steam ref based h2 plant ion

Steam Reformer Based Hydrogen Plant Optimisation

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by

Swen Fritsch

Krupp Uhde GmbH

Dortmund, Germany

Prepared for Presentation at

The International Conference

HYFORUM 2000

Munich, Germany

11 – 15 September 2000

Copyright © Krupp Uhde GmbH

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1. Introduction

In the last decade the refinery demand for hydrogen increased rapidly. Feedstock crudeoil has been getting heavier and more sour. More stringent environmental regulationsforce the refinery industry to reduce the content of sulphur, olefins and aromatics in thetransportation fuels. The combined effect of these trends is that refineries have becomesignificantly short of hydrogen.

2. Process Selection of Hydrogen Production

To produce the necessary hydrogen different solutions are possible (Figure 1).

Off gas fromaromatic production

from LPG feedstock

Hydrogen

Off gas fromcatalytic reforming

Steam reforming(SMR) from

natural gas upto naphtha

Thermal partialoxidation (POX) from

natural gas upto coal

Off gas fromdehydrogenation

(Ethylbenzene, Propane, Butane)

Off gas fromsteam cracker

Off gas fromcoker

Off gas fromFCC

Figure 1: Hydrogen Sources

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World-wide refinery hydrogen is mostly supplied as a by-product from the catalyticreforming unit which converts heavy straight-run naphtha into a high-octane (aromatic)gasoline blend component. Other refinery process units including fluidised catalyticcracking (FCC) units and thermal processes, such as coking, produce less hydrogenand the purity of the hydrogen of these units is low. Petrochemical processes such assteam cracking for olefins production, dehydrogenation of ethylbenzene, propane orbutane (e.g. to make isobutane for MTBE), and the production of aromatics fromliquefied petroleum gas (LPG) are excellent sources of refinery hydrogen.

The major hydrogen production technologies used in refineries are Steam-Methane-Reforming (SMR) and Thermal Partial Oxidation (POX).

SMR

This technology essentially consists of a fired heater containing catalyst-filled tubes(Figure 2). Steam is added to the feedstock, which can vary from refinery off-gasesover natural gas up to naphtha. The mixture is then heated to about 600 °C and passedthrough catalyst tubes. The catalyst converts the gas mixture to hydrogen, carbonmonoxide and carbon dioxide.The reaction is endothermic with heat supplied by burners located adjacent to the tubes.The reaction mixture exits the reformer at about 900 °C. Combustion gas heat isrecovered in the convection section.

FEED + STEAM600 °C

Thermal PartialOxidation

( POX )FEED

120 °C

O2 (99 %)

80 °C

1350 °CREF. GAS

Steam Reforming( SMR )

900 °CREF. GAS

Figure 2: Major Hydrogen Production Technologies

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POX

The partial oxidation reactor is a refractory lined vessel containing a burner, but nocatalyst. The feed is directly mixed with oxygen from a burner located near the top of thevessel. Partial oxidation and reforming reactions occur in a combustion zone below theburner. Gas exits at about 1,350 °C.

In contrary to SMR, POX requires oxygen as an additional expensive feedstock.

POX can be the choice if heavy feedstocks have no value and the gasifying solves adisposal problem, but the majority of the hydrogen plants world-wide is based on SteamMethane Reforming.

These hydrogen plants are built with capacities up to 130000 Nm³/h (12 t/h) per train.Even higher capacities in a single train up to capacities of 250000 Nm³/h hydrogen aretechnically possible.

A hydrogen plant has to operate as an utility plant with a reliability of as high as 99.8 %.If a hydrogen plant in a refinery is not available, mostly the downstream plants like ahydrocracker or a hydrotreater has to be shut down.

3. Process Concept

3.1 Process Units

The basic process steps for a hydrogen plant is: (Figure 3)

• Feed Gas Desulphurisation

• Optional Pre-Reforming

• Steam Reforming and Waste Heat Recovery

• CO-Conversion

• Synthesis Gas Purification (PSA)

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Desulphurization

Steam Reforming

CO Conversion

PSA Unit

Hydrogen

Feed

Fuel

Hydrogen

Offgas

Steam Export

Pre-Reforming

Figure 3: Block diagram hydrogen plant

3.2 Process Description

The hydrogen plant process diagram is shown in Figure 4. The feedstock varies fromhydrogen containing refinery off-gases over natural gas up to naphtha. Before enteringthe reformer, the feedstocks have to be desulphurized. Therefore they are preheated inthe reformer convection section to desulphurization temperature of about 370 °C. Thenatural gas contains organic sulphur, therefore hydrogenation using a Co/Mo catalyst isnecessary to convert the organic sulphur (mercaptans, thiophenes, etc.) to H2S beforedesulphurization.

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Desulphu-rizationDesulphu-

rization

Combustion air

BFW

Fuel

1 2 3 4 5

Convection bank coils1) Steam generator2) Steam superheater3) Feed / steam preheater4) Feed preheater5) Condensate partial evaporator6) Combustion air preheater

Convection bank coils1) Steam generator2) Steam superheater3) Feed / steam preheater4) Feed preheater5) Condensate partial evaporator6) Combustion air preheater

6

FeedHydrogen

Steam export

HT-shiftHT-shift PSAPSA

Process gascooler

Process gascooler

Steam reformerSteam reformer

C.W.

Figure 4: Flow Scheme Hydrogen Plant

Desulphurization of the feedstocks carried out with a zinc oxide bed is needed sincesulphur can poison the reformer catalyst. After purification the feedstock is mixed withprocess steam to a steam/carbon ratio of 2.5 to 3.5. The feed/steam mixture is furtherpreheated by reformer flue gas up to 650 °C before entering the reformer. The primaryreformer contains tubes filled with nickel reforming catalyst. The catalyst converts feedgas at temperatures up to 950 °C and pressures between 20 and 35 bar to anequilibrium mixture of hydrogen, methane and carbon oxides.

Most of the carbon monoxide in the reformer product is converted to carbon dioxide andadditional hydrogen in the HT-shift reactor. The HT-shift catalyst is based on iron oxide.

After cooling and separation of the process condensate, pure hydrogen of 99.9 % orhigher purity is recovered by a pressure swing adsorption (PSA) unit. The absorbentbeing a mixture of activated carbon and zeolithes removes all of the contaminants fromthe hydrogen product in a single step.

The PSA off-gas is utilised to provide the vast majority of the reformer fuel which isbalanced by additional fuel in the order of approx. 10 % of total heat release in thereformer.

Waste heat is available from two sources of the steam reformer: the heat in the reformerflue gases and the heat in the process gases coming out of the reformer.

This heat is used to generate steam, superheat steam, preheat feed/steam mixture andfeed, partial evaporation of process condensate and preheat air.

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4. Design Features

The following design features with respect to the above mentioned process units areremarkable in view of plant reliability, investment cost savings and lower operatingcosts.

4.1 Reforming

The Krupp Uhde top fired reformer including the patented Krupp Uhde cold headersystem is the heart of the hydrogen plant.

4.1.1 Reformer Box

Krupp Uhde's reformer design is a vertical tube down-fired box type furnace which hasbeen developed in the sixties and has been continuously improved during the last 30years.

The typical reforming section is shown in Figure 5. The tubes are arranged in rows withrows of down-firing burners in between. The tubes are penetrating ceiling and bottom ofthe radiant cell.

The reformer tubes are connected to the manifolds as follows:

• At top:Via flexible connections (pig tails) having a moderate operating temperature.

• At bottom:Directly welded to the nozzles of the outlet manifold system.

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Inlet manifold

Burner

Reformertubes

Catalyst grid

Outletmanifoldsystem Refractory

lining

Flue gas

Figure 5: Primary Reformer

The outlet manifold system is placed on foundations. The thermal expansion of thetubes is directed upwards and is compensated by flexible connections at the inletheader.

The furnace is equipped with a forced draft fan and an induced draft fan. The flue gas isdrawn off through tunnels arranged on the furnace bottom one between each two tuberows.

The Krupp Uhde reformer is a modular design type. A module consists of one row oftubes and the associated panels, structural elements, burner lane, manifolds and tunnel.Thus, the difference in furnace size first of all is dictated by the different number ofmodules.

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4.1.2 Reduction of Reformer Radiant Size

Krupp Uhde is continuously improving the material selection of the reformer tubes. Formany years HP modified has been the alloy commonly used to fabricate reformer tubes.

The material micro-alloy permits the increase of reformer tube diameter and the averageheat flux.

Compared with a hydrogen plant built in 1990, Krupp Uhde can nowadays design asignificantly smaller reformer box. This is achieved by using the micro-alloy tubematerial which allows to increase the tube inner diameter from 110 mm to 127 mm andthe average heat flux from approx. 70 kW/m² to approx. 80 kW/m². This results in areformer tube number reduction of about 25 %.

A further reduction of the reformer box size is planned in the future.

4.2 Outlet Manifold System

The Krupp Uhde Outlet Manifold System is a patented, unique design that has proven tobe extremely reliable during many years of operation.

There is no reformer design which can avoid to expose the reformer tubes to very hightemperatures under full pressure. They are working under creep condition and in spite ofall improvements of alloys, testing and manufacturing the tubes remain a component oflimited lifetime which is normally expected to be more than 10 years.

Many designs in the market connect the end of the reformer tube to the outlet manifoldvia high alloy pipe elements. There is also a design which has the manifolds locatedeven in the heated zone of the furnace.

The pipe connections at the tube outlets and the high alloyed outlet manifolds have tostand the maximum gas temperature at full gas pressure and are - in addition - designedto compensate thermal expansions of manifolds and tubes. They are subjected tointernal corrosion, creep, cycling and external forces.

These pig tails, often made from cast or forged Alloy 800-type material, are as much apotential source of failure as the reformer tubes. That means there are twice as muchcritical elements in such a design than necessary and twice as much sources of failure.

The Krupp Uhde design eliminates 50 % of the critical elements. It reduces their numberto the unavoidable minimum, thus to the number of reformer tubes.

Figure 6 shows the design. There is a removable catalyst grid at the bottom of thereformer tube. A thin Alloy 800 funnel and tube are welded gastight to the grid support.This tube conducts the reformer gas down into the refractory lined manifold. The tubecan expand free downwards. It is not exposed to any pressure or external stress. Theannular space between the gas conducting tube and the pressure bearing shell is filledwith ring shaped insulating bricks in order to isolate the pressure retaining outer tubefrom high temperature.

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300 600 900 1000°CT

Feed/steam

Furnace bottom

Furnace archReformer tubeFire box

Catalyst grid

Gas conducting tubeBellow

Skin temperature profileCarbon steelShop weld

Field weldOutlet manifold

Figure 6: Reformer tube to manifold connection with skin temperature profile

By separating high temperature service from pressure bearing service this tube tomanifold connection got the same reliability as any other ordinary pipe in the plant. Theskin temperature profile along the pressure bearing part of the lower reformer tubeunderneath the furnace is shown on the right side of Figure 6. This temperaturedecreases very fast down to a temperature suitable for carbon steel. The transition weldbetween the cast alloy and the carbon steel end is a shop weld. The assembly weld ofthe reformer tube to the nozzle of the manifold is a simple carbon weld being performedat site.

The penetration of the reformer tube through the furnace bottom is completely sealedgas-tight by a flexible bellow. It prevents any ingress of leak into the heated zone of thefurnace and allows for the small thermal expansion of the manifold. The operating skintemperature of the manifold is as low as approx. 200 °C.

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4.3 Convection Section

Reformer furnaces are relatively inefficient, only about half of the heat released in theradiant section is adsorbed in the reaction tubes. Thus, a large amount of waste heat iscarried by the hot flue gases (approx. 1,050 °C) into the convection section of thefurnace where most of the heat is recovered by heat exchange from various users:

• Generation of HP steam

• Superheating of HP steam

• Preheating feed/steam mixture

• Preheating and evaporation of process condensate

• Preheating of feed

• Preheating of combustion air

The flue gas stack temperature is approx. 150 °C. The overall efficiency for reformer boxand convection section (excluding air heating) is about 93 %.

4.4 CO-Conversion

In the CO-shift conversion the CO-content of the reformed gas is reduced by theconversion of carbon monoxide and steam into carbon dioxide and hydrogen via anexothermic catalytic reaction.

Different CO-conversion options are possible: (Figure 7)

• HT-conversion (adiabatic)

• HT-conversion (adiabatic) and LT-conversion (adiabatic)

• Isothermal MT-conversion

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Steam

Reformed gasfrom PGC

To PSA

C.W.

Reformed gasfrom PGC

Steam

To PSA

C.W.

HT-shift LT-shift MT-shiftaddition

Figure 7: CO shift options

Adding a LT-shift (adiabatic reactor, inlet temperature 200 °C) to a HT-shift reactor(adiabatic reactor, inlet temperature 340 °C) would have the following results:

• Reduction of feed by 4 %

• Increase in fuel by 19 %

• Marginal increase in feed and fuel by 0.6 %

• Increase in steam export by 3 %

• Reduction of absorbed duty (i.e. reformer size) by 4 %

Substituting the HT and LT shift by an isothermal MT shift with a typical exit temperatureof 260 °C would produce a similar result as well.

It is obvious that only a relatively expensive feed and a cheaper fuel would justify theadditional investment of supplementing the HT shift by a LT shift or substituting them byan isothermal MT shift.

4.5 Hydrogen Purification

For the hydrogen purification a pressure swing adsorption (PSA) system meets highpurity requirements. The efficiency of the hydrogen recovery varies between 86 and92 % and the number of adsorbers vary depending on the capacity between 4 and 14vessels.

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5. Plant Optimisation

A steam reformer ( SMR ) based hydrogen plant usually produces as a by-productexport steam. The plant can be designed for different steam to hydrogen product ratios(H2O/H2) which vary from 0 ( no steam export ) up to 20 tons of steam export per ton ofhydrogen.

To reach the required H2O/H2 ratio different flow sheets and process parameters arepossible, which have varying consequences on feed and fuel consumption and on theinvestment cost for the hydrogen plant. For a fixed hydrogen product stream theselection of the economic steam export stream depends mainly on the maximumrequired steam export as well as on the cost for feed and fuel, on the credit for theexport steam and on the investment cost.

Figure 8 shows two hydrogen production efficiencies of hydrogen versus the specificsteam export. One curve is without giving credit for steam export, the second is withgiving credit for steam export. The hydrogen production efficiency without taking thesteam production into account is defined as the energy content of hydrogen divided bythe required feed plus fuel energy input. Both energies are based on the lower heatingvalue. The hydrogen production efficiency has the highest value of 79 % with no steamexport. The value decreases with increasing specific steam export to a value of 61 %.

t Steam / t Hydrogen

eta

[%]

5 10 15 2060

70

80

90

0

Hydrogen production efficiency with steam production

Hydrogen production efficiency with steam production

Hydrogen production efficiency without steam production

Hydrogen production efficiency without steam production

Figure 8: Hydrogen Production Efficiency versus Specific Steam Export

The second curve shows the hydrogen production efficiency with taking the steamproduction into account. This efficiency value is defined as the energy content of thehydrogen and the produced steam divided by the required feed plus fuel energy input.With this definition the hydrogen production efficiency has the lowest value of 79 % withno steam export. The value increases with increasing specific steam export up to valueof 89 %.

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A hydrogen plant has two kinds of energy losses. One is the flue gas from the reformerto the atmosphere at temperatures of approx. 130 – 160 °C and the second is the heatof the process gas from temperatures of about 120 ° C down to 40 °C, which is removedby air cooler and/or water cooler.

Existing hydrogen plants have a slightly lower energy efficiency of approx. 3 – 4 % asshown in the diagram. The hydrogen production efficiency could be increased due tomaterial and hardware developments as well as developments of the used catalysts.

Figure 9 illustrates the steam production efficiency versus the specific steamproduction. The steam production efficiency is defined as the produced steam exportenergy divided by the additional feed plus fuel energy input compared to the base casewith no steam export.


eta

[%]

5 10 15 20

200

0

400

600

800

1000

1200

1400

1600

1800

Steam production efficiencySteam production efficiency

Figure 9: Steam Production Efficiency versus Specific Steam Export

The steam production efficiency drops from almost infinity at low steam export values toapprox. 200 % at a specific steam export value of 8 ton per ton of hydrogen anddecreases further to a value of 130 % at high steam export values.

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This diagram explains clearly that steam export is a by-product of a hydrogen plantwhich can be produced with low additional energy input. The reason for the high steamproduction efficiencies is the huge amount of low temperature heat available in theprocess gas cooling section which can be used economically to heat boiler feed waterfrom ambient temperatures up to the boiling point. All the remaining low temperatureheat in the process gas cooling section which can’t be used for steam production of feedpreheating is waste heat which has to be removed by air or water coolers.


1 / t

Hyd

roge

n

2 10 12 20

110

0

120

130

140

150

160

170

180

4 6 8 14 16 18

Feed + fuel 100,steam 50






Figure 10: Specific Operating Cost versus Specific Steam Export

Figure 10 illustrates the specific operating cost against the specific steam production forthree different steam credit values. In the first curve the steam has a 30 % higher valuethan feed plus fuel, in the second curve the steam has a 10 % lower value than the feedplus fuel and in the third curve the steam has only half the value of the feed plus fuel.Starting with a specific operating cost of 151 in all three curves the specific operatingdecreases first with increasing steam export. Is the value of the steam higher than 90 %of feed and fuel cost, the operating cost decreases all the more so, as the steam exportincreases. With a steam value of 90 % of feed and fuel cost the operating costdecreases until a specific steam export of approx. 6 t steam per ton of hydrogen. Withhigher amounts of export steam the operating cost are unchanged.

Is the value of steam lower than 90 % of feed plus fuel cost, the operating costdecreases slightly from no steam export to a minimum located at approx. 5 t steam perton of hydrogen. Increasing further the steam export the operating cost increases aswell.

Based on a ten years net present value, the operating cost of a medium sized hydrogenplant (50000 Nm³/h H2) located in Europe with a natural gas price of 3 US$ per MMBTU(6.50 DM/GJ) are approx. three times higher than the capital cost. This comparisondemonstrates the importance of the operating cost and its impact on the long termeconomic of the plant.

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6. Conclusion

It is a necessity for the design of a hydrogen plant to consider beside the requiredhydrogen capacity the feed plus fuel cost as well as the steam credit and the maximumamount of steam which can be utilised at site.

Even if no additional steam export is required, the efficiency of the steam production ofthe steam reformer based hydrogen plant is so high, that the hydrogen plant cansubstitute the operation of existing package boilers for steam production.

Only if all that is recognised a tailor made plant with optimum consumption figures canbe designed.

steam ref based h2 plant ion

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