Getting the Most Out of Your Refinery Hydrogen Plant

Refinery Process Stream Purification Refinery Process Catalysts Troubleshooting Refinery Process Catalyst Start-Up / Shutdown
Activation Reduction In-situ Ex-situ Sulfiding Specializing in Refinery Process Catalyst Performance Evaluation Heat & Mass
Balance Analysis Catalyst Remaining Life Determination Catalyst Deactivation Assessment Catalyst Performance
Characterization Refining & Gas Processing & Petrochemical Industries Catalysts / Process Technology - Hydrogen Catalysts /
Process Technology – Ammonia Catalyst Process Technology - Methanol Catalysts / process Technology – Petrochemicals
Specializing in the Development & Commercialization of New Technology in the Refining & Petrochemical Industries
Web Site: www.GBHEnterprises.com
GBH Enterprises, Ltd.
Getting the Most Out of
Your Refinery Hydrogen
Plant
Process Information Disclaimer
Information contained in this publication or as otherwise supplied to Users is
believed to be accurate and correct at time of going to press, and is given in
good faith, but it is for the User to satisfy itself of the suitability of the Product for
its own particular purpose. GBHE gives no warranty as to the fitness of the
Product for any particular purpose and any implied warranty or condition
(statutory or otherwise) is excluded except to the extent that exclusion is
prevented by law. GBHE accepts no liability for loss, damage or personnel injury
caused or resulting from reliance on this information. Freedom under Patent,
Copyright and Designs cannot be assumed.

Contents Getting the Most Out of Your Refinery
Hydrogen Plant
Summary
1 Introduction
2 "On-purpose" Hydrogen Production
3 Operational Aspects
4 Uprating Options on the Steam Reformer
4.1 Steam Reforming Catalysts and Tube Metallurgy
4.2 Oxygen-blown Secondary Reformer
4.3 Pre-reforming
4.4 Post-reforming
5 Downstream Units
6 Summary of Uprating Options
7 Conclusions

"Getting the Most Out of Your Refinery Hydrogen Plant"
Summary
As demand for hydrogen in refineries has increased, so the operation of
existing hydrogen plants has increasingly come under the spotlight. It has
become essential to get the maximum output from an existing plant, in
order to prevent or minimize expenditure for any additional hydrogen.
This document describes in some detail the options which are available to
the operator of an existing hydrogen plant. This includes optimizing the
operating conditions, using latest generation catalysts and steam reformer
tube metallurgies, and retrofitting pre-reformers, post-reformers or
secondary reformers in order to maximize hydrogen production. Each
option is described, with an order-of-magnitude estimate of the typical
increase in throughput likely, and the cost involved.
1 Introduction
Hydrogen is produced for use in a number of industrial applications. The largest
of these is for use as Synthesis Gas, or Syngas, for use in the production of
ammonia and methanol. After that, the next largest use for hydrogen produced
"on-purpose" (that is, not as a by-product from another process, such as catalytic
reforming in refineries) is in refineries. About 85% of "on-purpose" hydrogen is
used in refineries, if hydrogen produced for use as Syngas is excluded.
Within the refinery, hydrogen is available as a by-product from a number of
refinery operations (principally the catalytic reformer), but is required for a
number of other refinery operations (principally hydrotreating and hydrocracking).
The balance between hydrogen production and hydrogen demand on the refinery
is usually termed the refinery hydrogen balance. If there is insufficient by-product
hydrogen to meet the hydrogen consumption demands on the refinery, then in
order to restore the hydrogen balance, extra hydrogen is needed. This
supplementary hydrogen can be purchased, or more typically, produced on site
using a dedicated "on-purpose" hydrogen plant.

The hydrogen balance is affected by the type of refinery. A "simple" refinery
contains no conversion processes, and so hydrogen is usually in surplus, with no
need therefore for any "on-purpose" hydrogen production.
Such refineries, however, are constrained to the production of larger amounts of
atmospheric residue or residual fuel oil. A "conversion" refinery has additional
units, typically fluid catalyst cracker (FCC) and coker units, that upgrade residual
fuel oil to lighter products, especially gasoline. Here, in the past, hydrogen was
just in balance.
Finally, refineries with hydrocrackers, which produce premium quality middle
distillates, almost always need supplementary hydrogen.
North America, with its high demand for gasoline, has the most refinery
conversion capacity, and therefore the most "on-purpose" hydrogen production.
Other regions such as Asia-Pacific and Europe have less "on-purpose" hydrogen
capacity. However, over the past decade, there have been significant changes to
the refinery hydrogen balance, particularly in North America. These changes
have been driven by the Clean Air Act Amendments (CAAA) of 1990, which
required low sulfur diesel and reformulated gasoline (RFG). This resulted in the
need for more hydrogen for hydrotreating, but also in the production of less by-
product hydrogen, as the catalytic reformer operation was modified to allow for
the lower levels of volatile organic compounds (VOC) permitted in RFG. As a
result, there was a surge in demand for additional refinery hydrogen in the early
1990s in North America.
Clean Air Act Amendments of 1990
Another set of major amendments to the Clean Air Act occurred in 1990
(1990 CAAA). The 1990 CAAA substantially increased the authority and
responsibility of the federal government. New regulatory programs were
authorized for control of acid deposition (acid rain) and for the issuance of
stationary source operating permits. The NESHAPs were incorporated into
a greatly expanded program for controlling toxic air pollutants. The
provisions for attainment and maintenance of NAAQS were substantially
modified and expanded. Other revisions included provisions regarding
stratospheric ozone protection, increased enforcement authority, and
expanded research programs.

Similar environmental legislation is now also taking effect in other parts of the
world. The European Union (EU) is introducing similar lower sulfur and lower
VOC specifications; Japan and some other parts of Asia-Pacific are also doing
the same; and India has significantly lowered the level of sulfur permissible in
diesel. In addition, although the demand for motor fuels is fairly flat in North
America and Europe, it is increasing rapidly in the developing nations, particularly
in Asia-Pacific. This means that the demand for hydrogen world-wide is set to
continue, though there is a geographic shift in demand.
The hydrogen requirements for the worldwide refining projected to 2030:
The following forecast also quantifies anticipated change in supply from gasoline
reforming versus on-purpose refinery / merchant production.
• The refining industry will require more than 14 trillion SCF of on-purpose
hydrogen to meet processing requirements between 2010 and 2030.
• Asia Pacific and the Middle East will represent nearly 40% of global
requirements.
• Despite completion of most ultra low sulfur on-road transportation fuel
requirements, North America and Europe will represent the largest portion
of incremental hydrogen demand during the first half of this decade.
• Nearly 2/3 of incremental refinery hydrogen demand will be for expanding
hydrocracking operations.
Much of the projected growth in North America has already happened; growth in
other regions is now occurring.
The basic choices for increasing refinery hydrogen availability are:
1. Increased by-product hydrogen formation. This, however, is usually
insufficient to meet the needs of the new refinery operating
conditions.
2. Purchase bulk hydrogen from Industrial Gas companies. This,
however, can be expensive unless there is access to a major
hydrogen pipeline network.

3. Recover by-product hydrogen from other refinery processes. This
can be done using Pressure Swing Adsorption (PSA), membrane
separation, or cryogenic separation. These options need to be
evaluated economically.
4. Expanding or building new "on-purpose" hydrogen capacity. This is
the area on which the rest of this document will focus.
2 "On-purpose" Hydrogen Production
The most widely used technology is steam reforming, often referred to as Steam
Methane Reforming (SMR), irrespective of the actual feedstock used. Although
some other technologies are used, such as Partial oxidation (POX), Autothermal
Reforming (ATR) and Integrated Gasification Combined Cycle (IGCC), these only
become economically attractive under certain conditions, and SMR technology
will continue to be the technology of choice for at least the next decade. This
document will therefore focus on SMR technology.
A typical process flow diagram for the production of "On-purpose" hydrogen is
shown in Figure 2 below.
Figure 2. "On-purpose" Hydrogen Plant
Hydrocarbon
Feed
Steam
Shift
Conversion
Steam
Reforming
Sulphur
Removal
PSA
Unit
Product
Hydrogen
Purge Gas to
Steam Reformer
Burners

The hydrocarbon feedstock, be it natural gas, shale gas, LPG, refinery off-gas or
naphtha, needs to be purified (principally to remove sulfur-containing and halide
compounds from the feed which would otherwise poison the downstream
catalysts.
The hydrocarbon feed is then mixed with steam and steam reformed to convert
the hydrocarbons to hydrogen, carbon oxides, unreacted hydrocarbon (now
principally methane) and steam. This product gas now needs to be purified,
which is typically done in a modern plant by using a High Temperature Shift
(HTS) reaction to convert the carbon monoxide to carbon dioxide, followed by a
PSA unit to purify the hydrogen to the desired level. The PSA purge gas is used
as fuel to fire the steam reformer (typically there is a requirement for some top-up
fuel such as natural gas to be used in conjunction with the PSA purge gas).
The heart of the process is the steam reformer, and it is here that most
opportunity lies for increasing the capacity of an existing unit. This document will
consider firstly the potential for increased throughput in an existing plant by
changing operating parameters, and then look at a range of other more complex
uprating options which may need plant modifications.
3 Operational Aspects
Steam reforming is an endothermic process, and large amounts of heat need to
be supplied in order to make the reaction occur at acceptable rates, even in the
presence of a catalyst. The steam reformer is, therefore, in simplest terms a heat
exchanger, with heat being transferred from the hot flue gas in the steam
reformer box to the cooler process gas within the catalyst-filled steam reformer
tubes. There are a number of differing process occurring. On the outside of the
tubes, heat is transferred by radiation together with chemical reaction in the form
of the combustion of the reformer fuel gas.
On the inside of the tube, there is a complex combination of heat and mass
transfer coupled with chemical reaction. Furthermore, since steam reformers
operate at high temperatures and pressures, the tube metallurgy becomes a key
issue. It is not surprising then that there is often scope to optimize the
performance of the hydrogen plant, and particularly the steam reformer.

The first step in understanding the operation of an existing hydrogen plant is to
gather sufficient process data and plant data to allow the construction of a plant
model. GBH Enterprises uses its in-house process simulator "VULCANIZER", in
conjunction with its detailed steam reformer modelling capabilities. Clearly, every
plant is different, and so each model will differ. The principal is the same, though;
model the process and in particular the steam reformer as accurately as
possible, and then see what scope there may be for increased throughput by
altering the operating parameters.
A detailed study has been conducted, of a modern hydrogen plant, which was
rigorously modelled and the results compared with plant operating data. It was
shown that modelling predictions were indeed borne out on the plant. In this
particular case, the modern design of the plant meant that there was in fact little
scope for significant increases in throughput. However, increases of a few
percent were seen, merely by making simple changes to the way in which the
plant was operated. An example for this specific plant is shown in Figure 3.
Figure 3. Effect of Steam Reformer Operating
Conditions on Hydrogen Plant Output
3.4
3.8
4.2
4.6
Methane Slip - %
3.4
3.6
3.8
4
4.2
4.4
Steam:Carbon Ratio
98
99
100
101
Relative
Plant Output - %
Note that whilst these trends may be similar for other plants, the absolute
numbers could well be higher.

One of the key parameters of steam reformer operation is the tube wall
temperature. The importance of this, and ways in which tube wall temperatures
can be measured accurately have been described elsewhere.
It is clear that the easiest and most commonly used methods for tube wall
temperature measurement are prone to significant error, in that they tend to
overestimate the tube wall temperature. This means that the operator believes
that he is operating to a tube wall temperature limit, when in fact he is not. The
result will be longer than expected tube lives, reducing the cost of re-tubing by
delaying it, often for many years. However, the plant throughput is being
artificially constrained, often by a significant amount (perhaps 10-15%). Before
any change to operating conditions is made, it is recommended that the steam
reformer operation is thoroughly reviewed.
4 Uprating Options on the Steam Reformer
In this section, a number of process options for increasing the throughput of an
existing hydrogen plant are reviewed. It should be noted, however, that the
increases in throughput quoted are "typical" only; each plant is different, and in
each case, due account of hydraulic and mechanical limitations must also be
taken; these are not considered in this document.
Thus, although an increase in capacity may be quoted, this may in reality not be
easily or economically achieved if there are other substantial plant limitations
(such as pressure drop, compression costs, flue gas coils and so forth) or
problems in plant integration (with for example the plant steam or fuel systems)
to overcome. Such matters can only be evaluated on a case-by-case basis.
4.1 Steam Reforming Catalysts and Tube Metallurgy
Since the 1930s, the standard shape for steam reforming catalyst has
been the simple Raschig ring. In recent years, however, it has become
clear that this shape was not ideal, and significant work has been done to
produce an optimized shape for steam reforming catalyst.
There are a number of parameters that need to be optimized. The steam
reforming reactions are endothermic, and to maximize process efficiency,
it is important to get as much heat as possible into the reactant gases
inside the steam reformer tubes as quickly as possible.

The limit of the steam reforming operation is often referred to as the
maximum heat flux that can be tolerated, though strictly speaking, this is
not the limiting factor.
The real limit is the tube wall temperature (TWT) which can be tolerated;
whether the flux is high or low is immaterial, as long as the permissible
TWT for the metallurgy under the process conditions is not exceeded.
Clearly, activity has a key role to play - the more reaction there is, the
cooler the tubes will be for a given firing rate, since the steam reforming
reaction is endothermic.
For the steam reforming reaction, the reaction takes place very much at
the surface, and so the activity is virtually directly proportional to the
geometric surface area (GSA) of the catalyst per unit volume. Thus, a
convenient and simple way to increase steam reforming catalyst activity
without moving away from the standard nickel catalyst chemistry which
has been the cornerstone of this technology since its inception is simply to
find a way to increase the GSA of the catalyst.
This can be done in a number of ways, but clearly there are other aspect
of catalyst design that need to be taken into account. The catalyst must be
physically strong, and must retain its strength during service; it must be
able to be loaded easily and uniformly into the narrow bore stem reformer
tubes; and the packed tube pressure drop should if possible be lower than
with the conventional rings.
These physical parameters will place constraints on the type of shape
designed; for example, shapes that are too open or fragile, or break easily
into many small fragments (thereby rapidly increasing the pressure drop)
will not be acceptable.
Similarly, shapes with external surface structure may be more difficult to
load, and their use should be carefully reviewed.
There is, however, one further parameter which until recently had been
very much overlooked, and that is the packed bed heat transfer coefficient
of the catalyst bed.

The heat transfer process from the steam reformer tube wall to the bulk
gas is limited by the heat transfer characteristics, particularly at the inside
tube wall gas film. It was noted that changes in the catalyst shape could
have a beneficial impact on the heat transfer characteristics. Furthermore,
it was recognized that the impact of relatively small changes here could
lead to significant changes in TWT - much more so than the equivalent
scale of change to the GSA would.
With that in mind, then, the optimum steam reformer shape is a complex
balance of these various parameters. A comparison of steam reformer
catalyst properties is given in the table below.
1 2 3 4
Relative
Pressure Drop
1 0.9 0.9 0.8
Relative HTC 1 1.3 1.1 1.0
Voidage 0.49 0.6 0.58 0.59
1 2 3 4
Heat transfer - Pressure drop
It can be seen that although all these shapes have better characteristic
than simple rings, some shapes are better than others. This will result in
significant differences in TWT for the same duty in a steam reformer;
lower TWTs will be seen with shapes than with rings, as shown below:

Top-fired Hydrogen Plant Reformer
70% Refinery Offgas Feed (70% C1-C5, 30% H2) + 30% LPG
Steam to carbon ratio = 4.5
Exit Temp./Pressure = 820°C(1508°F)/20 barg (290 psig)
1.
Ring
2.
4-Hole
3.
Multi-
hole
4.
Cross-
hatch
Pressure drop (bar) 1.7 1.5 1.5 1.4
Methane Slip (mol%,
dry)
3.4 2.9 3.0 3.2
Methane-Steam
Equilibrium
Approach (°C)
15 5 9 11
Max. Tube Wall
Temperature (°C)
885 860 876 880
It is possible to take advantage of this by increasing the through-put until
the TWT limit is again reached. For top-fired reformers, increases in
throughput of typically 10-15% have been seen; for side-fired reformers,
typically 5-8%.
In cases where this change has already been made, or further increases
in throughput are sought, re-tubing the steam reformer using modern
promoted metallurgies can result in further significant increases in
throughput.
An example is shown in the table below. At the Base Case (Case A), the
steam reformer operates at 100% throughput with a catalyst pressure drop
of 2.2 kg/cm² (31 psi), and a maximum tube wall temperature of 834°C
(1534°F), using old metallurgy. If the throughput were to be increased
without changing the tubes (Case B), then unacceptably high pressure
drops and tube wall temperatures would be seen.

However, a change to modern micro alloy tubes would allow the use of
thinner tubes; this gives more catalyst volume, and results in lower
pressure drop and tube wall temperatures (Case C). Here, an increase of
throughput to 120% would result in no increase in maximum tube wall
temperature compared with the Base Case (Case A), with only a slight
increase in catalyst pressure drop. As a further option, increasing the tube
outside diameter (OD) as well as the inside diameter (ID) would result in
even greater increases in plant throughput (Case D).
Case A Case B Case C Case D
Plant throughput (%) 100.00 120.00 120.00 120.00
Tube metallurgy IN519 IN519 HP50 HP50
Tube id mm (ins)
85.1
(3.4)
85.1
(3.4)
90.6
(3.6)
103.4
(4.1)
Tube od mm (ins)
109.8
(4.3)
109.8
(4.3)
109.8
(4.3)
122.6
(4.8)
Catalyst pressure drop (kg/cm²) 2.20 3.10 2.60 1.70
Catalyst pressure drop (psi) 31.00 44.00 37.00 24.00
Max. tube wall temperature (° C) 834.00 843.00 832.00 827.00
Max. tube wall temperature (° F) 1534.00
1550.0
0
1530.0
0
1521.00
4.2 Oxygen-blown Secondary Reformer
The concept of secondary reforming is well known in ammonia plants.
Here, a secondary reformer is installed immediately downstream of the
steam reformer (now often referred to as the "primary" reformer). This is a
large vessel, filled with catalyst. The gas exit the "primary" reformer is
further steam reformed, using air combustion to achieve very high
temperatures. This achieves two things on the ammonia plant: the
nitrogen required for the ammonia synthesis is introduced, and the use
of very high temperatures lowers the methane level exit the secondary
reformer to around 0.1% (dry).
A similar concept can be used on an existing hydrogen plant, though
clearly the combustion medium now must be oxygen rather than air. This
use of a secondary reformer allows some of the steam reforming load to
be moved from the "primary" to the secondary reformer.

Also, since very low levels of exit methane are achieved in the secondary,
some of the operational requirements in the "primary" (such as steam
reformer exit temperature and steam:carbon ratio) can be relaxed. This
allows the use of more feed gas without the need for extra firing in the
steam reforming section.
The plant configuration is shown in Figure 4.
Figure 4. Incorporation of an Oxygen-blown Secondary Reformer
Hydrocarbon
Feed
Steam
Shift
Conversion
Secondary
Reformer
Primary
Reformer
Sulphur
Removal
Oxygen
Steam
The secondary reformer is generally a refractory-lined carbon steel vessel
housing an oxygen burner assembly in the top part of the vessel. Recent
advances in the use of Computational Fluid Dynamics (CFD) for
secondary reforming have led to improvements in the design of vessels
and burners use with oxygen.
Note that the use of an oxygen-blown secondary will only be economically
attractive if there is a source of competitively-priced oxygen on the site.

4.3 Pre-reforming
Pre-reforming is the term that has been applied to the low temperature
adiabatic steam reforming of hydrocarbons. This process was originally
developed by British Gas in the 1960s for use in the production of Towns
Gas and SNG.
Although over 100 plants were built and operated for several decades in
many parts of the world, most of these have now been shut down as a
result of the widespread availability of natural gas. There are still a number
of plants operating in Japan, Hong Kong, China, Brazil and other regions
where there is limited availability of natural gas.
In the pre-reformer, all higher hydrocarbons are converted to give an
equilibrium mixture of steam, hydrogen, carbon oxides and methane. The
endothermic steam reforming reaction is followed by the exothermic water
gas shift and methanation reactions.
For a natural gas feed, the overall reaction is still endothermic; for butane
feeds, however, the overall reaction is almost thermo- neutral, and for
naphtha's, the overall reaction is exothermic. The process depends on the
use of a high nickel, highly active catalyst, produced by HAISO
TECHNOLOGY for GBH Enterprises, Ltd.
With light feedstocks such as natural gas or refinery offgas, the retrofitting
of a pre-reformer into an existing plant can offer some potential for
increased throughput. A typical plant configuration is shown in Figure 5.

Figure 5. Incorporation of a Pre-reformer
Steam
Pre-reformer
Reformed Gas
Desulphurised Feed
Steam
Steam Reformer
The steam reformer feedstock is purified and pre-heated, and mixed with steam.
Instead of now going to the steam reformer, however, this stream is first passed
through the pre-reformer. Here, due to the endothermic nature of the reactions,
the exit temperature is now lower than the inlet temperature.
This gas stream can now be re-heated to the original inlet temperature, using
low-grade flue gas duct heat rather than expensive fired heat in the steam
reformer. This typically will result in an increase in throughput of around 7-10%.
Note that there will be a reduction in the amount of steam generated now, since
heat from the flue gas has been used to do more steam reforming rather than
steam raising.
With heavier feedstocks, the overall reaction is thermo-neutral or exothermic, and
now the exit stream from the pre-reformer no longer contains heavy
hydrocarbons. This means that the steam reformer inlet temperature can be
raised significantly (assuming that the inlet metallurgy allows) which reduces the
steam reformer heat load, offering increases in plant capacity of the order of 25-
50%.

4.4 Post-reforming
In 1988, ICI commissioned two LCA (Leading Concept Ammonia) plants at
Severnside in England. These plants utilized a heat-exchange design of
steam reformer called the GHR (Gas Heated Reformer). This concept can
also be used to advantage on hydrogen plants as a retrofit, as shown in
Figure 6.
Figure 6. Incorporation of a GHR
Reformed
Gas
Steam
Steam Reformer GHR
Desulphurised
Feed
Steam
Pre-heat Coil
Bypass
Bypass
Here, purified and pre-heated feed gas is split into two streams: most goes into
the conventional steam reformer; the rest goes to the GHR, and uses heat from
the steam reformer exit gas to drive the steam reforming reaction. The gas
streams are then combined for purification. In this way, an extra 20% throughput
can typically be achieved without any increase in the conventional steam
reformer heat load, and without the need for extra fuel firing, with therefore no
increase in burner NOx levels.
Again, though, there will be a significant impact on the steam raising capability of
the plant.

5 Downstream Units
In modern hydrogen plants, the typical flowsheet has a single shift vessel with
high temperature shift (HTS) catalyst. The purpose of the shift stage of the plant
is to convert carbon monoxide to carbon dioxide, which is easier to remove
during the purification stage; also, some further hydrogen is generated. The shift
reaction is exothermic, and is catalyzed by an iron oxide catalyst.
More recently, such catalysts have been doped with various promoters such as
copper; these relatively new copper doped HTS catalysts offer significantly
higher activity than older iron-based un-promoted catalysts. However, as was the
case with steam reforming catalysts, there are a number of other variables to
consider when looking at HTS catalyst design. One of these is strength in
service; it is well-known that although all HTS catalysts are similar in terms of
fresh (unused) strength, there are significant differences with the reduced
strength, which is much lower than the fresh strength.
Often, the reason for a catalyst change-out can be due to pressure drop
problems caused by too low a reduced strength, or caused by plant upsets and
water-soaking of the HTS catalyst. The latest generation of HTS catalyst, has
not only higher activity, but also higher strength.
This can be seen in Figure 7, which shows the improvements in strength from
first generation catalyst, to that of first generation but with a modified
manufacturing route, to that of copper promoted HTS, and finally, to the latest
generation catalysts which have advanced pore structures.

Figure 7. HTS Catalyst Strength Comparisons
All crush strengths have been corrected for pellet size
Fresh Reduced Discharged
0
10
20
30
40
50
60
Mean Horizontal Crush Strength (kg/cm2)
Advanced pore structure
Promoted
Modified Manufacturing
First Generation
The advanced pore structure has also led to benefits in terms of improved
activity. This means that the volume of HTS catalyst required for a given duty has
steadily decreased with each catalyst development, and the latest generation,
advanced pore structure catalysts can perform an equivalent HTS duty with less
than half of the catalyst volume required by first generation HTS catalyst, as
shown in Figure 8.

Figure 8. Comparison of HTS Volumes
Typical 70 mmscfd Hydrogen Plant
A. First Generation B. Improved Manufacturing Process
C. Cu Promotion D. Advanced Pore Structure
A B C D
0
20
40
60
Catalyst Volume (m3)
In larger plants, the addition of a Low Temperature Shift (LTS) vessel may be
advantageous. This approach is well-established in the ammonia industry, but
less used in modern hydrogen plants. At these lower temperatures, iron-based
catalysts are not suitable; instead, higher activity copper/zinc catalysts are used.
The addition of an LTS vessel to the flowsheet of a hydrogen plant increases
throughput by around 5%. This is typically only attractive in terms of the
cost/benefit analysis for plants larger than say 30-35 MMscfd.
Another option is to replace the HTS catalyst with a Medium Temperature Shift
(MTS) catalyst; this is also a copper/zinc catalyst, but one that has been specially
formulated to be able to run at higher temperatures than the conventional LTS
catalysts.
This will be a cheaper option than installing a separate LTS vessel, but only
offers 2-3% increase in throughput.

For the PSA unit, there are a number of options available here for increasing the
capacity. One simple option is to review the required hydrogen purity level. If, for
example, 50 ppm carbon oxides were permissible rather than 10 ppm, then there
could be a significant increase in throughput with an existing PSA system. Other
options are:
1. Use of latest generation adsorbents. Early adsorbents are less
effective than current ones, and depending on the type used,
increases of up to 25% may be possible by changing adsorbents.
2. Improved control systems (from pneumatic to electronic) allow the
cycle time to be reduced, resulting in around 25% increased
throughput.
3. Modified cycles can also give up to an extra 20% throughput, but
this only applies for larger plants with more than 4 beds, using
electronic control systems.
4. Adding adsorbers from say 10 vessels to 12 vessels can
significantly increase the plant capacity, in some cases by as much
as 100%.
5. Installing larger vessels is not really a realistic option, since in
effect, the entire PSA unit is being replaced, and a large amount of
civil and piping work will be needed.

6 Summary of Uprating Options
The table below summarizes the various uprating options that have been
discussed.
The typically achieved increase in throughput is listed, together with an
approximate capital cost of installation.
This is based on North American economics for a "typical" 35 MMscfd
(39,000Nm³/h) hydrogen plant. The cost figure is the fraction of the cost
compared with the cost of building a new hydrogen plant (cost 1.0).
Note that costs will be additive if more than one modification is needed; note also
that costs do not include debottle-necking of mechanical or hydraulic limitations.
Retrofit Option
Increase in Plant
Rate
(% of plant rate)
Approx. Capital Cost
(relative to new
plant)
Comments
Replace all
catalysts
15-20 0.15-0.20 Note 1
Re-tube steam
reformer
20.00 0.10-0.20
Cost includes
catalyst
Oxygen-blown
secondary
25-50 0.20-0.40 Note 2
Pre-reformer
(natural gas)
7-10 0.10-0.15
Pre-reformer
(naphtha)
25.00 0.15-0.20 Note 3
Post-reforming 20.00 0.40
Replace HTS
with LTS
2-3 Virtually nil Note 1
Add LTS vessel 5.00 0.05-0.10
PSA
modifications
20-100 Note 4

Notes
1. The catalyst costs are negligible if phased with a normal catalyst
turnaround.
2. Operating costs strongly dependent on cost of oxygen. For a major
increase in throughput (50%), the total capital will undoubtedly be higher
because of other limitations.
3. Cost does not include any changes to steam reformer inlet metallurgy,
which would be significant.
4. These options should be discussed with PSA vendors to get comparative
costings.
7 Conclusions
A wide range of options exists for increasing the amount of hydrogen produced
on an operating unit. The first step should be a thorough and accurate
assessment of operating conditions, particularly in the area of steam reforming,
where there may be scope to effect significant increases in throughput by
changing the operating conditions.
This should be followed by a detailed analysis of the uprating options for the
specific plant. Generally, the older the plant, the more scope there is for
significant increases in output; modern plant designs using the latest catalysts,
adsorbents and tube metallurgies have very little "slack" in the design. In these
cases, pre- and post-reforming may offer the most attractive solutions.

Getting the Most Out of Your Refinery Hydrogen Plant

Recommended

Recommended

More Related Content

What's hot

What's hot (20)

Viewers also liked

Viewers also liked (20)

Similar to Getting the Most Out of Your Refinery Hydrogen Plant

Similar to Getting the Most Out of Your Refinery Hydrogen Plant (20)

More from Gerard B. Hawkins

More from Gerard B. Hawkins (20)

Recently uploaded

Recently uploaded (20)

Getting the Most Out of Your Refinery Hydrogen Plant