This document discusses key topics related to optimizing ammonia plant performance: 1) Carbon formation in primary reformers can be minimized by maintaining proper steam-to-carbon ratios and avoiding temperature excursions. Tube wall temperatures, pressure drop, and methane approach to equilibrium should be monitored. 2) Flow distribution in catalytic reactors is important to ensure proper catalyst utilization. Poor distribution can increase costs by reducing conversion efficiency. 3) Operational upsets and transient conditions like startups and shutdowns can lead to carbon formation if not carefully controlled. Sulfur poisoning from even small amounts in feed can significantly impact temperatures.

1. AmmoniA operAtions
Nitrogen+Syngas 321 | January-February 2013 1
T
he ammonia industry has seen sig-
nificant change in recent decades as
high turnover, leaner organisational
structures, and new operations have shifted
the demographics of the operations and
technical personnel staff. These changes
have resulted in a loss of long-term opera-
tional knowledge and can lead to incidents
of equipment and catalyst damage during
operation and upset conditions.
As catalyst technology continues to
improve, the products’ high performance
are best enabled when accompanied by a
thorough understanding of the chemical and
physical principles of catalytic reactions1.
Carbon formation in
primary reformers
An important step in the ammonia manu-
facturing process is steam reforming; the
conversion of hydrocarbons into hydrogen.
The primary reformer is the main piece of
equipment for this and represents the larg-
est expenditure in terms of capital and on-
going energy costs. Optimum performance
of the reformer and the installed reforming
catalyst is critical to ensuring high plant
productivity and efficiency. Poisoning, foul-
ing or incorrect operation can adversely
affect the catalyst’s performance and can
lead to costly equipment failure.
Primary reformer reactions
The main reactions in the primary reformer
are:
Steam reforming of hydrocarbons
CnH2n+2 + n H2O n CO + (2n+1) H2
∆H 0
Steam reforming of methane
CH4 + H2O CO + 3 H2
∆H = 88700 Btu/lb-mole
(206 kJ/mol)
Water-gas shift
CO + H2O CO2 + H2
∆H = -17,600 Btu/lb-mole
(-41 kJ/mol)
Overall, the process is endothermic. A lower
CH4 leakage is favoured by higher exit tem-
perature, higher steam-to-carbon ratio, and
lower exit pressure.
The standard primary reformer system
is comprised of a tubular furnace where
the feed stream passes over a catalyst
packed in multiple banks of externally
heated tubes. With improvements in tube
metallurgy, primary reformers operate at
pressures up to 6,000 kPa and tube wall
temperatures up to 1010°C. The aver-
age heat flux may be as high as 94,000-
110,000 W/m2.
Firing is usually controlled such that
tube wall temperatures are maintained at
values that will give a reasonable tube life.
By design and industry practice, maximum
allowable tube wall temperatures give an
in-service life of about 100,000 hours.
Carbon formation reactions
With no steam present and at normal
reformer operating temperatures, all hydro-
carbons will decompose into carbon and
hydrogen via the following reaction:
Thermal cracking
CH4 C + 2 H2
Cracking reactions are favoured thermody-
namically at high temperature. In the pres-
ence of steam, gasification reactions also
occur over the primary reformer catalyst as:
Gasification
C + H2O CO + H2
Operating the reformer at conditions that
drive this reaction can prevent the accumu-
lation of carbon deposits. When conditions
exist that favour hydrocarbon cracking,
the heavier hydrocarbons in the feed will
Keylessonsto
optimiseammonia
plantperformance
Understanding the basic principles of catalytic reactor operation in an ammonia plant can
help to avoid costly performance and operating problems. M. Anderson, and S. Osborne
of Clariant Corporation (formerly Süd-Chemie) focus on three particular topics that have
a significant impact on production performance: 1) chemistry, consequences and avoidance
of carbon formation in the steam reformer; 2) proper flow distribution in catalytic reactors,
and 3) the impact of improper plant operation on catalyst and equipment.

2. AmmoniA operAtions
2 Nitrogen+Syngas 321 | January-February 2013
1500
1550
1600
1650
1700
1750
0 5 10 15
temperature,°F
months on stream
Benchmark
Clariant
Fig 1: Comparative average TWT
surveys of two charges of
catalyst at a 1200 STPD
ammonia plant
Fig 2: IR thermographic image of a
row of reformer tubes
1500
1525
1550
1575
1600
1625
1650
EW
EE
DW
DE
CW
CE
BW
BE
AW
AE
1625-1650
temperature °F
1600-1625
1575-1600
1550-1575
1525-1550
1500-1525
North
1625-1650
temperature °F
1600-1625
1575-1600
1550-1575
1525-1550
1500-1525
top level: surface profiletop level: surface profile
top level: top profiletop level: top profile
North
EW
EE
DW
DE
CW
CE
BW
BE
AW
AE
Fig 3: Plotted TWT surveys showing
the heat distribution across the
reformer
crack first. As this occurs, the active sites
of the catalyst become masked, resulting
in less reforming reaction and higher gas
temperatures, which further increases the
tendency for cracking and coke deposition.
Common causes of carbon formation
and deposition are low steam/gas ratio
operation, heavy hydrocarbons in the feed,
sulphur poisoning, and temperature excur-
sions resulting from poor firing control.
Higher catalyst activity can reduce carbon
formation potential by lowering gas temper-
ature when methane concentration is high
enough for cracking, thereby reducing the
carbon formation potential.
The overall effect of carbon formation on
the primary reformer is reduced conversion,
increased pressure drop and increased tube
wall temperatures which, if severe, can be
seen as ‘hot bands’ on the tubes. Not only
does the carbon deposit block active sites,
its formation within the pores of the catalyst
can cause catalyst breakage.
Monitoring program
During routine operation of a primary
reformer, a monitoring program assists in
optimising the catalyst’s performance as
well as providing early detection of issues
that can lead to catalyst or tube failure.
Three performance parameters are typically
trended during operation; calculated CH4
approach to equilibrium (ATE), tube wall tem-
peratures (TWT), and relative pressure drop.
Approach to equilibrium
Reformer catalyst activity is commonly
expressed as CH4 ATE, which is the differ-
ence between the measured catalyst outlet
temperature and the calculated equilibrium
temperature for the observed methane
leakage. An increased ATE or methane slip
from the front-end results in a higher purge
rate from the synthesis loop and increased
energy consumption for the plant. While
operating the primary reformer at higher
exit temperature can mitigate this effect,
the cost of increased fuel requirements
and the impact of hotter tube wall tempera-
tures must be considered when choosing
to operate in that non-optimal mode.
Tube wall temperatures
Tube costs are a significant item in the
overall economics of a reforming plant.
The goal is to operate the reformer under
conditions that result in the lowest possi-
ble tube wall temperatures (TWTs) consist-
ent with satisfactory reformed gas quality.
Even a slight increase in the tube wall
temperatures will have a drastic impact on
the tube life. A TWT increase of as little as
10oC may result in up to a 30% shortened
life-time of the reformer tubes.
The use of a high activity/shape opti-
mised reforming catalyst results in lower
tube wall temperatures and a significant
cost and efficiency advantage to an opera-
tor. As an example, a 1,200 short t/d plant
in North America was able to realise such
an advantage by installing the high activity
ReforMax
®
reforming catalyst from Clariant.
The Kellogg-designed primary reformer of
this plant was initially designed to produce
600 short t/d of ammonia. After several
expansions, the reformer now operates at
a high heat flux of around 110,000 W/m2.
Figure 1 shows that the mean TWT is 55°C
cooler with catalyst from Clariant than with
the previous competitor charge.
Additionally, the maximum tube wall
temperature was also lowered and the
number of tubes with TWTs exceeding the
temperature limit of 927°C was cut from
five to zero, significantly reducing the risk
of potential tube failure.
Effective monitoring of the TWTs and
balancing of the firing of the reformer fur-
nace can extend the life of the tubes and
optimise the performance of the catalyst
(Figures 2 and 3).
Pressure drop
The impact of increasing primary reformer
pressure drop is reduced plant efficiency,
caused by the lowering of the suction pres-
sure at the synthesis gas compressor, and
the resulting higher horsepower require-
ments to maintain the desired loop pres-
sure. Depending on pressure relief valve
set-points, a higher reformer ∆P may also
require reduction of plant operating rates if
near those limits.
Using Clariant’s proprietary Cat-Trends
®
program, calculated ATE and relative
pressure drop can be trended over time.
Carbon deposition on the catalyst will
manifest itself as increasing ATE, increas-
ing relative pressure drop and increasing
TWTs. In severe cases, visual hot bands
on the tubes will also become evident.
Operational control factors for
minimising carbon formation
To prevent carbon formation, a minimum
steam-to-carbon (S/C) ratio must be main-
tained at all times. This value is contingent
on the composition of the feed gas, firing
rates, operating pressure and incoming
feed temperature. Most carbon formation

3. AmmoniA operAtions
Nitrogen+Syngas 321 | January-February 2013 3
0
10
20
30
40
50
60
70
0.0 1.0 2.0 3.0
sulphur in feed, ppmv
increaseintubewalltemp,°F
Fig 4: The impact of sulphur
concentration in feed gas on
the tube wall temperature
in the reformer
Fig 5: The reformer tube appearance before and after a decoking procedure
incidents occur, however, during transient
conditions such as start-up, operating trips
or shutdowns. Modern operating plants
typically have a low S/C trip interlock
that cuts out the feed-gas to protect the
reformer during upset conditions. Addition-
ally, a properly installed minimum stop on
the steam valve ensures that steam con-
tinues to flow through the reformer in an
emergency trip to sweep all feed-gas out
of the reformer. Complete isolation of the
feed-gas valve during a trip or shut-down is
necessary before steam cut-out.
Controlled start-up and shutdown pro-
cedures and emergency interlocks are
designed to avoid carbon formation. The
objectives are to:
l Prevent over-firing and thus tempera-
ture excursions during rate changes.
This typically considers the following
operating parameters: fuel flow, fuel
pressure, flue-gas temperature, proc-
ess gas pressure, steam flow, feed-gas
flow, combustion air flow, exit process
gas temperature and the number and
location of burners lit. Often a manual
start-up rate guide is used and in some
cases an algorithm with interlocks
guides the operator during the start-up.
l Maintain a high S/C ratio during rate
changes. This is typically achieved by
increasing the steam flow before the feed-
gas on start-up and reducing the feed-gas
before the steam on shutdowns.
l Ensure sufficient and even flow distri-
bution throughout the entire reformer.
Before feed gas is introduced to the
reformer during start-up, the steam
flow should be sufficient to ensure
that all tubes have similar operating
conditions i.e. even flow and even heat
through all tubes.
l Ensure sufficient steam flow exists for
adequate mixing of steam and feed-gas
well before the tube inlet manifold.
Visual monitoring of the tubes and burners
is critical to detect any issues during the
start-up process.
Sulphur poisoning as root cause for
carbon formation
Sulphur has an adverse impact on reform-
ing catalyst activity as it reacts with the
reforming catalysts active nickel compo-
nent and forms an equilibrium condition
over the active sites of the catalyst:
Ni + H2S NiS + H2
The sulphur must be removed from the
feed-gas in an upstream desulphurisation
step to levels less than 0.1 ppmv. Sulphur
is not a permanent poison for short-term
exposures and when removed from the
feed, the catalyst can regain the original
activity. However, sulphur poisoning can
deactivate the catalyst to the extent that
less reforming is done and elevated gas
temperatures cause high tube wall tem-
peratures. Even small amounts of sulphur
in the feed can significantly influence the
tube wall temperatures as shown in Fig. 4.
If the temperature rises to the point
where carbon formation occurs, pressure
drop will increase and catalyst perform-
ance will be further compromised.
In a 600 short t/d design ammonia plant,
the desulphurisation unit was bypassed due
to vessel design concerns. During this time,
the high performance ReforMax primary
reformer catalyst was exposed to inlet sul-
phur levels of up to 3 ppmv for nine months.
As expected, catalyst poisoning occurred
which resulted in activity loss in the 30-40%
top inlet tube location.
The plant was subsequently brought
down and a successful steam-air regenera-
tion procedure was done with the assistance
of Clariant technical support. On re-start, the
reformer experienced full recovery to pre-inci-
dent performance levels and tube wall tem-
peratures; visual success of the decoking
procedure is reflected in Fig. 5.
The ReforMax catalyst withstood the
intense conditions during this high sulphur
exposure time-frame and continued to pro-
vide the full expected design life-time.
Given the impact of sulphur poisoning
on the performance of the reformer, the

4. AmmoniA operAtions
4 Nitrogen+Syngas 321 | January-February 2013
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
1 10 14 17 19 25 34
time on-stream, months
CH4
leakage,%
Fig 6: Secondary reformer exit
methane slip over time
Fig 7: Photo of the surface of the secondary reformer showing mixing of catalyst
and inert balls and bed disruption
choice of high activity desulphurisation
catalyst should not be under-estimated.
Proper monitoring of the total sulphur slip
from the pre-treatment section ensures
that the reforming catalyst is protected.
Flow distribution in reactors
The design of a catalytic reactor considers
effective flow and distribution of the feed-
gas through the catalyst bed volume. Poor
distribution of gas flow can lead to bypass-
ing/channeling within the catalyst bed and
ineffective utilisation of the loaded cata-
lyst volume. This manifests itself in higher
calculated ATE in primary and secondary
reformers, higher CO leakage in HTS and
LTS reactors, higher COx from methana-
tors and reduced conversion from ammonia
converters. The impact of these non-opti-
mal conversions is less efficient and more
costly operation of the NH3 plant.
Using the results of a study carried
out to examine the impact of such opera-
tion, the cost impacts can be converted
to dollar values to demonstrate the large
impact they can have on overall plant prof-
itability. This study was based on a 1,500
short t/d ammonia plant with membrane
purge gas recovery2. Assuming a produc-
tion rate of 1,500 short t/d, $2.70 per
MMBtu natural gas and $450/t ammonia,
the table above shows the results in terms
of potential costs.
As shown in Table 1, the increase in
costs associated with lost production and –
if the lower production can be compensated
by a higher energy input – in energy costs
can be significant for these different sce-
narios. It is clear that proper initial design,
procedures and operation is imperative to
ensure stable and optimal catalyst perform-
ance from start of run (SOR) to end of run
(EOR). The potential causes of poor flow in a
catalyst bed include gas distributor issues,
vessel design issues, movement of catalyst
during operational upsets and initial loading
techniques.
Movement of catalyst at the top of a bed
caused by poor feed distributor designs,
burner issues in a secondary reformer or
plant upsets can create flow or channeling
issues, and result in changes in perform-
ance. Such an event was experienced
recently in a 1,750 short t/d ammonia plant.
During operation, the secondary reformer
experienced an increase in methane leak-
age from around 0.3% to over 0.5% (Fig. 6).
On shutdown it was found that there was
mixing of the catalyst and inert balls which
were pushed up toward the walls of the ves-
sel (Fig. 7). The root cause was traced back
to an operational upset that caused high
gas velocities and bed movement creating
unevenness/flow distribution issues in the
bed which subsequently compromised the
overall performance.
The purpose of any initial loading of cat-
alyst and support media is to achieve a uni-
form density throughout the bed to ensure
proper flow distribution. In fixed beds cross-
sectional uniformity is also desired.
The primary reformer is one of the most
intensive and time-consuming reactors to
load and carries severe penalties if tubes
are loaded improperly. With no physical
means for balancing the feed gas flow
through tubes, catalyst with non-uniform
loaded density results in non-uniform flow
rates which in turn cause some tubes
to operate hotter than the average tube
temperatures. Additionally, catalyst bridg-
ing can occur when particles interlock to
form a “bridge” across the tube cross-
section. These voids within the catalyst
layer result in localised overheating of the
tube. Hot tubes fail prematurely and can
be very costly. It has been estimated that
Change in performance Change in energy
cost ($/annum)
Lost production
($/annum)
Primary
reformer
5.6°C increase in ATE 15,000 170,000
Secondary
reformer
5.6°C increase in ATE 30,000 370,000
HTS 1% increase in CO leakage 56,000 590,000
LTS 0.1% increase in CO leakage 160,000 2,000,000
Table 1: Impact of changes in performance of selected front-end catalysts on
energy costs and lost production in a 1,500 short t/d ammonia plant
with membrane purge gas recovery

5. AmmoniA operAtions
Nitrogen+Syngas 321 | January-February 2013 5
Fig 8: Photo of reduced particle size of HTS catalyst after
exposure to extreme temperatures caused by air
introduction to the HTS reactor
Fig 9: Photo of fused inert ball, floating screen and catalyst
removed from the HTS following exposure to air and
extreme temperatures
a 2.5% lower than average tube flow rate
can result in an 4-6°C hotter tube wall tem-
perature3. Further experience has shown
that tube wall temperatures in the area of
a catalyst void are about 16-28°C hotter
than the surrounding area of the tube.
The impact of plant operations on catalyst
Many catalyst failures are the result of
poor operation or operational upsets. To
illustrate the significant potential negative
impacts, a number of different examples
for various catalytic services are discussed
below.
Reduced/pyrophoric catalyst
Catalysts are often in the reduced metal
state during operation in HDS, HTS, LTS,
methanator and ammonia synthesis reac-
tors. Exposure to air can result in spontane-
ous ignition with extremely high localised
temperatures causing catalyst failure,
fusing and equipment damage. For turna-
rounds and longer shut-down periods,
ensuring a continuous positive pressure
inert blanket is one routine monitoring item.
Discharging of pyrophoric catalyst without
first oxidising it must be done under nitro-
gen with precautions taken to avoid ingress
of air to the vessel by a chimney effect. In
cases where repair work is required around
vessels without adequate isolation, the risk
is particularly high and continuous nitrogen
purge is necessary. It should be noted that
without flow through the catalyst bed, tem-
perature measurements do not accurately
represent the bed temperatures nor do
they detect localised peak temperatures.
They can, however, give some indication or
warning of an issue.
Complete isolation of the air flow to the
secondary reformer is required in a sec-
ondary reformer trip to protect downstream
catalysts. In one incident at a 1,700 short
t/d ammonia plant in North America, the
HTS catalyst performance was greatly
affected following an upset during start-up
of the secondary reformer. A high pres-
sure air flow condition resulted in air to
the HTS for a short duration. The event
occurred relatively fast and other events
were occurring so it was not even noticed
by the operations team.
Shortly after stabilising plant rates,
however, the performance of the charge in
terms of CO leakage, while acceptable to
allow operation, was higher than expected.
After approximately a year on-stream, the
decision was made to unload the catalyst
during a shut-down. Inspection of the reac-
tor showed that the overall outage from the
manway had increased by over 0.9 m and
that in a particular area of the bed the cat-
alyst particle size was reduced from 6x6
mm to roughly 3x3 mm. The support balls
were also found fused to melted screen
pieces which in turn were fused to cata-
lyst. For this level of shrinkage and fusion
to occur it is estimated that temperatures
must have exceeded 870°C. It is important
to note that despite this significant temper-
ature exposure, the Clariant catalyst was
still able to achieve an acceptable level of
performance with no pressure drop issues.
(Figs 8 and 9)
Impact of water on catalyst
Water can affect the integrity of all cata-
lysts and in some cases can lead to cata-
strophic incidents. Water entrainment
during start-up of the primary reformer can
lead to instantaneous flashing at the hot
catalyst releasing enough energy to rupture
reformer tubes. Water entrainment can
occur when low point drains are plugged
or not used during shutdown and start-up.
HTS and LTS are particularly at risk due to
their location downstream of the HP boiler,
quenches and temperature controls. When
water comes in contact with catalyst, cata-
lyst breakage and increased pressure drop
results from water vaporising out of the cata-
lyst pores at a faster rate than the porosity
of the catalyst will allow. The HTS catalyst is
very susceptible to leaks from the upstream
HP boiler which are almost inevitable during
a plant’s life. Clariant’s HTS ShiftMax
®
cata-
lyst was able to survive major boiler leaks
at the 1,600 short t/d Terra plant4. In this
case, the HTS was flooded with water from
a major boiler failure and a dry out proce-
dure was successfully implemented under
the direction of Clariant. Besides the effect
of wetting the catalyst, accumulated boiler
solids from boiler leaks can also increase
the catalyst’s pressure drop and create poor
gas distribution issues. The impact of water
and boiler solid carryover can be minimised
by using a high voidage inert top layer.
During operational upsets, temperature
controller swings can lower the inlet tem-
perature of the shift converters below the
dew point and wet the catalyst. Operating
the LTS at an inlet temperature 17-22°C
above dew point can reduce this potential.
Boiler feed water quenches are normally
installed upstream of shift converters in
order to control inlet temperature as well
as lower the CO equilibrium. Atomisation
is critical to prevent wetting the catalyst

6. AmmoniA operAtions
6 Nitrogen+Syngas 321 | January-February 2013
extent that the resultant exotherm elevates
the methanator temperature to equipment
design limits, process gas should be imme-
diately shut off or vented upstream.
The operation and control of the CO2
removal system is critical to maintaining
the inlet levels to the methanator. Plants
that have been revamped or are running at
higher than design throughputs can push
the CO2 removal system to maximum load.
At these conditions, small fluctuations can
lead to gross carryover of solution and CO2
breakthrough. Some solutions can poison
and permanently deactivate the methana-
tor catalyst, or leave deposits on the top
of the bed which reduces catalyst activity
by blocking active sites and also leads to
increased pressure drop. Severe liquid
carryover of all solutions can also cause
physical damage of the catalyst further
increasing the pressure drop. A skim of the
catalyst is can sometimes be performed to
correct the situation.
Conclusion
In summary, understanding the basic prin-
ciples of catalytic reactor operation in an
ammonia plant can help to avoid costly per-
formance and operating problems. Opera-
tions and technical crews should routinely
review and analyse process monitoring
systems and procedures to ensure that the
consequences of mal-operation are under-
stood and best available procedures and
technology are in place. Understanding the
impact on catalyst and process equipment
from a safety perspective and an opera-
tional cost standpoint is important so that
potential adverse situations are avoided
or minimised in magnitude. With this com-
plete knowledge and understanding, high
performance catalysts can achieve their full
utilisation and maximise the efficiency and
productivity of the plant. n
References
1. Gebert S., Cai Y,. and Kniep B.: “Continuous
improvement in ammonia catalysts” Nitro-
gen + Syngas (Jan-Feb 2012).
2. Rice D.: “The value of catalyst performance
in ammonia plants” AIChE 2002.
3. Rice D.: “Loading of primary reformer cata-
lyst tubes” AIChE 1993.
4. Borzik K., Howerton R. and Archer T.: “Suc-
cessful recoveries from major WHB failures
experience” AIChE 2007.
5. Richardson J., Wagner J., Drucker R., Rajesh
H.: “Understanding hydrocarbon reactions
during LTS catalyst reductions” AIChE 1996.
and proper design and maintenance of
the quench valve/nozzles is imperative.
Quench valves also require complete iso-
lation during a shutdown to prevent water
buildup in low point piping areas that can
be entrained onto the catalyst on restart.
Water treatment
The steam quality in an ammonia plant is
typically monitored to ensure protection of
rotating equipment. However it is also an
important consideration for reforming cata-
lyst. Carry-over and deposition of silica and
sodium on reforming catalyst can act like a
permanent poison by physically covering the
active sites of the catalyst. This can cause
poor reformer performance, increased pres-
sure drop and increased TWTs at the top
of the tubes. These contaminants can also
be introduced to the process via boiler feed
water quenches upstream of the reformer.
A well controlled and monitored water treat-
ment program is necessary for effective
catalyst performance.
LTS reduction
LTS reduction is a procedure required to
activate the catalyst and is typically per-
formed every 3 to 4 years. It is a process
that can have catastrophic consequences
if not done carefully. The infrequency of
the procedure requires a thorough review
before performing the procedure.
The CuO in the LTS catalyst is reduced
to copper metal by an exothermic reac-
tion. To effect the reduction, hydrogen is
carefully and precisely added to a carrier
gas (natural gas or nitrogen) and the bed
temperatures are closely monitored. Meas-
ures taken to achieve adequate control of
hydrogen include:
l complete isolation/blinding of any
potential sources of hydrogen other
than reduction gas especially when at
higher pressure than the carrier gas;
l gas composition analysis of the reduc-
tion and carrier gases (Note: natural
gas carrier gas should be free of hydro-
gen recycle);
l reliable supply of reduction and carrier
gas flows, including reduction gas flow
meter calibration at constant conditions
of carrier gas flow and back pressure;
l Gas testing schedule of inlet and outlet
hydrogen concentrations.
Caution statements in a procedure describe
reducing or removing reduction gas at cer-
tain peak temperatures. Achieving optimum
catalyst performance and preventing tem-
perature run-away are both considered.
Studies have been done to understand
what happens at elevated temperatures
that cause run-away. In the case of a natu-
ral gas carrier, CuO reacts with methane at
246°C via the following reaction5:
CH4 + 4CuO → 4Cu0 + CO2 + 2H2O
This is also exothermic creating even
higher bed temperatures. At temperatures
near 370oC, hydrogen can be produced by
hydrocarbon cracking, which can lead to
temperature run-away. Under such circum-
stances vessel design temperatures can
be exceeded.
Methanator operation
Methanator operations are seldom given
much attention and COx concentrations are
typically not measured. However upstream
conditions affect its performance and can
have an adverse effect on plant efficiency,
production and even equipment.
During start-up, the front end of the
plant is brought up and the synthesis gas
is introduced into the synthesis loop. The
inlet temperature of the methanator is typi-
cally controlled by available steam and the
feed/effluent exchanger. The introduction
of low temperature make-up gas to the loop
can introduce higher levels of COx. These
higher concentrations increase the poten-
tial for ammonium carbamate formation
within the loop that fouls equipment and
increases the loop pressure drop, which
in turn increases the energy requirements
of the plant. In addition to ammonium car-
bamate formation, ammonia synthesis
catalyst is temporarily poisoned by oxygen-
ates and this will adversely impact plant
production and loop performance. Once
the COx levels are returned to normal,
production rates can take quite some time
before recovering. If ammonia synthesis
catalyst is exposed to high COx levels for
long periods of time, the poisoning affects
can be permanent.
In addition to operating conditions at
the methanator, upsets at the upstream
LTS reactor that cause increased inlet CO
to the methanator can overload the avail-
able capacity of the catalyst and lead to
higher COx leakage.
Upsets at the CO2 removal unit tend to
be more frequent and can have a greater
impact on the methanator. Also CO2 is
more difficult to methanate than CO, so
higher levels can overwhelm the methana-
tor and result in premature COx leakage. In
cases where the CO2 increases to such an