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

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GBH Enterprises, Ltd.

Getting the Most Out of Your Refinery Hydrogen

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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

PSAUnit

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 - %

3.43.6

3.84

4.24.4

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 4Relative Pressure Drop

1 0.9 0.9 0.8

Relative HTC 1 1.3 1.1 1.0Voidage 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% LPGSteam to carbon ratio = 4.5Exit 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 1.4Methane 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 1550.00

1530.00 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

ShiftConversion

SecondaryReformer

PrimaryReformer

SulphurRemoval

OxygenSteam

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

ReformedGas

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 sizeFresh Reduced Discharged

0

10

20

30

40

50

60Mean Horizontal Crush Strength (kg/cm2)

Advanced pore structurePromotedModified ManufacturingFirst 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 VolumesTypical 70 mmscfd Hydrogen Plant

A. First Generation B. Improved Manufacturing ProcessC. Cu Promotion D. Advanced Pore Structure

A B C D0

20

40

60Catalyst 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.

