HYDROGEN GAS PRODUCTION BY PARTIAL OXIDATION OF HEAVY FUEL OIL

1
CHAPTER ONE
INTRODUCTION
1.1 The Project
This is a report on a design of a plant to produce 20 million standard cubic feet per day (0.555 ×
106
standard m3
/day) of hydrogen (H2) of at least 95% purity from heavy fuel oil (HFO) with an
upstream time of 7680 hours/year applying the process of partial oxidation of the heavy oil
feedstock. Heavy fuel oil feedstock is delivered into the suction of metering-type ram pumps
which feed it via a steam preheater into the combustor of a refractory-lined flame reactor. The
feedstock must be heated to 2000C in the preheater to ensure efficient atomization in the
combustor. A mixture of oxygen and steam is also fed to the combustor, the oxygen being
preheated in a separate steam preheater to 2100C before being mixed with the reactant steam.
The problem statement required that material and energy balances be made for the entire process
and different parts of the process thereof. Further, it was required that various diagrams for
process description, process definition and equipment design purposes be made. Specifically,
these diagrams are the mass balance diagram for the entire process, an energy balance diagram
for the waste heat boiler, the process flow diagram, an instrumentation of control diagram for the
Carbon monoxide (CO) conversion section of the plant and engineering drawings for illustrating
recommendations for the mechanical design of CO catalytic converter.
All necessary information for the design of the project was obtained from the problem statement
and from cited literature as where stated.

2
1.2 Process Description
The process employed is the partial oxidation of oil feedstock. The entire process consists of two
stages:
1. The actual gasification stage consisting of the sub-processes of:
 preheating of starting raw materials, HFO and Oxygen (O2)
 gasification,
 heat/energy recovery and
 quenching
2. The purification stage, consisting of the following sub-stages:
 the Primary Hydrogen sulphide (H2S) removal stage,
 the CO conversion stage,
 the final H2S removal stage and
 the Carbon dioxide (CO2) removal stage
1.2.1 The Gasification Stage
The feedstock is heated to 200o
C in the preheater to ensure efficient atomisation in the
combustor. A mixture of oxygen and steam is also fed to the combustor, the oxygen being
preheated in a separate steam preheater to 210o
C before being mixed with the reactant steam.
The gasification process involves the partial oxidation of a hydrocarbon fuel. Partial oxidation is
a non-catalytic process which involves a combination of exothermic and endothermic reactions,

3
thermal cracking, and steam reforming. The net reaction is however exothermic and produces a
gas which is mainly CO and H2 according to the following overall net reaction.
2𝐶𝐻𝑛 + 𝑂2 ⟶ 2𝐶𝑂 + 𝑛𝐻2 (1 < 𝑛 < 4)
The crude gas contains relatively small amounts of CO2, H2O and H2S and impurities such as
CH4, N2 and NH3, the quantities being determined by the composition of the feed stock, the
oxidant and the actual gasification temperature (usually between 1300o
C-1400o
C but 1300o
C for
this particular process). A small amount of unconverted carbon is also present and ranges from
between 1.0 -1.5 % wt of liquid feed stock (1.5 % in this case).
The hot crude gas then passes to a special waste heat boiler (WHB) which is used to generate
steam, the gas being cooled in the process. The steam produced in the WHB is used as wholly as
reactant in the feed into the reactor with the remainder being sent to the CO conversion section to
be mixed with the feed stream in to the CO catalytic converter. The last step of the gasification
stage of the process is the direct quenching of the crude gas which results in the carbon being
removed as a suspension.
The most important advantage of partial oxidation is that it can process virtually any
hydrocarbon feedstock, from natural gas to petroleum residue and petroleum coke, even solid
feeds such as coal or even metallurgical coke. It should be noted that the hydrogen to CO ratio
primarily depends on the carbon to hydrogen ratio of the feedstock. Historically the SGP was
primarily used with fuel oil (as it is for this particular project case) and bunker C oil as
feedstocks. Overtime, however, the feed has become more concentrated and viscous, containing
higher levels of sulphur and heavy metals.

4
Presently, hydrocarbon fuels that can be used for this process include natural gas, asphalt,
refinery gas, vacuum residue, orimulsion and liquid waste.
1.2.2 Purification Stage
The quenched gas passes to an H2S removal stage where H2S is selectively scrubbed down to 15
parts per million with substantially nil removal of CO2. Solution regeneration in this process is
undertaken using the waste low-pressure steam from another process. The scrubbed gas, at 35o
C
and saturated, then undergoes CO conversion, final H2S removal, and CO2 removal to allow it to
meet the product specification. CO conversion is carried out over chromium-promoted iron oxide
catalyst employing two stages of catalytic conversion; the plant also incorporates a saturator and
desaturator operating with a hot water circuit. Incoming gas is introduced into the saturator (a
packed column) where it is contacted with hot water pumped from the base of the desaturator;
this process serves to preheat the gas and to introduce into it some of the water vapour required
as reactant. The gas then passes to two heat exchangers in series. In the first, the unconverted gas
is heated against the converted gas from the second stage of catalytic conversion; in the second
heat exchanger the unconverted gas is further heated against the converted gas from the first
stage of catalytic conversion. The remaining water required as reactant is then introduced into
the unconverted gas as steam at 600 psig (4140 kN/m2
gauge) saturated (a part of this supply
coming from the steam produced in the WHB) and the gas/steam mixture passes to the catalyst
vessel at a temperature of 370o
C. The catalyst vessel is a single shell with a dividing plate
separating the two catalyst beds which constitute the two stages of conversion. The converted
gas from each stage passes to the heat exchangers previously described and thence to the
desaturator, which is a further packed column. In this column the converted gas is contacted
countercurrent with hot water pumped from the saturator base; the temperature of the gas is

5
reduced and the deposited water is absorbed in the hotwater circuit. An air-cooled heat
exchanger then reduces the temperature of the converted gas to 40o
C for final H2S removal. Final
H2S removal takes place in four vertical vessels each equipped with five trays of iron oxide
absorbent. Gas leaving this section of the plant contains less than 1 ppm of H2S and passes to the
CO2 removal stage at a temperature of 35o
C. CO2 removal is accomplished employing Mono-
ethanolamine (MEA) solution. The feed gas is fed to the bottom of absorber and flows upward
countercurrent to the descending solvent. The rich solvent in which the acidic components are
dissolved leaves the bottom of the absorber and is sent to the stripper and regenerated there using
low pressure steam. The regenerated solvent is recycled back to the top of the absorber.
1.3 Justification of the project
This project comes at a critical time in the world’s evolving energy sector. Against the backdrop
of rising energy costs, soaring environmental concerns in the area of emissions from burning
fossils fuels, it would seem only a matter of time before a sustained shift towards more
acceptable energy sources ensues. The United States government seems to have taken the lead
among oil producing nations in this area, albeit years after countries like Japan have gone as far
as developing technologies based on cleaner fuels.
H2 has long been recognized as a potential cleaner replacement for fossil fuels and still remains
an important starting material for quite a number of important chemical products. Africa, indeed
Nigeria stands a far better chance of future global importance by tapping early into this
foreseeable future global trend. Hence, the review of processes such as this that utilizes fossil
fuels as starting material in order to produce potentially cleaner fuels like H2 is very crucial in
the course of our national development. Also, the competitiveness of the SGP as a means of

6
achieving the production of H2 is of crucial importance since as in all cases, motivation for
process industrial development, construction and operation is firmly rooted in economic
viability. These two areas (the demand for H2 as a cleaner fuel and the competiveness of the
SGP), are thus examined as justification of the current project.
1.3.1 Global demand for H2
Hydrogen is required now more than ever on an ever increasing scale globally. With the drive
towards a hydrogen economy wherein power and transportation systems will be run solely on
hydrogen as fuel or as a storage medium the demand for hydrogen will certainly escalate in the
years to come. This is of course coupled with the already high demand for hydrogen in the
chemical and petrochemical industries as a starting material in the synthesis of chemical raw
materials. Also the demand for hydrogen in refinery operations is growing constantly as attempts
are made to comply with global legislature dictating low sulphur contents. The motivation for all
of this is a cleaner, greener environment in which pollution is eliminated. Specifically, in the
automotive industry, vehichles run on hydrogen or hydrogen based technologies (fuel cells, for
example) produce little or no emissions of greenhouse gases. Also H2 high energy content — 1
kg corresponding to 3.5 litres of petroleum — highlights its importance as fuel in applications
where weight rather than volume is the important factor.
In Nigeria today due to the problems in the power sector, the need arises to develop modern long
term technologies for power generation. One of the means adopted is a in the diversification of
the sources of electricity to include both solar and wind sources as well as the traditional hydro
and thermal sources. In this area however, the current trend is to shore up the contribution of
these fluctuating generation electricity producers (Wind, Solar) with the aid of a storage

7
mechanism. Hydrogen presently has the most attractive properties as a "Storage Medium" of
electricity:
Compared with the storage of electricity in batteries, the material costs are many times lower.
Also, electricity generation using fuel cells could prove to be a very attractive alternative to the
conventional technologies.
Also in the long term, hydrogen may very well be on its way to becoming a pollution free fuel
for all traffic applications, be it ships, trains or airplanes.

8
CHAPTER TWO
LITERATURE REVIEW
2.1 Overview
Hydrogen is a colourless, odourless, tasteless, flammable gaseous substance that is the simplest
member of the family of chemical elements. The hydrogen atom has a nucleus consisting of a
proton bearing one unit of positive electrical charge; an electron, bearing one unit of negative
electrical charge, is also associated with this nucleus. Under ordinary conditions, hydrogen gas is
a loose aggregation of hydrogen molecules, each consisting of a pair of atoms, a diatomic
molecule, H2. The earliest known important chemical property of hydrogen is that it burns with
oxygen to form water, H2O; indeed, the name hydrogen is derived from Greek words meaning
―water former.‖
It occurs in vast quantities as part of the water in oceans, ice packs, rivers, lakes, and the
atmosphere. As part of innumerable carbon compounds, hydrogen is present in all animal and
vegetable tissue and in petroleum. Since hydrogen is contained in almost all carbon compounds
and also forms a multitude of compounds with all other elements (except some of the noble
gases), it is possible that hydrogen compounds are more numerous.
Elementary hydrogen finds its principal industrial application in the manufacture of ammonia (a
compound of hydrogen and nitrogen, NH3) and in the hydrogenation of carbon monoxide and
organic compounds.

9
2.2 Properties of Hydrogen
2.2.1 Melting and Boiling Point
Hydrogen has an extremely low melting and boiling points resulting from the weak forces of
attraction between the molecules. The existence of these weak intermolecular forces is also
revealed by the fact that, when hydrogen gas expands from high to low pressure at room
temperature, its temperature rises, whereas the temperature of most other gases falls. According
to thermodynamic principles, this implies that repulsive forces exceed attractive forces between
hydrogen molecules at room temperature—otherwise, the expansion would cool the hydrogen. In
fact, at −68.6°C attractive forces predominate, and hydrogen, therefore, cools upon being
allowed to expand below that temperature. The cooling effect becomes so pronounced at
temperatures below that of liquid nitrogen (−196°C) that the effect is utilized to achieve the
liquefaction temperature of hydrogen gas itself.
2.2.2 Mobility
Hydrogen is transparent to visible light, to infrared light, and to ultraviolet light to wavelengths
below 1800 Å. Because its molecular weight is lower than that of any other gas, its molecules
have a velocity higher than those of any other gas at a given temperature and it diffuses faster
than any other gas. Consequently, kinetic energy is distributed faster through hydrogen than
through any other gas; it has, for example, the greatest heat conductivity.
2.2.3 Reactivity of hydrogen
One molecule of hydrogen dissociates into two atoms (H2 → 2H) when an energy equal to or
greater than the dissociation energy (i.e., the amount of energy required to break the bond that

10
holds together the atoms in the molecule) is supplied. The dissociation energy of molecular
hydrogen is 104,000 calories per mole—104 kcal/mole (mole: the molecular weight expressed in
grams, which is two grams in the case of hydrogen). Sufficient energy is obtained, for example,
when the gas is brought into contact with a white-hot tungsten filament or when an electric
discharge is established in the gas. If atomic hydrogen is generated in a system at low pressure,
the atoms will have a significant lifetime—e.g., 0.3 second at a pressure of 0.5 millimetre of
mercury. Atomic hydrogen is very reactive. It combines with most elements to form hydrides
(e.g., sodium hydride, NaH), and it reduces metallic oxides, a reaction that produces the metal in
its elemental state. The surfaces of metals that do not combine with hydrogen to form stable
hydrides (e.g., platinum) catalyze the recombination of hydrogen atoms to form hydrogen
molecules and are thereby heated to incandescence by the energy that this reaction releases.
Molecular hydrogen can react with many elements and compounds, but at room temperature the
reaction rates are usually so low as to be negligible. This apparent inertness is in part related to
the very high dissociation energy of the molecule. At elevated temperatures, however, the
reaction rates are high.
Sparks or certain radiations can cause a mixture of hydrogen and chlorine to react explosively to
yield hydrogen chloride, as represented by the equation
H2 + Cl2 → 2HCl.
Mixtures of hydrogen and oxygen react at a measurable rate only above 300° C, according to the
equation
2H2 + O2 → 2H2O

11
Such mixtures containing 4 to 94 percent hydrogen ignite when heated to 550°–600°C or when
brought into contact with a catalyst, spark, or flame. The explosion of a 2:1 mixture of hydrogen
and oxygen is especially violent. Almost all metals and nonmetals react with hydrogen at high
temperatures. At elevated temperatures and pressures hydrogen reduces the oxides of most
metals and many metallic salts to the metals. For example, hydrogen gas and ferrous oxide react,
yielding metallic iron and water,
H2 + FeO → Fe + H2O;
hydrogen gas reduces palladium chloride to form palladium metal and hydrogen chloride,
H2 + PdCl2 → Pd + 2HCl.
Hydrogen is absorbed at high temperatures by many transition metals (scandium, 21, through
copper, 29; yttrium, 39, through silver, 47; hafnium, 72, through gold, 79); and metals of the
actinoid (actinium, 89, through lawrencium, 103) and lanthanoid series (lanthanum, 57, through
lutetium, 71) to form hard, alloy-like hydrides. These are often called interstitial hydrides
because, in many cases, the metallic crystal lattice merely expands to accommodate the dissolved
hydrogen without any other change.
2.2.4 Hydrogen bond
Some covalently bonded hydrides have a hydrogen atom bound simultaneously to two separate
electronegative atoms, which are then said to be hydrogen bonded. The strongest hydrogen
bonds involve the small, highly electronegative atoms of fluorine (F), oxygen, and nitrogen. In
the bifluoride ion, HF2−
, the hydrogen atom links two fluorine atoms. In the crystal structure of
ice, each oxygen atom is surrounded by four other oxygen atoms, with hydrogen atoms between

12
them. Some of the hydrogen bonds are broken when ice melts, and the structure collapses with
an increase in density. Hydrogen bonding is important in biology because of its major role in
determining the configurations of molecules. The helical (spiral) configurations of certain
enormous molecular chains, as in proteins, are held together by hydrogen bonds. Extensive
hydrogen bonding in the liquid state explains why hydrogen fluoride (HF), water (H2O), and
ammonia (NH3) have boiling points much higher than those of their heavier analogues, hydrogen
chloride (HCl), hydrogen sulfide (H2S), and phosphine (PH3). Thermal energy required to break
up the hydrogen bonds and to permit vaporization is available only at the higher boiling
temperatures.
The hydrogen in a strong acid, such as hydrochloric (HCl) or nitric (HNO3), behaves quite
differently. When these acids dissolve in water, hydrogen in the form of a proton, H+, separates
completely from the negatively charged ion, the anion (Cl−
or NO3−
), and interacts with the water
molecules. The proton is strongly attached to one water molecule (hydrated) to form the
oxonium ion (H3O+
, sometimes called hydronium ion), which in turn is hydrogen-bonded to
other water molecules, forming species with formulas such as H(H2O)n+
(the subscript n
indicates the number of H2O molecules involved). The reduction of H+
(reduction is the chemical
change in which an atom or ion gains one or more electrons) can be represented as the half
reaction:
H+
+ e−
→ 1/2H2.
The energy needed to bring about this reaction can be expressed as a reduction potential. The
reduction potential for hydrogen is taken by convention to be zero, and all metals with negative
reduction potentials i.e., metals that are less easily reduced (more easily oxidized; e.g., zinc)

13
Zn2+
+ 2e−
→ Zn − 0.763 volt
This can, in principle, displace hydrogen from a strong acid solution:
Zn + 2H+
→ Zn2+
+ H2.
Metals with positive reduction potentials (e.g., silver: Ag+
+ e−
→ Ag, + 0.7995 volt) are inert
toward the aqueous hydrogen ion.
2.3 Hydrogen Demand
There has been a continual increase in refinery hydrogen demand over the last several decades.
This is a result of two outside forces acting on the refining industry: environmental regulations
and feedstock shortages. These are driving the refining industry to convert from distillation to
conversion of petroleum. Changes in product slate, particularly outside the United States, are
also important. Refiners are left with an oversupply of heavy, high-sulfur oil, and in order to
make lighter, cleaner, and more salable products, they need to add hydrogen or reject carbon.
The early use of hydrogen was in naphtha hydro-treating, as feed pretreatment for catalytic
reforming (which in turn was producing hydrogen as a by-product). As environmental
regulations tightened, the technology matured and heavier streams were hydro-treated.
These included light and heavy distillates and even vacuum residue. Hydro-treating has also been
used to saturate olefins and make more stable products. For example, the liquids from a coker
generally require hydro-treating, to prevent the formation of polymers.

14
At the same time that demand for cleaner distillates has increased, the demand for heavy fuel oil
has dropped. This has led to wider use of hydrocracking, which causes a further large increase in
the demand for hydrogen.
2.4 Methods of Production
As hydrogen use has become more widespread in refineries, hydrogen production has moved
from the status of a high-technology specialty operation to an integral feature of most refineries.
This has been made necessary by the increase in hydro-treating, hydrocracking, and partial
oxidation of heavier feedstocks. Steam reforming is the dominant method for hydrogen
production. This is usually combined with wet scrubbing to purify the hydrogen to greater than
99 vol %.
2.4.1 Steam Reforming/Wet Scrubbing
In this case the feedstock is preheated and purified to remove traces of sulfur and halogens in
order to protect the reformer catalyst. The most common impurity is H2S; this is removed by
reaction with ZnO. Organic sulfur may also be present; in this case recycled product hydrogen is
mixed with the feed and reacted over a hydrogenation catalyst (generally cobalt/molybdenum) to
convert the organic sulfur to H2S. If chlorides are present, they are also hydrogenated and then
reacted with a chloride adsorbent. The feed is then mixed with steam, preheated further and
reacted over nickel catalyst in the tubes of the reformer to produce synthesis gas—an equilibrium
mixture of H2, CO, and CO2. The steam/carbon ratio is a key parameter, since high steam levels
aid in methane conversion. Residual methane in the synthesis gas will pass through the plant
unchanged (along with any N2 in the feed).

15
The synthesis gas passes through the reformer waste heat exchanger, which cools the gas and
generates steam for use in the reformer; the surplus is exported. The cooled gas [still at about
650°F (345°C)] is reacted over a fixed bed of iron oxide catalyst in the high temperature shift
converter, where the bulk of the CO is reacted, then cooled again and reacted over a bed of
copper zinc low-temperature shift catalyst to convert additional CO. The raw hydrogen stream is
next scrubbed with a solution of a weak base to remove CO2.
CO2 in the gas reacts reversibly with potassium carbonate to form potassium bicarbonate. The
solution is depressured and steam-stripped to release CO2, with the heat for the regenerator
reboiler coming from the hot synthesis gas. The regenerator overhead stream is then cooled to
condense water. The CO2 is available for recovery or can be vented. The raw hydrogen leaving
the CO2 removal section still contains approximately 0.5 percent CO and 0.1 percent CO2 by
volume. These will act as catalyst poisons to most hydrogen consumers, so they must be
removed, down to very low levels. This is done by methanation, the reverse of reforming. As in
reforming, a nickel catalyst is used, but as a fixed bed.
Typical final hydrogen purity is 97 vol %, with the remaining impurities consisting mainly of
methane and nitrogen. Carbon oxide content is less than 50 vol ppm.
Product hydrogen is delivered from the methanator at approximately 250 lb/in2 (17 bar) guage
and must generally be compressed before final use. This is done in a reciprocating compressor.
Centrifugal compressors are not feasible because of the low molecular weight; the pressure rise
per foot of head is too low, and too many stages would be required.
Chemistry

16
In steam reforming, light hydrocarbons such as methane are reacted with steam to form hydrogen
(Fuderer, 1982):
The reaction is typically carried out at approximately 1600°F (870°C) over a nickel catalyst
packed into the tubes of a reforming furnace. Because of the high temperature, hydrocarbons also
undergo a complex series of cracking reactions, plus the reaction of carbon with steam. These
can be summarized as
2.4.2 Partial Oxidation
Partial Oxidation (POX) involves the reaction of hydrocarbon feed with oxygen at high
temperatures to produce a mixture of hydrogen and carbon monoxide. It’s a process that involves
an intimate coupling of several complex chemical reactions which produce synthesis gas (CO
and H2). The partial oxidation reaction mechanism involves exothermic, partial combustion of a
portion of a hydrocarbon feed which supplies heat to the endothermic steam cracking of the
balance of the feed. Besides carbon monoxide, hydrogen, carbon dioxide, hydrogen sulfide, and
other trace impurities, partial oxidation produces soot in non-equilibrium amounts. The

17
composition of the products, particularly H2 /CO ratio, sulfur, and soot, are generally determined
by the type of feedstock, the oxygen/fuel ratio and the amount of steam used.
Since the high temperature takes the place of a catalyst, POX is not limited to the light, clean
feedstocks required for steam reforming. Partial oxidation is high in capital cost, and for light
feeds it has been generally replaced by steam reforming. However, for heavier feedstocks it
remains the only feasible method (Miller et.al., 1989).
2.4.3 Catalytic Partial Oxidation
Also known as auto-thermal reforming, catalytic partial oxidation involves the reaction of
oxygen with a light feedstock, passing the resulting hot mixture over a reforming catalyst. Since
a catalyst is used, temperatures can be lower than in non-catalytic partial oxidation, which
reduces the oxygen demand (Fuderer, 1982).
Feedstock composition requirements are similar to those for steam reforming: light hydrocarbons
from refinery gas to naphtha may be used. The oxygen substitutes for much of the steam in
preventing coking, so a lower steam/carbon ratio can be used. Since a large excess of steam is
not required, catalytic POX produces more CO and less hydrogen than steam reforming (Twigg,
1989). Because of this it is suited to processes where CO is desired, for example, as synthesis gas
for chemical feedstocks.
2.4.4 Advanced Cracking Reaction (ACR)
In this, crude oil feedstock values are converted by a thermal cracking reaction mechanism to
reaction products high in olefins. Superheated steam is generated by the burning of oxygen and
fuel (usually H2 and/or CH4) to produce combustion gases of about 2000° C. This is

18
supplemented by superheated steam generated externally from the reaction zone. The combined
streams form the so-called "heat carrier" or steam cracking medium (Kamma, 1979).
Downstream from the burner, crude oil distillates are injected into the high temperature stream
and rapidly vaporize. The vaporized feedstock and combustion gases are accelerated through an
orifice or throat into the diffuser or reaction chamber where the adiabatic cracking occurs in 10-
20 milliseconds residence time. The steam and reaction products are rapidly quenched and then
cooled in a unique wetted-wall heat exchanger which generates high pressure steam. A gas-liquid
phase separation takes place, pitch being discharged and the vapor going to the gasoline
fractionator which is followed by compression and acid gas removal. This results in an olefins
rich stream containing ethylene, acetylene, propylene, and the other cracking by-products.
2.4.5 Thermochemical (High-temperature) Water Splitting
Thermochemical (High-temperature) water splitting involves the use of High temperature heat
(500 – 2000 °C) to drive a series of chemical reactions that produce hydrogen. The chemicals
used in the process are reused within each cycle, creating a closed loop that consumes only water
and produces hydrogen and oxygen. The high-temperature heat needed can be supplied by next-
generation nuclear reactors under development (up to about 1000 °C) or by using sunlight with
solar concentrators (up to about 2000 °C).
2.5 Feedstocks for Hydrogen Production
The best feedstocks for steam reforming are light, saturated, and low in sulfur; this includes
natural gas, refinery gas, LPG, and light naphtha. The use of Heavy Fuel Oil also proves to be

19
efficient but less efficient than fuels with lower sulfur content. These feeds can be converted to
hydrogen at high thermal efficiency and low capital cost.
2.5.1 Natural Gas
Natural gas is the most common hydrogen plant feed, since it meets all the requirements for
reformer feed and is low in cost. A typical pipeline natural gas contains over 90 percent C1 and
C2, with only a few percent of C3 and heavier hydrocarbons. It may contain traces of CO2, with
often significant amounts of N2 which affect the purity of hydrogen of which can be removed in
the wet scrubbing unit if required.
2.5.2 Refinery Gas
Light refinery gas, containing a substantial amount of hydrogen, can be an attractive steam
reformer feedstock. Since it is produced as a by-product, it may be available at low cost.
Processing of refinery gas will depend on its composition, particularly the levels of olefins and of
propane and heavier hydrocarbons. Olefins can cause problems by forming coke in the reformer.
They are converted to saturated compounds in the hydrogenator, giving off heat. This can be a
problem if the olefin concentration is higher than about 5 percent, since the hydrogenator will
overheat. A recycle system can be installed to cool the reactor, but this is expensive and wastes
heat.
Heavier hydrocarbons in refinery gas can also form coke, either on the primary reformer catalyst
or in the preheater. If there is more than a few percent of C3 and higher compounds, a promoted
reformer catalyst should be considered, to avoid carbon deposits.
2.5.3 Liquid Feeds

20
Liquid feeds, either LPG or naphtha, can be attractive feedstocks where prices are favorable.
Naphtha is typically valued as low-octane motor gasoline, but at some locations there is an
excess of light straight-run naphtha, and it is available cheaply. Liquid feeds can also provide
backup feed, if there is a risk of natural gas curtailments.
The feed handling system needs to include a surge drum, feed pump, and vaporizer, usually
steam-heated. This will be followed by further heating, before desulfurization. The sulfur in
liquid feeds will be in the form of mercaptans, thiophenes, or heavier compounds.
2.5.4 Heavy Fuel Oil
This involves heating the feedstock to about 2000o
C in the preheater to ensure efficient
atomization in the combustor. Thereafter, a mixture of oxygen (being preheated in a separate
steam preheater to 2100o
C) and steam is passed to the combustor. The crude gas, which will
contain some carbon particles, leaves the reactor at approximately 1300o
C and passes
immediately into a special waste-heat boiler where steam at 600 psig (4140 kN/m2
gauge) is
generated. The crude gas leaves the waste heat boiler at 2500o
C and is further cooled to 500o
C by
direct quenching with water, which also serves to remove the carbon as a suspension.
Suitable separation techniques are thereafter applied to remove the CO, CO2, CH4 and H2S
content of the product of combustion.

21
Figure 1: SI Cycle Flowsheet
(Source: McHugh, 2005. ―Hydrogen production methods‖)

22
CHAPTER THREE
DISCUSSION
3.1 Scope of design Work Required
The following design work was carried out as stipulated by the problem statement:
1. Process design
2. Mechanical design and
3. Control (Instrumentation) system design.
3.2 Process Design
The following process design activities were carried out as required by the problem statement
(see appendix IV).
1. Preparation of a mass balance diagram for the entire plant,
2. Preparation of an energy balance diagram of the flame reactor and the associated waste
heat boiler,
3. Preparation of a Process flow diagram of the entire plant.
3.2.1 Mass Balance
Table 1 shows the summary of the various inlet and outlet streams identified by the stream
numbers involved in the process.

23
Table 1: Gasification/Purification Gas Flows, Compositions, Pressures and Temperatures of Hydrogen Gas Synthesis
Basis: 18.83Km3
/hr.
On Stream Time: 7680 hr/yr.
Stream No 1 2 3A 3B 3 4 5 5A
TAN
K
PREHEAT
ER
O2
PREHEAT
ER
STEA
M
MIXE
R
FLAME
REACTO
R
WH
R
STEAM
GENERAT
ED
QUENCHE
R
LIN
E
PRODUC
T
STREAM
C
H2
S
CO
CO2
O2
CH4
H2S
N2
16.02
2.071
7.532
-
-
-
-
-
-
16.02
2.071
7.532
-
-
-
-
-
-
-
-
-
-
-
21.843
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
21.843
-
-
-
-
-
-
-
-
21.843
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
-
8.963
-
7.927
1.563
-
0.0188
0.094
0.264
-
8.963
-
7.927
1.563
-
0.018
8
0.094
0.264
-
17.889
-
0.188
0.188
-
0.188
-
0.377
PRESSURE
(KN/M2
)
206.9 - - 4140 4140 - 4140 - - -
TEMPERATUTRE(
O
C)
50 200 210 - 210 300 250 - - - -

24
Material Balance Basis
In order to perform the material balance calculations, an assumed basis of 1000 kg of Heavy Fuel
Oil feed was selected. A weight basis was selected due to the nature of the information given on
the principal raw material –HFO.
Figure 2: The Converter Section of the Plant.
From the above, and as calculated on appendix 1, the production rate was found to be
4.041x108m3/hr. Hence, the rate of flow of CO, CO2, N2 into the conversion system were found
to be 4040918m3
/hr, 4040918m3
/hr and 8081836m3
/hr. respectively. While CO, CO2 coming out
of the quencher were calculated to 242,455080m3
and 47913742m3
respectively. Also, the
amount of CO converted is 238,414,162 m3
. Then, H2, CH4 out of the converter were found to be
271318780 m3
and 57727m3
respectively. Finally, the amount of H2 produced in the converter
was found to be 1.329 x 108m3
.
Process Stage by stage material balance
The mass balance sheet was divided into various sections according to the different stages in the
process along the route from the raw materials’ streams to the final product steam in which there
are changes in material quantities and compositions. The sections include:

25
1. Gasification,
2. Quencher,
3. Primary H2S Removal section,
4. CO Conversion Section,
5. Secondary H2S Removal stage and
6. The CO2 Removal Section,
7. Equipment Schedule for the C Conversion Section.
All Calculations in these sections typically employ the principles of mass and material balance:
conservation of mass, the tie component principle e.t.c. as shown by appendix I. Some peculiar
treatment required in the course of the mass balancing are however highlighted as follows:
1. Gasification, Quencher and Primary H2S Removal Stage
The first three stages constitute the backbone of pertinent information necessary to the material
balancing of the entire process from raw materials to product, as supplied by the problem
statement.
In the gasification section material quantities and compositions changes in the flame reactor
which is followed by stream component division in the quencher. Consequently a mass balance
is carried out around a material boundary which encompasses both units. The nature of the
information supplied in the problem statement necessitates the computing of atomic balances in
order to correctly estimate material flows in and out of this material boundary.

26
2. CO Conversion Section
In order to perform material balance for the CO conversion section, a combined mass and energy
balance as well as reaction equilibria were employed.
In the CO conversion section, material quantities changes in four units namely: the saturator, the
mixer, the catalytic converter and the desaturator. Specifically, the main process stream
undergoes changes in water content in the saturator (increases), mixer (increases) and the
desaturator (decreases). In the converter, CO and water are used up to produce more CO2 and H2.
In proceeding, the quantities of water to be added in the saturator and desaturator were fixed in
anticipation that units to achieve that level of heat and mass transfer can be easily designed. With
those quantities fixed, the amount of water necessary for reaction in the catalytic converter that
will be added as make-up through the mixer was then estimated. However preceding this, the
total amount of water needed in the catalytic converter in order to achieve:
i. target conversion based on the amount of CO left in the material stream ascompared
to the product stream CO content specification as well as,
ii. adiabatic operation of each stage of the CO catalytic converter unit, was estimated.
3. Secondary H2S Removal stage and the CO2 Removal Section.
Of interest here is the fact that since the gas stream is being cooled progressively as it flows
along (primarily through the air–cooled heat exchanger and subsequently as a result of
anticipated net losses), some water will condense out, thus giving rise to some form of two phase
flow. However, due to the nature of the problem statement, the effect of this was not of
immediate concern.

27
3.2.2. Energy Balance Diagram
The energy balance diagram for the flame reactor and waste heat boiler is shown in figure 3. The
diagram shows the material and energy flows, all the streams associated with the two units as
well as the heat load corresponding to each unit. Also indicated is the datum temperature used
for the computation.
Figure 3: Energy Balance Diagram for the Flame Reactor and the Associated Waste Heat
Boiler (WHB).

28
The values of the material streams indicated were extracted from the mass balance diagram
discussed earlier. As indicated in the diagram the mass of steam produced by waste heat boiler
was estimated to be 7930.95kg while energy of the crude gas leaving the waste heat boiler is
4.537 x 1011
kJ. The energy required to produce steam by the waste heat boiler was estimated to
be 22,325,624.25 kJ. See appendix III for detailed calculation on the energy balance of the
system.

29
3.2.3 Process Flow Diagram
3.2.3.1 The Preliminary Block Flow Diagram
Figure 4 is the diagram showing all the components of the plant and how they are interconnected for process delivery. Figure 5 shows
the detailed process flow diagram of the plant including all the components.
Figure 4: Block Flow Diagram for Hydrogen Synthesis Using Partial Oxidation of Heavy Fuel Oil.

30
Figure 5: Process flow diagram of the plant.

31
3.2.4 Equipment Schedule for the CO Conversion Section of the Plant
3.2.4.1 Introduction
In order to meet the production capacity as stated in the process description, the process is
scheduled to run all day long (24hrs of the day).
The following are the various utilities and items of equipment involved in the CO conversion
section of the plant:
 Utilities: Heat supply, Potassium Carbonate (K2CO3), Catalyst (chromium-promoted
iron oxide catalyst), and steam.
 Items of Equipment: A saturator, 2 desaturators, catalyst vessel (2 catalyst beds), trays
of iron oxide absorbent, 2 heat exchangers.
Figure 4 shows the Block Flow Diagram (BFD) of the equipment arrangement.
Activity Definition/Sequencing
1. Incoming Gas Preheating: Incoming gas is introduced into the saturator (a packed
column) where it is contacted with hot water pumped from the base of the desaturator;
this process serves to preheat the gas and to introduce into it some of the water vapour
required as reactant.
2. Catalytic Conversion: The gas then passes to two heat exchangers in series. In the first,
the unconverted gas is heated against the converted gas from the second stage of catalyst

32
conversion; in the second heat exchanger the unconverted gas is further heated against
the converted gas from the first stage of catalytic conversion.
3. Desaturation Process: The converted gas from each stage passes to the heat exchangers
previously described and thence to the desaturator, which is a further packed column. In
this column the converted gas is contacted countercurrent with hot water pumped from
the saturator base; the temperature of the gas is reduced and the deposited water is
absorbed in the hot-water circuit.
4. Air-Cooling: An air-cooled heat exchanger then reduces the temperature of the
converted gas to 400C for final H2S removal.
5. Final Conversion and CO2 Removal: Final H2S removal takes place in four vertical
vessels each approximately 60 feet (18.3 m) in height and 8 feet (2.4 m) in diameter and
equipped with five trays of iron oxide absorbent. Each vessel is provided with a locking
lid of the autoclave type. The total pressure drop across these vessels is 5 psi (35 kN/m2
).
Gas leaving this section of the plant contains less than 1 ppm of H2S and passes to the
CO2 removal stage at a temperature of 350C. This now leads to the removal of CO2
employing high-pressure potassium carbonate wash with solution regeneration. (see
figure 4)

33
Figure 5: Block Flow Diagram of CO conversion section

34
Activity Duration Estimation
1. Incoming Gas Preheating: This activity is estimated to take 30-45 minutes to complete
which is a function of the heating temperature.
2. Catalytic Conversion: The conversion of the unconverted gas at the second stage of
catalyst conversion takes approximately 1hour, and that of the first stage conversion also
takes 1hour independently of the first conversion process. At the first heat exchanger,
conversion is expected to last for less than 1 hour. Likewise the second heat exchanger.
3. Desaturation Process: This step is expected to last for less than 2 hours as described by
step 3 above.
4. Air-Cooling: This process is allowed to last for 3-4 hours to ensure proper cooling of the
air from the air-cooled exchanger.
5. Final Conversion and CO2 Removal: The final conversion and removal of H2S is
expected to last for 2 hours while CO2 removal with the aid of high-pressure potassium
carbonate wash with solution regeneration is estimated to last for 1 hour. The autoclave-
like lid helps to prevent heat escape which improves the efficiency of the process.
The whole process for conversion of CO is expected to last for 9-11 hours.
The scheduling including the equipment section is as shown on table 2.

35
Table 2: Equipment Schedule of CO Conversion Section of the Plant.

36
3.3 Mechanical Design of the Absorber
Only the mechanical design of the Absorber Section was considered. The necessary calculations
needed used in estimating the vessel diameter, shell thickness and height are given on appendix
II. Table 3 shows the breakdown of the system.
Table 3: Process Operation Conditions Specification.
Parameters Values
Temperature
Operating Temperature
Design Temperature
83C
91.3C
Pressure
Operating Pressure
Design Pressure
28.5 bar
31.35 bar
Column diameter
Column Height
Shell thickness
Shell shape
End plate Details
Material of construction
Mechanism of Liquid distribution
18.3m
66.6m
12mm+2mm corrosion allowance
Cylindrical
Hemispherical end is recommended
carbon steel
Liquid sprayer

37
Table 4: Size Specification for the Shell and Catalytic Converter Components.
For the shell Values
Diameter of the column
Height of the column
Outside pressure is 1atm
Allowable stress f
Welding joint efficiency
Modulus of elasticity E
18.3m
66.6m
0.1 MN/m2
= 98.1 MN/m2
0.85
1 x 105
MN/m2
Catalytic Converter
Outside pressure is 1atm
Design pressure
Allowable stress
Welding joint efficiency
0.1 MN/m2
0.105 MN/m2
98.1 MN/m2
0.85
The following considerations were proposed for the design of the vessel internals as required by
the problem statement:
1. Catalyst bed supports
The catalyst bed support proposed is shown in figure 5. It consists of two half circle frames made
of alloy steel held together at the centre by means of bolts to which wire gauze is fitted. The

38
gauze pores will be big enough to allow the passage of the gas but small enough to prevent the
catalyst particles from passing through it. Proposed dimensions are given in figure 6.
2. Gas distribution
In order to facilitate the proper distribution of the gas within the vessel the gas distributor of
figure 7 is proposed. It consists of a conical shaped receiver connected to an elbow pipe which
acts as a means of re-directing the gas. The high pressure at which the gas is to be pumped will
make this erstwhile simple design effective. The proposed dimensions are given in figure 8.
3. Facilities for discharging and charging the catalyst
The manholes provided by the design are to be used for charging the catalyst while the catalyst
discharge nozzle will facilitate in the catalyst discharge. The nozzle is designed to taper for ease
of collection of the catalyst particles. To push the particles out of the vessels, compressed air is
recommended.

39
Figure 6: CO Catalytic Convereter
Source: Osafehinti, E.K. (2008). ―Hydrogen production using Shell Gasification Process (SGP)
Scheme of heavy fuel oil.‖

40
Figure 7: Dimensioned CO Catalytic Converter.

41
Figure 8: Catalyst Bed Support.

42
Figure 9: Top Elevation of Catalyst Bed Support.

43
Figure 10: Gas Distributor.

44
Figure 11: Gas Distributor (Front and Side Elevation).

45
Figure 12: Gas Distributor.

46
4. Choice of construction materials
Due to the design of the vessel- the number of openings, branches and the temperature regime
expected the best choice for construction will be alloy steels.
5. Facilities for instrumentation
The instrumentation ports shown in figure 12 are to facilitate the need for instrumentation. They
have been located to the site at which measurements (gas temperature and pressure) are to be
taken.
3.4 Control (Instrumentation) System Design
The control of instrumentation diagram of the CO conversion section is shown in figure 12. In
order to control the liquid level in the hot water circuit an additional tank was introduced which
will act as a means to remove the inevitable build-up of water in the circuit. The control the
temperature levels in the converter, the flow of the gas stream into the unit is manipulated. In the
case of the failure of the control system, two solenoid valves are provided to shut down the entire
section by shutting down the flow of material into the converter and into the whole section itself.
Also provided are high and low level alarms for both the temperature control and the liquid level
control.

47
Figure 13: Control and Instrumentation Plan of the CO Conversion Section.
Source: Osafehinti, E.K. (2008). ―Hydrogen production using Shell Gasification Process (SGP) Scheme of heavy fuel oil.‖

48
3.5 Health, Safety and Environmental Considerations/Project Costing
3.5.1 Precautions for the Safe Handling of Hydrogen
Hydrogen is an extremely flammable gas that forms explosive mixtures with air and oxidizing
agents. Hence, the following precautions should be taken when handling hydrogen:
 Keep away from heat, hot surfaces, sparks, open flames and other ignition sources. There
should be no smoking near the plant. The use of only non-sparking tools is
recommended.
 Wear leather safety gloves and safety shoes when handling cylinders.
 Protect cylinders from physical damage; do not drag, roll, slide or drop.
 While moving cylinder, always keep in place removable valve cover. Never attempt to
lift a cylinder by its cap; the cap is intended solely to protect the valve.
 When moving cylinders, even for short distances, use a cart (trolley, hand truck, etc.)
designed to transport cylinders.

49
 Never insert an object (e.g., wrench, screwdriver, pry bar) into cap openings; doing so
may damage the valve and cause a leak.
 Use an adjustable strap wrench to remove over-tight or rusted caps. Slowly open the
valve. If the valve is hard to open, discontinue use and contact your supplier. Close the
container valve after each use; keep closed even when empty.
 Never apply flame or localized heat directly to any part of the container. High
temperatures may damage the container and could cause the pressure relief device to fail
prematurely, venting the container contents.
3.5.2 General Measures for Leakage Suspicion
Evacuate personnel to a safe area. Appropriate self-contained breathing apparatus may be
required. Approach suspected leak area with caution. Remove all sources of ignition. if safe to
do so. Reduce gas with fog or fine water spray. Stop flow of product if safe to do so. Ventilate
area or move container to a well-ventilated area. Flammable gas may spread from leak. Before
entering the area, especially a confined area, check the atmosphere with an appropriate device.
3.5.3 First Aid Measure
First-aid measures after inhalation: Remove victim to uncontaminated area wearing self-
contained breathing apparatus. Keep victim warm and rested. Call a doctor. Apply artificial
respiration if breathing stops.

50
First-aid measures after skin contact: Adverse effects not expected from this product.
First-aid measures after eye contact: Immediately flush eyes thoroughly with water for at
least 15 minutes. Hold the eyelids open and away from the eyeballs to ensure that all surfaces are
flushed thoroughly. Get immediate medical attention.
First-aid measures after ingestion: Ingestion is not considered a potential route of exposure.
Suitable extinguishing media: Carbon dioxide, dry chemical powder, water spray, fog.
3.5.4 Fire Hazard
It is an extremely flammable gas. The hydrogen flame is nearly invisible. Hydrogen has a low
ignition energy; escaping hydrogen gas may ignite spontaneously. A fireball forms if the gas
cloud ignites immediately after release. Hydrogen forms explosive mixtures with air and
oxidizing agents.
Explosion hazard: Extremely explosive gas forms explosive mixtures with air and oxidizing
agents.
Reactivity: No reactivity hazard other than the effects described in sub-sections below.
3.5.5 Firefighting Instructions
If venting or leaking gas catches fire, do not extinguish flames. Flammable vapors may spread
from leak, creating an explosive re-ignition hazard. Vapors can be ignited by pilot lights, other
flames, smoking, sparks, heaters, electrical equipment, static discharge, or other ignition sources
at locations distant from product handling point. Explosive atmospheres may linger.

51
Before entering an area, especially a confined area, check the atmosphere with an appropriate
device. Evacuate all personnel from the danger area. Use self-contained breathing apparatus
(SCBA) and protective clothing. Immediately cool containers with water from maximum
distance. Stop flow of gas if safe to do so, while continuing cooling water spray. Remove
ignition sources if safe to do so. Remove containers from area of fire if safe to do so. On-site fire
brigades must comply with OSHA 29 CFR 1910.156 and applicable standards under 29 CFR
1910 Subpart L—Fire Protection.
Protection during fire-fighting: Compressed gas: asphyxiant. Suffocation hazard by lack of
oxygen.
Special protective equipment for fire fighters: Standard protective clothing and equipment
(Self Contained Breathing Apparatus) for fire fighters.
3.5.6 Other Precautions for Handling, Storage, and Use
When handling product under pressure, use piping and equipment adequately designed to
withstand the pressures to be encountered. Never work on a pressurized system. Use a back flow
preventive device in the piping. Gases can cause rapid suffocation because of oxygen deficiency;
store and use with adequate ventilation. If a leak occurs, close the container valve and blow
down the system in a safe and environmentally correct manner in compliance with all
international, federal/national, state/provincial, and local laws; then repair the leak. Never place a
container where it may become part of an electrical circuit.

52
3.5.7 Exposure control
Appropriate engineering controls: An explosion-proof local exhaust system is acceptable. Local
exhaust and general ventilation must be adequate to meet exposure standards. Mechanic
(general) engineering controls: Use only in a closed system.
Eye protection: Wear safety glasses with side shields.
Respiratory protection: An air-supplied respirator must be used while working with this
product in confined spaces. The respiratory protection used must conform with OSHA rules as
specified in 29 CFR 1910.134.
Thermal hazard protection: None necessary.
Other information: Consider the use of flame resistant anti-static safety clothing. Wear safety
shoes while handling containers.
Hydrogen detector should be placed on the ceiling of the storage and production room and good
alarm system should be put in place.
Storage conditions: Store only where temperature will not exceed 125°F (52°C). Physical
storage of compressed hydrogen gas in high pressure tanks (up to 700 bar). Post ―No Smoking or
Open Flames‖ signs in storage and use areas. There must be no sources of ignition. Separate
packages and protect against potential fire and/or explosion damage following appropriate codes
and requirements (e.g., NFPA 30, NFPA 55, NFPA 70, and/or NFPA 221 in the U.S.) or
according to requirements determined by the Authority Having Jurisdiction (AHJ). Always
secure containers upright to keep them from falling or being knocked over. Install valve
protection cap, if provided, firmly in place by hand when the container is not in use. Store full

53
and empty containers separately. Use a first-in, first-out inventory system to prevent storing full
containers for long periods.

54
CHAPTER FOUR
CONCLUSION
From the results of the project, the following conclusion can be made:
I. The hydrogen to be produced according to the design work carried out will be of
approximately 96.6% purity which exceeds the production benchmark of 95% purity for
hydrogen product stream.
II. In order to achieve the production target of 20 million standard cubic feet of Hydrogen
per year at 96.6% purity, production capacity of 4.041x108
m3
/hr at the rate of
4.041x108
m3
/hr. of hydrogen, and oxygen (at 95% purity) required.
III. The amount of CO2 converted is 238,414,162 m3
and the energy requirement of the plant
is estimated to 4.537x 1011
kJ
It should be noted that the above figures are based on a schedule of 7680 operating hours per
year as given in the problem statement.

55
REFERENCES
1. A. Fuderer, ―Catalytic Steam Reforming of Hydrocarbons,‖ U.S. patent 4,337,170, June
29, 1982.
2. G. Q. Miller and J. Stoecker, ―Selection of a Hydrogen Separation Process,‖ NPRA
Annual Meeting, San Francisco, March 1989.
3. Kamm, G.R., Tanaami, K. ―Integrated process for the partial oxidation-thermal cracking
of crude oil feedstocks‖, United States Patent 4134824, January 16, 1979.
4. McHugh, K. (2005). ―Hydrogen Production Methods‖, Prepared for MPR Assiociates
Inc., MPR-WP-0001.
5. M. V. Twigg, Catalyst Handbook, 2d ed., Wolfe Publishing, London, July 14, 1989.
6. Osafehinti, E.K. (2008). ―Hydrogen production using Shell Gasification Process (SGP)
Scheme of heavy fuel oil.‖ Unpublished Process Design Project, Dept. of Chemical Engr.
OAU, Ile-Ife.

56
APPENDIX I
MATERIAL BALANCE
On-Stream Time: 7680hr/yr. = 18.83km3
/hr.
Production Capacity: 0.515 x l06
m3
/day = 0.02146x106
m3
/hr.
Production Rate: 0.02146x106
x18830
= 4.041x108
m3
/hr.
CO: 1.0% x 404091800 = 4040918
CO2: 1.0% x 404091800 = 4040918
N2: 2% x 404091800 = 8081836
CO IN  CO OUT 
Let A = Volume of Crude Gas Out of Quencher
Using Tie Component Nitrogen
Nitrogen out of Quencher = Nitrogen in Product
0.014A = 8081836
A = 577274000m3
CO out of Quencher = 0.42x577274000
= 242,455080m3
CO2 out of Quencher = 0.083x 577274000
= 47913742m3
CO out of Converter = CO in Product (no CO removed)
= 4040918m3
Amount of CO converted = CO into Converter – CO out of Converter
CONVERTER

57
= 242,455080 - 4040918
= 238,414,162 m3
The Gas Stream Mixture is 370C= 643K
By Interpolation, we have
Te (K) Kp
600 3.69*10-2
700 1.11*10-2
800 2.48*10-1
(643 – 600)(700-600) = (x-3.69x10-2
)/(1.11x10-2
- 3.69x10-2
)
X = Kp = 2.581x10-2
H2 OUT of quencher= H2 into converter = 0.47 x577274000 m3
= 271318780 m3
CH4 OUT of quencher = 1 x 10-3
x577274000
= 57727m3
H2 OUT of converter = 0.95 X4.041*10^8 m3
=383,895,000m3
H2 produced in converter = (4.041-2.712) x 108
m3
= 1.329 x 108
m3
CO + H2O  H2 + CO2
Amount of CO2 converted = 238,414,162 m3
Kp = 2.5806x10-2
Kp = (PCO x PH2O) / (PCO2 x PH2)
2.5806x10-2
= (238,414,162 m3
x PH2O) / (238,414,162 x1.329 x 108
)
PH2O = 3429,617.4 m3

58
APPENDIX II
Mechanical Design of the Absorber
Using a Cylindrical packed column with pall ring packing and tori-spherical head:
On-Stream Time: 7680hr/yr. = 18.83km3
/hr.
Production Capacity: 0.515 x l06
m3
/day = 0.02146 x 106
m3
/hr.
Production Rate: 0.02146x106
x18830
= 4.041x108
m3
/hr.
Recall,
CO2 out of Quencher = 0.083x 577274000
= 47913742m3
Amount of CO2 converted = 238,414,162 m3
Hence,
CO2 entering absorber = 47913742m3
+ 238,414,162 m3
= 48152156.16m3
Rate of CO2 entering absorber = 13375.60m3
/hr
CO2 leaving absorber = 8084.40m3
/hr
Total amount of gas leaving the absorber = 808440m3
/hr
Total amount of gas entering the absorber = 808440 + (13375.6-8084.40)
= 813,731.2m3
/hr
Feed Concentration = 13375.60/813,731.2
= 0.0164
Y1 = 0.0164/(1-0.0164)
= 0.0167
Feed gas rate on solute basis = 813,731.2x(1-0.0164)
= 800,386.01m3
/hr

59
Y2 = 8084.40/800,386.01 = 0.0101
Equation of equilibrium line
Y=1.32X
X1 = Y1/1.32 = 0.0127
Now the material balance equation is GS (Y1-Y2) = LS(X1-X2)
800,386.01 (0.0167-0.0101) =LS (0.0127-0)
LS=415,948.64mol/hr
Actual liquid rate is 1.25 times the original, LS= 1.25 x 24212
= 42818.45m3
/hr
The liquid at the bottom of the tower (L1) = LS + CO2 absorbed
= 42818.45 + (13375.60-8084.40)
= 48109.65m3
/hr
= 2579.76kmol/hr
Calculation of tray or column diameter
Data given:- ρG= 1.98 kg/m3
ρL= 2430 kg/m3
Now flow parameter (FLV) = (L1/G1) x (ρG/ ρL)0.5
= (2579.76/1859.34) x (1.98/2430)0.5
= 0.0396unit
Now calculate CSB= 0.23 from graph
Superficial velocity Usfl= CSB x [( ρL- ρG)/ρG]0.5 x (σ/20)0.2 ft/sec
= 0.23x1.004x19.19
= 4.5 ft/sec
Taking the operating velocity as 70% of the flooding velocity
US= 4.5 x 0.70
= 3.15 ft/sec
Active tray area (Aa) = volumetric flow rate of gas/ operating velocity
= 800,386.01 / (0.96 x 3600)
= 231.593m2
Tower cross section (AT) = Aa/(1-fd)
= 231.593/(1-0.2)
= 289.49m2
Tower diameter = (289.49x4/3.44)0.5

60
=18.3m
Tower height= 18.3/0.2756
=66.57m
=66.6m
From Coulson and Richardson Vol6
The minimum thickness for absorber of diameter 3.0 -3.5 is 12mm
Therefore the design thickness is taken to be 12mm
Based on the above values:
The operating temperature and pressure are obtained from Niaz Bahar et al (2013)
Design temperature and pressure is 10% higher than operating temperature and pressure
(Coulson and Richardson Vol6)
Operating Temperature: 83o
C
Design Temperature: 91.3o
C
Operating Pressure: 28.5 bar
Design Pressure: 31.35 bar
Column diameter:18.3m
Column Height: 66.6m
Shell thickness (using Coulson and Richardson, volume 6 as reference) was obtained to be
12mm+2mm corrosion allowance.
Shell shape: cylindrical
End plate Details: Hemispherical end is recommended considering the design pressure.
Thickness of end can be taken as 12mm+2mm corrosion allowance.
Material of construction: carbon steel
Mechanism of Liquid distribution: Liquid sprayer
For the shell:
Diameter of the column =18.3m
Height of the column= 66.6m
Outside pressure is 1atm= 0.1 MN/m2
Allowable stress f= 98.1 MN/m2
Welding joint efficiency j= 0.85
Modulus of elasticity E= 1 x 105
MN/m2

61
for the catalytic converter
Outside pressure is 1atm= 0.1 MN/m2
Design pressure Pd= 0.1x1.05
= 0.105 MN/m2
Shell is I.S 2825-1969
Allowable stress f= 98.1 MN/m2
Welding joint efficiency j= 0.85

62
APPENDIX III
ENERGY-BALANCE DIAGRAM FOR THE FLAME REACTOR AND FOR THE
ASSOCIATED WASTE-HEAT BOILER
Nitrogen balance (Tie component balance)
Nitrogen out of combustor = nitrogen in product
0.014 A = 0.02 x 0.02146x10^6
Volume of crude gas leaving the Combustor, A = 30,657.143m3
NITROGEN INTO REACTOR = NITROGEN OUT OF REACTOR
0.05B = 0.014 x 30,657.143
B = 8584.00m3
/hr. (feed rate of oxygen into reactor)
Density of oxygen = 1429Kg/m3
Mass of oxygen = Volume of oxygen x density of oxygen
= 8584.0 x1429 = 12,266.536kg
1kg of heavy oil is combusted against 1.16kg of O2
Mass of heavy oil feed = (mass of oxygen x1 / 1.16)
= 10,574.6 kg

63
Energy released by the combustor = Calorific value x mass of feedstock
= 42.9 x106
x10, 574.6
= 4.537x 1011
J
Energy of O2 entering the combustor = MCT
C= 0.910kJ/kg.K at 2100
C, T= 2100
C =(210+273)K =483K
Energy of O2 entering the combustor = 8.648 x1010
x0.910x483 Kj
= 3.801 x 1013
kJ
Steam produced = 0.75kg/kg of heavy oil feedstock
Mass of steam produced by waste heat boiler = 0.75 x mass of oil
= 0.75 x 10,574.6 kg
= 7930.95kg
Energy used to produce steam by waste heat boiler = MC T + ML
= 7930.95 x (4.2x (100-25) +2500)
= 22,325,624.25 kJ
Energy of crude gas leaving WHB= Energy of Crude gas entering the WHB-Energy
used to produce steam.
= 4.537x 1011
kJ - 22,325,624.25 kJ
= 4.537x 1011
kJ

64
APPENDIX IV
DEPARTMENT OF CHEMICAL ENGINEERING OBAFEMI AWOLOWO UNIVERSITY, ILE-IFE
CHE 505 – PROCESS DESIGN
October 24, 2014 Prof. Funso Akeredolu & Dr. O.J. Odejobi
INSTRUCTIONS
This problem will be treated as test of your ability to tackle a practical problem in the same way
as might be expected if you were required to report as chemical engineer on a manufacturing
proposal.
PLEASE BE NEAT AND SHOW ALL YOUR CALCULATIONS. THE ANSWERS TO THE
QUESTIONS SHOULD BE DERIVED BY APPLICATION OF FUNDAMENTAL
PRINCIPLES TO AVAILABLE PUBLISHED DATA. PARTICULAR CREDIT WILL BE
GIVEN FOR CONCISE ANSWERS TYPED DOUBLE-SPACED ON A4 PAPER.
Your graphs and drawings must be neat and initialed and dated by you. Your calculations should
be done in SI units. References must be given in details to all sources of published information
consulted. Reporting schedules are listed in this document. Accordingly, your final BOUND
report is due NOT LATER THAN March 23, 2015. The oral defense will hold between April 13
and 15, 2015.
The Project
A plant is to be designed to produce 20 million standard cubic feet per day (0.515 x l06
standard
m3
/day) of hydrogen of at least 95 per cent purity. The process to he employed is the partial
oxidation of oil feedstock (refs 1 to 3).
Materials Available
(1) Heavy fuel oil feedstock of viscosity 900s Redwood One (2.57 x 10-4
m2
/ s) at 1000
F with the
following analysis:
Carbon 85% wt
Hydrogen 11% wt
Sulphur 40 %wt
Calorific value 18,410 Btu/lb (42.9 MJ/kg)
Specific gravity 0.9435
The oil available is pumped from tankage at a pressure of 30 psig (206.9 kN m2
gauge) and at
500
C.

65
(2) Oxygen at 95 per cent purity (the other component assumed to be wholly nitrogen) and at
200
C and 600 psig (4140 kN/m2
gauge).
Services Available
(1) Steam at 600 psig (4140 kN/m2
gauge) saturated
(2) Cooling water at a maximum summer temperature of 250
C
(3) Demineralized boiler feed water at 20 psig (138 kN/m2
gauge) and 150
C suitable for direct
feed to the boilers
(4) Electricity at 440 V, three phase, 50 Hz, with adequate incoming cable capacity for all
proposed uses
(5) Waste low-pressure steam from an adjacent process.
On-Stream Time
See Table 1 for assigned individual value (hours per year).
Product Specification
Gaseous hydrogen with the following limits of purities:
CO 1.0% vol maximum (dry basis)
CO2 1.0% vol maximum (dry basis)
N2 2.0% vol maxi mum (dry basis)
CH4 1.0% vol maximum (dry basis)
H2S Less than1 ppm
The gas is to be delivered at 350
C maximum temperature, and at a pressure not less than 300 psig
(2060 kN/m2
gauge). The gas can be delivered saturated, i.e., no drying plant is required.
The Process
Heavy fuel oil feedstock is delivered into the suction of metering-type ram pumps which
feed it via a steam preheater into the combustor of a refractory-lined flame reactor. The feedstock
must be heated to 2000
C in the preheater to ensure efficient atomization in the combustor. A
mixture of oxygen and steam is also fed to the combustor, the oxygen being preheated in a
separate steam preheater to 2100
C before being mixed with the reactant steam.
The crude gas, which will contain some carbon particles, leaves the reactor at
approximately 1300 C and passes immediately into a special waste-heat boiler where steam at
600 psig (4140 kN/m2
gauge) is generated. The crude gas leaves the waste heat boiler at 2500
C
and is further cooled to 500
C by direct quenching with water, which also serves to remove the
carbon as a suspension. The analysis of the quenched crude gas is as follows:

66
H2 47.6 % vol (dry basis)
CO 42. 1 % vol (dry basis)
CO2 8.3 % vol (dry basis)
CH4 0. 1 % vol (dry basis)
H2S 0.5 % vol (dry basis)
N2 1.40 % vol (dry basis)
100.0 % vol (dry basis)
For the primary flame reaction steam and oxygen are fed to the reactor at the following rates:
Steam 0.75 kg/kg of heavy fuel oil feedstock
Oxygen 1. 1 6 kg/kg of heavy fuel oil feedstock
The carbon produced in the flame reaction, and which is subsequently removed as carbon
suspension in water, amounts to 1.5% by weight of the fuel oil feedstock charge. Some H2S
present in the crude gas is removed by contact with the quench water.
The quenched gas passes to an H2S removal stage where it may be assumed that H2S is
selectively scrubbed down to 15 parts per million with substantially nil removal of CO2. Solution
regeneration in this process is undertaken using the waste low-pressure steam from another
process. The scrubbed gas, at 350
C and saturated, has then to undergo CO conversion, final H2S
removal, and CO2 removal to allow it to meet the product specification.
CO conversion is carried out over chromium-promoted iron oxide catalyst employing two
stages of catalyst conversion; the plant also incorporates a saturator and desaturator operating
with a hot water circuit.
Incoming gas is introduced into the saturator (a packed column) where it is contacted
with hot water pumped from the base of the desaturator; this process serves to preheat the gas
and to introduce into it some of the water vapour required as reactant. The gas then passes to two
heat exchangers in series. In the first, the unconverted gas is heated against the converted gas
from the second stage of catalyst conversion; in the second heat exchanger the unconverted gas
is further heated against the converted gas from the first stage of catalytic conversion. The
remaining water required as reactant is then introduced into the unconverted gas as steam at 600
psig (4140 kN/m2
gauge) saturated and the gas steam mixture passes to the catalyst vessel at a
temperature of 3700
C. The catalyst vessel is a single shell with a dividing plate separating the

67
two catalyst beds which constitute the two stages of conversion. The converted gas from each
stage passes to the heat exchangers previously described and thence to the desaturator, which is a
further packed column. In this column the converted gas is contacted countercurrent with hot
water pumped from the saturator base; the temperature of the gas is reduced and the deposited
water is absorbed in the hot-water circuit. An air-cooled heat exchanger then reduces the
temperature of the converted gas to 400
C for final H2S removal.
Final H2S removal takes place in four vertical vessels each approximately 60 feet (18.3
m) in height and 8 feet (2.4 m) in diameter and equipped with five trays of iron oxide absorbent.
Each vessel is provided with a locking lid of the autoclave type. The total pressure drop across
these vessels is 5 psi (35 kN/m2
). Gas leaving this section of the plant contains less than 1 ppm of
H2S and passes to the CO2 removal stage at a temperature of 350
C.
CO2 removal is accomplished employing high-pressure potassium carbonate wash with
solution regeneration (ref. 4).
Data
I. Basic Data for CO Conversion Section of the Pant
(a) Space Velocity
The space velocity through each catalyst stage should he assumed to be 3500 volumes of gas
plus steam measured at NTP per volume of catalyst per hour. It should further he assumed that
use of this space velocity will allow a 100
C approach to equilibrium to be attained throughout the
possible range of catalyst operating temperatures listed below.
(b) Equilibrium Data for the CO Conversion Reaction
For
Kp = PCO x PH2O
PCO2 x PH2
Temperature (K) Kp
600 3.69 x 10-2
700 1.11 x 10-2
800 2.48 x 10-1
(c) Heat of Reaction
CO + H2O == CO + H2 ∆H = -9.84 kcal
II. Basic data for CO2 Removal Using Hot Potassium Carbonate Solutions

68
The data presented in ref. 4 should be employed in the design of the CO2 removal section of the
plant. A solution concentration of 40%wt equivalent K2 CO3 should he employed.
Preliminary Literature Search on Hydrogen
Review the literature on the current and future industrial uses, production, storage, infrastructure
and market for hydrogen. The report must be concise and provide a justification for the design
project about to be done by you.
Scope of Design Work Required
1. Process Design
(a) Calculate, and prepare a diagram to show, the gas flows, compositions, pressures and
temperatures, at each main stage throughout the processes of gasification and
purification.
(b) Prepare a mass-balance diagram for the CO conversion section of the plant including
the live steam addition to the unconverted gas. Basic data that should be employed for the
CO conversion process are presented in Coulson and Richardson Volume 6 Appendices
(c) Prepare an energy-balance diagram for the flame reactor and for the associated waste-heat
boiler. Is the energy use optimal according to your study of Pinch Technology?
(d) Prepare a process flow diagram showing all major items of equipment. This should be done
with the aid of a named commercial process engineering simulation package. Your input
should give an indication of the internal construction of each item (with the exception of the
flame reactor, waste-heat boiler and quench tower). The primary H2S removal stage need not
be detailed.
(e) Prepare an equipment schedule for the CO conversion section of the plant, specifying major
items of equipment.
2. Chemical Engineering Design
(a) Prepare a detailed chemical engineering design of the absorber in the CO removal stage.
(b) Prepare a chemical engineering design for the saturator in the CO conversion Section.
3. Mechanical Design
Make recommendations for the mechanical design of the CO removal absorber, estimating the
shell and end-plate thickness and showing, by means of sketches suitable for submission to a
design office, how:
(a) The beds of tower packing are supported
(b) The liquid is distributed.
Develop a detailed mechanical design of the CO conversion reactor, paying particular
attention to the choice of alloy steels versus refractory linings, provisions for thermal expansion,
inlet gas distribution, catalyst bed-support design, facilities for charging and discharging catalyst
and provisions for instrumentation.
4. Control
Prepare a full instrumentation flow-sheet of the CO conversion section of the plant,

69
paying particular attention to the methods of controlling liquid levels in the circulating water
system and temperatures in the catalyst beds. Derive the unsteady-state equations which would
have to he employed in the application of computer control to the CO conversion section of the
plant.
5. Health, Safety and Environmental Considerations/Project Costing
Prepare a full safety brief on the materials involved in the process being designed, Comment
specifically on the hydrogen storage method you would recommend for adoption.
Prepare an approximate cost of the project stating your assumptions clearly
References
1. J. H. GARVIE, Chem. Proc. Engng, Nov. 1967, pp. 55 65. Synthesis gas manufacture.
2. Hydrocarbon Processing Refining Processes Handbook. Issue A, Sept. 1970, p. 269.
3. S. C. SINGER and L. W. TER HAAR, Chem. Eng Prog., 1961, 57, pp. 68 74.
Reducing gases by partial oxidation of hydrocarbons.
4. H. E. BENSON, J. H. FIELD and W. P. HAYNES, Chem. Eng Prog., 1956, 52,
pp. 433 438. Improved process for CO2 absorption uses hot carbonate solutions.

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