
a
mechanistic
level
because
of
the
overlapping
of
mass
and
heat
transfer
in
at
least
two
dimensions
with
catalytic
reactions.
Gas-phase
concentration
and
also
tem-
perature
can
strongly
vary
within
a
millimeter
in
radial
and
axial
directions.
Surface
coverages
may
strongly
vary
as
well
in
flow
directions
[22,28,35,36].
In
modeling
of
these
reactors,
simplifying
assumptions
are
usually
made
to
describe
the
heat
and
mass
trans-
port
effects
[37,38].
These
simplifications
can
be
crucial
in
deriving
the
intrinsic
kinetics
as
soon
as
parallel
and/or
sequential
catalytic
reactions
as
well
as
gas-phase
reactions
occur.
As
an
alternative
to
tubular
flow
reactors,
the
TAP
reactor
is
used
to
investigate
the
reaction
kinetics
under
isothermal
conditions.
The
only
transport
process
here
is
realized
by
Knudsen
diffusion.
However,
the
TAP
reactor
is
usually
operated
under
low-pressure
with
small
amount
of
reactive
mixtures
[33,39].
The
stagnation-flow
reactor
(SFR)
(Fig.
1)
is
another
useful
con-
cept
to
investigate
catalytic
kinetics.
Even
though
the
flow
field
of
the
entire
reactor
is
two-dimensional
(2D)
with
axial
and
radial
spatial
coordinate,
a
potential
flow
can
be
established
leading
to
a
one-dimensional
(1D)
boundary-layer
over
the
catalyst
[40,41].
Hence,
the
entire
catalyst
(except
at
the
edges)
is
exposed
to
the
identical
gas-phase
leading
to
no
lateral
variations
of
the
cata-
lyst
surface
coverage
in
general.
A
number
of
groups
have
studied
catalytic
kinetics
in
stagnation-flow
reactors
to
better
understand
catalytic
combustion
[42–46],
partial
oxidation
and
steam
reform-
ing
on
noble
metals
[47–49]
as
well
as
diamond
growth
[50].
Recently,
McGuire
et
al.
[49]
studied
dry
reforming
of
methane
over
Rh
supported
on
strontium-substituted
hexaaluminates.
They
used
a
microprobe
sampling
technique
which
enables
a
resolution
of
the
concentration
profiles
in
the
gas-phase
boundary-layer
adjacent
to
the
catalytic
surface.

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 33
Fig.
2.
Tubular
flow
reactor
configuration.
In
this
paper,
we
present
an
experimental
and
modeling
study
on
WGS
reaction
kinetics
over
a
Rh/Al2O3catalyst
by
using
a
stagnation-flow
reactor
and
a
tubular
flow
reactor
with
catalyt-
ically
coated
disks
and
honeycomb
monoliths,
respectively.
This
work
continues
our
recent
study
on
H2-O2reactions
over
Rh/Al2O3
[51].
In
the
present
study,
we
extend
the
surface
reaction
mecha-
nism
by
COxspecies
to
describe
reactions
of
the
CO/CO2/H2/H2O/O2
system
over
Rh/Al2O3.
WGS,
RWGS,
and
preferential
oxidation
(PROx)
of
CO
are
experimentally
studied
in
a
stagnation-flow
reac-
tor
at
600–800 ◦C.
The
data
obtained
are
then
used
to
develop
a
thermodynamically
consistent,
multi-step
reaction
mechanism
which
is
able
to
describe
low
and
high-temperature
WGS
and
RWGS
reactions.
The
wider
applicability
of
the
reaction
mechanism
is
fur-
ther
tested
by
conducting
experiments
in
a
laboratory-scale
tubular
flow
reactor
with
honeycomb
coated
catalyst
in
the
temperature
range
of
200–900 ◦C.
Both
experiments
are
numerically
simulated
using
detailed
chemistry
and
flow
and
transport
models
to
distin-
guish
between
kinetic
and
transport
effects
on
the
overall
reaction
rate.
2.
Experimental
2.1.
Catalytic
reactors
In
this
section,
we
shortly
describe
both
the
lab-scale
reactors
used
in
our
kinetic
study.
Detailed
descriptions
of
the
stagnation-
flow
reactor
and
the
tubular
flow
reactor
are
given
in
our
previous
publications
[24,51,52].
2.1.1.
Stagnation-flow
reactor
The
stagnation-flow
reactor
can
be
operated
at
100–1100
mbar
and
temperatures
between
25
and
900 ◦C.
The
pressure
is
con-
trolled
by
a
butterfly
valve
(MKS,
T3BIA).
Gaseous
reactants
such
as
O2,
CO,
H2,
CH4,
C2H6,
and
C3H8and
vaporized
liquids
such
as
water,
alcohols,
and
hydrocarbons
can
be
fed
to
the
reactor.
The
reaction
chamber
is
made
of
stainless
steel
and
is
isolated
from
the
ambient
by
an
ethylene
glycol
(40 ◦C)
circulation
to
keep
the
reactor
wall
at
constant
temperature.
Gas
lines
are
also
heated
(150 ◦C)
to
prevent
the
condensation
of
the
liquids.
Gases,
here
Ar,
CO2,
CO,
H2,
and
O2,
are
dosed
via
mass
flow
controllers
(MFC,
Bronkhorst).
Liquid
H2O
is
dosed
via
a
liquid
mass
flow
controller
(LFC,
Bronkhorst)
and
is
evaporated
by
using
a
micro-structured
nozzle
evaporation
technique
[53].
Gases
are
premixed
before
they
enter
the
reaction
zone.
The
gas
mixture
passes
through
a
porous
cylindrical
frit,
which
creates
axial
velocity
at
the
inlet,
which
is
uni-
form
in
radial
direction
(Fig.
1).
The
flow
configuration
is
oriented
upwards
so
that
the
buoyancy
effect
on
the
stagnation-flow
field
is
diminished.
The
gases
finally
flow
around
the
catalytic
disk
and
are
exhausted
through
an
annular
pipe
on
the
backside
of
the
catalytic
disk.
The
reactor’s
operating
pressure
is
selected
to
be
500
mbar
in
this
study
in
order
to
obtain
an
optimum
boundary-layer
thickness
for
the
system
under
consideration.
2.1.2.
Tubular
flow
reactor
with
catalytic
monolith
A
quartz
tube
of
19.5
mm
inner
diameter
(ID)
serves
as
flow
reactor,
which
is
equipped
with
a
commercial
Rh/␥-Al2O3coated
honeycomb
monolith,
600
channels
per
square
inch
(cpsi)
(Fig.
2).
The
monolith
of
19
mm
outer
diameter
(OD)
is
10
mm
in
length.
Upstream
and
downstream
the
catalytic
monolith,
a
foam
mono-
lith
(␥-Al2O3,
85
pore
per
inch
(ppi))
and
a
honeycomb
monolith
(␥-Al2O3,
400
cpsi),
respectively,
are
placed
close
to
the
catalyst.
Both
these
monoliths
are
uncoated,
1
cm
in
length,
and
serve
as
heat
shields,
holders
for
thermocouples,
and
flow
straightener
(foam).
The
K-type
(front)
and
N-type
(back)
thermocouple
are
placed
close
to
the
catalyst.
The
reactor
is
placed
in
an
electrically
heated
fur-
nace.
Reactive
gases
(H2(purity
5.0),
CO
(purity
4.7),
CO2(purity
5.0),
O2(purity
5.0))
and
N2as
dilution
(N2(purity
5.0))
are
dosed
via
mass
flow
controllers
(Bronkhorst).
Water
supply
controlled
by
a
liquid
flow
controller
(LFC,
Bronkhorst)
is
evaporated
(evaporator:
Bronkhorst
CEM)
and
diluted
by
N2before
entering
into
the
reactor.
Lines
are
heated
up
to
190 ◦C
to
avoid
condensation.
2.2.
Catalyst
preparation
2.2.1.
Catalytic
disk
of
SFR
A
stagnation
disk
(5.5
cm
in
diameter)
is
manufactured
using
a
high-temperature
castable
ceramic
resin
and
hardener
(Cotron-
ics
Corporation).
An
R-type
(rhodium–platinum)
thermocouple
of
0.2
mm
thickness
(Omega
Newport)
is
placed
in
the
center
of
the
disk
during
casting.
The
surface
is
first
dried
at
130 ◦C
for
2
h
and
then
cured
at
600 ◦C
for
further
2
h
before
coating.
The
resulting
ceramic
structure
consists
of
99
wt.%
Al2O3.
Aqueous
solution
of
rhodium
(III)
nitrate
(Umicore)
(9
wt.%
Rh)
and
boehmite
(AlOOH)
solutions
(20
wt.%
boehmite)
are
mixed
to
obtain
a
5
wt.%
Rh/Al2O3composition.
The
suspension
is
diluted
with
water
and
provided
to
the
disk
by
spin-spray
technique
to
ensure
a
homogeneously
distributed
active
phase
on
the
surface.
The
stagnation
disk
is
held
on
a
rotary
support
which
spins
at
1000
rpm.
The
suspension
is
sprayed
by
compressed
air
via
a
spray
gun.
The
surface
is
dried
at
130 ◦C
for
10
min,
and
the
procedure
is
repeated
until
the
desired
coating
thickness
of
100–130
␮m
is
reached.
The
coated
stagnation
disk
is
then
calcined
at
700 ◦C
for
2
h
in
ambient
air.
Two
discs
are
coated
with
this
technique
described
to
assure
reproduction
of
the
catalytic
disk.
Prior
to
the
measurements,
the
surface
is
oxidized
in
a
flow
of
5
vol.%
O2in
an
Ar
atmosphere
at
500 ◦C
for
2
h.
The
result-
ing
rhodium
oxide
phase
is
reduced
under
Ar-diluted
5
vol.%
H2
at
500 ◦C
for
2
h.
2.2.2.
Catalytic
monolith
In
the
tubular
flow
reactor,
a
commercial
1.2
wt.%
Rh/␥-Al2O3
coated
honeycomb
(600
cpsi)
monolith
made
out
of
cordierite
is
used
without
modification.
Prior
to
the
measurements,
the
catalyst
is
oxidized
at
600 ◦C
by
21
vol.%
O2diluted
in
N2and
reduced
at
400 ◦C
by
1
vol.%
H2diluted
in
N2for
45
min.
2.3.
Catalyst
characterization
The
coating
thickness
and
the
homogeneity
of
the
coating
layer
on
the
stagnation
disk
are
investigated
by
means
of
a
light
microscopy
(LM:
Rechert
MEF4A).
For
the
investigation
of
Rh
particles
and
the
washcoat
structure,
scanning
electron
microscopy
(SEM:
Hitachi
S570)
is
applied
in
combination
with
energy-dispersive
X-ray
spectroscopy
(EDX)
and
high
resolution
transmission
electron
microscopy
(HR-TEM:
Philips
CM200
FEG).
Metal
dispersions
of
the
catalysts
(stagnation
disk
and
mono-
lithic
catalyst)
are
measured
before
the
kinetic
study
by
CO
chemisorption
using
temperature-programmed
desorption
(TPD)
operated
under
atmospheric
pressure
and
continuous-flow
[54].

34 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
2.4.
Measurements
procedure
2.4.1.
Stagnation-flow
reactor
A
series
of
measurements
on
WGS
and
RWGS
kinetics
are
carried
out
in
the
stagnation-flow
reactor
under
varying
feed
composi-
tion
and
temperature.
Preferential
oxidation
of
CO
with
varying
H2
and
O2mixtures
is
also
examined.
A
summary
of
all
experimental
conditions
is
given
in
Table
1.
The
concentration
profiles
of
the
species
are
measured
within
the
boundary
layer
as
a
function
of
distance
from
the
catalytic
disk
using
a
microprobe
sampling
technique.
The
microprobe
is
made
out
of
quartz
(3
mm
OD,
1
mm
ID)
and
exhibits
a
50
␮m
opening
on
the
tip
(OD
<
0.4
mm)
with
a
bend
angle
of
15◦to
achieve
close
contact
to
the
catalyst
surface
without
interference.
The
probe
is
positioned
at
a
radial
distance
of
nearly
0.8
cm
from
the
center
of
the
catalyst
to
disturb
the
flow
field
as
little
as
possible
(Fig.
1).
The
microprobe
is
positioned
by
a
step
motor
controller
(Thermionics
Northwest,
Inc.).
In
the
beginning
of
the
profile
measurement,
the
microprobe
is
positioned
on
the
surface,
and
the
contact
between
probe
and
surface
is
determined
visually.
Starting
from
the
zero
position,
the
probe
is
moved
downward
through
the
boundary
layer
in
steps
of
0.5
mm.
The
concentration
profiles
of
the
species
are
detected
at
steady-state
temperature.
A
quadrupole
mass
spectrometer
(MS)
(Airsense
500,
V&F)
is
used
for
detection
of
O2,
an
electron
pulse
ionization
mass
spec-
trometer
(H-Sense,
V&F)
is
used
for
H2detection,
and
a
FTIR
(MKS)
for
CO,
CO2,
and
CH4.
All
analytical
tools
are
used
simultaneously.
H2O
concentration
is
calculated
via
oxygen
mass
balance.
At
each
axial
measurement
point,
sampling
is
repeated
four
times
with
a
pause
of
at
least
4
min
to
ensure
repeatability.
Preferential
oxidation
of
CO
is
studied
in
the
presence
of
H2
at
varying
O2concentrations
at
600 ◦C
(Cases
1–3
in
Table
1).
Ar-
diluted
gas
mixtures
are
fed
to
the
reactor
with
a
total
flow
rate
of
15.5
standard
liter
per
minute
(slpm:
defined
as
the
standard
at
20 ◦C
and
1
atm).
The
inlet
flow
rate
results
in
a
homogenous
axial
flow
velocity
at
the
gas
inlet
of
51
cm/s
at
500
mbar.
Boundary-layer
profiles
of
the
species
are
measured
at
steady
state.
The
WGS
reaction
is
studied
at
600,
735,
and
800 ◦C
with
molar
steam/carbon
(S/C)
ratio
of
1.1
for
all
cases
(Cases
4–6
in
Table
1).
The
calculated
flow
velocity
is
74
cm/s
at
the
inlet
of
the
reac-
tion
chamber
at
500
mbar
and
an
inlet
temperature
for
150 ◦C.
The
RWGS
is
studied
at
600
and
700 ◦C
(Cases
7
and
8
in
Table
1).
Ar-
diluted
CO2and
H2gases
with
a
molar
CO2/H2ratio
of
unity
are
fed
to
the
reactor
at
an
inlet
temperature
of
40 ◦C
at
500
mbar.
2.4.2.
Tubular
flow
reactor
with
catalytic
monolith
WGS
and
RWGS
are
studied
in
the
tubular
flow
reactor
using
a
commercial
Rh/Al2O3catalyst
(Cases
9–12
in
Table
1)
CO,
CO2,
H2O,
CH4and
H2concentrations
in
the
outlet
gas
are
detected
by
means
of
an
FTIR
(MKS)
and
an
electron
pulse
ionization
mass
spec-
trometer
(H-Sense,
V&F),
respectively.
The
reaction
is
carried
out
in
a
temperature
range
of
200–900 ◦C
by
applying
a
heating
rate
of
7◦C/min.
The
co-feeding
effects
of
the
products
CO2and
CO
for
the
WGS
and
RWGS
reaction,
respectively,
are
also
investigated.
All
measurements
are
carried
out
at
atmospheric
pressure
with
a
total
flow
rate
of
5
slpm
that
corresponds
to
a
gas
hourly
space
velocity
(GHSV)
of
100,000
h−1.
3.
Modeling
approach
3.1.
Flow
models
3.1.1.
Stagnation-flow
reactor
model
The
flow
impinges
on
the
flat
catalytic
disk
and
forms
an
axisym-
metric
potential
flow
field
with
a
one-dimensional
boundary-layer
of
thickness
ıs(Fig.
1)
as
long
as
the
inlet
velocity,
the
distance
between
inlet
and
disk,
and
the
inlet
tube
and
disk
diameters
are
chosen
in
the
right
dimensions.
Then,
the
flow
field
of
the
SFR
can
accurately
be
modeled
by
a
1D
approach
[41,55,56],
in
which
the
temperature,
concentration,
and
velocity
depend
only
on
the
spatial
coordinate
in
axial
flow
direction
(z),
and
on
time
for
transient
processes.
With
this
transformation,
the
computational
solution
of
reactive
flows
over
catalytic
disks
can
be
reduced
to
a
one-dimensional
(z)
ordinary
differential
equation
boundary-value
problems,
including
fluid
mechanics,
multi-component
diffusive
transport,
and
elementary
chemical
kinetics
resulting
in
the
governing
equations
for
mass
(2),
axial
momentum
(3),
radial
momentum
(4),
energy
(5),
chemical
species
(6)
being
well-
documented
[41,57].
∂(u)
∂z +
2V
=
0.
(2)
u ∂u
∂z +∂p
∂z −
2∂V
∂z −4
3
∂
∂z ∂u
∂z +4
3
∂
∂z (V).
(3)
u ∂V
∂z +
V2−∂
∂z ∂V
∂z +

=
0,

=1
r
∂p
∂r .
(4)
cpu∂T
∂z −∂
∂z ∂T
∂z −
u∂p
∂z +

Ng

k=1
(cpYkVk)∂T
∂z
+
Ng

k=1
hkWk˙ωk=
0.
(5)
u ∂Yk
∂z +∂
∂z (YkVk)
−
Wk˙ωk=
0
k
=
1,
.
.
.,
Ng.
(6)
Symbol
nomenclature:
temperature
T,
species
mass
fractions
Yk,
axial
velocity
u,
scaled
radial
velocity
V
=
v/r
in
which
v
is
the
radial
velocity
and
r
is
the
radial
coordinate,
molecular
viscosity
,
ther-
mal
conductivity
,
mixture-specific
heat
cp,
molecular
weight
Wk,
enthalpy
of
formation
hk,
molar
reaction
rate
˙ωkof
species
k
and
the
number
of
gas
phase
species
Ng·
Vkrepresents
the
diffusion
velocity
of
species
k
in
the
axial
direction
by
Vk=1
Xk¯
W
Ng

k
/=
j
WjDkj
∂Xj
∂z −
DT
k
Yk
1
T
∂T
∂z .
(7)
with
Xj,
mole
fractions
of
species
j;¯
W,
mean
molecular
weight;
Dkj,
multi-component
diffusion
coefficients;
DT
k,
thermal
diffusion
coefficient
of
species.
Boundary
conditions
at
the
gas-catalyst
interface
(z
=
0)
are:
u
=
0
V
=
0
T
=
Ts(8)
The
boundary-conditions
for
species
governing
equations
on
the
catalyst
surface
are
discussed
in
Section
3.2.
The
boundary
conditions
at
the
inlet
of
the
reactor,
in
this
study
at
z
=
3.9
cm,
are
given
as:
T
=
Tin u
=
uin V
=
0
Yk=
Yk,in (9)
The
numerical
solution
of
this
set
of
differential
and
algebraic
equations
is
accomplished
using
the
CHEMKIN
SPIN
codes
[58]
and
DETCHEMSTAG[59].
3.1.2.
Tubular
reactor
model
The
model
assumes
that
radial
temperature
profiles
over
the
honeycomb
monolith
can
be
neglected
and
radially
uniform
tem-
perature,
composition,
and
velocity
profiles
are
established
in
the

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 35
Table
1
Experimental
conditions
for
preferential
oxidation,
WGS
and
RWGS
reactions.
Reaction
Temperature
(◦C)
H2(vol.%)
CO
(vol.%)
O2(vol.%)
H2O
(vol.%)
CO2(vol.%)
Preferential
oxidation Case
1
600
6.00
5.20
5.20
–
–
Case
2 600
6.00 5.20
2.65
–
–
Case
3
600
2.57
2.57
4.00
–
–
WGS Case
4
600
–
4.75
–
5.18
–
Case
5
735
–
4.75
–
5.18
–
Case
6
800
–
4.75
–
5.18
–
RWGS Case
7
600
5.20
–
–
–
5.20
Case
8 700
5.20 –
–
–
5.20
WGSaCase
9
200–900
11.20
11.60
2.00
Case
10
200–900
10.08
10.49
RWGSaCase
11
200–900
10.40
2.04
10.88
Case
12
200–900
10.03
9.52
aReactions
are
conducted
in
tubular
flow
reactor.
front
of
the
catalytic
monolith
due
to
the
foam
placed
in
front
of
the
catalyst.
Then,
all
channels
of
the
honeycomb
behave
alike
and
only
one
single
channel
needs
to
be
analyzed.
Numerical
simulation
of
this
single
channel
is
based
on
the
2D
boundary-layer
model
[60],
for
details
on
the
governing
equations
we
refer
to
the
literature
and
the
manual
of
the
software
applied.
The
monolith
channel
has
a
hydraulic
diameter
of
1.0
mm.
The
axial
uniform
flow
velocity
at
the
standard
conditions
is
calculated
to
be
28.3
cm/s.
The
software
DETCHEMCHANNEL is
used
for
the
numerical
simulation
of
the
single
channel
flow
[59,61].
3.2.
Modeling
the
catalytic
reaction
rate
Since
the
temperatures
(<900 ◦C)
and
the
residence
times
at
this
temperature
(few
milliseconds)
are
rather
low
in
this
study,
gas-phase
reactions
are
not
likely
to
occur
and
catalytic
reactions
only
need
to
be
considered
in
the
model.
The
mean-field
approx-
imation
is
applied,
i.e.,
the
surface
is
characterized
by
coverages,
which
depend
on
the
axial
position
(flow
direction)
in
the
channel
configuration
but
have
no
local
variations
in
the
stagnation
flow
configuration
due
to
a
radially
uniform
boundary
layer.
The
molar
reaction
rates
(˙
sk)
of
adsorbed
and
gas-phase
species
are
then
a
function
of
the
local
coverage
and
gas-phase
concentrations
at
the
gas-catalyst
[57,62]:
˙
sk=
Ks

j=1
vkjkfj
Ng+Ns+Nb

n=1
cvnj
n(10)
Here,
Ksis
the
number
of
surface
reactions
and
cnare
the
species
concentrations,
given
in
mol/m2for
Nssurface
species
and
in
mol/m3for
Nggas
phase
species
and
Nbbulk
species.
The
con-
centration
of
adsorbed
species
ciis
related
to
the
coverage,
i,
by
the
surface
site
density

(ci=
I*

).
The
surface
site
density,
i.e.
the
number
of
adsorption
sites
per
surface
area
of
the
Rh
particle,
is
estimated
to
be
2.72
×
10−9mol
cm−2.
kfj are
the
forward
rate
coefficients
of
reaction
j
with
stoichiometric
coefficients
of
vkj.
The
net
production
rate
of
each
chemical
species
in
the
gas
phase
is
balanced
with
the
diffusive
flux
of
that
species
in
the
gas
phase
at
steady-state
conditions
by
assuming
that
no
deposition
or
ablation
of
chemical
species
occurs
on/from
the
catalyst
surface:
YkVk=
Fcat/geo ˙
skWkk.
(11)
The
term
Fcat/geo,
which
is
derived
from
the
CO
chemisorption
mea-
surements,
is
introduced
as
a
scaling
factor
for
the
active
catalytic
surface
area
Acatalyst and
the
geometric
surface
area
Ageometric of
the
stagnation
disk
or
the
channel
wall,
Fcat/geo =Acatalyst
Ageometric
.
(12)
Internal
mass
transport
limitations
inside
the
washcoat
layer
is
accounted
for
by
introduction
of
the
effectiveness
factor
(k)
into
Eq.
(11).
The
effectiveness
factor
is
modeled
in
terms
of
a
Thiele
modulus,
in
which
the
effective
diffusion
coefficient
accounts
for
Knudsen
diffusion
and
molecular
diffusion
depending
on
the
pore
size
distribution
[59,63].
For
a
catalyst
with
a
slab
geometry,
the
washcoat
is
assumed
to
be
thick
enough
to
sustain
the
zero
con-
centration
gradient
at
the
deepest
point
of
the
washocat
(L).
On
these
assumptions,
the
effectiveness
factor
is
defined
as
k=tan
˚k
˚k
and
˚k=
Lkf
Defk
.
(13)
˚kis
defined
to
be
the
Thiele
modulus,
while
kand
˚kare
calcu-
lated
for
the
user-defined
species
(CO
is
chosen
in
this
study).
The
effective
diffusion
coefficient
(Defk)
can
be
calculated
by
taking
into
account
the
Knudsen
diffusion
coefficient
(DKnudi)
and
the
molec-
ular
diffusion
coefficient
(Dmolk)
of
the
species
k
in
the
mixture
[41,63].
Defk=εp
TDk.
(14)
1
Dk
=1
Dmolk
+1
DKnudi
.
(15)
In
the
above
equations,
εpand
T
represent
the
porosity
and
tor-
tuosity
respectively.
For
more
details
on
modeling
internal
mass
transfer
limitation
in
stagnation
flow
reactors
it
is
referred
to
our
recent
study
[79].
4.
Results
and
discussion
4.1.
Catalyst
characterization
The
in-house
made
catalyst
used
in
the
stagnation-flow
reac-
tor
experiments
is
characterized
after
the
catalytic
measurements.
The
light
microscope
study
of
the
catalyst
reveals
the
formation
of
a
uniform
∼100
␮m
thick
catalyst
layer
on
the
stagnation
disk.
The
SEM
images
reveal
a
porous
alumina
structure
with
Rh
parti-
cles
approximately
100
nm
in
diameter.
Additionally,
smaller
Rh
particles
with
a
diameter
of
15–50
nm
are
detected
in
HR-TEM
investigations
(Fig.
3).
Based
on
the
chemisorption
measurements
with
the
assumption
of
1:1
adsorption
stoichiometry
between
Rh
and
CO
molecules,
the
active
catalyst
surface
area
is
measured
to
be

36 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
Fig.
3.
SEM
(a)
and
HR-TEM
(b)
images
of
the
catalyst
(after
the
catalytic
reactions)
used
in
stagnation-flow
experiments.
0.21
m2/g.
With
this
information,
the
model
parameter
Fcat/geo for
the
in-house
made
catalyst
is
calculated
to
be
30.
The
effectiveness
factor
is
estimated
to
be
∼0.03
based
on
the
washcoat
parameters.
For
the
commercial
honeycomb
catalyst
used
in
the
tubular
flow
experiments,
5%
Rh
dispersion
is
derived
from
the
chemisorption
measurement
resulting
in
an
active
metal
surface
area
of
0.22
m2/g,
and
Fcat/geo =
19.
The
washcoat
thickness
is
measured
to
be
40
␮m
by
LM.
The
effectiveness
factor
was
computed
to
be
0.1.
Porosity
and
tortuosity
values
are
assumed
to
be
0.6
and
3,
respectively.
The
pore
diameter
is
measured
to
be
25
nm
[64].
The
reaction
temper-
ature
has
been
chosen
to
be
813
K,
at
which
the
reaction
is
neither
kinetically
nor
thermodynamically
limited.
4.2.
Surface
reaction
mechanism
In
a
previous
study
[51]
we
established
a
multi-step
surface
reaction
mechanism
with
the
associated
kinetics
for
the
oxidation
of
H2over
Rh/Al2O3.
Using
this
kinetics
without
modifications
we
now
extend
the
surface
mechanism
by
including
CO
and
CO2as
adsorbed
species
and
further
intermediates.
It
is
assumed
that
all
the
species
adsorb
only
on
the
active
metal,
i.e.,
the
alumina
support
does
not
function
as
an
active
site.
The
thermodynamic
consistency,
at
both
the
entalphic
and
the
entropic
level,
of
the
reaction
mecha-
nism
has
been
ensured
by
using
a
numerical
adjustment
procedure
of
the
rate
coefficients
as
discussed
in
our
recent
papers
[65].
The
tubular
flow
reactor
experiments
show
that
CH4is
formed
only
at
ppm
level
during
the
WGS
and
RWGS
reactions.
No
CH4is
detected
in
the
stagnation-flow
experiments.
Therefore,
reactions
of
CH4are
not
included
into
the
reaction
mechanism.
The
mechanism
consists
of
30
reactions
among
five
gas-phase
and
eight
surface
species
and
is
given
in
Table
2
including
the
rate
coefficients.
The
reaction
rates
are
modeled
by
a
modified
Arrhenius
expression
including
a
temperature
dependence
of
the
pre-exponential
factor,
k
=
ATˇexp −Ea
RT (16)
The
nominal
values
of
the
pre-exponential
factors
are
assumed
to
be
1013NA/
(cm2/mol
s)
where
NAis
Avagadro’s
number
and

is
the
surface
site
density
(1.637
×
1019 site/cm2for
Rh
110).
The
nominal
value
of
1013 is
the
value
calculated
from
transition
state
theory
(kBT/h)
with
kBis
being
Boltzmann’s
constant
and
h
is
Plank’s
constant
[65].
For
adsorption
reaction
(initial)
sticking
coefficients
are
used,
which
are
related
to
an
uncovered
surface.
The
effect
of
the
coverage
on
the
adsorption
rate
is
accounted
for
by
the
concen-
trations
of
adsorbed
species
according
to
Eq.
(10).
The
adsorption
of
the
gas
species
is
considered
to
be
non-activated
[24,29,66].
The
initial
sticking
coefficients
of
CO
and
CO2are
implemented
to
be
4.97
×
10−1[24]
and
4.80
×
10−2[49],
respectively.
RWGS
are
in
particular
sensitive
to
the
adsorption–desorption
equilib-
rium
of
CO2as
discussed
below
and
shown
in
Fig.
6.
If
both
the
rate
constants
for
adsorption
and
desorption
of
a
species
reveal
high
sensitivity
coefficients
with
opposite
sign,
then
the
adsorption–desorption
equilibrium
matters.
In
our
former
kinetic
studies
on
partial
oxidation
of
hydrocarbons
over
Rh
[24,28],
the
CO2adsorption–desorption
kinetics
were
not
very
sensitive
for
conversion
and
product
selectivity,
i.e.,
the
kinetic
data
given
for
CO2adsorption
and
desorption
could
not
be
critically
evaluated,
and
therefore
needed
further
validation.
Recently,
McGuire
et
al.
[49]
studied
dry
reforming
of
CH4over
Rh/Al2O3.
They
found
CO2
adsorption–desorption
equilibrium
is
much
more
on
the
adsorp-
tion
side
than
literature
suggested
[24,28]
and
proposed
a
sticking
coefficient
of
4.80
×
10−2.
In
the
present
study
we
also
use
this
stick-
ing
coefficient
for
CO2adsorption.
The
pre-exponential
factor
of
CO2desorption
was
slightly
modified,
so
the
model
can
predict
the
former
and
currently
presented
experimental
data.
CO
is
easily
adsorbed
on
Rh
that
may
even
lead
to
blockage
of
adsorption
of
other
species
such
as
oxygen
and
water
coverage
and,
hence,
prevents
the
ignition
of
CO
oxidation
or
WGS
reac-
tions
[5].
The
onset
of
conversion
in
our
experiments,
however,
implies
a
decrease
of
the
activation
energy
for
desorption
at
highly
CO
covered
surfaces
[24,66–68].
Based
on
comparison
of
model
predictions
and
experimental
observation,
coverage
dependent
activation
energy
for
CO
desorption
of
(134.07
−
47CO)
kJ/mol
was
introduced.
The
reference
values
for
activation
barriers
of
R8-R20
are
taken
from
our
former
kinetic
studies
[24,28]
and
they
are
slightly
mod-
ified
within
the
thermodynamic
consistency
range.
Dissociation
of
the
CO(s)
(R19,
R30
in
Table
2)
is
introduced
in
the
mechanism
to
take
into
account
the
possible
carbon
formation
from
CO.
Reactions
of
CO(s)
and
OH(s)
species
used
in
the
reac-
tion
mechanism
are
illustrated
in
Fig.
4.
Several
proposals
have
been
made
to
explain
the
WGS
reaction
pathway
over
noble
metal
supported
catalysts
as
discussed
in
Section
1,
e.g.
[19,28,32,69,70].
Although,
as
discussed
in
the
introduction,
there
is
a
disagreement
in
the
literature
as
to
whether
the
redox
formate
or
carboxyl
mech-
anism
is
the
dominant
reaction
path
for
WGS,
recent
studies
favor
the
carboxyl
mechanism
[2,3,71],
in
which
the
reaction
between
the
adsorbed
CO
and
OH
yields
a
carboxyl
species
(COOH)
with
a
very
small
activation
barrier
[3].
An
analysis
of
the
formate-based
mechanism
for
the
water-gas
shift
reaction
over
several
catalysts
including
Rh
has
been
carried
out
by
Burch
et
al.
[69].
It
was
shown
that
with
very
few
exceptions,
the
published
results
cannot
be
used
to
provide
any
mechanistic
evidence
either
for
or
against
a
formate
model.
However,
the
authors
emphasize
that,
the
contribution
of

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 37
Table
2
Proposed
surface
reaction
mechanism.
Reaction
A
(cm2/mol
s)
ˇ
Ea(kJ/mol)
R1
H2+
Rh(s)
+
Rh(s)
→
H(s)
+
H(s)
3.000
×
10−2Stick.
coeff.
R2
O2+
Rh(s)
+
Rh(s)
→
O(s)
+
O(s)
1.000
×
10−2Stick.
coeff.
R3
H2O
+
Rh(s)
→
H2O(s)
1.000
×
10−1Stick.
coeff.
R4
CO2+
Rh(s)
→
CO2(s)
4.800
×
10−2Stick.
coeff.
R5
CO
+
Rh(s)
→
CO(s)
4.971
×
10−1Stick.
coeff.
R6
H(s)
+
H(s)
→
Rh(s)
+
Rh(s)
+
H25.574
×
1019 0.239
59.69
R7
O(s)
+
O(s)
→
Rh(s)
+
Rh(s)
+
O25.329
×
1022 −0.137
387.00
R8
H2O(s)
→
H2O
+
Rh(s) 6.858
×
1014 −0.280 44.99
R9
CO(s)
→
CO
+
Rh(s)
1.300
×
1013 0.295
134.07
−
47.00CO
R10
CO2(s)
→
CO2+
Rh(s)
3.920
×
1011 0.315
20.51
R11
H(s)
+
O(s)
→
OH(s)
+
Rh(s)
8.826
×
1021 −0.048
73.37
R12
OH(s)
+
Rh(s)
→
H(s)
+
O(s)
1.000
×
1021 0.045
48.04
R13
H(s)
+
OH(s)
→
H2O(s)
+
Rh(s)
1.743
×
1022 −0.127
41.73
R14
H2O(s)
+
Rh(s)
→
H(s)
+
OH(s) 5.408
×
1022 0.129
98.22
R15
OH(s)
+
OH(s)
→
H2O(s)
+
O(s)
5.735
×
1020 −0.081
121.59
R16
H2O(s)
+
O(s)
→
OH(s)
+
OH(s)
1.570
×
1022 0.081
203.41
R17
CO2(s)
+
Rh(s)
→
CO(s)
+
O(s)
5.752
×
1022 −0.175
106.49
R18
CO(s)
+
O(s)
→
CO2(s)
+
Rh(s) 6.183
×
1022 0.034
129.98
−
47.00CO
R19
CO(s)
+
Rh(s)
→
C(s)
+
O(s)
6.390
×
1021 0.000
174.76
−
47.00CO
R20
C(s)
+
O(s)
→
CO(s)
+
Rh(s)
1.173
×
1022 0.000
92.14
R21
CO(s)
+
OH(s)
→
COOH(s)
+
Rh(s)
2.922
×
1020 0.000
55.33
−
47.00CO
R22
COOH(s)
+
Rh(s)
→
CO(s)
+
OH(s)
2.738
×
1021 0.000
48.38
R23
COOH(s)
+
Rh(s)
→
CO2(s)
+
H(s)
1.165
×
1019 0.160
5.61
R24
CO2(s)
+
H(s)
→
COOH(s)
+
Rh(s)
1.160
×
1020 −0.160
14.48
R25
COOH(s)
+
H(s)
→
CO(s)
+
H2O(s) 5.999
×
1019 −0.188
33.55
R26
CO(s)
+
H2O(s)
→
COOH(s)
+
H(s)
2.258
×
1019 0.051
97.08
−
47.00CO
R27
CO(s)
+
OH(s)
→
CO2(s)
+
H(s)
3.070
×
1019 0.000
82.94
−
47.00CO
R28
CO2(s)
+
H(s)
→
CO(s)
+
OH(s)
2.504
×
1021 −0.301
84.77
R29
C(s)
+
OH(s)
→
CO(s)
+
H(s)
4.221
×
1020 0.078
30.04
R30
CO(s)
+
H(s)
→
C(s)
+
OH(s) 3.244
×
1021 −0.078 138.26
−
47.00CO
The
rate
constants
are
given
in
the
form
of
k
=
ATˇexp(−Ea/RT);
adsorption
kinetic
is
given
in
form
of
sticking
coefficients;
the
surface
site
density
is

=
2.72
×
10−9mol
cm−2.
Electronic
versions
of
the
mechanism
in
DETCHEM
and
CHEMKIN
format
can
be
downloaded
from
www.detchem.com.
IR-observable
formate
to
the
production
of
CO2is
of
only
minor
importance
for
all
catalyst
which
show
high
WGS
activity.
Grabow
et
al.
presented
experimental
data
and
a
micro-kinetic
model
for
the
WGS
over
Pt
at
temperatures
from
250
to
300 ◦C
for
various
gas
compositions
[71]
and
concluded
that
the
most
significant
reac-
tion
channel
proceeds
via
a
carboxyl
(COOH)
intermediate
and
the
formate
(HCOO)
acts
only
as
a
spectator
species.
We
initially
included
both
pathways
in
our
model
but
found
that
all
experi-
mental
data
we
looked
at
can
satisfactorily
described
the
using
the
carboxyl
path
only.
Even
without
any
attempt
to
draw
here
a
final
Fig.
4.
Reaction
pathways
for
formation
of
CO2.
conclusion,
we
neglected
the
formate
path
in
further
consider-
ations
in
our
model.
In
our
kinetic
scheme,
the
formation
of
COOH(s)
proceeds
as
a
reaction
between
the
CO(s)
and
OH(s)
species
and
CO(s)
+
OH(s)
→
COOH(s)
+
Rh(s)
(R21)
COOH(s)
can
either
form
CO(s)
and
H2O(s)
reacting
with
H(s)
or
form
CO2(s)
and
H(s)
COOH(s)
+
H(s)
→
CO(s)
+
H2O(s)
(R23)
COOH(s)
+
Rh(s)
→
CO2(s)
+
H(s)
(R24)
The
dissociation
of
CO2(s)
by
H(s)
is
also
included
in
the
reaction
paths
(R28),
because
it
was
reported
that
hydrogen
enhances
the
CO2(s)
dissociation
[20,72].
The
reaction
steps
R21–R30
were
taken
from
the
studies
of
Shusturovic
et
al.
[73,74],
where
they
calcu-
lated
the
activation
barriers
using
the
unity
bond
index-quadratic
exponential
potential
(UBI-QEP)
method.
However,
the
activation
barriers
listed
in
Table
2
are
slightly
modified
within
the
entalphic
consistency.
Figs.
5
and
6
show
the
sensitivity
of
the
gas-phase
concentra-
tions
of
CO2and
H2O
close
to
the
catalytic
disk
for
WGS
and
RWGS
reactions,
respectively,
on
the
rate
coefficients.
The
sensitivity
anal-
yses
are
performed
at
four
temperatures
200 ◦C,
400 ◦C,
500 ◦C,
and
700 ◦C.
CHEMKIN
SPIN
software
[57]
is
used
for
the
sensitivity
anal-
yses.
Normalized
sensitivity
coefficients
(Ski)
for
each
reaction
are
calculated
in
the
form
of
partial
derivatives
in
which
Xkpresents
the
mole
fraction
of
the
gas-phase
species
k
close
to
the
surface
and
Aiis
the
pre-exponential
factor
for
the
reaction
of
i.
Ski =Ai
Xk
∂Xk
∂Ai
(17)
The
results
show
that
for
all
temperatures,
gas-phase
CO2and
H2O
concentrations
are
highly
sensitive
to
H2O
adsorption
and

38 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
Fig.
5.
Sensitivity
analyses
of
CO2gas
phase
concentration
for
WGS
reaction
at
different
temperature
points.
Inlet
mole
composition
is
chosen
to
be
4.76
vol.%
CO
and
5.18
vol.%
H2O
in
Ar
dilution.
dissociation
as
well
as
to
CO2adsorption
and
desorption,
i.e.,
the
adsorption–desorption
equilibrium
of
those
species
matter.
The
formation
rate
of
COOH
(R21)
also
significantly
effects
the
gas-
phase
concentration
of
CO2at
low
temperatures
(Fig.
5).
This
is
in
accordance
with
the
data
from
the
sensitivity
analyses
of
Maestri
et
al.
[23]
in
which
a
reaction
mechanism
of
Maestri
et
al.
[75]
was
used
to
model
methane
steam
and
dry
reforming
on
a
Rh
catalyst.
Maestri
and
Reuter
[14]
performed
semi-empirical
reaction
path
analysis
on
the
reaction
mechanism
of
Maestri
et
al.
[75]
showed
that
the
rate
determining
step
is
water
dissociation
that
forms
OH
and
reacts
with
CO.
Other
dominant
reaction
steps
are
in
partial
equilibrium.
However,
at
high
temperatures
(>450 ◦C)
this
surface
reaction
is
no
longer
dominant,
because
the
surface
coverage
is
rather
low.
Unlike
the
semi-empirical
analysis,
first
principles
refinement
through
DFT
calculations
predicts
that
reaction
between
CO
and
OH
(R27)
is
not
an
elementary
reaction.
Instead,
minimum
energy
path
suggest
that
first
a
cis-COOH
forms
and
then
it
isomers
to
trans-
COOH.
Lastly,
formation
of
CO2and
H
occurs
via
decomposition
of
trans-COOH
[14].
In
the
reaction
scheme
presented
here,
formation
of
CO2via
the
reaction
between
CO
and
OH
is
assumed
as
a
lumped
reac-
tion
step;
its
reverse
reaction
is
also
included
in
the
kinetic
scheme
(R28)
However,
the
sensitivity
analyses
show
that
it
does
not
have
a
significant
effect,
neither
on
WGS
nor
R-WGS
(Figs.
5
and
6).
On
the
other
hand,
sensitivity
analyses
for
RWGS
show
that
the
formation
of
gas-phase
H2O
strongly
depends
on
the
CO2adsorption–desorption
(R4,
R10)
equilibrium
and
the
rate
of
CO2dissociation
(R17)
and
COOH(s)
formation
(R24)
and
decomposition
reactions
(R23)
(Fig.
6).
The
adsorption–desorption
equilibrium
(R5,
R9)
of
CO
also
influences
H2O
formation,
because
concentration
of
COOH(s)
is
related
to
the
concentration
of
the
CO(s).
Rates
of
COOH(s)
reactions
have
a
minor
effect
on
the
for-
mation
of
H2O
at
high
temperatures.
4.3.
Stagnation-flow
reactor
experiments
As
mentioned,
the
mechanism
and
kinetics
of
the
H2/O2system
are
taken
from
our
former
study
[51]
without
any
modification.
As
basis
for
the
kinetics
of
the
other
reactions
(R4,
R5,
R9,
R10,
R17–R30)
several
literature
source
were
used
to
set-up
a
tentative
kinetic
scheme,
which
was
then
“fine-tuned”
by
comparison
of
the
species
profiles
obtained
by
the
measurements
in
the
stagnation
flow
reactor
and
the
ones
computed
using
the
models
discussed.
Afterwards,
the
kinetic
scheme
was
tested
by
modeling
tubular
flow
reactor
experiments
(Section
4.4)
and
tubular
flow
reactor,
under
varying
inlet
conditions.
4.4.
Preferential
oxidation
of
H2in
CO/O2mixtures
Preferential
oxidation
of
CO
is
studied
at
a
moderate
temper-
ature
of
600 ◦C
at
the
conversion
is
considerably
high
but
not
yet
equilibrated.
Figs.
7–9
show
a
comparison
of
the
experimen-
tal
(symbols)
and
model
predicted
boundary-layer
concentration
profiles.
The
C
balance
closes
always
above
95%
in
the
experiment.
An
inlet
composition
of
5.2
vol.%
CO,
5.2
vol.%
O2,
and
6.0
vol.%
H2
serves
as
baseline
case
(Case
1
in
Table
1).
A
5
mm
thick
concentra-
tion
boundary-layer
is
formed
(Fig.
7).
The
model
predicts
a
slightly
thicker
boundary-layer
for
H2which
is
attributed
to
interplay
of
the
fast
diffusion
of
H2with
the
sampling
accuracy.
In
the
second
case
(Table
1),
the
oxygen
mole
fraction
is
decreased
leading
to
fuel
(CO
+
H2)
rich
conditions
resulting
in
a
decreased
formation
rate
of
CO2as
well
as
H2O.
No
significant
selective
CO
oxidation
is
observed
(Fig.
8).

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 39
Fig.
6.
Sensitivity
analyses
of
H2O
gas
phase
concentration
for
RWGS
reaction
at
different
temperature
points.
Inlet
mole
composition
is
chosen
to
be
5.2
vol.%
CO2and
5.2
vol.%
H2in
Ar
dilution.
Lastly,
the
oxygen
concentration
in
the
gas
mixture
is
increased
by
50
vol.%,
thus,
enough
O2is
supplied
for
total
oxidation
of
both
H2and
CO
(Case
3
in
Table
1).
No
preferential
oxidation
behavior
is
observed
again
(Fig.
9).
However,
both
H2and
CO
are
almost
totally
consumed
on
the
catalyst
surface.
The
mol
fraction
of
water
on
the
0
0.01
0.02
0.03
0.04
0.05
0.06
0.07
76.565.554.543.532.521.510.50
Mole fractions
Distance from the surface [mm]
CO2
O2
H2O
H2
CO
Fig.
7.
Boundary-layer
mole
fractions
of
preferential
oxidation
of
CO
in
H2/O2mix-
tures
at
600 ◦C
with
the
inlet
velocity
of
51
cm/s
and
inlet
composition
of
6.0
vol.%
H2,
5.20
vol.%
CO
and
5.20
vol.%
O2diluted
in
Ar.
catalyst
surface
is
calculated
to
be
0.054
via
C
and
oxygen
mass
balance,
which
is
in
good
agreement
with
the
predicted
value
of
0.052.
Even
though
the
detected
surface
concentrations
of
the
species
are
in
good
agreement
with
the
experimental
values,
the
0
0.01
0.02
0.03
0.04
0.05
0.06
0.07
76.565.554.543.532.521.510.50
Mole fractions
Distance from the
surface [mm]
H2
CO
CO2
O2
H2O
Fig.
8.
Boundary-layer
mole
fractions
of
preferential
oxidation
of
CO
in
H2/O2mix-
tures
at
600 ◦C
with
the
inlet
velocity
of
51
cm/s
and
inlet
composition
of
6.0
vol.%
H2,
5.2
vol.%
CO
and
2.65
vol.%
O2in
Ar
dilution.

40 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
0.000
0.010
0.020
0.030
0.040
0.050
0.060
6
5.554.543.532.521.510.50
Mole fractions
Distance from
the
surface
[mm]
CO2
O2
H2O
H2
CO
Fig.
9.
Boundary
layer
mole
fractions
of
preferential
oxidation
of
CO
in
H2/O2mix-
tures
at
600 ◦C
with
the
inlet
velocity
of
51
cm/s
and
inlet
composition
of
2.57
vol.%
H2,
2.55
vol.%
CO
and
4.0
vol.%O2in
Ar
dilution.
concentration
profile
of
H2within
the
boundary
layer
slightly
dif-
fers
from
the
predicted
values.
It
is
worth
mentioning
that,
here,
the
gas
concentrations
and
the
temperatures
selected
for
the
preferential
oxidation,
are
not
motivated
by
technical
applications
but
rather
serve
the
purpose
of
the
model
evaluation.
Our
previous
study
showed
that,
in
case
of
H2oxidation,
the
reaction
is
carried
out
under
stoichiometric
conditions;
ignition
occurs
at
184 ◦C
and
results
in
a
total
H2conversion
at
adiabatic
con-
ditions
[76].
Besides,
recent
studies
showed
that
CO
can
be
totally
oxidized
at
125 ◦C
under
oxygen-rich
conditions
(1
vol.%
CO
and
1
vol.%
O2)
[77].
For
this
reason,
we
conclude
that,
no
preferential
oxidation
behavior
between
H2and
CO
is
observed
at
the
tempera-
ture
point
of
600 ◦C,
at
which
the
reaction
is
not
kinetically
limited.
However,
both
reactants
(CO,
H2)
have
an
ignition-inhibiting
effect
on
each
other.
Since
O2is
the
limiting
reactant
(Case
2
in
Table
1),
neither
a
total
CO
nor
total
H2oxidation
is
observed.
The
reac-
tion
proceeds
in
the
path
of
both
CO
and
H2oxidation
at
an
equal
reaction
rate.
4.5.
WGS-
and
RWGS
reactions
WGS
activity
of
the
catalyst
is
tested
at
three
different
surface
temperatures
(Cases
4–6,
Table
1)
with
an
inlet
composition
of
4.75
vol.%
CO
and
5.18
vol.%
H2O
diluted
in
Ar.
The
inlet
flow
veloc-
ity
is
calculated
to
be
74
cm/s
at
an
inlet
temperature
of
150 ◦C
at
500
mbar.
Fig.
10
reveals
a
boundary-layer
thickness
of
4.5
mm.
Increasing
the
temperature
to
735 ◦C
significantly
augments
the
WGS
activity
(Fig.
11),
whereas
the
effect
on
product
distribution
is
small
at
800 ◦C
(Fig.
12).
RWGS
activity
is
tested
at
surface
temperatures
of
600 ◦C
and
700 ◦C
with
an
inlet
gas
composition
of
5.2
vol.%
CO2and
5.2
vol.%
H2diluted
in
Ar.
The
inlet
gas
temperature
is
40 ◦C
resulting
in
an
inlet
velocity
of
51
cm/s
at
500
mbar.
The
boundary-layer
con-
centration
profiles
of
the
species
are
shown
in
Figs.
13
and
14.
No
significant
RWGS
activity
is
detected
at
600 ◦C
(Fig.
13).
The
boundary-layer
thickness
is
quite
small
(4
mm).
The
maximum
level
of
CO
is
measured
to
be
0.52
vol.%
close
to
the
catalyst
sur-
face.
The
overall
reaction
rate
is
kinetically
controlled.
Usually
the
reactants
reach
their
minimum
concentration
on
the
catalyst
sur-
face.
However,
here
the
large
thermal
diffusion
of
hydrogen
close
to
the
hot
catalyst
has
a
significant
effect
on
the
H2concentration
pro-
file
and
the
maximum
H2concentration
is
observed
at
the
catalytic
surface
(z
=
0).
A
similar
trend
is
observed
at
700 ◦C.
0.000
0.005
0.010
0.015
0.020
0.025
0.030
0.035
0.040
0.045
0.050
0.055
00.
5
11.
5
22.
5
33.
5
44.5 5 5.
5
6
Mole fractions
Distance
from
the
surface
[mm]
CO
CO2
H2O
H2
Fig.
10.
Comparison
of
measured
boundary-layer
mole
fraction
profiles
with
the
model
prediction
for
WGS
reaction
carried
out
with
molar
H2O/CO
ratio
of
1.1
and
at
the
surface
temperature
of
600 ◦C.
Calculated
inlet
velocity
is
74
cm/s
at
the
inlet
temperature
of
150 ◦C.
0.000
0.005
0.010
0.015
0.020
0.025
0.030
0.035
0.040
0.045
0.050
0.055
00.
5
11.
5
22.
5
33.
5
44.
5
55.
5
6
Mole fractions
Distance
from the
surface
[mm]
CO
CO2
H2
H2O
Fig.
11.
Comparison
of
measured
boundary-layer
mole
fraction
profiles
with
the
model
prediction
for
WGS
reaction
carried
out
with
molar
H2O/CO
ratio
of
1.1
and
at
the
surface
temperature
of
735 ◦C.
Calculated
inlet
velocity
is
74
cm/s
at
the
inlet
temperature
of
150 ◦C.
0.000
0.005
0.010
0.015
0.020
0.025
0.030
0.035
0.040
0.045
0.050
0.055
00.
5
11.5 2 2.
5
33.
5
44.
5
55.
5
66.
5
7
Mole fractions
Distance from
the
surface
[mm]
CO
CO2
H2
H2O
Fig.
12.
Comparison
of
measured
boundary-layer
mole
fraction
profiles
with
the
model
prediction
for
WGS
reaction
carried
out
with
molar
H2O/CO
ratio
of
1.10
and
at
the
surface
temperature
of
800 ◦C.
Calculated
inlet
velocity
is
74
cm/s
at
the
inlet
temperature
of
150 ◦C.

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 41
0
100
200
300
400
500
600
700
0.000
0.010
0.020
0.030
0.040
0.050
0.060
0.070
0.080
012
3
4
5
6
Temperature[oC]
Mole fractions
Distance
from
the
surface
[mm]
H2
CO
CO2
Temperature
Fig.
13.
Comparison
of
measured
boundary-layer
mole
fraction
profiles
with
the
model
prediction
for
RWGS
reaction
with
CO2/H2ratio
of
1.0
and
at
the
surface
temperature
of
600 ◦C
together
with
the
predicted
temperature
profile
along
the
boundary-layer.
0
100
200
300
400
500
600
700
800
0.000
0.010
0.020
0.030
0.040
0.050
0.060
0.070
0
1
2
3
4
5
6
Temperature [oC]
Mole fractions
Distance from
the surface
[mm]
H2
CO2
CO
Temperature
Fig.
14.
Comparison
of
measured
boundary-layer
mole
fraction
profiles
with
the
model
prediction
for
RWGS
reaction
with
CO2/H2ratio
of
1.0
and
at
the
surface
temperature
of
700 ◦C
together
with
the
predicted
temperature
profile
along
the
boundary-layer.
At
700 ◦C,
the
RWGS
activity
is
expected
to
increase.
However,
there
is
no
difference
in
measured
CO
fraction
at
the
catalyst
surface
(0.53
vol.%
at
700 ◦C).
Since
the
mass
transport
of
H2toward
the
catalyst
surface
is
enhanced
by
increasing
the
temperature
due
to
thermal
diffusion,
the
mole
fraction
of
the
CO
on
the
catalyst
surface
decreases.
0.000
0.020
0.040
0.060
0.080
0.100
0.120
0.140
200300400500600700800900
Mole fractions
Temperature [oC]
H2O
CO2
Equil. CO2
Fig.
15.
WGS
reaction
carried
out
in
a
flow
reactor
with
inlet
mole
composition
of
11.6
vol.%
H2O
and
11.2
vol.%
CO
diluted
in
N2.
Comparison
of
experimental
(sym-
bols)
and
predicted
H2O
and
CO2mole
fractions
as
a
function
of
temperature.
0.000
0.020
0.040
0.060
0.080
0.100
0.120
20030
040
050
060
070
080
090
0
Mole fractions
Temperature [oC]
CO2
H2O
Equil. CO2
Fig.
16.
CO2co-feeding
effect
in
WGS
reaction
which
is
carried
out
in
flow
reactor
with
the
inlet
mole
composition
of
10.49
vol.%
H2O,
10.08
vol.%
CO
and
2.0
vol.%
CO2
in
N2dilution.
Comparison
of
experimental
(symbols)
and
predicted
H2O
and
CO2
mole
fractions
as
a
function
of
temperature.
4.6.
Tubular
flow
reactor
experiments
The
reaction
kinetics
are
evaluated
by
the
experiments
in
the
tubular
flow
reactor
with
the
Rh/Al2O3coated
honeycomb
mono-
lith
under
steady-state
conditions
and
at
atmospheric
pressure.
WGS
and
RWGS
reactions
are
carried
out
in
a
temperature
range
of
200–900 ◦C.
The
co-feeding
of
CO2and
CO
on
WGS
and
RWGS
reactions,
respectively,
is
also
examined.
Comparisons
of
the
exper-
imental
values
and
the
model
predictions
are
presented.
CO2and
H2O
mole
fractions
in
the
outlet
gas
stream
are
given
as
function
of
temperature
in
Fig.
15
for
an
inlet
mole
composition
of
11.6
vol.%
H2O
and
11.2
vol.%
CO
diluted
in
N2(Case
9
in
Table
1).
The
maximum
concentration
of
CH4is
100
ppm
at
540 ◦C.
WGS
activity
starts
at
400 ◦C
and
the
reaction
reaches
the
equilibrium
at
770 ◦C.
The
model
predicts
the
WGS
activity
in
good
agreement
with
the
experimental
results
over
a
wide
range
of
temperatures.
In
Case
10,
2.0
vol.%
CO2is
added
to
the
10.49
vol.%
H2O
and
10.08
vol.%
CO
mixture
diluted
in
N2.
No
significant
effect
is
observed
on
the
rate
of
the
WGS
reaction
(Fig.
16),
even
though
the
reaction
reached
the
equilibrium
at
a
slightly
higher
temperature
(810 ◦C).
The
WGS
reaction
proceeds
via
CO
oxidation
by
OH
radicals,
and
the
rate-determining
step
is
water
dissociation.
0.00E+00
1.00E-02
2.00E-02
3.00E-02
4.00E-02
5.00E-02
6.00E-02
7.00E-02
1.00E-09
1.00E-08
1.00E-07
1.00E-06
1.00E-05
1.00E-04
1.00E-03
1.00E-02
1.00E-01
1.00E+00
300400500600700800
Rh(s)
C(s)
CO(s)
H(s)
H2O(s)
O(s)
COOH(s)
OH(s)
CO2(s)
CO2(gas p
hase)
CO2mole fraction
surface coverage
Tem
perat
ure [oC]
Fig.
17.
Numerically
predicted
surface
coverage
and
CO2mole
fractions
as
a
function
of
reaction
temperature
(300–800 ◦C)
for
WGS
reaction
(Case
9
in
Table
1)
at
the
reactor
exit
(z
=
10
mm).

42 C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44
0.000
0.020
0.040
0.060
0.080
0.100
0.120
20030
040
050
060
070
080
090
0
Mole fractions
Temperature [oC]
H2
CO
Equil.H2
Fig.
18.
Comparison
of
experimental
and
predicted
values
of
H2and
CO
for
RWGS
reaction
carried
out
in
flow
reactor.
Inlet
mole
composition
is
chosen
to
be
10.40
vol.%
H2and
10.88
vol.%
CO2in
N2dilution.
0.000
0.020
0.040
0.060
0.080
0.100
0.120
20030
040
050
060
070
080
090
0
Mole fractions
Temperature [oC]
H2
CO
Equil.
H2
Fig.
19.
Effect
of
CO
co-feeding
in
RWGS
reaction
carried
out
in
flow
reactor
with
the
inlet
mole
composition
of
10.03
vol.%
H2and
9.52
vol.%
CO2and
2.04
vol.%
CO
in
N2
dilution.
Comparison
of
experiments
(symbols)
and
predicted
(lines)
mole
fractions
as
a
function
of
temperature.
0.00E+00
5.00E-03
1.00E-02
1.50E-02
2.00E-02
2.50E-02
3.00E-02
3.50E-02
4.00E-02
4.50E-02
5.00E-02
1.00E-09
1.00E-08
1.00E-07
1.00E-06
1.00E-05
1.00E-04
1.00E-03
1.00E-02
1.00E-01
1.00E+00
30040
050
060
070
080
0
Surface coverage
Tempera
ture
[oC]
Rh(s)
C(s)
CO(s)
H(s)
H2O(s)
O(s)
COOH(s)
OH(s)
CO2(s)
H2O (gas phase)
H2O mole fraction
Fig.
20.
Numerically
predicted
surface
coverage
and
H2O
mole
fractions
as
a
func-
tion
of
reaction
temperature
(300–800 ◦C)
for
RWGS
reaction
(Case
11
in
Table
1)
at
the
reactor
exit
(z
=
10
mm).
R² = 0.99
0.0E+
00
1.0E+04
2.0E+04
3.0E+04
4.0E+04
5.0E+04
6.0E+04
7.0E+
04
8.0E+04
9.0E+
04
1.0E+05
0.0E+
00 2.
0E+
04
4.0E+
04 6.
0E+
04
8.0E+
04 1.
0E+
05
Calculated TORH2
Measured TORH2
Fig.
21.
Comparison
of
calculated
and
measured
TOR
values
of
H2for
WGS
reaction
carried
out
in
the
flow
reactor.
An
addition
of
CO2is
expected
to
increase
the
RWGS.
Dry
reform-
ing
should
leads
to
the
consumption
of
CH4.
Consequently,
the
maximum
experimentally
observed
amount
of
CH4is
only
80
ppm
detected
at
530 ◦C.
The
numerically
predicted
surface
coverage
of
the
catalyst
at
the
reactor
exit
(z
=
10
mm)
is
plotted
as
a
function
of
temperature
in
Fig.
17
for
Case
9
(Table
1).
At
the
catalyst
exit,
the
surface
is
mainly
covered
by
CO(s),
H(s),
and
C(s).
The
mole
fraction
of
COOH(s)
and
CO2(s)
are
only
in
the
range
of
1
×
10−8and
1
×
10−7respectively.
At
low
temperature
(<600 ◦C)
the
formation
of
CO2(s)
increases
with
increasing
COOH(s)
concentration.
However,
at
high
temperature
(700–800 ◦C),
CO2selectivity
decreases
due
to
thermodynamics
explaining
the
slight
decrease
in
the
COOH(s).
No
direct
effect
of
surface
oxygen
concentration
on
CO2(s)
is
observed.
This
behavior
is
consistent
with
Fig.
5
since
formation
of
COOH(s)
is
favored
at
low
temperature
regimes.
For
studying
RWGS,
an
inlet
composition
of
10.40
vol.%
H2and
10.88
vol.%
CO2diluted
in
N2was
chosen
(Case
11
in
Table
1).
Low
activity
is
observed
below
500 ◦C,
and
increasing
the
reaction
tem-
perature
results
in
a
linear
increase
in
CO
formation
(Fig.
18).
The
product
composition
is
close
to
equilibrium
at
890 ◦C.
The
effect
of
CO
co-feeding
is
also
examined
by
the
addition
of
2.04
vol.%
CO
to
the
mixture
of
10.03
vol.%
H2and
9.52
vol.%
CO2
diluted
in
N2(Case
12
in
Table
1).
The
addition
of
CO
has
no
signif-
icant
effect
on
the
product
distribution
as
function
of
temperature
R² = 0.98
0.E+
00
1.E+04
2.E+
04
3.E+04
4.E+
04
5.E+04
6.E+04
7.E+
04
8.E+
04
9.E+
04
1.E+05
0.0E+
00 2.
0E+
04
4.0E+
04 6.
0E+
04
8.0E+
04 1.
0E+
05
Calculated TORH2O
Measured TORH2O
Fig.
22.
Comparison
of
calculated
and
measured
TOR
values
of
H2O
for
RWGS
reac-
tion
carried
out
in
the
flow
reactor.

C.
Karakaya
et
al.
/
Applied
Catalysis
A:
General
470 (2014) 31–
44 43
(Fig.
19)
similar
to
the
study
Graven
et
al.
[78],
which
showed
that
the
overall
reaction
rate
of
the
formation
of
CO
follows
a
first-order
dependence
on
CO2concentration
and
a
0.5th-order
dependence
on
H2concentration.
Similar
to
WGS
reaction,
numerically
predicted
surface
coverage
of
the
catalyst
(z
=
10
mm)
is
plotted
as
a
function
of
temperature
for
Case
11
(Table
1)
in
Fig.
20.
The
main
surface
species
are
CO(s),
C(s)
and
H(s)
and
the
mole
fraction
of
CO2(s),
OH(s)
and
COOH(s)
are
less
than
1
×
10−5.
Simulation
reveals
decreasing
amount
of
COOH(s)
with
increasing
OH(s).
This
behavior
can
be
attributed
to
decompo-
sition
of
COOH(s)
by
reaction
R22
as
it
was
shown
as
an
important
reaction
steps
for
formation
of
gas-phase
H2O
(Fig.
6).
Furthermore,
there
is
a
correlation
between
CO2(s)
and
O(s)
species.
Thus,
the
surface
oxygen
is
formed
via
decomposition
of
CO2(R17).
5.
Conclusion
WGS
and
RWGS
reaction
kinetics
over
Rh/Al2O3catalysts
are
investigated
experimentally
and
numerically
in
two
reactor
con-
figurations,
a
stagnation
flow
reactor
with
a
catalytic
disk
and
a
tubular
flow
reactor
with
a
honeycomb
monolith
as
catalyst.
While
the
first
one
leads
to
a
one-dimensional
boundary
layer
and
zero-
dimensional
surface,
the
latter
exhibits
a
two-dimensional
flow
field
and
a
one-dimensional
surface.
In
the
stagnation
flow
reactor
a
microprobe
sampling
technique
is
used
to
measure
the
gas-phase
composition
in
the
boundary
layer
adjacent
to
the
catalyst
surface.
In
the
tubular
flow
reactor,
the
outlet
composition
is
analyzed.
The
reactions
are
studied
at
rather
high
temperatures
to
reveal
the
sig-
nificance
of
WGS
on
partial
oxidation
and
steam/dry
reforming
of
hydrocarbons.
Based
on
a
previously
established
reaction
kinetics
for
the
sys-
tem
H2/O2over
Rh/Al2O3,
a
thermodynamically
consistent
reaction
kinetics
was
derived
for
the
CO/CO2/O2/H2/H2O–Rh/Al2O3system
using
the
experimental
data
of
this
study.
The
mechanism
is
able
to
predict
all
experimental
observations
of
the
WGS,
RWGS,
and
preferential
oxidation
of
CO
over
Rh/Al2O3.
A
major
improvement
to
formerly
published
reaction
mechanisms
[24,28]
is
the
predic-
tion
of
CO2formation
at
low
temperatures
due
to
the
inclusion
of
a
carboxyl
(COOH)
reaction
path.
At
high
temperatures,
direct
oxida-
tion
of
CO(s)
by
O(s)
is
favored.
The
rate
determining
step
is
found
to
be
the
dissociation
of
H2O(s)
and
gas-phase
CO2concentration
is
highly
effected
by
the
reaction
of
H2O(s)
dissociation
as
well
as
the
adsorption–desorption
equilibrium
of
CO2.
The
main
reaction
path
for
RWGS
is
the
formation
and
dissociation
of
COOH(s).
The
measured
and
the
model-predicted
profiles
agree
quite
well
giving
confidence
in
the
applicability
of
the
mechanism
for
a
wide
range
of
conditions.
Exemplarily,
all
data
of
the
WGS
and
RWGS
study
derived
from
the
tubular
flow
reactor
are
compared
in
terms
of
cumulative
turnover
rate
(TOR)
over
the
honeycomb
catalyst
in
a
parity
diagrams
(Figs.
21
and
22).
Exemplarily,
definition
of
TOR
for
H2is
shown
in
Eq.
(18).
TOR
=XH2∗˙
ntotal

∗
Ageometric ·
Fcat/geo ·
(18)
with ˙
ntotal as
total
molar
species
flux
and

as
residence
time.
Acknowledgements
This
work
was
supported
by
Deutsche
Forschungsgemeinschaft
(DFG).
We
thank
S.
Tischer
and
H.
Karadeniz
(KIT)
for
their
sup-
ports
in
using
DETCHEMTM and
R.J.
Kee
(Colorado
School
of
Mines)
for
collaboration
during
the
development
of
the
stagnation
flow
reactor
and
many
fruitful
discussions.
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