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DESIGN OF AN AMMONIA SYNTHESIS SYSTEM FOR PRODUCING SUPERCRITICAL STEAM IN THE CONTEXT OF THERMOCHEMICAL ENERGY
STORAGE
Chen Chen Mechanical and Aerospace Engineering Dept. University of California, Los Angeles, CA, U.S.
H. Pirouz Kavehpour Mechanical and Aerospace Engineering Dept. University of California, Los Angeles, CA, U.S.
Keith Lovegrove IT Power
Canberra, Australia
Adrienne S. Lavine
Mechanical and Aerospace Engineering Dept. University of California, Los Angeles CA, U.S.
ABSTRACT Concentrating solar power plants typically incorporate
thermal energy storage, e.g. molten salt tanks. The broad
category of thermochemical energy storage, in which energy is
stored in chemical bonds, has the advantage of higher energy
density as compared to sensible energy storage. In the
ammonia-based thermal energy storage system, ammonia is
dissociated endothermically as it absorbs solar energy during
the daytime. The stored energy can be released on demand (for
electricity generation) when the supercritical hydrogen and
nitrogen react exothermically to synthesize ammonia. Using
ammonia as a thermochemical storage system was validated at
Australian National University (ANU), but ammonia synthesis
has not yet been shown to reach temperatures consistent with
the highest performance modern power blocks such as a
supercritical steam Rankine cycle requiring steam to be heated
to ~650°C. This paper explores the preliminary design of an
ammonia synthesis system that is intended to heat steam from
350°C to 650°C under pressure of 26 MPa. A two-dimensional
pseudo-homogeneous model for packed bed reactors previously
used at ANU is adopted to simulate the ammonia synthesis
reactor. The reaction kinetics are modeled using the Temkin-
Pyzhev reaction rate equation. The model is extended by
accounting for convection in the steam to predict the behavior
of the proposed synthesis reactor. A parametric investigation is
performed and the results show that heat transfer plays the
predominant role in improving reactor performance. INTRODUCTION
Today’s commercial concentrating solar power plants can
provide dispatchable power by integrating thermal energy
storage, e.g. molten salt tanks. The broad category of
thermochemical energy storage (TCES), in which energy is
stored in chemical bonds, has the advantages of higher energy
density and minimal energy losses during the storing operation
[1]. Shakeri et al. [2] indicated that an ammonia-based thermal
energy storage system was more efficient for long-term storage
than compressed air energy storage, pumped hydroelectric
energy storage, vanadium flow battery, or thermal energy
storage. In the ammonia-based thermochemical energy storage
system (Figure 1), ammonia (NH3) is dissociated
endothermically as it absorbs solar energy during the daytime.
The stored energy can be released on demand when the
supercritical hydrogen (H2) and nitrogen (N2) react
exothermically to synthesize ammonia. The released thermal
energy can be used to generate electricity. Ammonia-based
energy storage for concentrating solar power systems has been
studied and investigated since 1974 [3]. Lovegrove et al. [4, 5]
developed and experimentally validated a two-dimensional
pseudo-homogeneous steady state model of catalytic ammonia
reactors. With the same model, Kreetz et al. [6] analyzed the
exergy of an ammonia synthesis reactor in a solar
thermochemical power system. Kreetz proposed that either very
small diameter reactors or adiabatic reactors were preferable for
minimizing exergy loss [6]. Lovegrove et al. built and tested a 1
kWsol closed loop ammonia-based TCES system [7] and a 15
kWsol ammonia-based TCES system for dish power plants [8].
Their systems demonstrated ammonia dissociation on a dish
concentrator and subsequent energy recovery at temperatures
high enough for electricity generation, but did not demonstrate
heating of a working fluid. Furthermore, ammonia synthesis
has not been shown to reach temperatures consistent with
Proceedings of the ASME 2015 Power Conference POWER2015
June 28-July 2, 2015, San Diego, California
POWER2015-49190
1 Copyright © 2015 by ASME
modern power blocks such as a supercritical steam Rankine
cycle requiring steam to be heated to ~650°C.
This paper investigates the preliminary design of a pilot-
scale catalytic ammonia synthesis reactor operating at 30 MPa.
The objective of the reactor is to heat 12.5 g/s of steam at 26
MPa from 350°C to 650°C [9], corresponding to 25 kWt. We
investigate the effect of various parameters on the reactor
performance. The optimum design demands low cost electricity
generation. While the paper does not perform a
thermoeconomic analysis, it does consider the material usage
required for tube wall material and catalyst, two of the main
system costs.
For ease of expression we refer to the water, which may
be supercritical or liquid, as “steam.” The N2, H2, NH3
mixture is referred to as “the gas,” even though the species are
all supercritical within the reactor and NH3 is liquid in the
lower temperature portion of the system.
SYNTHESIS REACTOR SYSTEM Figure 2 is a schematic of an entire synthesis and steam
heating system consisting of a reactor and two heat exchangers.
As shown in this figure, the cold incoming synthesis gas enters
at the bottom right and is initially heated in a counterflow heat
exchanger from temperature T0 to temperature T1 by the gas
exiting the reactor. This “initial preheater” is only sufficient to
bring the incoming gas close to the reactor exit temperature
(T3), not to the desired reactor inlet temperature, T2. Therefore,
additional heat transfer is needed to further preheat the
incoming gas to T2. This “secondary preheating” can be
accomplished in various ways; in this paper we consider a
Figure 1. Schematic of ammonia dissociation, storage, and synthesis system [3]
Figure 2. Schematic of an ammonia synthesis system with a reactor and two heat exchangers (dashed boxes)
2 Copyright © 2015 by ASME
design in which the incoming gases are further preheated in a
gas-steam heat exchanger shown on the far left; “extra” steam
is heated in the reactor for this purpose. The total flow of steam
entering the reactor (at the top right) is heated from Tsi = 350ºC
to Tso = 650ºC by the exothermic reaction. The desired “base
steam” flow rate of 12.5 g/s exits the system to be used for
electricity generation. Then the “extra steam,” at Tso = 650ºC, is
used to preheat the incoming gas from T1 to T2.
Two possible configurations have been considered for the
design of the reactor itself. The first, shown in Figure 3, is a
concentric tube configuration with the reacting gas flowing in
the inner tube and steam in counterflow in the outer annulus.
The second, shown in Figure 4, is a shell-and-tube
configuration with the reacting gas in the tubes and steam in
counterflow in the shell. For purposes of preliminary design,
the shell-and-tube reactor is viewed as N small concentric tube
reactors, as previously illustrated in Figure 3. The fixed
parameters are listed in Table 1.
Table 1. The fixed dimensions and properties of the reactor
Fixed variables Value Steam tube wall thickness, W
s (cm) 2
Gas mass flow rate, gm (g/s) 20
Steam mass flow rate, sm (g/s) 21.6
Inlet steam temperature, Tsi (ºC) 350
Outlet steam temperature, Tso
(ºC) 650
Inlet gas temperature, T2 (ºC) 597
Steam pressure, Ps (MPa) 26 Gas mixture pressure, Pg (MPa) 30 Pre-exponential factor, ko,m (kg∙m
-3∙s
-1) 3.83×10
13
Activation energy, Ea (J/mol) 1.83×105
REACTOR MODEL As described in the preceding section, gas flows in a
tubular packed-bed catalytic reactor which exchanges heat with
steam through the outer wall. The pseudo-homogeneous steady
state model is very commonly employed for designing packed
bed reactors [10]. The model used here was originally
developed by Richardson et al. at the University of Houston for
tubular packed-bed methane reforming reactors [11]. It has
been modified and validated by Lovegrove for ammonia
dissociation via experiments on an electrically heated 1 kWchem
ammonia dissociation reactor [5].
In the pseudo-homogeneous model, the catalyst bed is
treated as a continuum with averaged properties, i.e. effective
conductivity (keff) and effective diffusivity (Deff). Also, the
model assumes: (1) negligible axial conduction, (2) gas
pressure is constant for purposes of evaluating properties and
reaction rate, (3) the average velocity and density are only
functions of axial location, (4) the radial gradients of specific
heat, effective conductivity, and effective diffusivity are
negligible. The governing equations for energy conservation
can therefore be expressed as:
�
�
Figure 4. Schematic of a tube bundle reactor. N tubes are
viewed as N concentric tube reactors.
�
�
Figure 3. Geometry and dimensions of a concentric tube reactor
3 Copyright © 2015 by ASME
1p eff
T Tvc k r r H
x r r r
(1)
The boundary conditions are [11]:
( , 0)in
eff w
wr
T r x T
Tk q
r
(2)
where rw is the inner wall radius, ( )w w cw s
q U T T , and
and w cw
U T are the overall heat transfer coefficient and catalyst
temperature at the wall, as defined in Richardson et al. [11].
The governing equations for mass conservation are:
3 3
1NH NH
eff
f fv D r r
x r r r
(3)
The boundary conditions for mass conservation are:
3 3,
3
( , 0)
0
NH NH in
NH
wr
f r x f
f
r
(4)
The Temkin-Pyzhev intrinsic rate expression is used for the
ammonia reaction [12]:
0,
123
1 (1 ) (1 )32
2 2 3
3 2
expa
m
u
NHH
p N o o
NH H
Er k
R T
ppK p p p
p p
(5)
Here po is the standard state pressure of 1 atm, explicitly
included to make the dimensions consistent; α is an empirical
parameter taken to be 0.5.
The total pressure has been assumed constant in the
model. This is typically a valid assumption; using the Ergun
equation, the pressure drop is found to be at most 3 MPa under
the conditions considered here, which is negligible compared to
the system pressure of 30 MPa.
The temperature of the steam is calculated from a cross-
sectionally averaged energy balance:
( )s
p s w
dTmc q P
dx (6)
The heat transfer coefficient for the steam flow is
calculated with the following equation:
0.8 0.4Nu
, Nu 0.023Re Prs
s
h
kh
D
(7)
where Nu is the Nusselt number, ks is the average thermal
conductivity of the steam, Dh is the hydraulic diameter, Re is
the Reynolds number based on the hydraulic diameter, and Pr is
the Prandtl number.
PARAMETRIC STUDY Effect of steam heat transfer coefficient. Similar to a
heat exchanger, the required surface area to heat the steam is
expected to decrease as the overall heat transfer coefficient, U,
increases. The overall heat transfer coefficient is related to the
thermal resistances of the catalyst bed, wall, and steam flow:
1
ln ( 2 ) /1 1
2 ( 2 )
g w s
g g g
g g w s g g
UP R R R
D W D
h D k h D W
(8)
where P is tube perimeter. In this sub-section, we consider the
effect of the steam heat transfer coefficient, hs. The heat transfer
coefficient for the steam flow can be manipulated by changing
the outer diameter of the steam tube, Ds, (see Figure 3); under
laminar conditions the Nusselt number is constant, but for
sufficiently small gap dimension the flow can be made
turbulent and the Nusselt number then increases with
decreasing gap spacing. Table 2 shows the values of the steam
tube diameter, Ds, that were selected to give steam heat transfer
coefficients hs from 33.3 to 100 W∙m-2
∙K-1
while fixing other
properties and dimensions (listed in Table 1).
Table 2. Outlet conditions and required reactor length for
different steam heat transfer coefficients
Dg
(cm)
Wg
(cm)
Ds
(cm)
hs (W∙
m-2
∙K-1
)
T3
(ºC) fNH3,out
L
(m)
Vw
(m3)
10 1 16.8 33.3 459 0.500 38.2 0.58
10 1 14.6 66.7 457 0.497 22.1 0.31
10 1 13.8 100.0 454 0.495 16.8 0.22
Figure 5 shows the bulk (cross-sectionally averaged)
temperature profiles of steam (dashed curve) and gas (solid
curve) along the x-direction (the direction of flow through the
reactor). After being preheated by steam, the stoichiometric
synthesis gas mixture (N2+3H2) enters the reactor (solid curve)
at x = 0 and begins flowing in the positive x-direction. As soon
as the gases enter the reactor, they undergo a very abrupt
temperature increase; this is because of the rapid reaction rate
caused by the elevated temperature. The steep temperature
increase brings the gases close to equilibrium conditions and
4 Copyright © 2015 by ASME
the reaction slows, but does not cease. Thereafter, the
temperature of the reacting gases decreases along the reactor
length as they lose heat to the steam. The steam (dashed curve)
enters at the right at 350°C and exits at the left at 650°C. The
slope change in the region roughly around 400°C corresponds
to the pseudocritical regime of the steam, with very large
specific heat. As the steam heat transfer coefficient hs increases,
the required reactor length to heat the steam to 650°C
decreases, as expected.
In Figure 6, the solid curve is the mass fraction of
ammonia in the gas stream as it flows through the reactor. The
dashed curve is the equilibrium mass fraction corresponding to
the bulk reactor temperature at each axial location. The
ammonia mass fraction reaches the equilibrium curve almost
immediately at the reactor inlet and then increases downstream
along with the equilibrium curve. It can be observed that under
the conditions investigated here, the synthesis process is
essentially heat-transfer-limited. That is, the reaction rate is fast
enough to maintain the reaction almost at equilibrium. The rate
at which the reaction proceeds is then dictated by the rate at
which heat is transferred to the steam, which reduces the gas
temperature and increases the corresponding equilibrium
ammonia mass fraction. The ammonia mass fraction deviates
from the equilibrium curve near the outlet; the lowered
temperature decreases the reaction rate so that it is no longer
effectively infinite. It can be seen in Figure 6 that the deviation
from the equilibrium ammonia mass fraction increases with
increasing steam heat transfer coefficient hs since greater heat
transfer tends to make the process more reaction-rate-limited
than heat-transfer-limited.
As shown in Table 2, both the outlet temperature T3
and the outlet ammonia mass fraction fNH3,out decrease slightly
as the steam heat transfer coefficient hs increases. Most
importantly, the required length decreases significantly, which
decreases material costs. The three main material costs are: the
catalyst which fills the gas tube; the wall separating the gases
and steam (inner wall), which sees a modest pressure
differential of 4 MPa; and the wall separating the steam from
the environment (outer wall), which sees a 26 MPa pressure
differential. In this case, with diameter Dg held fixed, both the
catalyst volume and inner wall material volume decrease in
proportion to the reactor length. The outer wall material volume
is given in the last column in Table 2 and is seen to decrease
rapidly as Ds decreases, due to both the reduced length and
diameter. The only disadvantage to reducing the outer diameter
is that it would also increase the pressure drop of the steam
flow. However, this pressure drop is estimated to be
significantly smaller than the pressure drop in the catalyst bed,
and would therefore not impose too great a cost for pumping
power. The conclusion is that the design should achieve a large
steam heat transfer coefficient hs by reducing the steam cross-
sectional area.
0 5 10 15 20 25 30 35 40 45 50
300
350
400
450
500
550
600
650
700
750
Tem
per
atu
res,
T (
oC
)
Reactor length, x (m)
Ds=16.8 cm: gas
Ds=16.8 cm: steam
Ds=14.6 cm: gas
Ds=14.6 cm: steam
Ds=13.8 cm: gas
Ds=13.8 cm: steam
Figure 5. Temperatures of the gas and steam along the reactor
length
0 5 10 15 20 25 30 35 40 45 50
0.0
0.2
0.4
0.6
Am
mo
nia
mas
s fr
acti
on
s, f
NH
3
Reactor length, x (m)
Ds=16.8 cm
Ds=16.8 cm: equilibrium
Ds=14.6 cm
Ds=14.6 cm: equilibrium
Ds=13.8 cm
Ds=13.8 cm: equilibrium
Figure 6. The ammonia mass fraction and corresponding
equilibrium ammonia mass fraction along the reactor length
Effect of reactor tube diameter. In this sub-section, we
seek to elucidate the impact of inner tube diameter, Dg, on
reactor performance. Varying Dg would have multiple effects,
e.g. it changes both of steam and gas thermal resistances, the
gas velocity (for fixed mass flow rate), and the catalyst volume.
The impact of steam thermal resistance was established in the
previous sub-section, by varying hs while holding Dg fixed. In
this sub-section, we hold steam thermal resistance fixed, so that
we can explore the other effects of varying Dg. Since steam
thermal resistance s
R depends on the product of the steam
heat transfer coefficient and surface area, in this sub-section we
hold the product hs(Dg + 2Wg), fixed. In order to hold this
product fixed while varying Dg, we simultaneously vary Ds to
change hs. The derivation uses the standard definition of
hydraulic diameter for the outer annulus, Dh = 4Ac/Pw where Pw
= Dg + 2Wg + Ds is the wetted perimeter. Then, with fixed
steam mass flow rate, and assuming turbulent steam flow, the
5 Copyright © 2015 by ASME
following relationships can be used to find the required Ds for a
given Dg:
1 0.8
0.8
~ , ~
2( 2 ) ~ const.
2 2
s h
s w s s
s
g g
s g g
g g s s g g
h DRe P Nu Re
k
D Wh D W
D W D D D W
Table 3 lists three combinations of Dg and Ds that give fixed
hs(Dg + 2Wg) and corresponding model outputs. The last two
columns in this table are the wall material volume and catalyst
volume, respectively.
Figure 7 shows the gas temperature (solid curve) and
steam temperature (dashed curve) along the reactor length.
Referring to Table 3 and Figure 7, it can be seen that as Dg
decreases, the required reactor length decreases slightly. If the
process were limited by reaction kinetics, we would see that the
required volume would stay constant (i.e. residence time or
space velocity would remain constant), necessitating an
increase in length as diameter decreased. Instead, the
behavior is predominantly a heat transfer effect due to the
changing thermal resistance within the catalyst bed. In Table
3, the thermal resistance within the catalyst bed is seen to
decrease as Dg decreases, which largely explains the shorter
length required.
Figure 8 shows the ammonia mass fraction and the
corresponding equilibrium ammonia mass fraction. Similar to
the behavior seen previously, decreasing Dg improves heat
transfer and increases the deviation from equilibrium.
Table 3 shows that thinner reactors require less tube wall
material and catalyst (last two columns). One downside to
decreasing the reactor diameter is that the pressure drop would
increase, requiring greater pumping power. While pressure
has been held constant in the model for purposes of evaluating
properties and reaction kinetics, the pressure drop has been
calculated using the Ergun equation. Table 3 shows the
increase in pressure drop of the reacting gas as the reactor
become thinner.
0 5 10 15 20 25 30
300
350
400
450
500
550
600
650
700
750
Tem
per
ature
s, T
(oC
)
Reactor length, x (m)
Dg=13 cm: gas
Dg=13 cm: steam
Dg=10 cm: gas
Dg=10 cm: steam
Dg=7 cm: gas
Dg=7 cm: steam
Figure 7. Temperature of gas and steam along reactor length
0 5 10 15 20 25 30
0.0
0.2
0.4
0.6
Am
monia
mas
s fr
acti
ons,
fN
H3
Reactor length, x (m)
Ds=13 cm
Ds=13 cm: equilibrium
Ds=10 cm
Ds=10 cm: equilibrium
Ds=7 cm
Ds=7 cm: equilibrium
Figure 8. Ammonia mass fraction and corresponding
equilibrium ammonia mass fraction along reactor length
Tube bundle reactor. In this subsection, the use of a tube
bundle reactor is investigated. As shown in Figure 4, gas flows
in N reacting tubes while the steam flows in the opposite
direction within the gap. The number of reacting gas tubes is
varied while the cross-sectional areas for gas and steam are
fixed. Also, the wall thickness of the reactor tube, Wg,N, is fixed
to be 0.5 cm. Table 4 lists combinations of Dg,N, Ds,N, and Ds
that give fixed cross-sectional areas for different numbers of
tubes, N.
Table 3. Comparison of outlet conditions, required reactor length, and material usage for different Dg and Ds combinations
Dg
(cm)
Wg
(cm)
Ds
(cm) s
R
(W-1
∙cm∙K)
gR
(W-1
∙cm∙K)
T3
(ºC)
fNH3,out
ΔP
(MPa)
L
(m)
Vw
(m3)
Vc
(m3)
13 1 18.2 4.7 1.8 461 0.502 0.3 26.2 0.45 0.35
10 1 15.0 4.7 1.4 458 0.498 0.8 24.9 0.35 0.20
7 1 11.8 4.7 1.0 447 0.486 3 24.6 0.28 0.10
6 Copyright © 2015 by ASME
0 5 10 15 20 25 30 35
300
350
400
450
500
550
600
650
700
750
Tem
per
ature
s, T
(oC
)
Reactor length, x (m)
N=10: gas
N=10: steam
N=8: gas
N=8: steam
N=6: gas
N=6: steam
N=1: gas
N=1: steam
Figure 9. Gas and steam temperatures along reactor length
0 5 10 15 20 25 30 35
0.0
0.2
0.4
0.6
Am
mo
nia
mas
s fr
acti
on
s, f
NH
3
Reactor length, x (m)
N=10
N=10: equilibrium
N=8
N=8: equilibrium
N=6
N=6: equilibrium
N=1
N=1: equilibrium
Figure 10. Mass fraction of ammonia and corresponding
equilibrium mass fraction along reactor length
Figure 9 shows the gas temperature (solid curve) and
steam temperature (dashed curve) along the reactor length.
Referring to Table 4 and Figure 9, the required reactor length
decreases significantly as N increases. Once again, this is
caused by improved heat transfer: The thermal resistances,
especially the steam thermal resistance, are decreasing with
increasing N. As N increases, the surface area for heat transfer
increases and the steam heat transfer coefficient hs increases
because the gap between tubes decreases.
Figure 10 shows the ammonia mass fraction and the
corresponding equilibrium ammonia mass fraction. Increasing
the number of tubes, which increases the heat transfer surface
area, improves heat transfer and increases the deviation from
equilibrium. As shown in the last column of Table 4, the wall material
usage decreases as N increases. In this case, with the gas cross-
sectional area held fixed the catalyst volume decreases in
proportion to the reactor length. Furthermore, the gas pressure
gradient doesn’t change with tube number N since the cross-
sectional area is fixed, therefore pressure drop also decreases in
proportion to length. The preliminary conclusion is that a
bundle of tubes is preferable to a single tube reactor, and that
performance improves as N increases. However, this
conclusion will require further investigation using a more
accurate model of the steam flow and heat transfer in the shell.
CONCLUSIONS The ammonia synthesis reactor design study demonstrates
the viability of using ammonia as a thermochemical storage
system for the direct production of supercritical steam at 650°C.
The supercritical steam can be used with modern power blocks
for electricity generation.
A parametric study was made to investigate the effect of
geometry for optimum design. Under the high temperature
reaction conditions explored here, the ammonia synthesis
reaction is predominantly heat-transfer-limited. As heat transfer
is improved, the required reactor length decreases. Tube wall
material and catalyst volumes, which are associated with
reactor tube diameter and length, can be reduced by improving
heat transfer. The effects of geometry and dimensions of the
reactor on heat transfer have been studied. Either decreasing
tube reactor diameters or using tube bundles can improve heat
transfer within the reactor. This study provides guidance for
designing a reactor that can achieve the desired steam heating
while minimizing cost. Future research will establish the cost of
an ammonia-based thermochemical energy storage system.
NOMENCLATURE cp Heat capacity of the reacting gas mixture, J/(kg∙K)
D Diameter, m
Deff Effective diffusivity, m2/s
Ea Activation energy, J/mol
3NHf Ammonia mass fraction
h Heat transfer coefficient, W/(m2∙K)
H Heat of reaction, J/kg
Table 4. Comparison of outlet conditions and material usage for different tube bundle reactors
N
Dg,N
(cm)
Ds,N
(cm) s
R
(W-1
∙cm∙K)
gR
(W-1
∙cm∙K)
Ds
(cm)
T3
(ºC)
fNH3,out
Δp
(MPa)
L
(m)
Vw
(m3)
1 9.49 14.14 6.05 1.23 14.14 457 0.498 1.1 30.0 0.35
6 2.87 6.22 2.94 0.28 15.3 452 0.494 0.5 13.5 0.19
8 2.35 5.50 2.10 0.22 15.6 449 0.489 0.4 9.83 0.14
10 2.00 5.00 1.59 0.19 15.8 444 0.484 0.3 7.63 0.12
7 Copyright © 2015 by ASME
k Thermal conductivity, W/(m∙K)
ko,m Pre-exponential factor, kg/(m3∙s)
Kp Equilibrium constant
L Reactor length, m
m Mass flow rate, kg/s
Nu Nusselt number
P Perimeter of tube, m
p Pressure, Pa
pj Partial pressure, where j indicates species, Pa
q Heat transfer rate, W
wq Heat flux, W/m
2
r Radial dimension, m
r Rate of ammonia synthesis, kg/s
Ru Universal gas constant, J/(mol∙K)
T Temperature, °C
U Overall heat transfer coefficient, W/(m2∙K)
V Volume, (m3)
v Velocity of the reacting gas mixture, m/s
W Wall thickness, m
x Axial dimension, m
Greek letters
Density of the gas mixture (kg/m3)
Subscripts
c catalyst
eff effective
g gas mixture (N2+3H2 and NH3)
i inner
in inlet
o outer
out outlet
ph preheat
s steam
si inlet of steam stream
so outlet of steam stream
w wall
ACKNOWLEDGMENTS The information, data, or work presented herein was
funded in part by the Office of Energy Efficiency and
Renewable Energy (EERE), U.S. Department of Energy, under
Award Number DE- EE0006536.
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