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5/6/2016
EBTAX: The Conversion
of Ethane to Aromatics
via Catalytic Conversion
Saud Alshahri, Aaron Cheese, Bridger Martin, Emily Schwichtenberg CHE 4080 PROCESS DESIGN II
1
Table of Contents I. Table of Tables..................................................................................................................................... 2
II. Table of Figures ................................................................................................................................. 4
III. Executive Summary (Emily)............................................................................................................ 6
IV. Scope of Work (Bridger) .................................................................................................................. 8
V. Introduction (Saud) .......................................................................................................................... 10
VI. Description of Base Case ................................................................................................................ 11
Section 100: Feed processing, reaction, and initial separation (Bridger) ....................................... 14
Section 200: Lights Separation Section (Emily) ............................................................................... 23
Section 300: Separation and Recovery of BTX and Heavy Aromatics Products (Saud) ................. 29
Section 400/500: Propane and Ethylene Refrigeration (Aaron) ...................................................... 33
VII. Design Alternatives (Bridger) ...................................................................................................... 42
Possible Reactor Alterations .............................................................................................................. 42
Product Recovery (Section 300) alternative designs ......................................................................... 43
Continuous Catalyst Regeneration .................................................................................................... 45
Fuel gas reallocation, C2 through C4 repurposing .......................................................................... 46
VIII. Permitting and Environmental Concerns (Emily) ................................................................... 47
IX. Safety and Risk Management (Emily) .......................................................................................... 51
X. Project Economics (Aaron) ............................................................................................................. 53
Equipment and Capital Cost .............................................................................................................. 53
Pricing, Revenue and Production Cost ............................................................................................. 61
Cash Flow Analysis ............................................................................................................................ 63
Sensitivities ......................................................................................................................................... 64
XI. Global Impacts (Saud) ................................................................................................................... 67
XII. Conclusions and Recommendations (Bridger) ........................................................................... 71
XIII. Future Work (Aaron) .................................................................................................................. 73
XIV. Acknowledgements (Saud) .......................................................................................................... 74
XV. References ...................................................................................................................................... 75
XVI. Appendices ................................................................................................................................... 78
2
I. Table of Tables
Table 1: List of reactions, their respective conversions, and the heats of reaction/ ................................... 18
Table 2: Pressure, temperature and enthalpy data for propane .................................................................. 37
Table 3: Pressure, temperature and enthalpy data for ethylene ............................................................... 38
Table 4: Various operating conditions specified for the different refrigeration cycles involved along with
the Aspen unit operation or stream that it corresponds to. .......................................................................... 40
Table 5: This table shows the emissions of thermal NOx with and without control measures. The emission
limit for needing a permit from the EPA is 100 tons/year, which can be obtained with control measures in
this process. ................................................................................................................................................. 49
Table 6: Specific information involved in sizing and costing compressors. .............................................. 54
Table 7: Specific information involved in sizing and costing turbines ...................................................... 54
Table 8: Specific information involved in sizing and costing furnaces ..................................................... 55
Table 9: Specific information involved in sizing and costing heat exchangers ......................................... 56
Table 10: Specific information involved in sizing and costing air coolers ............................................... 57
Table 11: Specific information involved in sizing and costing vessels ...................................................... 58
Table 12: Specific information for costing the PSA unit ........................................................................... 58
Table 13: Specific information for sizing and costing the amount of catalyst used ................................... 59
Table 14: Specific information for sizing and costing the amount of catalyst used ................................... 60
Table 15: Specific information for distillation column and tray sizing and costing .................................. 60
3
Table 16: Fixed Capital Investment for the various equipment involved in the process along with the
resulting total .............................................................................................................................................. 61
Table 17: Income or cost of each of the materials consumed or produced ................................................ 62
Table 18: Cost of utilities ........................................................................................................................... 63
Table 19: Various fixed costs associated with the design .......................................................................... 63
Table 20: Results of the cash flow analysis conducted on this design ....................................................... 64
Table 21: Sensitivities run, along with the resulting IRR .......................................................................... 65
4
II. Table of Figures
Figure 1: Overall Process Flow Diagram. Part A: Section 100, 200, and 300 up to M301. Part B:
Remaining portion of section 300. Part C: Section 400 and 500. ........................................................... 13
Figure 2 : Feed and reactor (Section 100). Feed ethane is mixed with hydrocarbon and hydrogen
recycles. The presence of hydrogen significantly reduces catalyst coking. The flash tank, D101, sunders
gaseous C1-C5 to section 200 and C6-C9 to section 300. .......................................................................... 14
Figure 3: Section 200, the Lights Separation Section. This section consists of a mixer to combine the
vapor stream from the flash drum and a recycle from the product recovery section, two distillation towers
and a pressure swing adsorption (PSA) unit to separate the product and recycle streams, one splitter to
allow for the hydrogen sale stream to be separated, and two compressors to pressurize the recycle streams
to appropriate pressures to be mixed with the feed stream ......................................................................... 23
Figure 4: Section 200 Up To T201. From the flash drum, the vapor stream is mixed in M201 with a
vapor recovery stream in the product recovery section before being sent to a distillation tower (T201) to
remove hydrogen and methane from the product stream as the vapor distillate (S203), with the remainder
exiting the tower in the bottoms stream (S206). ......................................................................................... 24
Figure 5: Section 200, T201 Distillate Path After T201. The hydrogen and methane stream is sent to a
heat exchanger to warm it up to room temperature before being sent to the pressure swing adsorption
(PSA) unit, where the methane is removed to a fuel gas stream and the hydrogen stream is sent to a
splitter (S201). This splits the hydrogen stream into a sale product stream and a recycle stream, which will
be sent compressed and sent back to the reactor to prevent coking of the catalyst..................................... 25
Figure 6: Section 200, T201 Bottoms Path. The bottoms of T201 is sent to T202, where C2 and C3
hydrocarbons are distilled off and sent to a compressor before being recycled back to the recycled to
increase the overall conversion of the reactor. The bottoms stream is sent to a mixer in the product
recovery section to recover any BTX products that could have been lost .................................................. 27
Figure 7: Heavy Separation (Section 300). From flash tank D101, the heavy stream is separated
remaining light hydrocarbons. The remaining heavies are separated into Benzene, Toluene, and Xylene
product and TMB byproduct. ...................................................................................................................... 29
Figure 8: Heavies Separation (Section 300). From flash tank D101, the liquid stream is separated fed
into the first distillation column (T301) to remove the remaining light hyrdrocarbons as well as recover
TMB as a product. The aromatic rich stream is sent on for further processing by S302. .......................... 30
5
Figure 9: Purge Stream. BTX rich streams are fed into tower T302, one of which comes from the lights
separation section, and one of which come from the previous tower, T301. T302 separates out any
remaining lights and purges them from the system. The bottoms of the tower is sent on for product
recovery as it mainly consists of BTX. ....................................................................................................... 31
Figure 10: Benzene Recovery. Benzene is recovered from the BTX rich stream leaving T302. The
bottoms of the tower is sent on to recover the remaining Toluene and Xylene. ......................................... 32
Figure 11: Toluene and Xylene Recovery. T304 separates toluene from para-xylene that is fed to the
tower from T303. ........................................................................................................................................ 33
Figure 12: Propane and Ethylene Refrigeration. Section 400 consists of two propane refrigeration cycles
operating at different pressures. Section 500 consists of only one ethylene refrigeration cycle. For each
cycle the refrigerant is compressed, condensed, expanded, and evaporated in order to complete the
cycle. Propane refrigeration is used in condensing the process fluid in T202 along with condensing the
ethylene in section 500. Ethylene refrigeration is only required to condense the process fluid in T201. 34
Figure 13: Tornado Diagram. This plot chose the change on IRR based on different variation of various
uncertain parameters. .................................................................................................................................. 66
Figure 14: Industry Rivalry. This figure illustrates the possible industry pressure associated with a new
competitor ................................................................................................................................................... 68
6
III. Executive Summary (Emily)
Team EBTAX was formed with the goal of taking ethane from natural gas refineries and
processing it into various aromatics, including benzene, toluene, and para-xylene (BTX). The market for
BTX chemicals is fairly stable, as they can be converted into larger molecules that are critical components
in polymer synthesis, producing plastics, textiles, and other consumer goods. The glut of natural gas in
the US has caused ethane prices to drop to around half of what they were at the beginning of 2014. This
makes it an ideal feedstock for our process, which includes a catalyzed reaction and several separation
units to separate the reaction products into pure component products. The plant will be located on an
existing oil refinery in the gulf coast area. This will provide easy allocation of products, as well as access
to the oil refineries and chemical plants that would purchase and further process the products.
To catalyze the reactor, a platinum-zeolite catalyst was chosen for its high selectivity toward
BTX compared to similar catalysts. US Patent 7745675 B2 only provides conversion and reaction
information for this catalyst from lab scale tests. The unfamiliarity with this catalyst and the lack of
information at diverse conditions led to several assumptions when modeling the process. Firstly, the
amount of catalyst required for the process scales linearly and ideally with the reactor inlet. Secondly, the
conversion of the reactions would not depend strongly on pressure. The patent also provides a functional
pressure for the catalyst specified from 20 to 2000 psia without providing correlations for pressure and
conversion. Catalyst lifetime use is assumed to outlive the life of the plant, and regeneration is assumed to
recover 100% of the catalyst. Regeneration alone will keep the catalyst active for the lifetime of the
project.
The process is broken down into five sections. To begin, the reactor section (Section 100),
consists of an ethane feed stream and two recycles from the separation sections. The feed is mixed with
the recycles and then heated before entering a gas-phase, fixed-bed, catalytic reactor at 1150°F. The
reactor houses 26 separate reactions. C2 and C3 hydrocarbons (HCs) undergo multiple equilibrium-based
reactions to form C1 through C5 linear and C6 through C9 aromatic HCs and hydrogen. After the stream
7
has reacted, it is cooled before being sent to the separation sections. An initial flash tank separates the
reactor effluent into light and heavy HC streams, which are sent to sections 200 and 300, respectively. In
the lights separation, two distillation columns and a hydrogen pressure swing adsorption (PSA) unit are
used. The four product streams from the light separation section are high purity hydrogen (99.5 mol%),
methane fuel gas, and a C2-C3 HC recycle. The hydrogen stream will be split into a sale stream and a
recycle stream, which will prevent coking of the catalyst. Sections 400 and 500 are propane and ethylene
refrigeration, respectively, which will be used to cool the condensers in the lights separation to allow the
small hydrocarbons to condense. The liquids from the initial flash tank are pumped to the product
separation (Section 300). Four distillation towers are used in this section. The first tower (T301) functions
as a light HC recovery unit, with the remaining light vapors entering section 200. The liquids enter the
next distillation tower (T302) to remove TMB. The last two towers separate the BTX into its components
to be sold.
The fixed capital investment for this project is $320 million, which is largely due to the many
compressors needed for the refrigeration section. With a 20 year project life, this project yields an IRR of
30.3%, with a payback period of 2.6 years.
8
IV. Scope of Work (Bridger)
The preliminary mission on team EBTAX is to design an industrial plant that takes advantage of
the abundant natural gas supplies in order to produce benzene, toluene, and xylene. Research into the
natural gas industry, along with common natural gas refinery processes, revealed ethane to be the most
probable feedstock. This was due to a number of contributing factors, including the recent practice of
ethane rejection, where ethane is allowed to flow with methane into the pipe gas stream. Ethane rejection
further lowers the cost of inexpensive ethane feeds.
US Patent US20130324778A1 was provided to team EBTAX as a starting point for the
conversion of ethane into valuable aromatics, including benzene, toluene, and xylene. This catalyst
became the basis for the plant design and created many constraints that had to be met by the plant.
Constraints
The primary design constraints are equilibrium constraints within the reactor. The catalyst
conversion is highly specific and creates the high volume reflux of C2 and C3 hydrocarbons. The catalyst
operating conditions also set the reactor temperature at 1150°F. These high temperatures can also ignite
the HCs if oxygen is present in the system. All process streams are run above atmospheric to prevent
oxygen from entering the system in the case of a leak.
The reactions taking place are exothermic. This creates the opportunity for a runaway reaction if
released heat and built up pressures go beyond controllable conditions. This creates additional safety
constraints, such as the inclusion of cooling systems, which will also be considered in more detailed
designs.
Thermodynamic constraints are present outside the reactor as well. To separate the C1-C9 HC
stream, extremely low temperatures are required for the lighter components. Two types of refrigeration
were included to reach the low temperatures. These refrigeration systems will be discussed in the Section
400/500 below.
9
Climate restrictions were also taken into account. The Gulf Coast air and water temperatures will
be higher than temperatures elsewhere in the country, particularly during the summer. The water
temperature was estimated at 105°F and the air temperature was estimated at 95°F.
The majority of other operating constraints are the purities of the products we intend to sell. If
hydrogen, benzene, toluene, or p-xylene are not at the correct purity, they will not be able to be sold at as
high of a price. These could also be considered economic constraints since they directly affect the
economic income of the plant.
By incorporating the above design requirements, the EBTAX team created a preliminary design
for a profitable, industrial scale plant. Team EBTAX verified the plant design using Aspen+ modeling
design. Capital costing was primarily performed by hand using cost graphs in Peters and Timmerhaus.
Compiled economic analyses were performed via Microsoft Excel to very that the plant design is
profitable. A full scale industrial plant was designed to fulfill these requirements from the initial design
concept, through the preliminary design analysis.
10
V. Introduction (Saud)
Recently, there has been a glut of natural gas in the United States. This has driven the price of
natural gas, including ethane, to almost half of what it was at the beginning of 2014. This remarkably low
cost of ethane has led to natural gas refineries no longer separating ethane from the natural gas, and
simply rejecting the ethane to pipeline natural gas that gets sent to residential homes. While the natural
gas plant does save money from rejecting ethane, there is significant potential loss and waste considering
ethane can be converted to valuable BTX products. Due to the large excess of natural gas liquids in the
United States, ethane prices are currently very low. From this information, team EBTAX was charged
with the task of researching and designing a way to capitalize on this low-priced feedstock. A recent
patent, US8772563 describes a platinum-zeolite catalyst which converts ethane into valuable aromatics,
primarily benzene, toluene, and para-xylenes (BTX). The plant is based on this catalyst, and is still in the
design phase, with the goal of producing 700 MMlb/yr of the benzene, toluene, and para-xylene products.
Some assumptions were needed to assist in modeling our catalyst and the plant. The catalyst was
chosen to achieve the highest selectivity of BTX product. This conversion data has only been tested at the
lab scale. US Patent 20130324778A1 stated that the catalyst could be used with pressures ranging from
20 to 2000 psia. The correlation between pressure and conversion will need to be determined
experimentally because to date, no data has been recorded for this catalyst. Although the single
regeneration life of the catalyst is unknown, similar catalysts have been shown to need regeneration every
one to six months. A continuous regeneration process will eliminate the plant shutting down due to time
needed for catalyst regeneration. Economic calculations assume that the overall catalyst life is more than
20 years, the entire length of the project, so it will never need to be replaced.
11
M101 H101CF101
H101H
H102 D101
R101
X101
T301
T201
M201
PR201
SP201
C101
C102
T202
H201
M301
S110
S103S104
S105 S106
S107
S201
S301
S102
S101
S303
S206
S120
S204
S202
S205
FUELGAS
H2SALE
S210
S220
S207
S203
TMBPROD
S302
FEED
VI. Description of Base Case Part A: Sections 100, 200, and 300 Up to the Mixer (M301)
12
M301
S304
T302
PURGE
T303
BENZPROD
S307
TOLPROD
T304 XYPROD
S305
V301 S306
Part B: Section 300 After the Mixer (M301)
13
M401
C401
H401 V401 SP401
V402H501C
C402
C501
H501HV501
H202C
H204C
HPFEED
S402
S403 S404 S405
S406S407S408
LPFEED
S401
ETHYLENE
S501
S502
S503S504
S409
S410
Part C: Refrigeration
Figure 1: Overall Process Flow Diagram. Part A: Section 100, 200, and 300 up to M301. Part B: Remaining portion of section 300.
Part C: Section 400 and 500.
14
Section 100: Feed processing, reaction, and initial separation (Bridger)
Figure 2 : Feed and reactor (Section 100). Feed ethane is mixed with hydrocarbon and hydrogen
recycles. The presence of hydrogen significantly reduces catalyst coking. The flash tank, D101,
sunders gaseous C1-C5 to section 200 and C6-C9 to section 300.
The EBTAX plant design begins with a stream of feed ethane purchased from a local natural gas
refinery. With the EBTAX plant located in the Houston area, purchasing the feedstock from a local
refinery will reduce shipping and processing cost associated with plant operation. The cost of feed ethane
is further reduced due to the recent practice of ethane rejection. Ethane that would now be attributed to
the household natural gas stream for little to no profit can instead be purchased by EBTAX at a very low
cost for conversion into valuable products.
The purchased ethane has a minimum purity standard of 95% ethane with the primary impurities
of propane and carbon dioxide. The maximum allowable range for propane content is 0% to 5%.
(Lonestar) These small amounts of propane create no detrimental effects within the system. Propane,
being a light hydrocarbon, also reacts via catalysis to form the valuable BTX products. Propane gas
actually has a higher conversion to our most valuable product, benzene, than ethane; however, ethane
remains the ideal choice of feedstock due its low cost and availability.
15
The maximum allowable range for carbon dioxide content is 0% to 0.1%. (Lonestar) Very few
changes occur throughout the system due to varying carbon dioxide concentrations. Carbon dioxide gas
remains inert within the reactor, so no additional reactions take place. No new products are formed and
the product selectivities remain constant. This is an assumption, but ethane is not available without trace
carbon dioxide. Additional problems should have been included in the literature, but this may need to be
verified. The carbon dioxide present in the reactor effluent continues through the separations process.
The flash separator, D101, diverts 100% of contaminant carbon dioxide to the lights separation process
(Section 200). During lights separation, 100% of the carbon dioxide is separated via the pressure swing
adsorption unit, PR201. The carbon dioxide remains in the methane gas stream, which is used to fuel the
heating processes for the EBTAX plant. The small carbon dioxide contents will only slightly contribute
to the overall emissions from the reactor furnace, F101, and reboiler furnace, F301.
Feed processing
The feed ethane begins at high pressure, 835 psia, at a rate of 186000 pounds per hour. The
ethane flow was determined to produce 700 million pounds of desired aromatics; benzene, toluene, and
xylene, annually at 8250 operating hours per year. This raw feed is passed through an expander, X101, to
reduce the pressure to 320 psia. This produces 1460hp of energy that can be rerouted to other energy
intensive unit operations in the process. Lowering the pressure allows the raw feed to be mixed with the
recycle streams near operating pressure.
The low pressure feed stream, S101, mixes with two recycle streams, S110 and S120, in mixer
M101 before being heated to operating temperature. Stream S110 consists of a 99.5 mol% hydrogen with
the remainder being methane. Hydrogen is recycled at a 1 to 10 mole ratio with the hydrocarbons (HC)
entering the reactor, including the HC recycle S120. This amounts to 2100 pounds per hour at standard
operating conditions. The recycled hydrogen reduces catalyst coking within the reactor. This decreases
the catalyst coking rate, allowing for more time to pass between catalyst regeneration cycles.
16
Recycle stream S120 contains the HC recycle stream from the light separation section (Section
200). The primary component in S120 is unreacted ethane at 91.2%. This stream also contains 4.5%
propane, and trace C4 and C5 HCs. The propane will react to desired aromatics using similar
mechanisms and selectivities as ethane, with a trend towards more benzene, our most valuable aromatic
product. Conversions for each reaction can be seen in table 1. The trace C4 and C5 HCs do not react
further and have no detriment to the reaction system. To prevent a build-up of inert components, a purge
stream in the lights separation (Section 200) is used to vent excess C4 and C5. These mid-range HCs are
used as fuel gas to heat the furnaces. This purge is used as fuel gas before use of excess produced
methane in order to minimize waste streams and associated processing and handling costs.
After mixing with the recycle streams, the complete feed is heated to the operating pressure of
1150°F through two units, a counter-current, cross-reactor heat exchanger, H101, and furnace F101. The
H101 heat exchanger utilizes our hot, 1150°F, reactor effluent as the heating fluid to simultaneously raise
the reactor inlet fluid temperature and cool the reactor effluent to prepare for separation. The hot effluent
is located on the tube side and the cool reactor inlet is on the shell side. This exchanger heats the inlet
fluid to a temperature of700°F, and cools the reactor effluent to 703°F.
The reactor feed flow is further heated through the furnace F101. The operating temperature for
the reactor is set to 1150°F. To achieve this temperature, the furnace operates using the C4 and C5 purge
stream as fuel gas. The purge stream contains other light HCs as well, which burn normally within the
system. Furnace F101 also uses the excess methane produced in the plant as a fuel source. The PSA
separation unit, PS201, diverts enough methane from the process that no additional methane or fuel is
required for purchase.
Reactor mechanics
At this point, the reactor inlet, S104, is at the proper inlet conditions. The temperature is 1150°F
and the pressure is at 300psia. The stream then enters two identical parallel reactors. The operation
17
utilizes three reactors. Two reactors are operated at any given time while the third is being used for
catalyst regeneration. Each reactor is composed of a glass-lined stainless steel vessel with a packed bed
of catalyst. Stainless steel is used in all components in contact with a hydrogen stream, particularly the
reactor, to protect against hydrogen embrittlement and possible explosion risk. The glass lining within
the reactors is used to protect against the chlorine gas used in the regeneration process.
The catalyst is composed of a zeolite base with germanium inserted into the structure. This
germanium-enhanced zeolite is used as an anchor for platinum, which provides the catalyst its reactivity.
A platinum content of .0441wt% is used according to the conversion and selectivity data available in US
patent US8772563. The single pass conversion for this catalyst at these operating conditions is 46%. The
overall selectivity towards desired aromatic products is 61%. A full list of reaction conversions can be
seen in Table 1.
18
Table 1: List of reactions, their respective conversions, and the heats of reaction
Reactions Conversion
ΔHrxn
(BTU/lb)
C2H6 + H2 --> 2 CH4 0.075 -3.81
C2H6 --> C2H4 + H2 0.05904 8.04
3C2H6 --> 2C3H6 + 3H2 0.00718 5.76
3C2H6 --> 2C3H8 + H2 0.00685 0.905
3C2H6 --> C6H6 + 6H2 0.17372 5.92
4C2H6 --> C7H8 +CH4 + 6H2 0.09250 4.60
4C2H6 --> C8H10 +7H2 0.01934 4.61
5C2H6 --> C9H12 + CH4 + 7H2 0.02982 7.56
C2H4 + 2H2 --> 2CH4 0.18710 -13.3
3C2H4 --> 2C3H6 0.01770 -2.56
3C2H4 + 2H2 --> 2C3H8 0.01690 -8.03
3C2H4 --> C6H6 + 3H2 0.42870 -2.38
4C2H4 --> C7H8 + 2H2 + CH4 0.22830 -3.87
4C2H4 --> C8H10 + 3H2 0.04770 -3.86
5C2H4 --> C9H12 + 2H2 + CH4 0.07360 -0.543
2C3H8--> C6H6 + 5H2 0.18153 3.08
3C3H8 --> C7H8 + C2H6 + 5H2 0.23330 2.21
3C3H8 --> C8H10 + CH4 + 5H2 0.10370 1.70
C3H6 + H2 --> CH4 + C2H4 0.32248 4.86
2C3H6 --> C6H6 + 3H2 0.21086 -1.07
3C3H6 --> C7H8 + C2H6 + 2H2 0.27106 0.104
3C3H6 --> C8H10 + CH4 + 2H2 0.12050 -0.803
3C3H8 --> C9H12+ 6H2 0.06466 -1.34
3C3H6 --> C9H12 + 3H2 0.07509 1.97
2C2H6-->C4H10+H2 0.00163
3C2H6-->C5H12+CH4+H2 0.00014
19
An operating temperature of 1150°F is the essential variable which controls the extent of the
reaction and gives the desired selectivities. Adjusting this temperature lowers the single pass conversion
and adjusts selectivity away from the desired aromatic products.
The operating pressure is a more flexible variable. Conversion and selectivity data in US patent
US8772563 is provided for lab conditions at atmospheric pressure, and the patent lists functional pressure
from 20-2000 psia. The chosen pressure, 300psia, reduces the size of the reactor significantly. According
to La Chatlier’s, principle, this increased pressure should reduce conversion due to a higher number of
moles in the products of many reactions taking place. However, when lowering pressure to the patent
tested values, the reactor becomes too large, and catalyst based capital costs become too high. The large
functional pressure range, combined with a lack of tests for pressure correlation in US patent US8772563,
creates the largest assumption that must be verified at the pilot scale and allows for multiple design
options; refer to Design Alternatives VII and Future Work XIII.
At the high temperatures occurring within the reactor, the catalyst undergoes coking as the
reactions occur. After approximately several months the catalyst in the reactor will become coked enough
to affect reaction conversions and selectivities. When coking becomes significant, a catalyst regeneration
procedure can be performed to return the catalyst to its previous, fully active state.
The regeneration procedure follows a 4 step procedure outlined by patent US20080154079. First,
coke is removed via high temperature oxidation. Temperatures in the reactor and during coke removal
are sufficient so that sintering of the catalyst occurs. The sintering causes catalyst particles to group up,
reducing catalyst efficiency and selectivities. The second step of regeneration is to redisperse the
platinum over the catalyst surface using a gas stream containing chlorine gas, oxygen, and steam. The
glass lining protects the stainless steel reactor from the chlorine gas. The chlorine gas is then removed
from the stream and steam continues to flow. The steam without chlorine allows the platinum to rebind to
the surface of the zeolite. The final step in catalyst regeneration is the reduction of the catalyst using
hydrogen. This counteracts the original oxidation and returns the catalyst to its original, active state.
20
Once the catalyst regeneration process is completed, the catalyst returns to its original activity.
The estimated lifetime for a single load of catalyst is estimated to be 20 years, the life of the project.
Sensitivities were performed to account for the possible replacement of catalyst one time during the 20
year plant life. These showed very little impact in the economic analysis; refer to Economics.
Catalyst regeneration procedures are the reason that three parallel reactors are used. During one
reactor’s regeneration cycle, the other reactor remains operational under normal operating conditions.
This allows the plant to maintain a continuous product stream, even while a reactor undergoes
maintenance. Once regeneration procedures are complete, the offline reactor can be started up to
continue normal operations.
Inside the reactor, the ethane, along with recycled ethylene, propane, and propylene, react while
interacting with the catalyst surface. These reactants react in 26 parallel reactions, converting light C2
and C3 HCs into a mixed stream of C1 through C9 HCs and hydrogen. Each of these reactions is
exothermic, producing heat and creating the possibility of a runaway reaction. Precautions regarding the
reactor, including cooling and other measures, will be installed to minimize the associated risks; refer to
Safety IX. A comprehensive reactions list and reaction enthalpies can be found in table 1.
Initial separation
The desired products consist of C6 through C8, benzene, toluene, and para-xylene. The catalyst
in use is unique in its ability to create p-xylene exclusively, without alternative conformations. The heavy
C9 by-product is 1, 3, 5 trimethylbenzene (TMB). These heavy components are later processed in
Section 300, product recovery. The light C1 through C5 products consist of straight chain paraffins and
olefins. The C2 and C3 HCs are recycled into the reactor via stream S120, and can react further to form
the desired products. The C1 methane, C4, and C5 HCs cannot react again and are isolated for purge in
the lights separation (Section 200). These gasses are used to entirely fuel the furnaces F101, and F301.
21
Hydrogen gas is purified from the mixed stream and can be sold. The necessary hydrogen for recycle via
stream S110 is removed from the 99.5mol% pure hydrogen stream via splitter SP201.
The streams from the parallel reactors reconverge in reactor effluent stream S105. Stream S105,
containing the full mixture of gasses, is cooled in the cross-reactor heat exchanger H101 to 703°F, as
mentioned above. In H101, reactor effluent is used as the heating fluid and is located on the tube side of
the exchanger.
The still gaseous stream is then cooled dramatically in heat exchanger H102 to facilitate the
condensation of heavy aromatic products. The product stream is lowered to a temperature of 70°F. The
pressure is dropped to 290psia. The process stream is located on the tube side of heat exchanger H102,
which is actually broken down into two different heat exchangers. Normal cooling water modeled to
enter at 95°F cools the process stream down to 105°F. After this, cooling water cooled in section 200 of
the process is used to cool the process stream down the remaining 35°F. This cooling water is set to enter
at 60°F and exit at 80°F.
Once the reactor effluent has been cooled, the separations processes can begin, allowing for the
separation of all valuable components. The first step in isolating the products is the initial splitting of the
stream to form a light HC rich stream and a heavy HC rich stream. The light HC stream, which still
contains some heavy components due to only having a single stage flash, is routed to the light separation
section (Section 200). The heavy HC stream, along with some light components, is routed to the product
separation section (section 300). These initial split streams have their trace contaminants separated and
diverted back to their respective purification sections. The extra separation processes in these sections
maximizes the product and feed recovery, as well as allow for maximum product purity.
The mixed effluent enters a flash drum, D101, to initialize this splitting. From the flash vessel,
gaseous light HCs are expelled from the top in stream S201 to the light separation (Section 200). This
stream contains 98% light hydrocarbons, C1 through C5 and hydrogen, and is contaminated with 2%
22
benzene, and trace heavier components. The liquid stream from flash drum D101, stream S301, begins
the heavy separation section (Section 300). This stream is composed of 82% desired aromatics; 49%
benzene, 26.4% toluene, and 7% p-xylene. Some light hydrocarbons remain in this liquid at 7% C2
through C5. The remaining impurity consists of TMB at 9%
23
Section 200: Lights Separation Section (Emily)
Figure 3: Section 200, the Lights Separation Section. This section consists of a mixer to combine
the vapor stream from the flash drum and a recycle from the product recovery section, two
distillation towers and a pressure swing adsorption (PSA) unit to separate the product and recycle
streams, one splitter to allow for the hydrogen sale stream to be separated, and two compressors
to pressurize the recycle streams to appropriate pressures to be mixed with the feed stream.
Section 200, the lights separation section, starts with the vapor stream exiting the flash tank from
section 100 (D101) and mixing with a vapor recovery stream from the 300 section, or the product
recovery section. After mixing, these streams enter a distillation column that separates methane and
hydrogen from the mixture. The hydrogen and methane mixture is heated and sent to a pressure swing
adsorption (PSA) unit where they are separated. The hydrogen product stream is split between a stream
that will be sold and a stream that is recycled back to the reactor to prevent coking in the catalyst. The
methane that is recovered can be used as fuel for the furnace before the reactor, but it can also be sold
based on its heating value. The components left after the hydrogen and methane have been taken out
continue on to another distillation column, where the components of benzene, toluene, para-xylene and
TMB are sent to the 300 section for further product recovery. The hydrocarbons heavier than methane but
lighter than benzene either continue through a recycle back to the reactor or get sent to the furnace to be
burned to aid in the pre-heating of the process stream before the reactor.
T201
M201
PR201SP201
C101
C102
H201
S204
S202
S205
FUELG AS
H2SALE
S210
S220
S203
T202
S206
S120
S110
S201
S303
S207
24
Figure 4: Section 200 Up To T201. From the flash drum, the vapor stream is mixed in M201
with a vapor recovery stream in the product recovery section before being sent to a distillation
tower (T201) to remove hydrogen and methane from the product stream as the vapor distillate
(S203), with the remainder exiting the tower in the bottoms stream (S206).
The distillate from D101 exits at 290 psia and 70°F. This vapor stream is then mixed with the
vapor distillate of the first product recovery distillation tower (T301) in order to collect all of the light
HCs, ideally those lighter than benzene, which is our lightest product from this process. This separation is
not perfect, so two recycle streams between Sections 200 and 300 have been incorporated in the design to
maximize product recovery and purity. After both the stream from the flash tank and the vapor distillate
from T301 have been mixed, they are sent to the first distillation tower (T201) that separates the stream
such that methane and hydrogen exit as the distillate at 270 psia and -109°F, and the remaining HCs exit
in the bottoms at 276 psia and 19°F. The reason behind separating methane and hydrogen first is that they
are the most abundant compounds in the beginning of this section. Removing them before any other
separations are made significantly lowers the amount of material flowing through the rest of the 200
section, which can help the rest of the section achieve more successful separations. Because very low
temperatures are required to condense ethane, the outlet temperature of the distillate is -109°F.
D101
S201
M201
S303
S202
T201
S107
S203
S206
25
Figure 5: Section 200, T201 Distillate Path After T201. The hydrogen and methane stream is sent
to a heat exchanger to warm it up to room temperature before being sent to the pressure swing
adsorption (PSA) unit, where the methane is removed to a fuel gas stream and the hydrogen
stream is sent to a splitter (S201). This splits the hydrogen stream into a sale product stream and a
recycle stream, which will be sent compressed and sent back to the reactor to prevent coking of
the catalyst.
The methane and hydrogen are then heated to 65°F, which is just under room temperature, and
sent to a pressure swing adsorption (PSA) unit where they are separated into a hydrogen stream and a fuel
gas stream using the concepts of adsorption and desorption and how they are related to changes in
pressure. A PSA unit was chosen for this part of the process mainly because operating temperatures for
this technology are generally near or at room temperature. Even though the amount of gas adsorbed to the
adsorbent depends on both pressure and temperature, a PSA unit “swings” the pressure from high to low
and back again theoretically with the temperature change of the unit being negligible. This makes the
whole unit inherently safer, as there will not be extreme temperatures in the PSA unit like those required
in the reactor, and there should be no reason why the temperature would rise suddenly, to pose other
safety risks. Any noticeable temperature change could be seen as the PSA unit not operating correctly,
and actions could be taken to correct the problem immediately.
Under normal operating conditions, PSA units go through a two-phase cycle, which can be
repeated without maintenance required at the end of each cycle. Because of this cyclical nature, the PSA
unit will consist of several adsorbent vessels with staggered cycle timing so as to provide a constant and
continuous flow of the product stream to continue through the process. Each vessel simply contains an
adsorbent, much like our reactor will contain a catalyst. The difference between a reactor and a PSA
vessel is that there are no chemical reactions that take place in a PSA vessel. The adsorbent being used
S203
H201
S204PR201
FUELGAS
S205
SP201
H2SALE
S210
C101
S110
26
has not been specifically designed, but will probably be either zeolite or activated carbon, as both of these
have very high surface area to volume ratios, which is an important factor in the amount of gas that can be
adsorbed per mass of adsorbent.
During the adsorption phase, the pressure is raised so that the methane will adsorb and the
hydrogen will pass through, thereby creating a product stream that is almost pure hydrogen. Once the
adsorbent is saturated with methane, the feed to the PSA unit will be drastically lowered to allow the
pressure to decrease almost to atmospheric pressure. The swing from a high pressure to atmospheric
pressure allows the methane to desorb and exit the system as a fuel gas stream with little hydrogen lost
overall. This desorption phase also “regenerates” the adsorbent and allows it to be used in the next cycle
with little adsorbance capacity lost between cycles. The desorption phase also allows the PSA unit to
operate without additional maintenance after every cycle. The hydrogen product stream from the PSA
unit will contain hydrogen at 99.5 mol% purity in a 6,000 lb/hr stream. Small amounts of methane will
also be present due to adsorption not being a perfect separation method.
After being purified in the PSA unit, the hydrogen product stream is divided between a recycle
stream and a sale stream, with 15 mol%, or 1,050 lb/hr, of the total hydrogen stream being recycled back
to the reactor to prevent coking in the reactor catalyst. Before reaching the reactor, this hydrogen will go
through a compressor to raise the pressure of the stream to match that of the other streams being mixed to
enter the reactor. It will then enter the heat exchanger and furnace that the feed stream enters to attain the
1150°F operating temperature of the reactor. The remaining hydrogen will be sold back to the oil refinery
that this chemical process is located on. A global market for hydrogen is basically nonexistent, because
hydrogen is difficult to store without it escaping easily due to its small size. Despite the lack of a global
market, refineries often use large quantities of hydrogen, since it is a valuable feedstock for many of the
refinery’s processes. These processes often include hydrogenating large hydrocarbons to break them into
smaller pieces that will be used in other processes or simply sold as fuels. Selling the hydrogen in this
stream back to the refinery is a natural economic decision, as it increases the revenue of this process
27
beyond that from merely the main separated benzene, toluene, and para-xylene products. The separated
methane will be released from the PSA unit into a fuel gas stream in the quantity of 74,000 lb/hr. The fuel
gas stream will contain roughly 50mol% methane, with impurities of hydrogen, ethane and ethylene.
Propane, propylene and C4 hydrocarbons will also be present, but in negligible amounts. This fuel gas
stream will be burned in one of the furnaces in the system, either the furnace before the reactor to heat the
reactants to the temperature required for the reactor or the furnace used in the reboiler for a distillation
column (T301) in the 300 section.
.
Figure 6: Section 200, T201 Bottoms Path. The bottoms of T201 is sent to T202, where C2 and
C3 hydrocarbons are distilled off and sent to a compressor before being recycled back to the
recycled to increase the overall conversion of the reactor. The bottoms stream is sent to a mixer in
the product recovery section to recover any BTX products that could have been lost
The bottoms stream from T201 is pumped to the next distillation tower (T202) to separate the
ethane, ethylene, propane, and propylene from the C4 and C5 HCs. The C2 and C3 HCs will be recycled
back to the reactor at a rate of 151.7Mlb/hr to achieve a higher overall conversion of the system. This
outlet stream will be at 320 psia and 57.9°F straight out of the tower, so it will have to be recycled to the
heat exchanger and furnace before the reactor to be heated to reactor conditions once again. Before it
S120
C102
S220
T202
S207
S206
T201
M301
S304
28
reaches the heat exchangers, a compressor (C102) will be used to pressurize the stream to match the
pressure of the feed stream. The bottoms from T202, which is at 206 psia and 229°F and is mostly C4
and C5 HCs, will be sent to Section 300, the product separation section. Just over 72 mol% of this
particular stream is composed of benzene, toluene, and para-xylene products that can and should be
recovered to maximize our total product stream, which in turn helps to maximize the economics of this
process.
As mentioned previously, the low temperatures being used to condense these light HCs requires
heat removal beyond the capabilities of cooling water in the condensers of both distillation towers. To
solve this problem, two refrigeration systems were introduced to the process to allow for extremely low
temperatures in the process. Ethylene and propane were chosen as refrigerants because of their capability
of working together to achieve the very cold temperatures that the separation processes requires in the
condensers. Both refrigeration systems work in tandem similarly to how they work in LNG plants and are
described in much more detail later in the 400 and 500 sections of this report. The -109°F temperature in
the condenser of T201 has been addressed with ethylene refrigeration. The propane refrigeration system
will primarily be used in the condenser of T202 to drop the temperature to 1.7°F. Propane refrigeration is
also required to condense the ethylene used in T201. The reboilers in T201 and T202 use cooling water
as their heating fluid. This allows for the cooling water to reach lower temperatures than normal and is
integrating into cooling in H102 in the 100 section and in the condenser in T302 in the 300 section.
29
Section 300: Separation and Recovery of BTX and Heavy Aromatics Products (Saud)
Figure 7: Heavy Separation (Section 300). From flash tank D101, the heavy stream is separated
remaining light hydrocarbons. The remaining heavies are separated into Benzene, Toluene, and
Xylene product and TMB byproduct.
The process of heavy separation of the BTX product starts at the flash drum. The liquid effluent
mainly contains aromatics and a small amount of light hydrocarbons. In order to separate this heavy
product, the stream is fed into a distillation column (T301). This distillation column will separate the
effluent into a BTX mixture (S302), a 1,3,5-trimethylbenzene (TMBPROD) product and will recover the
light hydrocarbons. In this step, the 1,3,5-trimethylbenzene (TMBPROD) stream is separated and will be
sold without further purification. The BTX mixture stream (S302) will be mixed in a mixer (M301) with
the light hydrocarbon stream from Section 200 that contains some of the escaped BTX (S207). The new
stream (S304) will then enter another distillation column (T302) for further separation. The separated
light hydrocarbons will be directed to the purge stream which will be used for utilities. As for the BTX
mixture (S305), it will go into a valve (V301) to drop the operating pressure. Lowering the pressure will
help to separate the BTX mixture into its components. The low pressure stream (S306) will then go into
another distillation column (T303) for further separation. In this distillation column, benzene is separated
from the stream. As for the toluene and xylene mixture (S307), it will go into another distillation column
(T304) for further separation. Toluene is then separated from the xylene, and all products will be sold
without further purification.
30
T301: TMB recovery and light separation:
Figure 8: Heavies Separation (Section 300). From flash tank D101, the liquid stream is separated
fed into the first distillation column (T301) to remove the remaining light hydrocarbons as well as
recover TMB as a product. The aromatic rich stream is sent on for further processing by S302.
Starting from the flash drum (D101), the liquid effluent stream will be fed to a distillation column
(T301) at a pressure of 290 psia and temperature of 70 F. This stream contains mainly aromatics, but it
still contains a fair amount of light hydrocarbons. In order to recover these light hydrocarbons, the stream
is fed into a distillation column (T301). This distillation is designed to divert all remaining light
hydrocarbons back up to Section 200 (to T201) in stream (S303) while still recovering aromatics and
separating the BTX from the by-product, 1,3,5-trimethylbenzene. The liquid distillate stream (S302)
containing the BTX mixture will be carried to another distillation column (T302). The by-product 1, 3, 5-
trimethylbenzene (TMBPROD) stream will be sold afterwards without further purification. The variations
of operating pressure and temperature inside the distillation column is between 280 to 289 psia, and the
temperature is 280 to 619 F with a flow rate of 79429 lb/hr for the BTX liquid distillate stream (S302).
The distillation column (T301) is made of stainless steel and has 51 actual stages.
31
T302: Purges and Remaining Light Recovery:
Figure 9: Purge Stream. BTX rich streams are fed into tower T302, one of which comes from the
lights separation section, and one of which come from the previous tower, T301. T302 separates
out any remaining lights and purges them from the system. The bottoms of the tower is sent on
for product recovery as it mainly consists of BTX.
The liquid distillate of T301 (S302) contains a BTX mixture and small fractions of light
hydrocarbons. Stream (S207) contains some of the BTX mixtures that escaped during the separation in
distillation column (T301). The stream is at a pressure of 206 psia and a temperature of 229 F. To ensure
that the escaped BTX is accounted for, both streams will then be mixed in the mixer (M301). The
combined stream (S304) will then go into the distillation tower (T302) for further separation. During this
process, this hydrocarbon stream is separated at a pressure of 280 psia and a temperature of 280 F. The
distillation tower has 65 stages and made out of stainless steel. The reason for using stainless steel instead
of carbon steel is the presence of hydrogen in this separation. The tower is designed to divert 99.9mol%
of all the remaining light HCs into the Purge stream. The purge stream will then be used for utilities,
specifically the furnaces in sections 100 and 300 (F101 and F301). The tower will also recover 99.9mol%
of the aromatics (S305) which is then directed to another distillation column (T303). The variations of
32
operating pressure and temperature inside the distillation column (T302) is between 200 to 210 psia, and
the temperature is 72 to 418 F. A temperature of 72°F is only obtained through the use of the cooled
cooling water created in the 200 section. The bottoms BTX stream undergoes a pressure of 210 psia and a
temperature of 418 F.
T303: BTX recovery:
Figure 10: Benzene Recovery. Benzene is recovered from the BTX rich stream leaving T302.
The bottoms of the tower is sent on to recover the remaining Toluene and Xylene.
The liquid effluent stream (S306) then enters the distillation column (T303). After a valve (V301)
to lower the pressure, the stream will be at a pressure of 50 psia and a temperature of 285 F. This
distillation column is designed to separate benzene from the BTX mixture. The liquid distillate stream
(BENZPROD) contains 99.6 wt% pure benzene at a pressure of 35 psia and temperature of 233 F with a
flow rate of 52,080.2 lb/hr. The liquid bottoms stream (S307) contains toluene and para-xylene at a
pressure of 43 psia, a temperature of 318 F, and a flow rate of 32,499 lb/hr. This stream (S307) will then
be carried into another distillation column (T304) for further separation. This column (T303) has 37
stages and is made out of carbon steel. The absence of hydrogen in this separation makes carbon steel a
viable building material for this tower.
33
T304: Toluene, and Xylene recovery
Figure 11: Toluene and Xylene Recovery. T304 separates toluene from para-xylene that is fed to
the tower from T303.
The liquid bottoms stream (S307) from T303 then enters the last distillation column (T304). This
inlet stream is at a pressure of 43 psia and a temperature of 318 F. The column is designed to separate the
TX mixture into toluene and xylene products. The liquid distillate stream (TOLPROD) contains 99.6 wt%
pure toluene at a pressure of 20 psia, a temperature of 254 F, and a flow rate of 25,937.5 lb/hr. The liquid
bottoms stream (XYPROD) contains para-xylene with a flow rate of 6561.38 lb/hr, a pressure of 29 psia
and a temperature of 331 F. This distillation column (T304) has 57 stages and is made out of carbon steel.
Carbon steel can also be used for this tower because of the absence of hydrogen in this separation.
Section 400/500: Propane and Ethylene Refrigeration (Aaron)
As mentioned in the discussion about the lights recovery section (Section 200), the condensers in
both distillation towers require refrigeration. The first tower in section 200 (T201) cools the vapor
distillate down to -109°F while the second tower (T202) cools the vapor distillate down to 2°F. A
34
T201 Condenser
Section 500
T202 Condenser
Section 400
temperature of -109°F requires the use of ethylene refrigeration (section 500) and a temperature of 2°F
requires the use of propane refrigeration (section 400). Propane refrigeration is also needed in order to
condense ethylene and so it is modeled as multi-stage refrigeration in order to obtain the proper
temperatures. The Aspen flow diagram for these sections is shown in Figure 12. The process is also
modeled using the Redlich-Kwong Wilson property method. Optimal operating conditions were
determined through the use of pressure-enthalpy data of both the propane and ethylene.
Figure 12: Propane and Ethylene Refrigeration. Section 400 consists of two propane
refrigeration cycles operating at different pressures. Section 500 consists of only one ethylene
refrigeration cycle. For each cycle the refrigerant is compressed, condensed, expanded, and
evaporated in order to complete the cycle. Propane refrigeration is used in condensing the
process fluid in T202 along with condensing the ethylene in section 500. Ethylene refrigeration
is only required to condense the process fluid in T201.
35
Process Description
Both refrigeration loops are modeled using a generic refrigeration process. In both sections, a
refrigerant is compressed to a desired pressure, one that brings the refrigerant into a temperature range
that allows it to be condensed through the use of the cooling method available. For the propane
refrigeration, the pressure was chosen so that air cooling would be capable of condensing the propane.
The ethylene refrigeration was compressed such that it could be condensed from propane refrigeration
(H501). After condensation, the refrigerant goes through adiabatic expansion in a Joule-Thompson valve
in order to decrease the pressure while simultaneously decreasing the temperature to the value necessary
to be used in the evaporators, or in other words the process heat exchangers in need of refrigeration.
These temperature values are found by giving a 10°F difference between the refrigerant temperature and
the process condenser temperature. The refrigerant is then sent to the heat exchanger that required the
refrigeration and it is evaporated during the process. Since evaporation is an endothermic process it
requires the intake of energy and it is this process that works as the actual refrigeration. The vapor then is
recompressed and the cycle is repeated.
Since the propane refrigeration needs to be multi-stage refrigeration, there are a couple of
differences. In this case, there are two cycles for two different operating conditions that are integrated in
order to save on overall compression and energy costs. Both cycles still however follow the same basic
refrigeration concept outlined above where the refrigerant is compressed, condensed, expanded, and
evaporated. The lower loop that can be traced corresponds to the lower pressure cycle, which is also the
refrigeration for the evaporator (H501C) used to condense the ethylene in section 500 (H501H). Starting
out the low pressure cycle, the propane is compressed (C402) up to the pressure of the high pressure
propane and the streams are combined. The combined streams are then compressed to a determined
pressure and condensed in an air cooler as described previously. The Joule-Thompson valve then that
follows (V401) drops the pressure of the combined streams down to that of the higher pressure loop and
the amount needed for the condenser of T202 (H204C) is split and diverted to that evaporator. The
36
remaining refrigerant is sent to another Joule-Thompson valve (V402) to drop the pressure further to what
is required for the evaporator used in section 500 (H501C).
Operating Conditions
Determining optimal operating conditions for refrigeration is actually very important. When
initially designing this section, values were not optimally chosen and resulted in very expensive
compression. In fact, it required approximately $75,000,000-100,000,000 more in FCI just due to an
excess of required compression. In order to choose optimal operating conditions for refrigeration,
pressure, temperature, and enthalpy data needs to be consulted for the refrigerant being used. The first
step in optimizing operating conditions started with finding the temperature that the refrigerant should be
at in order to evaporate at an acceptable temperature. A basis of at least 10°F temperature difference
between the process fluid being condensed and the refrigerant was chosen, because less than this can
cause control issues.
The high pressure refrigeration cycle in section 400 is designed for the condenser in T202
(H204C), which operates at a temperature of 2°F. A temperature of the propane was chosen to be -10°F.
This means that the corresponding pressure in the evaporator becomes approximately 30 psia, which is
the pressure for the high pressure cycle. The ethylene refrigeration cycle is designed for the condenser in
T201 (H202C), which operates at a temperature of -109°F. A temperature of -125°F was chosen meaning
that the pressure of the ethylene in the evaporator then becomes approximately 35 psia. The low pressure
refrigeration cycle in section 400 is designed for condensing the ethylene in the ethylene refrigeration
section (H501C). This temperature is decided by what the ethylene is compressed to. The colder propane
is, the lower the required compression is in the ethane refrigeration. In order to hopefully minimize the
compression, the temperature of the propane in this section was chosen to be -44°F as this corresponds to
saturated propane at 15 psia, or just above atmospheric pressure. It is important to keep the pressure in
the process above atmospheric so that if there is a break anywhere in the process line, propane will flow
37
out of the system instead of oxygen rushing in. See Table 2 and Table 3 for pressure, temperature, and
enthalpy data [Bühner].
Table 2: Pressure, temperature and enthalpy data for propane
Propane
Pressure Temperature Enthalpy
P [bar] P [psi] T [°C] T [°F] hL [J/g] hV [J/g] Δh [J/g] 0.1 1.1 -85.0 -121.0 -698.4 -158.8 539.6
0.1 1.6 -80.0 -112.0 -672.3 -152.8 519.5
0.2 2.4 -75.0 -103.0 -649.4 -146.8 502.6
0.2 3.3 -70.0 -94.0 -628.9 -140.7 488.2
0.3 4.4 -65.0 -85.0 -610.3 -134.6 475.7
0.4 5.9 -60.0 -76.0 -593.4 -128.6 464.8
0.5 7.7 -55.0 -67.0 -577.6 -122.5 455.1
0.7 9.9 -50.0 -58.0 -562.8 -116.4 446.4
0.9 12.6 -45.0 -49.0 -548.8 -110.4 438.4
1.1 15.8 -40.0 -40.0 -535.3 -104.3 431.0
1.4 19.6 -35.0 -31.0 -522.3 -98.3 424.0
1.7 24.0 -30.0 -22.0 -509.5 -92.4 417.2
2.0 29.1 -25.0 -13.0 -497.0 -86.4 410.6
2.4 35.1 -20.0 -4.0 -484.6 -80.6 404.1
2.9 41.9 -15.0 5.0 -472.3 -74.8 397.5
3.4 49.6 -10.0 14.0 -459.9 -69.0 390.9
4.0 58.4 -5.0 23.0 -447.6 -63.4 384.2
4.7 68.3 0.0 32.0 -435.1 -57.8 377.3
5.5 79.4 5.0 41.0 -422.5 -52.4 370.2
6.3 91.7 10.0 50.0 -409.8 -47.0 362.7
7.3 105.5 15.0 59.0 -396.8 -41.8 355.0
8.3 120.7 20.0 68.0 -383.7 -36.8 346.9
9.5 137.5 25.0 77.0 -370.3 -31.9 338.4
10.7 155.9 30.0 86.0 -356.7 -27.3 329.4
12.1 176.1 35.0 95.0 -342.7 -22.9 319.9
13.7 198.2 40.0 104.0 -328.5 -18.7 309.8
38
Table 3: Pressure, temperature and enthalpy data for ethylene
Ethylene
Pressure Temperature Enthalpy
P [bar] P [psi] T [°C] T [°F] hL [J/g] hV [J/g] Δh [J/g]
0.2 2.3 -130.0 -202.0 -729.0 -206.0 523.0
0.2 3.4 -125.0 -193.0 -714.4 -200.5 513.9
0.3 5.0 -120.0 -184.0 -700.8 -195.1 505.7
0.5 7.2 -115.0 -175.0 -687.9 -189.9 498.0
0.7 10.0 -110.0 -166.0 -675.4 -184.9 490.5
0.9 13.7 -105.0 -157.0 -663.1 -180.0 483.1
1.3 18.3 -100.0 -148.0 -650.9 -175.2 475.7
1.7 23.9 -95.0 -139.0 -638.8 -170.7 468.1
2.1 30.9 -90.0 -130.0 -626.7 -166.4 460.3
2.7 39.3 -85.0 -121.0 -614.5 -162.3 452.2
3.4 49.3 -80.0 -112.0 -602.3 -158.5 443.7
4.2 61.2 -75.0 -103.0 -589.8 -155.0 434.9
5.2 75.0 -70.0 -94.0 -577.3 -151.7 425.6
6.3 91.0 -65.0 -85.0 -564.5 -148.7 415.9
7.5 109.4 -60.0 -76.0 -551.6 -146.0 405.6
9.0 130.3 -55.0 -67.0 -538.4 -143.7 394.7
10.6 154.0 -50.0 -58.0 -524.9 -141.8 383.1
12.5 180.8 -45.0 -49.0 -511.2 -140.3 370.8
14.5 210.7 -40.0 -40.0 -497.0 -139.3 357.7
16.8 244.0 -35.0 -31.0 -482.5 -138.9 343.6
19.4 280.9 -30.0 -22.0 -467.5 -139.1 328.4
22.2 321.8 -25.0 -13.0 -451.9 -140.1 311.8
25.3 366.7 -20.0 -4.0 -435.6 -142.0 293.6
28.7 416.1 -15.0 5.0 -418.4 -145.1 273.3
32.4 470.1 -10.0 14.0 -400.0 -149.9 250.2
36.5 529.3 -5.0 23.0 -379.9 -157.0 223.0
After finding the different evaporators’ temperatures and pressures, compressor discharge
pressures were determined. These values were determined by what is being used to condense the specific
refrigerant. For the propane refrigeration cycle, air cooling was chosen. It is possible that using cooling
water could be more economic, but for now air cooling (H401) will be considered. A base temperature of
the air being used was set to be around that of air in the gulf coast of 100°F, which means the temperature
39
of the condensed propane was determined to be around 115°F. The pressure-temperature data in Table 2
shows that this temperature corresponds to a pressure of approximately 230 psia. A pressure drop of 5 psi
across the air cooler was assumed and so the main compressor (C401) was determined to need a discharge
pressure of 235 psia. For the ethylene refrigeration cycle, as mentioned previously, propane refrigeration
is used. The temperature for the propane in the evaporator (H501C) has already been determined to be -
44°F. This means that the ethylene is condensed (H501H) at approximately -34°F, which corresponds to
a saturation pressure of 235 psia. Taking a pressure drop of 5 psi through the condenser, the discharge
pressure of the ethylene compressor (C501) was determined to be 240 psia.
Considering now that every heat exchanger has a pressure drop of 5 psi, a lot of the needed
operating conditions are known. Most of the remaining conditions are set by the refrigeration cycle itself.
In all three evaporators, H204C, H501C, and H202C, the discharge was set to be at the dew point.
Similarly, in all of the condensers, H401 and H501H, the discharge was set to be at the bubble point. The
only thing remaining to be specified is the flow rate required for all of the evaporators in order to meet the
process needs. Starting with ethylene refrigeration, the duty of the condenser in T201 is calculated in
Aspen and found to be 72 MMBtu/hr. Using the difference in the enthalpy between the input and output
streams of the ethylene in the evaporator, the mass flow rate of ethylene in the refrigeration cycle was
determined. Taking the condenser duty and dividing it by this difference, the flow rate required was
determined to be 530,000 lb/hr. This process was then repeated for the propane refrigeration section
using the condenser duty of T202 and the condenser in section 500 (H501H) along with the
corresponding enthalpy of the inlet and outlet. This calculation resulted in a propane flow of 27,000 lb/hr
in the high pressure loop and 1,100,000 lb/hr in the low pressure loop. A summary of this data including
flowrates and duties is shown in Table 4.
40
Table 4: Various operating conditions specified for the different refrigeration cycles involved
along with the Aspen unit operation or stream that it corresponds to.
Operating
Condition
High Pressure Propane
Cycle (Section 400 -
Upper Loop)
Low Pressure Propane
Cycle (Section 400 -
Upper Loop)
Ethylene Cycle
Unit
Operation or
Stream
Spec'd
Value Spec'd
Unit
Operation or
Stream
Spec'd
Value Spec'd
Unit
Operation or
Stream
Spec'd
Value Spec'd
Refrigerant Flowrate
(Mlb/hr) HPFEED 33 LPFEED 1455 ETHYLENE 563
Evaporator Duty
(MMBtu/hr) H204C 3 H501C 119 H202C 76
Evaporator Pressure
(psia) H204C 30 H501C 15 H202C 35
J-T Valve Discharge
Pressure (psia) V401 35 V402 20 V501 40
Compressor
Discharge Pressure
(psia)
C401 235 C402 30 C501 240
As far as modeling this process in Aspen, the only needed piece of information that hasn’t been
cover is the split fraction for the total flow as to how much propane is diverted to the low pressure and
high pressure cycles. This value was calculated by taking the flow rate in the low pressure cycle and
dividing it by the combined flow rate of both cycles. This value gives the split fraction that is diverted to
the lower pressure cycle. This calculation along with that of each of the mass flows is implemented into
different calculators making the simulation more robust to changes made to the process.
Overall the process being used is pretty sound. Aspen results for pressure, temperature, and
enthalpy data matched up with data that was found in literature (Bühner). It also is quite similar to
refrigeration used in natural gas plants. Some things that could be done that might further reduce capital
include finding ways to decrease the amount of needed refrigeration in the process, analyzing different
refrigerants, and considering water cooling as an alternative to air cooling. A location to search for the
best reduction in capital cost and utility cost is in the compression. When optimizing temperatures and
41
pressures throughout the process the overall compression was dropped from a total of 70,000 hp down to
44,000 hp. This, while resulting in a fair drop in utility costs, dropped the FCI by $75,000,000-
100,000,000. This process as is, is likely not perfect and can be improved but is likely quite good.
42
VII. Design Alternatives (Bridger)
Throughout the initial design phase, the production plant for conversion ethane to aromatics underwent
much iteration. Several of these design plans were still able to yield successful, economic plants under
many different circumstances. These changes vary in the number of units, schematics, operating
conditions, and level of product separation. The primary design was created with the intention of
maximizing IRR while making the plant as applicable as possible.
Possible Reactor Alterations
The largest design flexibility lies in the reactor section (Section 100). Some of these available
reactor alternatives may likely be used, as pilot testing may reveal unexpected catalyst properties. US
patent, US20130324778A1, lists operating condition ranges for catalytic functionality of the zeolite-
germanium-platinum catalyst. While the functional range for pressure is listed from 20psia to 2000psia,
the stoichiometry of the reactions, seen in table 1, reveals that more moles of gas are present in the
products than in the reactants. According to La Chatlier’s principle, increasing the pressure of a reaction
with more moles in the product will slow the reaction, and therefore reduce conversion of the feedstock.
This was not reported in the patent, as lab conversion tests were performed solely at atmospheric pressure,
but is suspected to occur, even slightly, once a larger scale plant is constructed.
Ideally, the reactor pressure of 300psia was chosen to reduce the reactor’s size since the catalyst
shows no indication of significant conversion losses with increasing pressure. At atmospheric pressure,
the reactor becomes so large it becomes wholly infeasible both physically and economically. The cost of
filling the high reactor volume exceeded all plant costs. However, the reactor/s can still operate at any
pressure between these two points. This allows the pressure to be dropped significantly without reaching
gigantic proportions. These changes are primarily dependent on further lab tests, or optimization at the
pilot scale.
43
Economically, lowering the pressure creates larger, or a larger number of, reactors, thereby
increasing the overall capital costs. These reactors must be filled with a larger amount of catalyst, further
increasing the cost. At a constant conversion, the highest-pressure reactor is the most economically
viable since it require the least materials and smallest equipment to produce the same amount of product.
If further research tests reveal that increased pressure reduces conversion, an optimization will have to be
performed that take into account the costs of increasing the reactor size as well as the conversion losses.
It is important to note that due to the C2 and C3 recycle, stream S120, the feed is still utilized to 100%.
The majority of losses associated with the loss of single pass conversion are associated with larger recycle
stream and increased equipment size, along with additional heating and cooling. An optimized reactor
sizing will likely remain within the process specifications and retain viability.
Product Recovery (Section 300) Alternative Designs
Another key area with the possibility of alternative design is the product recovery section
(Section 300). In the proposed plant, benzene, toluene, and xylene are each distilled to their pure
chemical standards, typically above 99.5% purity. In the current design, this allows for the maximum
profit yield. Previous design schematics forewent the final separations and sold a simple, mixed benzene,
toluene, and xylene (BTX) stream. Various pros and cons are present within the mixed stream sale
design, but it was eventually foregone for a more robust process which can applied to many more
opportunities and remain stand alone as a producer in the market.
In previous design iterations, a mixed BTX stream was sold. Selling a mixed BTX stream
reduces the costs associated with separation. This includes the capital cost of multiple towers (T303 and
T304) as well as associated labor, maintenance, and utilities. The original idea was to sell this mixed
stream at a reduced chemical price to a BTX distillery. Many operations that involve BTX are equipped
with the proper units to process and distill the BTX to its pure components. By selling the produced BTX
44
at a slightly reduced price, this allows the partner distillation company to make a small profit while
processing the products.
This approach was foregone primarily due to the idea that finding a company to purchase mixed
BTX at a defined price was much more niche than creating a stand-alone plant. To provide a more secure
basis for the company, assumed associations, complications, and potential falling points were removed
from the design. The mixed stream sale is also difficult to estimate using economic analyses. The prices
of individual components are set based on the market, while selling a mixed stream to a distillery is based
on the individual contract, as well as market conditions. If the BTX stream is treated as a fuel component,
its individual chemical value is lost. The fuel market is the primary consumer of BTX products, in which
case the sale price falls significantly. If a company is found to purchase mixed BTX near the chemical
price, reducing the equipment capital cost is one potential method to improve economically.
Using the mixed BTX strategy still proved economically viable, even before the economic
improvements made in other sections. This shows a high potential for a mixed BTX selling plant if the
proper conditions can be met in terms of sale price, etc. The previous plant IRR was roughly 25% before
extensive optimization of the refrigeration section. The IRR for the designed plant is 32.7%. It comes to
reason that an optimized plant with a mixed BTX stream should have an IRR between these values.
Another design alternative concerning the product stream is to include the first of the two product
separation distillation columns, T303 and T304. The first column (T303) separates chemical grade
benzene, which can be sold for 44 cents per pound. Benzene is the most valuable aromatic product,
which allows for high value sales without the separation of toluene and xylene. The toluene and xylene
stream can be sold similarly to the mixed BTX stream above, either to be distilled, or as a fuel additive.
The toluene and xylene separation tower is fairly small, but still contributes significant capital cost. This
design scheme could be best utilized if a reliable company cannot be found for high price mixed
aromatics sales, and there are capital cost constraints that must be met.
45
Continuous Catalyst Regeneration
One alternative design that could be applied to improve the functionality of the catalyst
regeneration system is to include continuous regeneration. Currently, three reactor vessels are used, and
alternated to run two reactors at any given time while the third undergoes regeneration. This method is
fairly efficient, as the catalyst will need to be regenerated often and takes a short time to go through the
four regeneration steps. By running two reactors, the plant remains at full capacity at all times.
The primary drawback to having three reactors is the capital cost associated with both the extra
reactor vessel and catalyst, particularly in a unit that is not used for one third of operating hours. By
implementing a continuous catalyst regeneration cycle, the extra reactors and catalyst can be removed.
This dramatically reduces capital costs associated with the reactor section, but will also add a complex
continuous unit. More research would need to be performed regarding the catalyst regeneration to create
a continuous process.
In a continuous regeneration operation, the catalyst is continuously removed from the reactor and
reinserted in its regenerated state. This allows fresh catalyst to continuously be present inside the reactor.
This is a benefit compared to the batch regeneration because in batch, the catalyst slowly decreases in
activity before reaching the regeneration threshold. With the continuous presence of fresh catalyst, the
entire system will be able to reach steady state. This will provide constant conversions and selectivities
compared to a deactivating batch process.
A continuous catalyst regeneration process will simplify the management of the reactor system.
The parallel reactors approach requires the switching of reactors every several months, along with the
operations of the regeneration cycle. Switching and manipulating a series of gas streams repeatedly
creates increased opportunity for operator error and hazards. A continuous process, although more
complicated to design, is an inherently safer process.
46
Fuel gas reallocation, C2 through C4 repurposing
Another opportunity for an alternative design is the reallocation of the C4 and C5 purge stream.
Although this stream is primarily used to vent the low levels in inert C4 and C5 in the system, this stream
also contains a number of other light hydrocarbons. The C2 and C3 components in the purge stream
make up only a small percentage of the feed, but these HCs still have the potential to react into higher
value products.
Similarly to mixed BTX, mixed light HCs could be sold to a company prepared to process the
stream into its components for either reaction or sale. This would provide more revenue for the stream as
opposed to using it as a fuel gas. This small percentage of waste HCs were not processed in the current
design due to excessively low temperatures and therefore high refrigeration costs.
This HC stream is very small relative to the other product streams. The economic impact of
changing this stream design appears negligible, but is once again difficult to estimate due to the
unreliability of selling a pre-product to the further refined. The prices are not readily available and a
specific price with the purchasing company would have to be negotiated.
47
VIII. Permitting and Environmental Concerns (Emily)
Environmental issues associated with the design of this chemical plant include those of accidental
as well as operational releases of any of the process chemicals. Accidental releases include any incidences
of the process fluid escaping the process in a place where and/or when it is not designed to leave the
system. These events could be as small as a leak in a joint of two pipes or as large as an explosion, and
include everything in between such as a pressure relief valve opening due to a buildup of pressure. These
events could correlate to a range of releases from almost negligible to extremely large and everything in
between. Accidental releases usually can’t be predicted, and often occur in an emergency situation when a
major part of the process is failing and safety is the most pressing issue due to the concerns of keeping the
workers safe and trying to save the process from being destroyed. Environmental impacts of such releases
are hard to predict, as many hydrocarbons could be released and they may respond differently to being
released into the environment. Very small releases may simply dilute very quickly and react or combust
on a very small scale as to almost be not noticeable. Very large releases may explode or start a fire if
contained within a building, or they may disperse in the ambient air and the larger molecules could
potentially settle out of the atmosphere before reacting and affect soils or bodies of water near the release.
Operational releases for this particular process will simply be the products of combustion
associated with burning the fuel gas stream in the furnaces of the process. Two separate furnaces will be
used to heat the reactor inlet stream before it is reacted and to act as the reboiler of T301. The furnace
before the reactor is required because of the very high temperature of 1150°F required in the reactor that
can’t be feasibly reached with superheated steam in a heat exchanger. The reboiler of T301, which
requires a temperature of 618°F, is also hot enough to make superheated steam in a heat exchanger
become unfeasible. All combustion reactions in these furnaces are expected to produce CO2, but are also
assumed to produce thermal NOx due to the high temperatures of the flames and the fact that air, which is
mostly nitrogen, will be used to provide oxygen to the flame to continue combustion. CO2 emissions can
be estimated using a simple mass balance: any carbon that goes in to be burned must come out, assumedly
48
in CO2. NOx emissions are harder to predict since nitrogen is not purposefully being burned, and there is a
large quantity of it available in the air being used to supply oxygen to the flames. The United States
Environmental Protection Agency (EPA) is in charge of keeping track of how much and what kinds of
pollutants are emitted by industrial processes, and to do this, they devised an equation that could predict
emissions based on the what is being combusted, how the combustion unit is configured, and what control
methods are being used to lower the emissions. Equation 1 is the equation given by the EPA in AP 42,
Compilation of Air Pollutant Emission Factors, to estimate emissions (“Emissions Factors”). Equation 1
is commonly used in estimation calculations, because it simplifies the interactions between the
specifications of the design so that the general population can use the equation. Many charts make up the
rest of AP 42, and they contain values for some of the variables, such as emission factor and overall
emission reduction efficiency, based on measurements of actual emissions from different combustion
units.
E = A x EF x (1-ER/100) Equation 1
Where:
E = emissions, tons/year;
A = activity rate, MMBtu/year;
EF = emission factor, lb/MMBtu; and
ER =overall emission reduction efficiency, %
Initial estimates show that the furnace fuel streams will contain no more than 252.2 MMBtu/hr of
thermal energy combined. These streams are made up of mainly methane, which is very similar to natural
gas, so the EPA natural gas combustion estimates and factors will be used for the emissions calculations
for reasons of simplicity. Because of the large capacity of both furnaces, they will both be considered
Large Wall-Fired Boilers to accurately decide which factors should be used. Without control measures,
49
these estimates provide a calculation resulting in 205 tons/year of thermal NOx being emitted, which is
over the limit of 100 tons/year, since NOx is a criteria pollutant, for needing a permit from the EPA to
operate the furnace.
Possible control measures include using low NOx burners either alone or with flue gas recirculation
(FGR). According to AP 42, low NOx burners achieve this result by breaking up the combustion process
into stages. The smaller stages draw out the process of combustion over a longer period of time, which
results in a cooler flame than normal combustion and suppression of the formation of thermal NOx. AP
42 also says that low NOx burners generally reduce emissions by 40 to 85%. FGR works with recycled
flue gas to dilute the combustion air. This mainly reduces combustion temperatures, but also reduces the
oxygen concentration in the primary flame zone. Both of these factors contribute to the formation of
thermal NOx, so the more they are reduced, the more the NOx emissions are reduced. The combination of
low NOx burners with FGR can be estimated to reduce emissions by 60 to 90%. From the tables in AP 42,
the Emission Factor for uncontrolled boilers is 190 lb/106 scf. In the same units, the Emission Factors for
low NOx burners alone and low NOx burners combined with FGR are 140 and 90, respectively
(“Emissions Factors”). Using these numbers in the equation above, the emissions can be lowered to less
than 100 tons/year of thermal NOx, which is within the acceptable limits, as shown in Table 5.
Table 5: This table shows the emissions of thermal NOx with and without control measures. The
emission limit for needing a permit from the EPA is 100 tons/year, which can be obtained with
control measures in this process.
Emissions (tons/year) Pre-Reactor Furnace Reboiler Furnace Total
Uncontrolled 106.4 99.3 205.8
Low NOx burners 47.1 43.9 91.0
Low NOx burners and
Fuel Gas Recirculation
22.4 20.9 43.3
50
An air permit from the EPA may be required because the uncontrolled emissions are higher than
those allowed without a permit, but control measures will be implemented to lower the emissions to
acceptable rates. Analyzing the Best Available Control Technology (BACT) is slightly difficult, as it
takes into account the cost of the control technology and compares it to how well it works at controlling
emissions. Many options will need to be considered for this analysis, as certain controls may work in one
furnace but not the other. Both of the control methods explained above rely on lowering the flame
temperature, which may be quite suitable for the reboiler furnace, since it only has to reach a temperature
greater than 618°F to transfer heat to the process fluid at that point. The furnace before the reactor needs
to have a flame temperature greater than 1150°F in order to heat the feed and recycle streams to the
reactor’s operating temperature. If the control methods lower the flame temperature below 1150°F,
thermodynamic laws will have to be violated to raise the temperature of the process stream to the
operating temperature of the reactor. To avoid violating thermodynamic laws, other control methods will
have to be considered that operate on principles other than lowering the flame temperature.
51
IX. Safety and Risk Management (Emily)
When operating as designed, this design should pose no safety issues. However, issues will arise
from time to time simply because life is unpredictable. This process involves hydrogen as well as many
different hydrocarbons, which means that the risk of fire and explosion can be extremely high if all
chemicals are not handled properly. Any leaks, even if they are small, can produce hazardous working
conditions for the operators and maintenance crews, as well as anyone else who happens to be working
near this process. Benzene leaks are especially hazardous, as benzene is a known carcinogen. Safety data
sheets for each chemical present in this process are available in Appendix 4 for further information on
their safety concerns.
Besides being hazardous to the environment and the workers near it, leaks are also hazardous to
the process itself, as a leak was not designed for, and the system may not be able to react to such an
unexpected situation. As seen in the HAZOP analysis in Appendix 3, there are many plausible situations
that could result in unknown problems to the rest of the process. These situations include temperatures
and pressures outside of normal operating conditions, as well as improper mixing in the several mixers
throughout the design. The many unit operations included in the process design present many
opportunities for malfunctions or failures to occur, as they all include pipes that go in and out, as well as
valves that control the flow in these pipes, in addition to the configuration of the equipment itself and the
control technology that it includes. Some unit operations have inherently safe design considerations, such
as the PSA, which operates at ambient temperatures, but even these more robust pieces of equipment
cannot save the process from other pieces of equipment not working properly. The problems that are
predicted to happen on a regular basis, such as the catalyst deactivating and heat exchangers scaling, can
be addressed with regular maintenance, and should not cause catastrophic consequences as long as these
problems are taken care of before they create emergency situations. The reactor is the least inherently safe
unit operation in the design simply because it operates at very high temperatures and contains extremely
flammable materials. Because many problems can be predicted from this unit operation, it will be under
52
careful supervision to ensure than none of these problems arise. For example, normal operating conditions
are such that the flammable materials in the reactor are at concentrations larger than the upper explosive
limit. Large deviations in pressure or temperature in the reactor could lead to runaway reactions, use of a
pressure relief valve, and even an explosion if oxygen is allowed to enter the system through a leak or if
the process gas is allowed to leak out into the ambient air. Although reactor failure may have the most
severe consequences of all the unit operations, almost any malfunction or failure could cause the whole
process to be shut down simply because of the recycle streams and fuel gas streams that connect each
section to one another, and also because these unit operations will have to occupy a small total
geographical space to accommodate the requirement that this process be built on or near an existing oil
refinery.
53
X. Project Economics (Aaron)
Equipment and Capital Cost
The costs for all of the individual unit operations, except for the PSA unit and pumps, were
determined through the use of Peters and Timmerhaus [Peters]. These costs were determined as FOB
costs from 2002. The prices of the equipment were scaled through time using average chemical
engineering cost indices from 2002 and 2015 to obtain the cost in 2015 dollars [Economic Indicators 1
and 2]. This FOB cost was then converted into fixed capital investment. The FCI for a PSA unit was
found separately in 2011 dollars and was also scaled into 2015 dollars using cost indexes [Economic
Indicators 1 and 2]]. Taking an estimate of OSBL being 10% of ISBL, and knowing that FCI is equal to
ISBL plus OSBL, ISBL and OSBL were back calculated.
Pumps
Pumps are generally rather small pieces of equipment. Since this process mostly exists in the
gaseous phase, the only pumps needed in the process are those needed for reflux in all distillation towers
along with pumps to move cooling water around. This makes pumps somewhat difficult to model but a
general rule of thumb exists that says pumps account for around 5% of the FOB cost of all equipment so
this assumption will be used for this case [Myers].
Compressors
Since this process mainly exists in the gas phase and requires refrigeration, compressors account
for a large chunk of the overall capital cost. All compressors were modeled in Aspen as polytropic
compressors with an efficiency of 0.75 using the ASME method. Sizing compressors depends on the
inlet actual volume flow and the power used. Compressors were considered to be centrifugal and driven
by a motor. Carbon steel was used for all compressors that didn’t contain significant hydrogen. If
significant hydrogen was present, they were sized as stainless steel. Some compressors required more
54
power than what are commonly available. For these cases, the compressor modeled in Aspen is broken
up into multiple units where in each unit falls within the range of commonly available compressors. For
specifics on compressor sizing and costing, see Table 6.
Table 6: Specific information involved in sizing and costing compressors.
Aspen
ID Description
Inlet Actual
Volume
Flow
[Mcf/hr]
KW
Used
Number
of Units KW/Unit
FOB Cost
MM$
C101 Hydrogen Rec 22 84 1 84 $ 0.14
C102 Light Rec 98 687 1 687 $ 0.40
C401 Prop Ref HP 5404 27264 3 9088 $ 14.10
C402 Prop Ref LP 5290 7561 1 7561 $ 3.80
C501 Ethylene Ref 1945 12606 2 6303 $ 6.40
Total $ 24.84
Turbines
Only one turbine is used in this process and it was sized using the same information as that of
compressors; inlet actual volume flow and power produced. The turbine was modeled in Aspen as an
isentropic turbine with an isentropic efficiency of 0.75. The turbine was designed to be made out of
carbon steel since there isn’t any hydrogen present. For specifics on turbine sizing and costing, see Table
7.
Table 7: Specific information involved in sizing and costing turbines
Aspen ID Description
Inlet Actual
Volume Flow
[Mcf/hr]
KW
Produced
Number
of Units KW/Unit
FOB Cost
MM$ (2002)
X101 Feed Pres Drop 35 1045 1 1045 $ 0.2
55
Furnaces
Due to the excess fuel gas made in this process, furnaces are reasonably available methods to heat
streams that need to be heated to fairly hot temperatures. Furnaces also present an opportunity to create
steam that is to be used around the plant. Since the heat is made from combustion, the hot gases created
during the process can be used as a heat source to generate the steam. This integration has not been taken
into account as of yet. These furnaces were designed as box type with horizontal radiant tubes. For
specifics on furnace sizing and costing, see Table 8.
Table 8: Specific information involved in sizing and costing furnaces
Aspen ID Description MMBtu/min Material FOB Cost MM$
(2002)
F101 Feed Preparation 2.17 Stainless Steel $ 3.10
F301 T301 Reboiler 2.03 Carbon Steel $ 1.70
Total $ 4.80
Heat Exchangers
All heat exchangers were considered to be fixed-tube-sheet heat exchangers with 0.75 inch OD x
1 inch square pitch and 16 – 20 ft bundles. Pressure factors were taken into consideration. Stainless-steel
was used for all heat exchangers that contained a significant amount of hydrogen. Values for the overall
heat transfer coefficient, U, were found by using common values for that between the process fluids
involved [Myers]. Steam was considered to be saturated at 525°F.
Generally, cooling water was modeled to enter at 95°F and exit at 115°F unless it was used to
heat where in the cooling water was modeled to enter at 95°F and exit at 60°F. In a few instances, 95°F
was not cold enough and so the cooling water that was used to heat was then integrated to be used to cool
as well. In order to justify the ability to do this, look at the energy transferred to the water in the reboilers
and the energy needed to be transferred using this special cooling water. Around 40 MMBtu/hr is
56
removed from the cooling water in the reboilers of T201 and T202 while only 10 MMBtu/hr is needed
from the special cooling water and so there is a sufficient amount of special cooling water. Some heat
exchangers that were sized fell outside of the range of the normal size of heat exchanger. If this was the
case, it was broken up into separate units until each unit fell within the common size range. For specifics
on heat exchanger sizing and costing, see Table 9.
Table 9: Specific information involved in sizing and costing heat exchangers
ID Description Material Duty,
[MMBtu/hr]
U
[Btu/hr-
F-ft2]
LMTD
[°F]
Pres
[psia]
Area
[ft2] Units
Area
per
unit
[ft2]
FOB
[M$]
(2002)
H101 Reactor Feed-Effluent SS 134.3 120 536.63 315 2085 1 2085 $ 39.74
H102-1
Reactor Out Cooler
(General CW) SS -152.7 110 141.94 295 9782 1 9782 $ 176.86
H102-2
Reactor Out Cooler
(Special CW) SS -8.6 110 16.02 295 4869 1 4869 $ 83.23
H201 Hydrogen Heater SS 14.6 110 80.33 270 1649 1 1649 $ 27.90
H202 T201 Condenser SS -76.3 120 41.14 270 15455 2 7727 $ 279.02
H203 T201 Reboiler CS 25.7 110 58.72 276 3979 1 3979 $ 28.98
H204 T202 Condenser CS -3.0 120 126.60 200 200 1 200 $ 4.97
H205 T202 Reboiler CS 26.6 170 346.41 206 452 1 452 $ 7.31
H301 T301 Condenser CS -95.1 110 251.27 280 3442 1 3442 $ 26.94
H303 T302 Condenser CS -1.8 110 23.44 200 682 1 682 $ 8.82
H304 T302 Reboiler CS 8.2 170 109.32 210 441 1 441 $ 7.32
H305 T303 Condenser CS -24.0 110 133.01 35 1640 1 1640 $ 12.58
H306 T303 Reboiler CS 17.0 170 207.93 43 481 1 481 $ 7.28
H307 T304 Condenser CS -1.7 110 156.63 20 98 1 98 $ 2.02
H308 T304 Reboiler CS 8.3 170 194.03 29 252 1 252 $ 4.83
H501
Ethylene Cooled by
Propane CS 119.3 120 51.79 415 19198 2 9599 $ 111.67
Total [MM$] $ 0.83
Air Coolers
The only air cooler required in the process is the propane condenser in the refrigeration section.
This particular piece of equipment may be replaced by just a regular heat exchanger but for now this type
57
of exchanger works just fine. It was designed to be made out of carbon steel. Air was assumed to enter
the exchanger at 95°F and exit at 105°F. This approximates conditions in the gulf coast on a fairly warm
day. For specifics on air cooler sizing and costing, see Table 10.
Table 10: Specific information involved in sizing and costing air coolers
Aspen
ID Description Material
Duty
[Btu/hr]
U
[Btu/hr-
F-ft2]
LMTD
[°F]
Pres
[psia]
Area
[ft2] Units
Area
per
unit
[ft2]
FOB
[MM$]
(2002)
H401 Air Cooler CS -241.2 15 37.76 195 425765 1 425765 $ 0.34
Vessels
Vessels involved in this process include the flash tank in section 100, along with all of the reflux
drums that are required. The size of the tank was determined to be the amount of volume that would be
filled by 10 minutes of the flow to the tank. It was then assumed that the length is three times the
diameter and so resulting in the diameter and the length dimensions. The flash tank was considered to be
vertical which was corrected for by a 10% increase to the FOB cost. Pressure factors and material factors
were taken into account where the material was chosen to be carbon steel if there isn’t a significant
amount of hydrogen involved. If there is significant hydrogen, the material was chosen to be stainless
steel. For specifics on vessel sizing and costing, see Table 11.
58
Table 11: Specific information involved in sizing and costing vessels
Aspen
Id Description
Horiz
or Vert
Total
Liquid
Vol Flow
[Mcf/hr]
Drum
Total
Volum
e [ft3]
L
[ft] D [ft]
Pres
[psia]
Materia
l
FOB [M$]
(2002)
D101 Light/Heavy Sep Vert 1.9 315.7 15.3 5.1 290.0 SS $ 52.52
D201 T201 Reflux Drum Horiz 9.8 1630.7 26.5 8.8 270.0 SS $ 116.03
D202 T202 Reflux Drum Horiz 0.6 104.0 10.6 3.5 200.0 CS $ 6.96
D301 T301 Reflux Drum Horiz 12.2 2040.3 28.5 9.5 280.0 CS $ 47.08
D302 T302 Reflux Drum Horiz 0.3 50.6 8.3 2.8 200.0 CS $ 4.35
D303 T303 Reflux Drum Horiz 3.0 508.3 18.0 6.0 35.0 CS $ 10.00
D304 T304 Reflux Drum Horiz 1.8 300.7 15.1 5.0 20.0 CS $ 7.00
Total [MM$] $ 0.24
PSA unit
Pressure swing adsorption units are not as common of equipment. In order to cost and size this
unit a similar unit was found that had been already been costed and sized [Analysis of Natural Gas-to
Liquid Transportation Fuels via Fischer-Tropsch]. The particular PSA unit found sized was done in
2011 dollars and had a hydrogen production of 7091 lb/hr. Using the Six-Tenths Factor Rule [Myers], the
cost of the PSA unit found was adjusted to the size of PSA for the EBTAX process. For specifics on PSA
sizing and costing, see Table 12.
Table 12: Specific information for costing the PSA unit
Aspen
ID Description
EBTAX H2
Production
[Mlb/hr]
Similar Costed
PSA Production
[Mlb/hr]
Similar
Costed
PSA FCI
[MM$]
EBTAX
PSA FCI
[MM$]
(2011)
SP201 Hydrogen Recovery 7.09 12.43 $ 12.47 $ 17.46
59
Catalyst
For this particular case, catalyst is assumed to last the life of the project and is a one-time cost. A
patent for the catalyst chosen gives a gas hourly space velocity for lab tests that were conducted. Using
the volumetric flow of gas to the reactor, the volume of catalyst was calculated. In order to allow
regeneration of the catalyst to occur without changing production, three reactors are used, containing half
of the calculated volume of catalyst each. It is difficult to price the catalyst as it is a rather specific make-
up but generally these catalysts cost 80-120 $/lb and so the price was chosen to be 100 $/lb [Myers]. The
density of similar catalysts was found to be 49-68 lb/ft3 and so the density was chosen to be 59 lb/ft3. For
specifics on sizing and costing catalyst used, see Table 13.
Table 13: Specific information for sizing and costing the amount of catalyst used
Description GHSV [1/hr] Total Vol. Flow To
Reactor [Mcf/hr]
Catalyst Mass
Needed [lb] FOB [MM$] (2015)
Catalyst 1000 677.08 40119.0 $ 6.02
Reactor
As described in the catalyst section, three reactors are used in parallel. Each reactor uses half of
the needed catalyst for the desired production. The size of each reactor is taken to be twice the volume of
catalyst used to account for support, screening, and distribution of feed. The reactor itself was sized the
same way as vessels from here. The length was determined to be three times the diameter. Pressure and
material factors were taken into account where in the material for the three reactors is glass lined
stainless-steel, since there is a significant amount of hydrogen present and chlorine is used in the
regeneration process. The vessels were also all taken to be vertical and so the cost was increased by 10%.
For specifics on sizing and costing the reactors, see Table 14.
60
Table 14: Specific information for sizing and costing the amount of catalyst used
Aspen ID Description Rx Total
Volume [cf] L [ft] D [ft]
Pres
[psia]
FOB [MM$]
(2002)
R101 Reactor 677.08 19.76 6.59 300 $ 0.24
Tower and Trays
Distillation towers require two main components, a shell and trays. The height of each of the
columns was determined by the number of real stages. Assuming 75% efficiency for each stage, the
number of theoretical stages was converted into the number of real stages for each tower. The height of
the column was then determined by assuming a 2 ft tray spacing and adding an additional 14 ft for various
tower needs such as support and distribution. The diameter of the tower and trays was calculated in
Aspen by designing the tower at 80% flood. An additional 5% was added to the cost of the shell to
account for manways and other various column needs. Pressure factors and material factors were taken
into account for the shell where in the towers containing significant amounts of hydrogen were considered
to be made out of stainless-steel. For specifics on sizing and costing results of trays and distillation
columns, see Table 15.
Table 15: Specific information for distillation column and tray sizing and costing
Aspen ID Description
Number of
Actual
Stages
Tower
Height [ft]
Aspen
Diameter
[ft]
Total FOB
[MM$]
(2002)
T201 Demethanizer 11 36 12.84 $ 0.62
T202 Purges 9 32 6.77 $ 0.11
T301 TMB Recovery/Lights Sep 51 116 18.76 $ 0.23
T302 Purges Remaining Lights 65 144 5.45 $ 0.89
T303 BTX Recovery 37 88 6.70 $ 0.12
T304 TMB/Xylene Recovery 57 128 5.58 $ 0.17
Total $ 2.14
61
After compiling this costing information, the FCI for the overall process in 2015 dollars was
determined. First, all costs were scaled to 2015 dollars if they needed to be. After that, all costs that were
FOB needed to be converted into FCI so that all values could be summed together to result in the FCI for
the design. To do this, they were multiplied by 1.1 to account for delivery and then multiplied by 5.04 to
convert it into FCI [Peters]. These results are compiled into Table 16.
Table 16: Fixed Capital Investment for the various equipment involved in the process along with
the resulting total
Equipment Quantity Total FCI (2015 MM$)
Pumps N/A $ 17.34
Compressors 7 $ 206.36
Turbines 1 $ 1.66
Furnaces 2 $ 39.88
Heat Exchangers 33 $ 6.89
Air Coolers 1 $ 2.86
Vessels 7 $ 2.03
PSA Unit 1 $ 17.44
Reactor 3 $ 1.99
Catalyst 1 $ 33.36
Tower+Trays 4 $ 17.75
Total 60 $ 347.55
ISBL N/A $ 315.96
OSBL N/A $ 31.60
Pricing, Revenue and Production Cost
Finding prices for raw materials and chemicals was rather difficult, but reasonable estimates were
found. Ethane was found to be around $3.75/MMBtu [Brown]. Benzene, toluene, and xylene were found
to be priced at $3.45/gal, $2.8/gal, and $2.84/gal respectively in 2008 dollars [Chemicals A-Z]. This
price was scaled into 2015 dollars through the use of economic indexes [Economic Indexes 1 and 2].
TMB was considered to be sold as a gasoline additive and so its price was chosen to be $2.50/gal to
reflect a price slightly higher than gasoline itself. Hydrogen was difficult to price since it doesn’t really
have a market, and most people that need hydrogen just make it themselves. Since this is the case,
62
hydrogen is priced in terms of how much it costs to make it, which turns out to be $0.65/lb [James].
Lastly, the purge stream and fuel gas stream are just burned for energy and so it is priced as natural gas at
$2.40/MMBtu [U.S. Natural Gas Wellhead Price]. Densities and heating values were gathered in Aspen
in order to convert all prices to a c/lb basis in 2015 dollars. A summary of these values and the resulting
cost or income is shown in Table 17.
Table 17: Income or cost of each of the materials consumed or produced
Product or Raw
Material Price [c/lb] Flowrate [Mlb/hr]
Income or Cost
[MM$/yr]
Ethane 7.66 186.17 -117.65
Benzene 44.63 52.34 192.74
Toluene 36.50 25.94 78.09
Xylene 37.40 6.56 20.25
TMB 34.54 9.42 26.83
Hydrogen 65.00 12.43 66.65
Fuel Gas 5.07 73.99 30.94
Purge 4.83 55.14 21.97
After this, the remaining costs of utilities and fixed costs were determined. Steam was assumed
to be priced at $8.00/MMBtu while cooling water was assumed to be priced at $0.40/MMBtu [Myers].
Natural gas was already found to be priced at $2.40/MMBtu [U.S. Natural Gas Wellhead Price].
Lastly electricity was assumed to be at a price of 4 c/kWh. Table 18 shows the summary of utility costs.
Fixed costs were all based off of assumptions made by John Myers in his economic notes [Myers]. For
labor, this design was assumed to have four men per shift, each making $50/hr with a 60% increase in
cost to account for overtime. Maintenance was assumed to be 4% of FCI. Laboratory costs were
assumed to be 10% of labor. Plant overhead was assumed to be 30% of labor, maintenance, and lab costs
combined. Lastly, taxes and insurance were assumed to be 3% of FCI. A summary of fixed costs can be
found in Table 19.
63
Table 18: Cost of utilities
Utility Price [$/MMBtu or
c/kWh]
Energy
Requirement
[MMBtu or kWh]
Cost [MM$/yr]
Steam 8.00 59.37 3.92
Cooling Water 0.40 510.51 1.68
Natural Gas 2.40 265.71 5.26
Electricity 4.00 35493.20 11.71
Table 19: Various fixed costs associated with the design
Fixed Cost Price Basis Yearly Cost
(MM$/yr)
Labor 4 men/shift @ 50$/hr *1.6*8760hr/yr 2.80
Maintenance 4% FCI 12.79
Laboratory 10% labor 0.28
Plant Overhead 30% of (labor, maintenance, lab) 4.76
Taxes and Insurance 3% FCI 9.59
Cash Flow Analysis
The cash flow that resulted for this design used assumptions found in Peters and Timmerhaus
[Peters]. Start-up cost was determined to be 10% FCI and working capital was determined to be 89% of
installed FOB. A build up period of two years was used where the production rate would start at 75%
nameplate, move up to 90% nameplate the following year, and finally reaching 100% in the third year. A
construction period of 3 years was used where 25% of the FCI was spent in the first year, 50% of the FCI
was spent in the second year, and the rest was spent in the third year. The project life was considered to
be 20 years and no scrap value was taken into consideration. A tax rate of 35% was used along with
MACRS5 depreciation. The minimum annual rate of return, or MARR, was decided to be 25%. The
results of the cash flow are compiled into Table 20. For more detail on any sizing, costing, and general
economics such as the cash flow and the production cost estimate, see Appendix 1.
64
Table 20: Results of the cash flow analysis conducted on this design
FCI (MM$) 347.6
NPV0 (MM$) 2838.3
NPV10 (MM$) 926.8
PBP (yrs) 2.6
IRR 30%
Sensitivities
Various sensitivities were run for uncertainties in the design itself. Prices for products and
reactants were varied, prices for utilities were varied, and FCI was varied. Some very specific
sensitivities were also ran for things like reactor size, purchasing more catalyst throughout the project life,
and feed composition. The results for this is compiled into Table 21 and was plotted to form a tornado
diagram, Figure 13. Even though the parameters were varied quite a bit, the only parameters that were
capable of dropping the IRR below or close to the MARR were the FCI, ethane price, and the benzene
price. Decreasing all prices simultaneously, which could happen since all materials scale with natural gas
and oil price, also lowered the profitability quite a bit. This, however, makes sense since these prices and
costs play the largest contribution to the cash flow.
65
Table 21: Sensitivities run, along with the resulting IRR
Parameter Variation Min Max Base Range
± 40% FCI 21.78% 46.41% 30.27% 24.63%
Ethane Price 4-16 c/lb 14.80% 35.54% 30.27% 20.75%
± 40% Benzene Price 21.94% 37.55% 30.27% 15.61%
All Products and Reactants Price Increase or Decrease 19.20% 32.38% 30.27% 13.18%
Reactor Size at 30 psia 22.82% 30.27% 30.27% 7.45%
± 40% Toluene Price 27.05% 33.33% 30.27% 6.28%
± 40% Hydrogen Price 27.53% 32.89% 30.27% 5.35%
± 50% NG price 28.97% 31.55% 30.27% 2.58%
Replacing the Catalyst every 5 years 28.63% 30.27% 30.27% 1.64%
± 40% Xylene Price 29.45% 31.08% 30.27% 1.62%
± 50% Electricity Price 29.73% 30.80% 30.27% 1.07%
± 30% Catalyst Price 30.14% 30.96% 30.27% 0.82%
± 50% Steam Price 30.07% 30.47% 30.27% 0.39%
± 50% Cold Water Price 30.19% 30.36% 30.27% 0.17%
Feed Composition with minimum propane 30.27% 30.33% 30.27% 0.06%
66
Figure 13: Tornado Diagram. This plot chose the change on IRR based on different variation of
various uncertain parameters.
0.00% 5.00% 10.00% 15.00% 20.00% 25.00% 30.00% 35.00% 40.00% 45.00% 50.00%
± 40% FCI
Ethane Price 4-16 c/lb
± 40% Benzene Price
All Products and Reactants Price Increase or Decrease
Reactor Size at 30 psia
± 40% Toluene Price
± 40% Hydrogen Price
± 50% NG price
Replacing the Catalyst every 5 years
± 40% Xylene Price
± 50% Electricity Price
± 30% Catalyst Price
± 50% Steam Price
± 50% Cold Water Price
Feed Composition with minimum propane
IRIRR
67
XI. Global Impacts (Saud)
The recent popularity of fracking among oil companies in the USA has been increasing the
availability of ethane, especially in the Gulf Coast area, as a feedstock. The BTX market is estimated to
have a total value of $81 billion in sales per year. Although the petrochemicals manufacturing companies
are more profitable and are expected to grow with an average rate of 3% over the next five years to $98
billion in 2020. At this point the BTX market is considered to be developing due to the production and the
demand in the market. The BTX market is still considered a good market for business because of the
shortage in supply due to an increase in demand.
The production of BTX emits significant amounts of CO2 and NOx which contributes towards
global warming which have direct impact on the environment. That is why this process requires
government permits in accordance with the federal and state regulations. Currently the plant does not
required permit for emitting CO2. The reason for that it is not regulated by EPA. For the NOx the EPA
required permit.
To understand the petrochemicals market economically, we must consider a basis to analyze the
level of competition within an industry by making a business strategy. Introducing this five forces
analysis, also called the Porter Five Forces, is a structure which used widely in industry to evaluate the
competitive forces that must be considered when the investor is willing to enter a prospective market. The
founder of this idea, Michael Porter, made this tool to analyze the five market forces that determine the
competitive industry and study the target market. The five forces are the competitive rivalry, the power of
suppliers, the power of the buyers, the threat of substitution, and the threat of new entry.
68
Figure 14: Industry Rivalry. This figure illustrates the possible industry pressure associated with
a new competitor
1-The competitive rivalry:
There are relatively few sources, other than fossil fuels, from which hydrocarbons can be
produced, for this reason there is a low potential in the market for new inventions for producing BTX.
Some markets are saturated with the amount of BTX being produced. For example, East Asia has a large
excess of benzene, which is being addressed by sometimes cutting the production to avoid significant
price drops in the market. The producer of the BTX are dependent on the price of the crude oil to indicate
their profit, therefore an alternative method should be considered, such as natural gas, which is more
economically beneficial. Because of that, the expected amount of the competitive rivalry is high.
2-The power of suppliers:
The natural gas that is required for our process, is produced in abundance in the Gulf coast
regions that is why the suppliers’ ability to affect the industry is less significant. One of the reasons for
this is that the price of natural gas is determined by the natural gas market and also because the expense of
transportation of natural gas from an offshore site is fairly large. Since there is a large number of
refineries in the region, which means more suppliers are in the market, hence it will be more convenient
69
to switch over to another supplier, compared to the suppliers of electricity and the catalyst which may
have another challenging factors for the project. And we are not forgetting the supplier of the equipment
needed for this process, the designing have large significant power that must be considered for making
our economic decision.
3-The power of the buyer:
The biggest challenge for the project is the hesitant incorporation of the buyer which means the
cutting of the demand from the producer of BTX. For example Coca-Cola have invested in renewable
para-xylene for their bottles, if the price of their suppliers increases, automatically the price of Coca-Cola
will increase to balance the increase in price from the suppliers, therefore the demand for Coca-Cola will
decrease, since the customers will be price sensitive. The BTX is identical to the other companies which
cause kind of threat to the project if switching to the other competitor. However, since we are threating
ourselves and our own buyers for the BTX these concerns have been lightened.
4-Threat of substitution:
The production of BTX in the market are not facing any threat. Because of the products of the
BTX which are benzene, toluene and para-xylene. For benzene which cannot be easily substituted in the
industries process because of the chemical structure which contains six membered aromatics substances
and they rely on it in many chemical process to form a product. For the toluene price is more sensitive,
because it is used as solvent for other products. And for the yield of the para-xylene usually become the
most threat for the project, because it is used in plastic industries. And the plastic industries are trying to
find an alternative method for producing renewable plastic. Overall the prices for BTX are economically
stable because it is hard to find a substitute source.
5-Threat of new entry:
In the last 18 months the price of the crude oil continues to decrease, hence there will be an
increase in the price and the interest in manufacturing of the BTX. For example the Coca-Cola Company
70
made an investment in Gevo, Virent and Avantium partnerships in aiming to produce para-xylene for
their plastic bottle in most advance renewable way. The summary of that investment is to representing the
new kind of way to threatening for their competitive player in the same industry which might affect the
increasing price of all BTX products. There are many barriers entirely in this industries such as the cost of
the constructions which can be range between couple hundred million to billions of dollars. And because
of that the industries deal with chemical waste and emission and the company are subject to many
regulations such as OSHA and EPA. In conclusion, the production of BTX as a project for the city and
the community is more profitable and beneficial since it will help the community by creating jobs and
supporting the family workers by providing good standard of living.
71
XII. Conclusions and Recommendations (Bridger)
The team EBTAX plant design for the catalytic conversion of ethane into valuable aromatic
compounds resulted in a successful process with potential for industrial application. The EBTAX plant,
which begins with 177000 pounds per hour of feed ethane, results in the production of 700,000 pounds
per year of desired aromatic products; benzene, toluene, and xylene. These pure products are sold as
precursor molecules for a multitude of plastics and other products.
The process, which consists of a simple catalytic packed bed reactor (Section 100), lights
separation (Section 200), product recovery (Section 300), and two refrigeration sections (Section 400 and
500), has a fixed capital investment of 319.7 million dollars. Over the 20 year project life, this plant is
expected to create an IRR of 32.7%. This is above the projected MARR of roughly 25%. This is an
economical process that is expected to perform very well in the projected market.
The official recommendation of team EBTAX is to proceed with further catalyst research before
moving on to the next design phase and a pilot plant. The lack of reliable pressure correlations on the
conversion and selectivity specifications, creates a contradiction between the 20psia to 2000psia listed
operation range and the La Chatlier’s principle. These tests should be very simple to perform at the lab
scale and are not expected to add significant time into the design phase. It is likely these effects may be
insignificant, but the plant should not move forward with costly commitments until this can be verified.
After ensuring that there are no critical detriments to the system due to an increase in pressure,
team EBTAX fully supports the continuation of plant design and initial pilot plant testing. The low cost
of feed ethane due to a large boom in natural gas production makes the process particularly viable. The
cost of natural gas is not expected to rise over the next several years. Even if a rise in prices occur, the
markets of natural gas and BTX products are closely related, and should correspond with increased selling
prices as well.
72
The designed plant schematics and economics are fairly robust, as proven by the sensitivities
analysis. If the price of ethane more than doubles, the IRR remains near 20%. This is the worst case
sensitivity for the plant, and also does not include any compensation via the corresponding increase in
product prices.
Due to the robust, economically viable nature of the plant, continuing research into the potential
drawbacks of the current design is strongly recommended by team EBTAX. The future findings are
expected to confirm the plants profitability and allow for a scale up to a pilot size plant. Once pilot tests
are completed, the full scale process can be designed and an industrial plant can be produced. With the
required small scale research and a build-up of two years, the ethane conversion plant is expected to
become operational near 2020 and remain operational for 20 years until 2040.
73
XIII. Future Work (Aaron)
In moving forward with this project, the first step is to verify and gather all possible information
on the catalyst itself. The catalyst needs to be tested at pressure to see if there are any significant effects
on conversion and selectivity. This was a large assumption that was made, in that conversions weren’t
really affected by pressure and so before moving on, this needs verified. The catalyst regeneration also
needs tested. The time it takes to regenerate the catalyst needs to be determined in order to potentially
improve the current regeneration scheme, or verify that it is good as is. The lifespan of the catalyst also
needs to be verified to know how often catalyst may need to be purchased.
The current design could also use a few finishing touches. Heat integration can be implemented
among heat exchangers to reduce utility cost and capital cost. Small alternatives such as replacing the air
cooler with just another cooling water heat exchanger, and experimentation with reactor pressure can be
explored. Different refrigerants should also be explored in the possibility of further reducing
compression. It would from here be worth analyzing whether or not the current production rate is at a
reasonable value for the current market and whether or not it could be more profitable. Pricing for all
components involved in the process should also be further looked into. Pricing forecasts of ethane and
benzene especially should be found since they play the biggest impact on profitability. The catalyst is
also priced at a generalized value. This price needs to be verified.
After successfully verifying everything unknown about the catalyst and fully optimizing the
current design, process and instrumentation diagrams can be made for each piece of equipment. Control
systems will also need to be designed. Once this is done, a pilot plant should be designed and tested in
detail to further verify the validity of the design and its potential profitability. If the pilot plant then
proves to be economic and reasonable, the design could be implemented at an existing petroleum refinery
at industrial scale.
74
XIV. Acknowledgements (Saud)
We would like to thank the following people for their invaluable guidance in the
successful completion of this Senior Design: our mentor, Dr. David Bell for providing us
assistance and guidance for the completion of the project; Professor John Myers for helping us
through the ASPEN; and Dr. Michael Sommer in the chemistry department for contributing
towards our project by guiding us through the difficult times. We would also like to thank the
people that provided moral support to each member of our group, including but not limited to,
our parents, who encouraged us from the very beginning to dream big and to attempt to
accomplish our dreams; our engineering professors, who taught us the foundational theory
behind every part of the senior design process; our non-engineering professors, who taught us
that a break from engineering is not always a bad thing; and our peers, who shared our burdens
and our accomplishments and provided useful feedback for our unique challenges in this design.
75
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XVI. Appendices
1. Aspen Simulation
2. Master Economics and Sizing
3. HAZOP
4. MSDS
