Hydrogen Production by Steam Reforming of Hydrocarbos

Niels R. Udengaard Haldor Topsoe Inc.

17629 El Camino Real, Suite 300

Houston, Texas 77058

Introduction

The use of hydrogen for petrochemicals, fertilizers and as energy carrier in connection with renewable energy production will increase substantially in the next 5-10 years as even more stringent environmental legislation is enforced. Low sulfur gasoline and diesel fuels will become mandatory and harmful emissions will be reduced drastically. Hydrogen will be required by refiners and specialty chemical manufacturers to meet the global need for cleaner products. The growing fuel cell market will be dependant on hydrogen as a primary fuel source.

Hydrogen requirements vary widely. For small capacities below 0.1 MM SCFD, supply in cylinders or production by electrolysis may be preferable. Hydrogen production from methanol or ammonia cracking is suitable for small, constant or intermittent requirements as used in the food, electronics and pharmaceutical industries. For larger capacities hydrocarbons are used primarily as feedstock in the steam reforming process for production of hydrogen and synthesis gas.

The current production of hydrogen in North America is about 6000 MM SCFD, and the annual growth is expected to be above 4%. Most of this growth is for refineries for their production of ultra-low sulfur diesel and gasoline. In spite of efforts to produce hydrogen by processes involving solar energy, wind energy, nuclear energy and biofuels, fossil fuels remain the most feasible feedstock in the near term, and for commercial scale production of pure hydrogen, steam reforming remains the most economic and efficient technology for a wide range of hydrocarbon feedstocks.

Steam Reforming Technologies

A typical layout of a hydrogen plant based on steam reforming including the following steps:

Natural gas feed is preheated in coils in the waste heat section of the reformer, and sulfur is removed over a zinc oxide catalyst. Process steam is added, and the mixture of natural gas and steam is further preheated before entering the tubular reformer. Here, conversion to equilibrium of hydrocarbons to hydrogen, carbon monoxide and carbon dioxide takes place over a nickel based reforming catalyst.

The gas exits the reformer and is cooled by steam production before entering the shift converter, typically a medium temperature shift. Over the shift catalyst more hydrogen is produced by converting carbon monoxide and steam to carbon dioxide and hydrogen. The shifted gas is cooled further to ambient temperature before entering the PSA unit. High purity hydrogen product is obtained, and the off-gas from the PSA unit is used in the reformer as fuel supplemented with natural gas fuel.

Combustion air for the tubular reformer burners can be preheated in coils in the reformer waste heat section. Part of the steam produced in the hydrogen plant is used as process steam, the excess steam is exported.

In many situations when natural gas is not available, higher hydrocarbons become the preferred feedstock for the reforming process. Many refineries also can benefit from flexibility in feedstock, taking advantage of the surplus of various hydrocarbon streams in the refinery.

Fired tubular reforming is generally the most competitive technology for capacities of up to more than 100 MM SCFD hydrogen.

Topsoe’s latest development in steam reforming process technology is the advanced steam reforming process. The characteristics of this process are:

• High reformer outlet temperature

• Low steam to carbon ratio.

• High combustion air preheat (optional).

• Adiabatic prereforming (optional).

• High heat flux reformer.

Low steam to carbon ratios, typically 2.5, in hydrogen plants, reduce the mass flow through the plant and thus the size of equipment. The lowest investment is therefore generally obtained for plants designed for low steam to carbon ratios. However, a low steam to carbon ratio also increases the methane leakage from the reformer. This can be compensated for by increasing the reformer outlet temperature to typically 1690 ºF in hydrogen plants. Furthermore, operating at a low ratio requires the use of non-iron containing catalyst i.e. a copper-based medium temperature shift catalyst in order to eliminate production of by-products in the shift section.

The installation of an adiabatic prereformer upstream of a tubular reformer has been found to be very advantageous in naphtha based plants and plant operating on fuel gases with higher concentrations of higher hydrocarbons. Since all higher hydrocarbons are converted over the prereformer catalyst, the inlet temperature of the gas inlet in the reformer can be increased to 1200 ºF and the reformer can be designed for higher heat fluxes. This reduces the size of the tubular reformer, resulting in direct capital cost reduction.

High combustion air preheat temperatures results in reduced fuel consumption and reduced steam production. The combustion air temperature can be used to adjust the steam export to a desired level. Temperatures of up to 1020 ºF are industrially proven in a radiant wall reformer.

The Topsoe convection reformer (HTCR) is a new concept which combines the radiant and waste heat section of the conventional reformer in one relatively small piece of equipment, thereby optimizing heat transfer to the steam reforming reaction and avoiding the surplus energy (often seen as steam export) inherent in conventional plants.

HTCR based layout is an attractive option for smaller sized hydrogen plants.

Efficiency and Production Costs

With no steam export the theoretical energy consumption is 300 BTU/scf H2 on LHV (lower heating value) basis. The industrial value for natural gas based plants is about 320 BTU/scf H2 corresponding to 94% of the theoretical efficiency. At locations with high natural gas prices, the energy efficiency becomes critical. For a natural gas price of 4 USD/MM BTU, the feedstock and utility costs makes about 65% of total operating costs.

Catalysts and the Mechanism of Steam Reforming The steam reforming process may appear straightforward from an overall consideration as the product composition is determined by simple thermodynamics, but in reality it is a complex coupling of catalysis, heat transfer and mechanical design.

In recent years, there has been progress in steam reforming technology resulting in less costly and more efficient plants, in part because of better materials for reformer tubes, better control and understanding of carbon limits, and better catalysts and process concepts with high feedstock flexibility. This progress has been accompanied by a better understanding of the reaction mechanism.

The typical steam reforming catalyst contains nickel. The catalyst properties are dictated by the severe operating conditions in the reformer with high temperatures and steam partial pressures. Sintering is an important cause of deactivation of nickel-containing steam reforming catalysts. The most important parameters are the temperature and the atmosphere in contact with the catalyst. The catalyst support can affect the sintering in various ways by loss of surface area. The sintering ceases when the nickel particle size exceeds a given size. The maximum size increases with temperature.

The catalyst activity is rarely a limiting factor. The catalyst volume (space velocity) is fixed from the tubular reformer design. The equilibrium conversion at high reforming temperatures is achieved at even very high space velocities when extrapolating the intrinsic rates. In practice, however, the utilization of the activity (as expressed through the effectiveness factor) is smaller than 10% because of transport restrictions. It can be shown by computer simulations that the catalyst is not the limiting factor for the design of a tubular reformer. An increase of the heat flux and the load at a given exit temperature by a factor of two results in an increase in methane leakage by only 10%.

Recent studies of the fundamentals of the steam reforming reactions have lead to a more consistent understanding of the mechanism of the main reactions and the competing reactions for carbon formation. The dissociation of methane on nickel surfaces has been investigated extensively, and several details of the reaction pathway are known from fundamental studies, in-situ high resolution electron microscopy and theoretical calculations.

In-situ high resolution electron microscopy has provided new information on sintering mechanisms and for the importance of steps in nucleation of whisker carbon. Density Functional Theory calculations have quantified the energetics of methane activation and shown that activation barriers are smaller on surface steps where also carbon is the most stable surface species.

Conclusions

The demand for hydrogen is growing in many industries, particularly in the chemical and refining industries. Topsoe has been active in this field for more than 50 years and offers a widerange of advanced technologies for small and large capacities meeting the demand for low operating and investment costs.

For small plants, convection reforming in the form of compact, skid-mounted plants is available. This is a low investment alternative to conventional concepts using tubular reforming.

For medium and large-scale plants the predominant technology is steam methane reforming of hydrocarbons. Advanced steam reformer design allows high temperatures and low steam to carbon ratios for hydrogen production resulting in high energy efficiency and lower cost plants.

For very large hydrogen and syngas plants the oxygen reforming technology may be applied, such as autothermal reforming and oxygen fired secondary reforming. These technologies are especially attractive when cheap oxygen is available.

References

(1) Rostrup-Nielsen, J.R.; Sehested, J.; Norskov, J.K.; Hydrogen and Syngas by Steam Reforming, Academic Press, 2002.

(2) Dybkjaer, I.; Winter Madsen, S., Hydrocarbon Engineering, December/January 1997/98.

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Burner Operating Characteristics

Burners are critical for the successful operation of industrial furnaces. Presented here is a set of equations that can be used to calculate characteristics of burner operation, including flame length, flame diameter, ignitability and flameout conditions. Equations are based on pre-mix burners operating at atmospheric pressure and firing natural gas only. Premix burners create short and compact flames compared to raw gas burners, and are designed to function with fuel-gas mixtures that have consistent specific gravity and composition.

Burner requirements

For direct-fired heaters to function correctly, burners must be capable of providing sufficient heat liberation from the fuel to meet heater processing requirements — based on the lower heating value (LHV) of the fuel. A fuel’s LHV can be defined as the amount of heat produced by combusting a specified volume, and returning the combustion products to 150C. For the heater to operate at the design process flowrate, the burners need to provide the heat necessary to maintain process fluid temperature and meet vaporization requirements at the heating coil outlet.

• The number, size and placement of burners must allow each coil to operate at the same design outlet temperature
• Design tube-metal temperature cannot be exceeded at any point on the coils
• Burner size must allow an outlet velocity that does not result in malfunction over the range of flow conditions
• Burner flame length should be less than firebox height (for vertical cylindrical heaters) or less than firebox length (for end-wall-fired heaters)
• Excessive flame height and diameter should be avoided to prevent flame impingement on tubes
• Burner spacing should be sufficient to allow burner-to-burner, as well as burner-to-tube clearance

The following equations can help establish optimal burner diameter:

Burner clearance

Establishing burner-to-burner clearance and burner spacing should be based on maximum burner flame diameter. Further, burner flame diameter should be evaluated at maximum burner-flame length. Sufficient burner-to-burner, outside diameter clearance should take into account the placement of structural elements between burners.

Sufficient burner-to-burner clearance prevents interference between the flame bodies and unburned fuel cores generated by adjacent burners, which results in the absence of unburned fuel within the burner flame when maximum flame length is reached. Burner center-to-center spacing should be at least one fully combusted flame diameter.

Clearance between the burner-flame (at maximum diameter) and the outside diameter of tubular heating surfaces should be set such that burner-to-tube flame impingement is avoided. Doing so will prevent tube damage due to overheating and will make best use of heating surfaces.

Flameout

At high burner velocities, flame loss can occur if the heat gain due to burner ignition is less than the heat loss from the burner flame. Burner velocities may be pushed well above that used in normal heater operation in an effort to achieve higher heater capacity. Aside from flame loss while the heater is in operation, flameout can also be characterized by difficulty maintaining a stable flame at startup, or an inability to ignite the burner. The following equations can help predict the circumstances under which flamout conditions might occur:

Flame velocity

The heat generated by combustion is dependent on the flame propagation velocity. In a situation with 0% excess air, the ratio of fuel-to-fuel+air is about 0.1. In that case, evaluation of the flame propagation velocity is straightforward. However, at fuel-to-fuel+air ratios higher or lower than 0.1, it is more difficult. The following equations can help predict flame propagation velocity in those cases:

NOMENCLATURE

Q_{lib heater}= Heater liberation, Btu/h

N_b= Number of burners

D_b = Burner diameter, ft

V_b= Burner exit velocity, ft/s

C_fuel = Fuel, ft³

LHV = Lower heating value of fuel, Btu/lb

C_air+fuel = Volume of air and fuel mixture, ft³

SV_fuel = Specific volume of fuel, ft³/lb

D_{f max} = Maximum flame diameter, ft

L_f = Flame length, ft

SV_flame= Specific volume of flame, ft³/lb

V_f = Flame propagation velocity, ft/s

Q_gain = Burner heat gain, Btu/h

Q_loss = Burner heat loss, Btu/h

A_s = Flame front area, ft2

(HTC)_c (HTC)_f, (HTC)_r = Natural convective, forced- convective, and radiative heat transfer coefficients, respectively, Btu/h-ft²-F

T_flame = Flame temperature, R

T_surr = Surrounding temperature, R

E_g = Flame emissivity

C_p = Gas specific heat, Btu/lb-F

A = Frequency factor in the Arrhenius equation

H = Heat of activation, Btu/lb-mol R

R= Gas constant, 1.987 Btu/lb-mol R

T= Gas Temperature, R

dC_m/dt= Fuel concentration change, mol per ft³/s

K = Reaction velocity constant, s–1

W_f= Fuel, lb/h

*The text was adapted from the article “Fired-Heater Burner Performance,” by Alan Cross. It appeared in the April 2008 issue of Chemical Engineering.

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Hydrogen Production By Steam Reforming

Management of the gas is critical for petroleum refiners

Ray Elshout Energy, Systems Engineering

Steam reforming of natural gas at petroleum refining facilities is the predominant means of producing hydrogen in the chemical process industries (CPI). Areas where hydrogen is heavily consumed include ammonia production, the cryogenics industry and methanol production (Table 1). Because hydrogen needs within various sectors of the CPI are at their highest levels in history, and are continuing to grow, an understanding of this method of hydrogen production and purification can be useful.

A major percentage of hydrogen used in the CPI goes toward production of ammonia, which continues find greater demand in the chemical fertilizer industry. On the other hand, methanol usage is declining in connection with its use as a feedstock for making methyl tert-butyl ether (MTBE; by reaction of methanol with tertiary butylene). In the U.S., MTBE had been used as a gasoline blend stock until recently, when use of the chemical as a gasoline oxygenate was phased out in favor of ethanol.

In addition to being producers of hydrogen, largely through steam reforming, petroleum refineries are also large consumers of the gas. Consumption of hydrogen by petroleum refineries has increased recently due to clean-fuels programs, which require refiners to produce low-sulfur gasoline and ultralow-sulfur diesel fuel. Management of hydrogen is a critical concern for refiners because various processes require different hydrogen pressure levels and purity.

Hydrogen-using processes that require high pressures and high purity, including hydrocracking, use hydrogen above the 100 kg/cm² (1,500 psig) level. When a recycle gas system is used, the higher pressures are needed to maintain hydrogen partial pressure at the desired level as methane concentration in the hydrogen feed to a hydrocracker increases. Sufficient hydrogen partial pressure promotes the intended reactions without producing undesirable coke.

If the hydrogen partial pressure cannot be maintained, the recycle gas should be bled. With pressure swing adsorption (PSA) processes producing hydrogen of purity in the range of four-nines (99.99%), this is not a problem.

Other hydrogen users, like those engaging in milder hydrotreating, can use lower-purity hydrogen at lower pressures (600 psig or lower). One approach that makes sense is recovering hydrogen from the users requiring higher pressure and reusing it at the lower pressure levels.

Minimizing the hydrogen bled into the fuel gas can keep the hydrogen production levels manageable. However, the hydrogen plant feed usually includes some hydrogen that goes through for a “free ride,” except for the cost of heating it up to reformer temperature and ultimately cooling it back down to recovery level.

A recently employed practice in the industry is for the hydrogen to be produced for adjacent producers and sold to the user as “over-the-fence” hydrogen. This keeps the production costs off the books from the adjacent user and has found popularity not only in the U.S. but also in Europe.

Figure 1. Steam-methane reforming is still responsible for the bulk of hydrogen production in petroleum refineries

Steam-Methane Reforming

Refinery hydrogen comes primarily from two sources — catalytic reforming of byproduct gas from the dehydrogenation of naphthenes into aromatics and high-octane gasoline blend stocks, as well as from direct hydrogen manufacture. The bulk of direct hydrogen manufacturing in a petroleum refinery is still accomplished via either steam-methane reforming (Figure 1) or steam-naphtha reforming. Partial oxidation of heavier hydrocarbons is also used to a limited extent.

In the overall steam methane reforming (SMR) reaction, methane reacts with steam at high temperatures and moderate pressures in catalyst-filled tubes to generate synthesis gas, a mixture of hydrogen, carbon monoxide and some carbon dioxide.

The reactions for the two simultaneous SMR mechanisms are shown as Equations (1) and (2). Both are endothermic, as shown by the positive heat of reaction. The reaction requires heat transfer to maintain temperatures favorable to the equilibrium reactions.

As the molecular weight of the feedstock increases, such as when heavier hydrocarbons (such as ethane, propane or butane) are included in the feed, the reactions are shown by Equations (3) and (4), with the corresponding heat requirements [2].

Product gas from the steam reforming of the methane and naphtha contains equilibrium amounts of hydrogen, carbon dioxide, carbon monoxide and excess steam. The calculated effluent composition of a reformer always needs to be checked against the equilibrium constant equations to ensure that simulations agree with known values.

Excess steam above the theoretical requirements is maintained to prevent the reforming catalyst from coking. The temperature exiting the reformer furnace tubes is usually about 760^oC (1,400^oF), a level that provides maximum hydrogen production within the temperature limitation of the reformer tube metallurgy (discussed later).

Water-shift gas reactions

Additional hydrogen can be generated from the carbon monoxide byproduct following the reforming reaction. First, the reformer effluent gas is cooled in two steps to favor the equilibrium toward the right side of the reaction. The first cooling step is followed by the high-temperature shift reactor, and the second cooling step is followed by a low-temperature shift reactor. Shift reactions are promoted as effluent gas flows down through the fixed catalyst reactor containing a ferric oxide catalyst in accordance with the reaction in Equation (5). Note the water-shift reaction is exothermic, which results in a temperature increase across the reactors as water reacts with CO to form CO₂ and more H₂.

Water shift gas equilibrium is not affected by pressure, since there is no volume change. Reduced temperatures favor the conversion of CO to H₂, as might be expected by its exothermic nature. A variety of catalysts are available for the service.

Hydrogen Plant Process

Figure 1 shows a schematic of a conventional steam-reforming hydrogen plant [4]. The plant is based on a feed gas with high sulfur content, requiring plant operators to hydrotreat the feed before the zinc oxide removes the sulfur compounds. The H₂ purification at the end of the process is based on the removal of CO₂ with a pressure swing adsorber (PSA) system shown as the H₂ purification block. The reformer is shown as a vertical furnace type with side firing. The reformer furnace design alternatives will be discussed below.

Feed gas — usually a mixture of hydrogen, methane and other light hydrocarbons — is first compressed to about 300 psig. The initial compression has been found to provide product hydrogen at a pressure that can easily reach the desired hydro-processing pressure with a four- or five-stage reciprocating compressor. This equipment is not part of the hydrogen plant.

The feed gas is preheated with reformer effluent gas and hydrotreated to convert the various sulfur compounds (such as mercaptans, carbonyl sulfide and carbon disulfide) to hydrogen sulfide. The gas is then passed through desulfurization reactors, usually containing a zinc oxide catalyst, which adsorbs the hydrogen sulfide. Low-sulfur feeds may not require the hydrotreating step.

Reforming furnace

The sulfur-free gas is mixed with a fixed amount of superheated steam to maintain the desired steam-to-hydrocarbon ratio. The steam-to-hydrocarbon ratio is kept within a range that is high enough to prevent laydown of coke on the reforming catalyst, but low enough to avoid overloading the reformer duty. Typically for a methane feed, the ratio would be three, whereas the theoretical requirement is somewhat less.

The combination of hydrogen and steam is heated to about 760^oC (1,400^oF). Since all of the reforming reactions are endothermic, additional heat is required to maintain the reaction temperature as the mixture flows down through catalyst-filled reformer tubes.

A critical factor in the reformer heater design is keeping the tube-wall temperature uniform and hot enough to promote the reforming reaction. Two types of heater designs have been employed for this purpose. Figures 2 and 3 show schematic diagrams of the side-firing reforming furnace, and the roof-fired heater design approach is shown in Figures 2 and 4.

Figure 2. Maintaining a tube-wall temperature that is hot enough for the reforming reaction is a critical factor in reformer heater design

Figure 3. A typical reformer furnace could have over 300 burners

Figure 4. Hydrogen plants with single heaters and capacities up to 100,000 ft3/d have used a down-firing approach

Side-fired reforming heaters. The coil arrangement in a typical side-fired reformer furnace (Figure 3) consists of two parallel rectangular fire boxes connected at the top with horizontal duct work into the vertical convection stack. Two rows of vertical tubes arranged on a staggered pitch are present in each of the radiant boxes. Several (typically four) rows of burners are used to fire each side of the two radiant sections. This arrangement allows direct radiant fire to reach most of the tube wall. Platforms are provided to access the burners at each of the four burner levels. A typical reformer furnace could have over 300 burners. Reformer tubes typically have diameters of 5 in. (127 mm), walls 0.5-in. (13 mm) thick and about 34 ft (11.5 m) of wall exposed to the burners. The tube metallurgy is usually 25% chrome, 20% nickel or a high-nickel steel such as HL-40.

The inlet manifold at the top of the heater has “pigtails,” which uniformly transfer the feed gas to the top of the tubes. Another manifold at the bottom of the heater connects another set of pigtails to the outlet transfer line. The pigtails provide for thermal expansion as the heater goes from startup temperature to reaction temperature. The objective is to have an equal pressure drop across each tube, which produces uniform flow to each of the tubes. The convection section includes several different coils. The hottest coil is a steam generation coil that protects the other coils from radiant heat. Usually, there is also a steam superheat coil, a feed preheat coil and another steam generation coil. Above these coils, there may be a boiler feed water (BFW) pre-heater and deaerator preheat coil.

Typically an induced draft fan is used to keep the fire box pressure slightly negative. Some reformers also have an air pre-heater and a forced draft fan.

Top-fired reformer. This type of reformer heater is usually a rectangular box. The tubes are still vertical, and inlet and outlet pigtails are used to connect the inlet header and the outlet transfer line, respectively. Figure 4 shows a schematic diagram of a down-fired reformer furnace [9].

The tubes are spaced on a pitch, which allows the burners to fire down between the tubes. The burners have a special “pencil-shaped flame” design. All burners are located in the penthouse above the inlet manifold. The flame and the flow through the tubes travel in the same direction.

Hydrogen plants with single reformer heaters and capacities up to 100 million ft³/d have used the vertical, down-firing approach. Each burner’s radiant flame covers one-quarter of four adjacent vertical tubes (except for the outside burners, which cover half of the two adjacent tubes).

The radiant gases exit the box horizontally through a horizontal convection section. The horizontal convection section is located about 3 m above grade to allow enough height for passage. The horizontal convection provides for a simpler support structure than that of the side-fired unit.

Transfer-line steam generator

The outlet transfer line from the reformer is used to generate high-pressure (usually 650 psig) steam. The reformer effluent gas exits through the transfer line at about 1,400^oF and enters the tube side of a single-pass steam generator. BFW is fed through the shell side and becomes 650 psig steam. Depending on the size of the reformer, there may be two transfer lines exiting opposite ends of the reformer and feeding two steam generators. Figure 3 shows the two transfer line steam generators.

Feed preheat exchanger.Gas is cooled to about 650^oF and is moved out of the steam generator. It then enters the tube side of the feed preheat exchanger. Feed gas is preheated to about 600F using heat from the effluent gas. This temperature can be controlled by partial bypass of the effluent side to maintain the desired hot-shift gas reactor temperature.

Hot shift-gas reactor.Effluent gas containing carbon monoxide and steam is passed over the hot gas-shift catalyst, where the water-shift gas reaction shown in Equation (5) occurs. This reaction is slightly exothermic, resulting in a temperature rise across the reactor.

More steam generation.Additional medium-pressure steam is generated, reducing the hot-shift reactor effluent to a temperature of about 500^oF, which shifts the reaction equilibrium toward more hydrogen production.

Cold shift-gas reaction.Additional hydrogen is produced by the gas-shift reaction at the lower temperature. The shift reaction is exothermic, which results in a temperature rise across the reactor.

Condensate removal. Cold gas-shift effluent is cooled by heat exchange with BFW, deaerator feedwater, and cooling water to about 34^oC (100^oF). Condensate is separated from the gas in a vertical knockout drum.

Hydrogen purification

Hydrogen purification is generally carried out using one of two approaches — solvent-based systems or pressure-swing adsorption (PSA) processes.

Solvent systems Most older units remove carbon dioxide from the hydrogen rich gas using a solvent, such as Catacarb or amines, in a typical acid gas separation unit (Figure 5).

Figure 5. Most older units remove carbon dioxide from the hydrogen-rich gas with a solvent

Remaining carbon oxides (primarily carbon monoxide) are reacted with hydrogen in a methanator reactor to convert them to methane. Methane is an undesirable component in the makeup gas to a hydrocracker because it builds up in the recycle gas, requiring bleeding of the recycle gas to maintain the desired hydrogen partial pressure in the hydrocracker.

Most solution-type carbon dioxide removal systems are similar. Gas enters the bottom of the absorber, where it contacts lean solution. The carbon dioxide is absorbed from the gas, leaving the rest of the contaminants and hydrogen relatively untouched.

The rich solution is then heat-exchanged with lean solution and enters the top of the stripper. The stripper uses a steam reboiler to regenerate the solvent, stripping out the absorbed carbon dioxide. The overhead from the stripper goes through a condenser to condense solvent and then to an overhead drum, where the carbon dioxide is separated from the stripper reflux.

PSA unit.The newer PSA process produces a hydrogen stream of four-nines (99.99%) purity. It separates carbon monoxide, carbon dioxide and unconverted hydrocarbons. A bank of adsorbers operates in a cycle where the adsorbers are rotated through a higher-pressure adsorption portion, followed by a pressure reduction, which allows the contaminants to be released from the adsorber. The hydrogen gas passes through the adsorber as almost-pure hydrogen. The contaminants flow into a fuelgas surge drum.

Figure 6 shows a schematic diagram of such a system. The valve openings and closings are all controlled by the central processing unit.

Figure 6. A PSA unit separates carbon monoxide, carbon dioxide and unconverted hydrocarbons from hydrogen. Adsorbers operate in a high-pressure to low-pressure cycle to adsorb and then release contaminants

The fuel gas is relatively low-BTU carbon oxides. It is supplemented with natural gas or other fuels as feed to the reformer furnace burners.

Pre- and post-reforming

These are two techniques used to expand the capacity of exisiting plants where the reformer furnace is heat-transfer-limiting.

Pre-reforming Pre-reforming is used when spiking the feed with liquified petroleum gas, which is used to increase the capacity of the existing unit. Examining the reforming Equations (1), (2) and (4) reveals the advantage of a heavier feed that yields more hydrogen per feed mole. The pre-reformer reaction breaks down the heavier hydrocarbons (propane and butane) to methane ahead of the heat-intensive reforming reactions, essentially shifting part of the load upstream of the reformer heater as shown in Figure 7 [8].

Figure 7. A pre-reformer breaks down heavier hydrocarbons into methane ahead of the reforming reactions

Feed at 950^oF passes down through the pre-reformer reactor, where the breakdown reactions occur. Then the pre-reformed feed passes through another convection coil to reheat it to about 1,100^oF before entering the reformer.

Adding the pre-reformer as a retrofit to an existing facility presents two problems — one of space and one of compatibility. Physical space contraints may not allow adding a feed reheat coil within the convection section. Also, the metallurgy of the inlet pigtails may not be able to handle the higher feed temperature.

Post-reforming. Post-reforming is an attempt to provide additional reforming catalyst outside the reformer heater. A down-flow reactor is added between the outlet transfer line and the waste heat steam generator. This can present a space and piping problem. The additional post-reformer catalyst reduces the overall total space velocity of the combined reformer and post reformer, thus achieving additional reaction. This reduces the downstream shift-reaction requirements. ■ Edited by Scott Jenkins

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