March 13, 2015 Professor Justin Opatkiewicz Department of NanoEngineering, UC San Diego 9500 Gilman Drive La Jolla, CA 92093 Professor of Chemical Engineering Dear Professor Justin Opatkiewicz: In order for our group to deliver electricity to a 300 office suite building, a PEMFC (Proton Exchange Membrane Fuel Cell) was modeled, utilizing hydrogen as fuel. A plant was designed that synthesized hydrogen from a starting feed of liquid heptane. Simulations were ran through the ASPEN program in order to evaluate an efficient design. Once a working plant model was established, the economics of the plant was analyzed in order to model of the costs and profit of the plant. Future modifications to the plant will help optimize the production of hydrogen to the fuel cell in a more cost efficient manner. Modifications in the delivery of energy to the system and recycling streams were considered to be valuable factors for future simulations, and are discussed although not modeled. These factors will help enable a more cost efficient delivery of hydrogen to the fuel cell, thus providing more profit. Sincerely, Group B4 Brandon Sanchez Janet Mok

Liliana Busanez Saman Hadavand Department of NanoEngineering, Chemical Engineering

Chemical Plant & Process Design:Ceng 124A

Department of Nanoengineering ­ Chemical Engineering University of California, San Diego

Economic Evaluation of a Stationary Fuel Cell Process Plant Design

Group B­4

Primary Author Section

Saman Hadavand Letter of Transmittal, Title Page, Abstract, and Appendices

Janet Mok Introduction, Conclusion, and Table Contents

Brandon Sanchez Results/Discussion

Lilana Busanez Economic Analysis, References

Date Submitted: March 13, 2015

1

Abstract

The total cost of the plant was determined to be $4,823,734.00 +/­ 30%. The net profit of

the plant was unprofitable, losing $239,010,00.00 over a 25 year span. The plant model requires

2895 kW of energy to operate. The total hydrogen produced for the fuel cell feed was 156.3

kmol/hr. The fuel cell outputs a total of 2765 kW of energy. The current density chosen to

operate at was 775 mA/cm2 with a corresponding power density of 0.275 W/cm2. The surface

area of the fuel cell was then determined to be 1005 m2.

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Table of Contents

Abstract Page 2

Table of Contents Page 3

Introduction Page 4

Results/Discussion Heptane Auto­Reformer High Temperature WGS Low Temperature WGS PROX Reactor Fuel Cell Heat Exchangers Optimization

Page 6 Page 7 Page 9 Page 10 Page 11 Page 12 Page 13 Page 13

Economic Analysis Page 15

Conclusion Page 17

References Page 18

Appendices Page 19

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Introduction

Fuel Cells are currently being used as potential successors to internal combustion engines

and stationary power generators like steam turbines and diesel engines because of their high

system efficiencies and low emissions. PEM fuel cells are used for stationary power supplying

electricity and are believed to be an ideal long­term alternative to other processes that waste

energy at conversion interfaces, including the use of platinum/catalyst that are more resistant to

carbon monoxide and minimize effects of poisoning for PEM fuel cells¹.

The Proton Exchange Membrane (PEM) fuel cell was used to provide electricity to a 300

office suite building . The fuel delivery module is what governs the delivery of the hydrogen to

the fuel cell stack and is needed in fuel cell installations. Hydrogen purity is critical in fuel cell

technologies². For the fuel cell performance and operation, the fuel cell stack operates by the

oxidation of hydrogen at the anodes of the individual cells:

(1)₂ 2H⁺ 2eH = +

where protons are produced in this oxidation and are transported through the PEM to the

cathode. Oxygen in air reacts with electrons and protons transfer through the cell to make water,

₂ 4H⁺ e H₂OO + + 4 = 2 (2)

The design of the PEM fuel cell consist of the auto­reformer, the water gas shift reactor, and

PROX reactors. The heptane fuel is fed into the auto­reformer to prepare the hydrogen fuel

containing carbon dioxide, and other impurities, as well as unreacted fuel. The equations used

inside the auto­reformer are,

(3)EPTANE 7H O → 7CO 15HN −H + 2 + 2 (4)O 3H → H O METHANE C + 2 2 +

(5)O H O → CO HC + 2 2 + 2

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(6)ETHANE 2H O → CO 4HM + 2 2 + 2 The Water Gas Shift reactor equilibrium reaction is

(reversible)O H₂O CO₂ H₂C + = + (7)

The kinetics model for the Water Gas Shift(WGS) reactor is shown in the equation

(8)k[CO](1 )− rco = η − β

(9)n k n k /(RT )l = l o − E

where is the effectiveness factor accounting for intraparticle mass transport limitation, [CO] is η

the gas phase concentration of CO, , is the equilibrium CO ][H ]/[H O][CO]K β = [ 2 2 2 T KT

constant for the WGS reaction, and .n(K ) 577.8/T .33l T = 4 − 4

The PROX (PFR) reactor is a packed catalyst bed where the main and side reactions are,

CO O₂ CO₂ (main reaction)2 + = 2 (10)

H₂ O₂ H₂ (side reaction)2 + = 2 (11)

The PROX reactor kinetics model is shown in the equation,

(12)1 /Q ) XCO = 1 − ( − η * k1 * k2 Total1.66

(13).58 exp(− 522/T [K])k1 = 7 * 106 8

(14)6.2 y k2 = 2 * P0.4

CO,in−0.6

* λ0.82

*mPt

where XCO is fractional conversion, is the effectiveness factor of 0.5, k1k2=12 std cm3/min, P isη

the total pressure, yCO,in is the carbon monoxide mole fraction in the feed to PROX,

, mPt is the mass of Pt in the catalyst, and QTotal is the std cm3/min of feed to theO ]/[CO]λ = 2 * [ 2

PROX.

Heat Exchangers, turbines, compressors, and separators were used throughout our

simulation to cool the feed, expand the pressure in the feed, and compress or separate the gases

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in the reactor. The Log Mean Temperature Difference (LMTD) was used to solve for the heat

exchanger area.

(15)

(16)AΔT Q = U LMTD

The utilities used for this simulation were a fuel cell, auto­reformer for

gasoline(N­Heptane), a High Water Gas Shift reactor, Low Water Gas Shift reactor, PROX

reactor, a compressor/turbine, and heat exchangers. The fuel cell cost $270/m2, the auto­reformer

cost $53/kg catalyst, high water gas shift reactor cost $14/kg catalyst, low water gas shift reactor

cost $22/kg catalyst, PROX reactor cost $150/kg catalyst, the compressor/turbine cost $600 each,

and the heat exchanger cost $4/kg. Costly testing can be minimized with 2­dimensional

simulations to stimulate fuel­cell system performance. Further research on fuel cell stacks

continue to be optimized¹.

A complete economic analysis was conducted for the stationary fuel cell design in order

to determine the most cost­efficient design with the proper sizing. The simulation was optimized

to use heat energy from different reactors to power the system as well as using recycle streams to

optimize materials. Equipment sizing, total capital equipment costs, and yearly annual costs were

determined using Aspen.

Results/Discussion

Our current fuel cell design does not supply the entire electrical load to our processing

plant. The fuel cell currently outputs 2765 kW of available energy. The sum of the processing

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plant and auxiliary loads is 1995 kW. The average load for the office building is 900 kW.

Therefore, the current plant design required 2895 kW of energy, leaving a 130 kW of energy still

required for plant operation. Optimally, the fuel cell would power the entire processing plant,

leaving no electrical costs.

The entire process flowsheet is displayed in two parts as Figures A1 and A2 in the

appendix. The following sections involve discussing different segments of the plant beginning

from the liquid heptane feed to the fuel cell exhaust. Only major parameters of streams, i.e. H2

molar flow rates, and equipment will be noted, with the remaining parameters located in the

appendix tables.

Heptane Auto Reformer

Figure 1: Beginning of plant process flowsheet. Liquid heptane and water are vaporized and compressed to prepare for

the auto reformer. The furnace provides heat for the endothermic steam reformation of heptane by combusting heptane.

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The process begins with a liquid heptane feed entering at 30 kmol/hr and 298 K mixing

with superheated steam in the vaporizer. The vaporizer flashes the stream to 885 K; the

parameters are illustrated in Table A1. The stream is then compressed to 5 atm, as illustrated in

Table A2.

The auto reformer operates at 1023 K and 5 atm. Eqn’s 3­6 illustrate the reactions taking

place within the reformer. An external furnace, illustrated in Table A3, was utilized by

combusting the heptane fuel and directing the heat to the auto reformer, thus providing energy

for the endothermic steam reformation of heptane. Once the steam reformation had enough

energy, the remaining reformation reactions were able to proceed. By redirecting the 2636 kW

generated from the furnace, we were able to produce 22 kW from the auto reformer. Reformer

stream results are illustrated in Table A4. 143.64 kmol/hr H2 is produced from the auto reformer,

with a 60% conversion of heptane with respect to Eqn 3. Using the sizing data provided at .51

l/kW of fuel cell output, the auto reformer volume was determined to be 1.41 m3.

High Temperature WGS

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Figure 2: Process flowsheet continuing after auto reformer. Residual methane/heptane are removed from product

stream. Product is expanded from 5 atm to 1 atm and sent through heat exchanger to prepare for high temperature

WGS. Water is used as cooling liquid for heat exchanger.

After the auto reformer, the products are separated to be expanded and cooled to prepare

for the high temperature WGS. The separator takes out residual methane and heptane gas, at a

total of 60.4 kmol/hr. Recycle of these streams was not implemented in the flow sheet and is

discussed in the optimization section later on. The separator is illustrated in Table A5. The

turbine expands the product feed containing: H2O, CO2, O2, H2 and CO to 1 atm. The turbine is

illustrated in Table A6.

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The expanded feed at 784.8 K was then sent through a heat exchanger at 314.8 kmol/hr.

Water was used as the cooling liquid at 298 K and 100 kmol/hr. The output feed was then cooled

to 693 K for the first WGS reactor. The first heat exchanger is illustrated in Table A7.

The first WGS reactor was modeled as a multi­tube plug flow reactor consisting of 100

tubes, each 100m long and 5m wide. The reactor operated at 693 K and 1 atm. Eq 7 illustrates

the reaction taking place within the reactor. Ergun pressure drop correlations were utilized in

order to model catalyst information from the project statement. 10 Mg of catalyst was used. The

approximate volume was calculated to be 13.3 m3. The reactor produced 31.2 kmol/hr H2, which

is a 21.7% increase from the amount of H2 originally present. The first WGS reactor is illustrated

in Table A8.

Low Temperature WGS

Figure 3: Process flowsheet continuing after high temperature WGS. Product is cooled to prepare for low temperature

WGS. Water is used as cooling liquid for heat exchanger.

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The products from the first WGS reactor were then sent to a heat exchanger. The second

heat exchanger cooled the feed to 473 K. Water at 100 kmol/hr and 298 K was used as the

cooling liquid. The second heat exchanger is illustrated in Table A9.

The second WGS reactor was also modeled as a multi­tube plug flow reactor, but

consisting of 50 tubes, each 100m long and 2m wide. Ergun pressure drop correlations were

utilized in order to model catalyst information from the project statement. 10 Mg of catalyst was

used. The approximate volume was calculated to be 12.2 m3. 21.5 kmol/hr of H2 was produced,

which is a 12.3% increase from the amount of H2 originally present. The second WGS reactor is

illustrated in Table A10.

PROX

Figure 4: Process flowsheet continuing after low temperature WGS. Product stream and air are fed to PROX reactor.

Product is then separated to purge everything left over except hydrogen gas.

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The PROX reactor operated at 473 K at 4 atm. 81.63 kg of catalyst was used. The volume

of the PROX reactor was calculated to be .1 m3. The mixed product from the second WGS is fed

to the PROX along with air. 6.13 kmol/hr of CO entered the PROX to be reacted according to

Eqn. 8. CO left the PROX at 27.6 mol/hr, which is a 99.5% conversion of CO. H2 reacted with

oxygen according to Eqn. 9. The amount of H2 leaving the reactor was 156.3 kmol/hr. The

product from the PROX is separated in order to isolate H2 from the other components.

Fuel Cell

Figure 5: Process flowsheet continuing after PROX. H2 gas is expanded to 3 atm and cooled to prepare for fuel cell

feed. Water is used as cooling liquid for heat exchanger. Oxygen is fed to fuel cell to oxidize H2.

The pure H2 stream was expanded to 3 atm and 446.4 K. The second turbine is illustrated

in Table A13. The stream was then sent to the final heat exchanger using water entering at 23.43

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kmol/hr and 298 K for cooling liquid. The H2 leaves the exchanger at 343 K and sent to the fuel

cell. The third heat exchanger is illustrated in Table A14. The amount of H2 entering the fuel

cell is 156.3 kmol/hr.

We chose to run the fuel cell at a current density of 750 mA/cm2 and a pressure of 3 atm

in order to maximize power density at .275 W/cm2, according Figure A3. The fuel cell utilized

140.7 kmol/hr of hydrogen from the feed. In order to achieve the desired current density, a

surface area of 1005 m2 is required. The fuel cell outputs 2765 kW with our current design

parameters. The fuel cell is illustrated in Table A15.

Heat Exchangers

Three heat exchangers were utilized to cool vapor feeds for the high and low temperature

WGS reactors, and the fuel cell. All exchangers were modeled as shell and tube, with the hot

feed entering the shell side and cooling water entering the tube side. The energy required for the

heat exchangers was used to calculate the active heat transfer surface area using the LMTD

approach, according to Eqns 13 and 14. The surface areas for each exchanger 1,2 and 3 were

calculated to be 1.45, 5.84 and 4.96 m2, respectively.

Optimization

Energy

Multiple units in the plant require energy that must be supplied externally in order to

operate. These pieces of equipment are: the feed vaporizer, feed gas compressor, all three heat

exchangers and the heptane auto reformer. The net sum of the energy required to operate these

pieces of equipment is 9172 kW according to our simulation results. The feed vaporizer and the

auto reformer require the majority of the energy, being 4563 and 2615 kW, respectively.

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Fortunately, there are sources of thermal energy released from various locations throughout the

plant that may be directed and used to power the equipment, although not modeled in our

simulation.

Sources of energy generated from the plant, not including the fuel cell, come from: both

of the gas turbines, both of the WGS reactors, and the PROX reactor. The net amount of energy

generated from these pieces of equipment is 4563 kW. The majority of the energy is supplied

from the PROX reactor, being 3226 kW, which is expected from the multiple combustion

reactions. Because the feed vaporizer required 4563 kW to operate, it may be useful to direct all

the heat generated from the PROX reactor to supply energy to the vaporizer. Additional energy

can be supplied from the gas turbines which put out a total of 748 kW, and the WGS reactors

which put out a total of 567 kW.

Water was used as the cooling liquid in all shell and tube heat exchanger models. The

outlet streams of all three of the cooling streams were at 373 K, some being full vapor and some

being mixed phase. At our current design, these heated streams are not utilized in any way. It

would be useful to further optimize the plant by directing heat from these vapor streams to the

externally driven equipment. It could then be possible to reuse the condensed water to combine

with the heat exchanger inlet cooling streams, thus reducing the total amount of water used.

Recycle

There are multiple places in the plant that can utilize recycle streams, although not

modeled in our simulation. These streams are: the residual methane/heptane stream separated

from the auto reformer products and the fuel cell exhaust stream. The heptane and methane

steam reforming reactions and the methane water gas shift reaction can be manipulated by

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recycling methane and heptane into the reformer feed. This design would be in consideration of

Le Chatelier's principle. In knowing that the auto reformer operates at equilibrium, addition of

methane and heptane may push the reactions towards producing more H2. Further optimization

of this recycle design could reduce the total amount of heptane feed required in the plant.

The fuel cell exhaust contains: unreacted hydrogen from the feed, oxygen and water

produced from the reduction of hydrogen. Further optimization of the plant would consider

recycling this water to other stages in the plant, possibly for the cooling liquid for the heat

exchangers. It may be possible to optimize the plant is such a way that the heat exchangers

primarily use the water produced from the fuel cell, and only use external water if needed. The

unreacted hydrogen should be rerouted to other fuel cells on site or any processing plant in the

vicinity that could utilize the excess hydrogen gas.

Economic Analysis

Using the cost curve method, which relates capital cost to capacity, Table 7.2 is applied

to rough estimate the capital cost giving $4,823,734 ± 30% for the plant. Equipment will be

made from 304 stainless steel and carbon steel. The equipment list provides the item combined

costs, and lang factors, included for the material type for the cost distribution of equipment

summing up to this capital cost (Table A16).

Costs factors in the analysis include fixed factory expenses such as equipment

depreciation, utilities, and maintenance as well as direct costs such as material and labor.

However, because this analysis is intended to model manufacturing costs, number of components

that contribute to the original equipment manufacturer are not included in the modeling. The

following is not included in this analysis: one­time costs such as research, design, engineering,

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warranties, advertising, and sales tax. The plant was determined to be unprofitable after taxes

and revenue based off sale of electricity at $32.29 cents per kW hour as priced in Hawaii3.

Overall, the net profit was determined to be ­$239,010,000 over the 25 year plant operation.

The sale of electricity was based according to the U.S. Energy Information

Administration and average retail price of electricity to customers in commercial sector from

December 2014­ which has shown to decline from December 2013. So therefore, a decrease in is

observed energy prices are not accounted for or compared to 2015 rates, and an average retail

price of electricity is used. To take full advantage of energy markets, partnerships, as well as

experienced energy managers would offer strategic approaches.4

Modeling results for PEM fuel system capital costs are broadly consistent with

manufacturer values provided by PEMFC technology and application in the global market today.

According to Fuel Cell Industry Review 2013, investment total capital is at $1.2 billion, where

production capacity is increasing according to agreements with energy industry and companies.

Our plants capacity and capital investment costs does not include promotion of company

agreements and government incentives that would otherwise contribute to greater revenue and

profits.5

Our PEMFC system with electrical output At 2.765 MWh, have utility costs that dominate the

total plant costs since heat or electricity generation is needed for or process. Primary cost drivers are the

compressors and turbines, reactors, and furnace for heat supply, in that order. Besides the costs for

system, the PEM stack consists of reactor size used to calculate stack costs, where the lifetime of the

stack can be increased at the expense of increased cost through system oversizings.6 This cost analysis

aims to develop economic models for our system, including for capital cost, manufacturing cost and

investment cost by taking into account process units and utilities of the system.

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Conclusion

The goal of the experiment was to design the most cost­efficient and economic fuel processing

system for a 300 suite building. Fossil fuels are depleting on Earth and alternative methods are needed

to sustain the human population. Fuel cells are able to produce almost completely renewable energy

especially the Proton Exchange Membrane Fuel Cell which utilizes hydrogen as a fuel for the cell.

Heptane was converted and then used for the process synthesis of hydrogen fuel for the cell. The

PEMFC system that was designed had an electrical output of 2.765 MWh. The fuel cell utilized 140.7

kmol/hr of hydrogen from the feed. A surface area of 1005 m2 is required to achieve the desired

current density. The fuel cell outputs 2765 kW with the current design parameters. The utilities

that were used were a furnace, 3 heat exchangers, a vaporizer, a turbine, a compressor, an

auto­reformer for heptane, a low water gas shift reactor, a high water gas shift reactor, a PROX reactor,

and a fuel cell. The equipment was sized and, and the price of the system was evaluated using Aspen

Plus software. The total capital cost and net profit were determined and an economic analysis was

conducted. The estimated total capital cost for the plant was found to be $4,823,734 ± 30%. The net

profit was determined to be ­$239,010,000 over a span of 25 years. Through the economic analysis of

the plant, the design was deemed non­profitable after taxes and revenue due to the sale of electricity.

The plant design was optimized by using the thermal energy released from various locations

throughout the system that may be directed and used to power the equipment. Sources of energy

generated from the plant, not including the fuel cell, came from both of the gas turbines, both of the

WGS reactors, and the PROX reactor. The net amount of energy generated from these pieces of

equipment was 4563 kW. These heated streams are currently not utilized in any way, and it would be

useful to further optimize the plant by directing the heat from the vapor streams to the externally driven

equipment. The condensed water could then possibly be reused to combine with the heat exchanger

17

inlet cooling streams which would reduce the total amount of water used. The design could also be

optimized by utilizing recycle streams, although not modeled in the simulation. Further optimization of

this recycle design could reduce the total amount of heptane feed required in the plant, and optimize the

plant is such a way that the heat exchangers primarily use the water produced from the fuel cell, and

only use external water if necessary.

References

[1] Weider, John W., et al. “Engineering a Membrane Electrode Assembly” The Electrochemical

Society Interface (2003): 41­43. Print.

[2] Energy.gov. U.S. Department of Energy, Hydrogen and Fuel Cell Technology Basics, 2013.

Web. 10 Mar. 2015.

[3] U.S. Energy Information Administration, Form EIA­826, Monthly Electric Sales and

Revenue Report with State Distributions Report

[4] ARAMARK Energy Services, “Best Practices in Energy Procurement” Managed Heat Rate

[5] FuelCellToday, The Fuel Cell Industry Review 2013, Johnson Matthey PLC trading

[6] Kamarudin, S.K.“Technical design and economic evaluation of a PEM fuel cell system”

ScienceDirect 157.2 (2006): 641–649. Print.

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Appendices

Figure A1: Part 1 of 2. Process flowsheet up to the first WGS reactor.

19

Figure A2: Part 2 of 2. Process flowsheet continuing after first WGS and ending at fuel cell.

20

Figure A3: Fuel Cell data used to solve for cell surface area and power.

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Table A1 (Vaporizer)

Table A2 (Compressor)

Table A3 (Furnace)

Table A4 (Reformer)

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Table A5 (Separator)

Table A6 (Turbine)

Table A7 (Heat Exchanger)

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Table A8 (High WGS)

Table A9 (Heat Exchanger 2)

Table A10 (Low WGS)

Table A11 (PROX)

24

Table A12 (Separator 2)

Table A13 (Turbine 2)

Table A14 (Heat Exchanger 3)

Table A15 (Fuel Cell)

25

Table A16. Utilities and Costs based on F for carbon steel (3.74) and stainless steel (3.2)

26