2006 project tendernsw/chbe446/ammonia_prj.pdf · 104 exp 9.1 104 91,000 0.5 1.5 1.3 1010 exp 1.4...

054402 – Design and Analysis – Course Project Definition 2006

© D. R. Lewin 2006

2006 Project Tender

Your group is being offered the opportunity to take over the management of a design of a process for the manufacture of ammonia from natural gas for International Chemicals Incorporated (the other ICI), after the termination of the services of the previous company working for ICI, Emek Projects Ltd (EPL). We have been approached by Mr. Zvi Fuhrer of the Haifa Municipality to carry out a feasibility study for the possible manufacture of ammonia from natural gas (possibly from Egypt), with the idea that the plant be constructed in the Negev desert. Your design should produce ammonia with a purity of at least 98 mol %, satisfying a demand at the market price of 22 cents/kg, assuming a basis of 12,000 kg/hr of methane as feed, at a cost of 10 cents/kg. The main byproduct of the process is CO2, which has a market value of 10 cents/kg provided that it can be produced at a purity of at least 98 mol %.

The process must be designed to maximize its profitability, quantified in terms of the following measures: (a) Return on Investment (ROI); (b) Approximate Payback Period (PBP); (c) Venture Profit (VP), which estimates the annual profit above the return of 20% interest to our venture profit partners. Figures 1 and 2 show the design proposed by EPL, which has a ROI of 3.9%, a PBP of about 25 years and a VP of −$11.3M, that is an annual loss of over 11 million dollars. The EPL flowsheet can be simulated in UNISIM using the file NH3_PROCESS_V3.usc, which we are supplying to your group. EPL stated in their report that “This poor economic performance is due to the relatively low market price for ammonia. It is impossible to make the process more profitable without a significant increase in the price of ammonia.” We believe that it is possible to make a profit and that poor engineering practice is the reason for the EPL failure, and for that reason, EPL are no longer working with us.

This document provides detailed technical information about the ammonia process, lists the modifications to the flowsheet that are allowed, the product specifications, process constraints, costs of raw materials and selling prices of the products, and equipment and utility costs. Note that a “grass-roots” process is to be constructed, meaning that all costs need to be considered in your analysis.

Your design is to be submitted as a technical report, supported by UNISIM files, emailed to me no later than thursday, 8th February 2007, 12:00. The specifications for the report are also given in this document.

Sincerely,

Prof. Daniel R. Lewin

Vice President and Chief Technology Officer (CTO)

International Chemicals Incorporated

Project_2006_v10.doc 13-12-06

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The rest of this document provides the following information:

a) A complete description of the ammonia process.

b) Hints to get you started.

c) Modifications to the flowsheet that are allowed and modification that are not acceptable.

d) Definition of profitability measures.

e) Product specifications and revenues.

f) Costs of raw materials.

g) Estimated equipment purchase costs.

h) Utility costs.

i) References.

j) Deliverables.


© D. R. Lewin 2006

a) A complete description of the ammonia process. A possible route to ammonia is from natural gas (largely methane). The process involves two main parts: the synthesis gas generation section and the ammonia synthesis loop section. Synthesis gas generation (see Figure 1): The objective of this section is to produce as much synthesis gas as possible, and to ensure its purity. The specifications for synthesis gas are: (a) a molar ratio of hydrogen to nitrogen of 3 (ideally, this ratio needs to be 3:1 in the NH3 converter feed); (b) no water; (c) CO and CO2 under 1 ppm each; (d) minimum inerts (Argon and CH4). To achieve these objectives, the following steps are employed:

a) The methane is combined with reformer steam, preheated in the furnace E-101 and then fed to the reformer, in which most of the methane is converted to hydrogen. The reformer is actually a furnace, in which the reaction mixture flows in tubes arranged on the furnace wall. It is modeled in UNISIM as an isothermal PFR (the effluent temperature is set to be equal to the feed temperature). In the EPL design, the operating temperature is selected as 700 oC. The following reactions take place in the reformer:

4 2 2CH H O 3H CO+ + (1)

2 2CO H O CO H2+ + (2)

According to Parisi and Laborde (2001), reaction rates for these two reactions are as follows:

[ ] 2

4 4 2

33

1, 1exp [kgmol/m s]27464exp 30.707

CO HCH CH H O

P Pr k E RT P P

T

∞

⎛ ⎞⎜ ⎟⎜ ⎟− = ⋅ − ⋅ − −

−⎡ ⎤⎜ ⎟+⎜ ⎟⎢ ⎥⎣ ⎦⎝ ⎠

(3)

[ ] 2 2

2

32, 2exp [kgmol/m s]

4048exp 3.765

CO HCO CO H O

P Pr k E RT P P

T

∞

⎛ ⎞⎜ ⎟⎜ ⎟− = ⋅ − ⋅ − −

⎡ ⎤⎜ ⎟−⎜ ⎟⎢ ⎥⎣ ⎦⎝ ⎠

(4)

Note that in the above equations, the species partial pressures are expressed in atm, T is the temperature in K, and that Eq. (4) holds for T > 860 K. Parisi and Laborde (2001) provide kinetic parameters as follows: 1 2 16,000 kJ/kgmolE E= = , 3

1, 200 kgmol/m sk ∞ = − and 3

2, 100 kgmol/m s.k ∞ = −

b) The reformer effluent is combined with air and more steam in such a way as to try to ensure a 3:1 mixture of hydrogen and nitrogen in the resulting synthesis gas. This mixture is reacted in the oxidation reactor, often referred to a “secondary reformer,” modeled in UNISIM as an adiabatic PFR, where the oxygen in the air generates additional hydrogen. In addition to reaction (1) above, the following reaction also take place in the oxidation reactor:

4 2 2 2CH 2O CO 2H O+ → + (5)

According to Wolf et al (1997), the reaction rate for the above reaction takes the mathematical form:


© D. R. Lewin 2006

[ ]( )

4 2

4

4 4 2 2 2 2 2 2

3, 3 32

exp [kgmol/m s]

1CH O

CH

CH CH O O CO CO H O H O

k E RT P Pr

K P K P K P K P∞ ⋅ − ⋅

− = −+ + + +

(6)

Note that in the above equations, the species partial pressures are expressed in kPa and T is the temperature in K. Wolf et al (1997) provide kinetic parameters as follows:

( )( )( )

( )

4 4 4

2 2 2

2 2 2

2 2 2

6

2

4

1.1 10 , 32,200 kJ/kgmol

1.1 10 , 28,400 kJ/kgmol

1.5 10 , 32,900 kJ/kgmol

5.3 , 27,300 kJ/kgmol

CH CH CH

O O O

CO CO CO

H O H O H O

K E RT E

K E RT E

K E RT E

K E RT E

−

−

−

= × ⋅ − =

= × ⋅ − =

= × ⋅ − =

= ⋅ − =

The two remaining principal kinetic parameters were selected to be: and

3 32,000 kJ/kgmolE =3

3, 1,000 kgmol/m s.k ∞ = −

c) Since the first two reaction steps also generate CO, which would poison the ammonia synthesis catalyst, shift reaction steps are employed to convert the CO to CO2. Both of these reactors are modeled in UNISIM as adiabatic PFRs. In the EPL design, the first shift reactor, HT shift, is fed the oxidation reaction effluent. It is possible to install a heat exchanger to modify its inlet temperature, as is done in the EPL design for the LT shift, where the heat exchanger E-102 reduces the feed temperature to 500 oC. Note that E-102 is used to generate HT steam generating a credit (see the spreadsheet Economic Evaluation). The following reaction takes place in the shift reactors:

2 2CO H O CO H2+ + (2)The same kinetic form is used as before, but, as stated in Parisi and Laborde (2001), the kinetic parameters are slightly different, since the shift reaction is carried out at lower temperatures:

[ ] 2 2

2

32, 2exp [kgmol/m s]

4577exp 4.33

CO HCO CO H O

P Pr k E RT P P

T

∞

⎛ ⎞⎜ ⎟⎜ ⎟− = ⋅ − ⋅ − −

⎡ ⎤⎜ ⎟−⎜ ⎟⎢ ⎥⎣ ⎦⎝ ⎠

(7)

Note that in the above equation, the species partial pressures are expressed in atm, , T is the temperature in K, and that Eq. (7) holds for T < 860 K. Parisi and Laborde (2001) provide kinetic parameters as follows: 2 16,000 kJ/kgmolE = , and 3

2, 100 kgmol/m s.k ∞ = −

d) Any remaining CO is converted back to methane in the methanator, modeled in UNISIM as an adiabatic PFR. In the methanator, reaction in Eq. (1) takes place, with kinetics as given by Eq. (3). The operating temperature needs to be low enough to ensure that the reverse reaction dominates. In the EPL design, the methanator feed temperature is selected as 250 oC.

e) The water produced in the previous reaction steps is removed. One possible implementation, suggested by EPL, involves cooling using E-104 to 40 oC to condense the water in the stream, and its removal largely using the flash vessel, V-100. Residual water is removed using adsorbing beds, modeled in UNISIM using a separator, X-100.

f) The CO2 produced in the previous steps is removed. One possible implementation, suggested by EPL, involves simply cooling the effluent from X-100 using E-105 to -120 oC, condensing


© D. R. Lewin 2006

out most of the CO2 in flash vessel V-101 as by-product, and then removal of the residual CO2 using adsorbing beds, X-101.

Ammonia synthesis loop (see Figure 2): The objective of this section is to produce as much ammonia as possible and to ensure its purity (≥98 mol %). To achieve these objectives, the following steps are employed:

a) The make-up synthesis gas is compressed to the operating pressure of the synthesis loop, which for the EPL design is selected to be 150 bar.

b) The make-up synthesis gas is combined with the recycle stream taken from the vapor stream of the flash separating vessel, V-102.

c) The combined feed is split three ways, with the largest portion entering the ammonia converter via the integrated heat exchanger, E-106, where it is preheated to ignition temperature using the hot converter effluent. In this design, the reacting synthesis gas progresses through three adiabatic PFRs, with intercooling provided by two cold shots, flowing through streams CS-1A/B and CS-2A/B. The reaction that takes place in the adiabatic beds is:

2 20.5 N 1.5 H NH3⋅ + ⋅ (8)

The rate of reaction is given by the following kinetic expression (Seider et al, 2004):

2 2 2 3

91,0004 4 0.5 1.5 10 510 exp 9.1 10 1.3 10 exp 1.4 10RT

N N Hr RT P P−

⎡ ⎤ ⎡− = − × − × − ×⎣ ⎦ ⎣ NHRT P⎤⎦ , (9)

where is the rate of nitrogen disappearance in kmol/m2Nr−

3–s, T is the temperature in K, Pi are the partial pressures of the reaction species in atm, and the activation energies for the forward and reverse reactions are in kJ/kmol.

d) The hot converted effluent is cooled by exchange with the cold feed in E-106, and is further cooled in E-107. The effluent temperature (of stream S-32) is selected to be low enough to ensure a pure enough ammonia product. In the EPL design, this means that the cooler E-107 needs to be cooled with expensive methane refrigerant. The cooled converter effluent is flashed in V-102 to a liquid ammonia product and a vapor stream for further processing.

e) The vapor product of V-102 is split into a small purge stream, which in this design is a waste stream that needs to be treated, and the remainder, which is recycled.

b) Hints to get you started. Your general task is to answer the question: “How can the proposed flowsheet be modified to make it as profitable as possible?” To investigate this question, you need to carry out the following steps: STEP 1. Redesign the flowsheet, by instigating changes that improve profitability. There are a

large number of changes that could be made to the flowsheet, and changes that can be made to key parameter values. As you should already have some experience with this flowsheet, you should be able to make informed decisions, and by imposing appropriate specifications, obtain better results. In this step, do not make changes regarding the heat integration. Study Modifications to the Flowsheet that are Allowed and Disallowed – carefully!


© D. R. Lewin 2006

STEP 2. Design the heat integrated network (HEN) for maximum energy recovery (MER) based on the given ∆Tmin: a) Extract the stream-data. b) Obtain the problem table (manually). You may check this result using the LNG

module in UNISIM if you wish, but this is not mandatory. c) Design a network the meets the MER targets by hand. d) Implement your HEN in UNISIM using the real streams. All the stream-stream heat

exchangers in the HEN must be of the “Shell & Tube” type. The heat exchanger areas required should be computed as described in detail below (See Estimated Equipment Purchase Costs – Heat Exchangers).

e) Modify the design to account for trade-offs between equipment and utility costs to find the best economical solution (∆Tmin free to be chosen as you see fit).

Before beginning work on the project, you need to sign up your group with Mr. Eytan Filiba ([email protected]), who will assign you a group number. Each group number has an assigned value of ∆Tmin for STEP 2, parts (a)-(d), as per Table A.

Table A. Definition of ∆Tmin for STEP 2 by Group.

Group 1 2 3 4 5 6 7 8 9 10 ∆Tmin 5 oC 5.5 oC 6 oC 6.5 oC 7 oC 7.5 oC 8 oC 8.5 oC 9 oC 9.5 oC

Group 11 12 13 14 15 16 17 18 19 20 ∆Tmin 10 oC 10.5 oC 11 oC 11.5 oC 12 oC 12.5 oC 13 oC 13.5 oC 14 oC 14.5 oC

Getting help.

1. ICI staff (Mr. Eran Nahari and Mr. Eytan Filiba) have been allocated only for consulation on technical issues concerning UNISIM usage. In this regard, please note that they have both been instructed not to offer suggestions concerning engineering judgement, which is your responsibility. Note also that the multimedia support materials on UNISIM usage is available on site, and you are expected to review these materials before seeking assistance from Messrs. Nahari and Filiba.

2. Any questions regarding the rules of the project tender, should be addressed only to the CTO of ICI, in writing, by email to: [email protected]. Please see also Modifications to the Flowsheet that are Allowed and Disallowed.

mailto:[email protected]

mailto:[email protected]


© D. R. Lewin 2006

c) Modifications to the Flowsheet that are Allowed and Disallowed.

The following changes are allowed:

1. Changing the operating temperatures and pressures of all unit operations. Operating temperatures in excess of 1,000 oC will require permission in writing from Prof. Lewin.

2. Changing the dimensions of unit operations, and in particular, the vessel diameters and lengths of all PFRs, noting that the vessel sizing will effect their performance.

3. Adding additional heat exchangers, flash vessels, separations devices (distillation columns, membrane separation and HME units), accounting for their equipment and operating costs.

4. Changing the heat management (in STEP 2). The current design calls for seven heat exchangers, of which only one, E-106, performs heat exchange between a hot process stream and a cold process stream. Note that the current design sets a desired effluent temperature of 310 oC, which is the feed temperature of the first adiabatic bed in the ammonia convertor. Your design needs to ensure no temperature crossover in this as well as other heat exchangers, but need not limit this temperature to the same value. The other exchangers are: E-101, which heats the feed to the reformer (currently, this is a furnace); E-102, which cools the feed to the LT Shift reactor to 500 oC; E-103, which cools the feed to the Methanator to 250 oC; E-104, which cools the Methanator effluent to 40 oC; E-105, which cools the feed to V-101 to -120

oC; and E-107, which cools the feed to V-102 to -100 oC. The number of process-process heat exchangers can be changed, through the addition of new heat exchangers, and/or the removal of those in the current design.

Please note the following:

1. SI units are to be used throughout, with the exception of certain spreadsheets that perform costing of vessels, which use field units.

2. You are not permitted to change the number or types of reactors in the synthesis gas section of the process. More specifically, they all have to be adiabatic PFRs with the exception of the Reformer, which is actually a furnace and needs to be represented as an isothermal PFR. For the case of the ammonia converter, note that the actual converter shell includes adiabatic PFR sections with intermediate cold-shot cooling, as well as the heat exchanger E-106. You need to use this design, but can change the number of adiabatic beds (increase or decrease) and their dimensions, provided they all have the same diameter. Note that this authothermal design exhibits multiplicity! For more details see Estimated Equipment Purchase Costs – Ammonia Converter.

3. Pressure drops in all unit operations are neglected in this design.

d) Definition of profitability measures.

The gross profit is computed as follows:

, ,i RM i WA ii i i

GP R C C COS= − − −∑ ∑ ∑ (10)


© D. R. Lewin 2006

where GP is the gross profit in $/year, Ri is the revenue on product i in $/year, CRM,i is the cost of raw material i in $/year, CWA,i is the cost of disposing of waste stream i in $/year, and COS is the annual cost of sales, which accounts for the costs associated with utilities and labor:

,U ii

COS C L= +∑ (11)

where CU,i is the cost of utility i in $/year, and L is the total annual cost of paying salaries to staff. For our process, we estimate that six operators, one lab technician, and one control technician will be needed per shift, with five shifts required. Following Seider et al (2004), pp. 574-576, we estimate L = $ 2,775,000/year.

Obviously, we are interested in designing a process that maximizes GP, while minimizing the total capital investment (TCI), computed by summing the f.o.b. equipment purchace costs and multiplying by a factor, F, that includes the additional costs associated with equipment installation, storage and utility facilities, and other contigencies:

,CP ii

TCI F C CAT= ⋅ +∑ (12)

where CP,i is the f.o.b. purchase cost of equipment item i in $, the term CAT accounts for the one-time charges such as the cost of catalysts and membranes, and the factor F = 5.38 is selected following the example in Seider et al (2004), pp. 496-497, which applies directly to an ammonia process of comparable size.

The revenues and raw materials costs are estimated on the basis of 24 hr/day, 330 day/year operation of the process, and using the material balances obtained from UNISIM, and the itemized materials listed in sections (e) and (f) below. Methods for the estimation for equipment purchase costs are itemized in section (g). Finally, the cost of utilities is estimated using Table C in section (h), which lists the annual cost per 106 kcal/hr or kW for each utility.

The three profitability measures that will be used to assess the design are as follows:

a) ROI – Return on investment, computed as:

( )1100 [%]

GP tROI

TCI−

= (13)

where t is the tax rate (we will use a value of t = 0.25). Note that the above expression does not account for the time value of money. More accurate expressions could be used1. For the assessment of the design, the expression of Eq. (12) suffices. A value of 25% for ROI is considered promising.

b) PBP – Pay back period, approximated by the expression:

( )100 [years]

1TCIPBP

GP t ROI≈ =

− (14)

Note that a more accurate estimate for PBP accounts for depreciation of capital investment1. As stated in Seider et al (2004), it is unlikely that a project with a PBP of more than 4 years would be considered.

c) VP – Venture profit, approximated by:

( )1 [$/year]mVP GP t i TCI= − − (15)

1 See the course 054401 and Chapter 17 of Seider et al (2003) for more details.


© D. R. Lewin 2006

where is the minimum acceptable rate of return payable to venture capitalists that finance the project, take here as Clearly, we are interested in a positive value for VP, and the larger the better!

mi0.2.mi =

e) Products Specifications and Revenues.

Ammonia

Specifications: Ammonia should be supplied in the range 140-150 bar, with a purity of at least 98 mol%, and with a supply temperature under 30 oC.

Revenues: The market will pay $0.22/kg for as much ammonia as you can produce.

Carbon Dioxide (CO2)

Specifications: CO2 should be supplied at a pressure in the range 15-50 bar, with a purity of at least 98 mol%. here are no restrictions on its supply pressure, provided it is a liquid.

Revenues: There is no upper limit in the production rate that can be accomodated. The revenue on CO2 that meets the above specifications is $0.1/kg. In the event that any of the specifications are not met, the gas is considered a waste stream that needs to be safely disposed of, at a cost of $0.05/kg2.

Water

Specifications: Water can be reused in the process, provided its purity is greater than 99.5 mol%, and at at temperature less than 50 oC. There are no restrictions on its supply pressure, provided it is a liquid.

Revenues: There is no upper limit in the production rate that can be accomodated. While product that meets the above specifications does not generate revenue, note that in the event that any of the specifications are not met, this stream is considered a waste stream that needs to be safely disposed of, at a cost of $0.05/kg2.

f) Costs of Raw Materials.

Methane

Specifications: Up to 12,000 kg/h (288 T/day) of pure methane is available at 15 bar and 30 oC, at a cost of $0.1/kg.

Steam

Specifications: An unlimited supply of 15 bar saturated steam is available, at a cost of $0.02/kg. This needs to be supplied both to the reformer (Reformer Steam) and the oxidation reactor (Combustion Steam).

Air

Specifications: An unlimited supply of air is available for free, at 1 bar and 30 oC. The composition of air is: 78.08 mol % nitrogen, 20.95 mol % oxygen, 0.94 mol % argon and 0.03 mol % carbon dioxide.

2 Note that the waste disposal unit can accept streams at any temperature, pressure and composition.


© D. R. Lewin 2006

g) Estimated Equipment Purchase Costs

The purchase costs of the following items need to be included: furnaces, reactor vessels, all heat exchangers, separation columns, flash vessels and compressors. Note that the purchase costs of pumps, which are considered negligable, are not included in the capital investment estimate. The following pages give details on how the purchase costs each of the chargable items are estimated. Note that all of these costing expressions have been implemented in the UNISIM file NH3_PLANT_V3.usc, either in the spreadsheet Economic Evaluation or in other spreadsheets as indicated.

A word on materials of construction: Since the entire system contains hydrogen, it is clear that an alloy steel or stainless steel must be selected as the material of constuction throughout, since carbon steel is not resistant to attack by hydrogen. For simplicity, we shall use stainless steel for all items at this stage, noting that cheaper alloy steels may suffice in certain circumstances (see next semester).

Furnaces (in the current design – E-100, the reformer feed preheater, and the reformer itself)

These are gas-fired furnaces in which the process fluid is fed to an array of pipes arranged on the internal walls of the chamber. The purchase cost of a furnace is estimated using the expression:

0.810.677 [$ f.o.b.]FURNC Q= (16)

where Q is the furnace duty in Btu/hr. For the reformer, we need to also account for the cost of the catalyst, at $70,000/m3 catalyst.

Low pressure reaction vessels (in the current design - Oxidation reactor, HT shift, LT shift, and Methanator).

These are all packed bed reactors, and designed as horizontal vessels. The purchase costs include the cost of the pressure vessels and platforms and ladders (to allow operator access), as well as for the catalyst packing. The purchase cost for a horizontal pressure vessel is estimated using the expression:

( ) ( ) 2, exp 8.717 0.2330 ln 0.04333 ln [$ f.o.b.]V HOR MC W W⎡ ⎤= − ⎡ ⎤ + ⎡ ⎤⎣ ⎦ ⎣ ⎦⎣ ⎦

F (17)

where FM is the material factor (2.1 for stainless steel) and W (in lb) is the weight of the pressure vessel, which is estimated using:

( )( )0.8 [lb]i S i i SW D t L D t= π + + ρ (18)

where Di and Li are the vessel diameter and length, respectively (in inches), ρ is the density of steel (0.284 lb/in3) and tS is the shell thickness (in inches), computed using:

[inch]2 1.2

d iS

d

P DtSE P

=−

(19)

where the allowable stress is taken as S = 13,750 psi, the weld efficiency is taken as E = 0.85, and the design pressure, Pd (in psig), is estimated using the expression in Seider et al (2003), pg. 529:

( ) ( ) 2exp 0.60608 0.91615 ln 0.0015655 ln [psig]d OP P⎡ ⎤= + ⎡ ⎤ + ⎡ ⎤⎣ ⎦ ⎣ ⎦⎣ ⎦OP (20)


© D. R. Lewin 2006

where PO is the vessel operating pressure (in psig). Each vessel is also equipped with platforms and ladders, whose costs are estimated using:

0.202941,580 [$ f.o.b.]PLC D= (21)

where D is the vessel diameter (in feet). Finally, the cost of the catalyst packing in each reactor vessel (including the reformer) needs to estimated. Table B summarizes the cost of each type of catalyst.

Table B. Cost Data for Catalysts. Catalyst Type Cost ($/m3)

Reformer 70,000

Oxidation Reactor 7,000

Shift Reactors 7,000 Methanator 100,000

Ammonia Synthesis 100,000

See the spreadsheet LP Reactor Vessels Costing.

Ammonia Converter (in the current design, this comprises of E-106, PFR-100, PFR-101 and PFR-102, all of which are inserted in a single, high pressure shell).

Figure 3 shows two common designs for ammonia synthesis converters. As can be clearly seen, these are designed to be auto-thermal reactors, that is, they incorporate an integrated heat exchanger designed to transfer heat between the hot reactor effluent and the cold reactor feed. Most modern designs involve multiple adiabatic beds, with cold-shot cooling. The EPL design, involving three beds, is an example of this set up.

(a) Four-bed axial flow design (b) Haldor-Topsøe converter with radial flow

Figure 3. Typical commercial designs for ammonia synthesis converters.

The estimate of the purchase cost for the ammonia converter is carried out as fllows. First the height of the pressure vessel is computed:


© D. R. Lewin 2006

( )1 [ft]iPFR HEX

i

L f L L⎡ ⎤= + +⎢ ⎥

⎣ ⎦∑ (22)

where is the length of the ith adiabatic bed, iPFRL HEXL is the length of the heat exchanger (assumed to

be 10 m = 19.69 ft) and f is a factor to allow for internal piping and other contructs (taken to be 0.3). For the EPL design using the above equation, the reactor height is L = 64.3 ft, and the diameter, D = 3.281 ft (1 m). The vessel thickness and weight are computed using Eqs. (19) and (18) respectively. The cost of the vertically positioned pressure vessel is computed using Seider et al (2004), pg. 527:

( ) ( ) 2, exp 6.775 0.18255 ln 0.02297 ln [$ f.o.b.]V VER MC W W⎡ ⎤= + ⎡ ⎤ + ⎡ ⎤⎣ ⎦ ⎣ ⎦⎣ ⎦

F

F

(23)

where FM is the material factor (2.1 for stainless steel) and W (in lb) is the weight of the pressure vessel. The converter is equipped with platforms and ladders, whose costs are estimated using:

0.7396 0.70684285.1 [$ f.o.b.]PLC D L= (24)

Finally, the cost of the total catalyst packing is estimated (using data in Table B). For complete details of the calculation, see the spreadsheet NH3 Converter Costing.

Allowing for deactivation. Ammonia synthesis catalyst undergoes slow deativation during the operating cycle of a n ammonia plant, and to compensate for this, we will assume that the pre-exponential factors of the forward and reverse rates can be reduced by as much as 50% of their intial values. The ammonia converter should be able to perform adequately also in these conditions.

Distillation Columns (none installed in the current design).

The purchase cost of these, and any additional columns required by your redesign, are estimated using Seider et al (2003), pg. 528:

( ) ( ) 2exp 7.0374 0.18255 ln 0.02297 ln [$ f.o.b.]COL MC W W⎡ ⎤= + ⎡ ⎤ + ⎡ ⎤⎣ ⎦ ⎣ ⎦⎣ ⎦

(25)

where CCOL is purchase cost of the distillation column in $ (excluding condenser and reboiler, whose purchase cost is estimated as with heat exchangers), FM is the material factor (2.1 for stainless steel) and W (in lb) is the weight of the pressure vessel, estimated using Eq. (18). The diameter of the column is estimated using the UNISIM Tray Sizing Utility. The column height in ft, is computed as the product of the number of actual trays, nta (see below), and the tray spacing, usually taken as 2 ft. The shell thickness accounts both for the operating pressure and wind stress (see below). The number of actual trays, nta, is computed by estimating the tray efficiency for the column, using the O’Connel correlation: nta, = E0·nt, where nt is the number of ideal trays and the tray efficiency, E0 , is given by:

, with the viscosity, ( ) 0.2450 0.492 LE −= µ α Lµ , and relative volatility, α, computed as geometric

averages between the top and bottom trays in the column. In addition to the cost of the pressure vessel for the tower, we also need to account for the cost of platforms and ladders and of the trays following Seider et al (2003), pp. 528 and 532):

0.63316 0.80161237.1 [$ f.o.b.] ( and in feet)PLC D L D L= ⋅ ⋅ (26)( )exp 0.1739 [$ f.o.b.] ( in feet)BT taC n D D= ⋅ (27)


© D. R. Lewin 2006

Shell thickness: To estimate the shell thickness for a tower, we need to account for both the effect of vessel pressure and the need to withstand wind loads. Following the procedure suggested by Mulet et al. (1981), for positive design pressures (i.e., not for vessels operating under vacuum), the wall thickness to withstand the design pressure is given in Eq. (19), with the vessel design pressure estimated using Eq. (20). For tall vertical vessels such as distillation columns, it is necessary to account for wind loads. The thickness necessary to withstand wind load is calculated assuming that the wind acts with a uniform intensity over the entire height of the vessel. Assuming a wind velocity of 140 miles/hr, the required thickness to withstand the wind load is:

( ) 2

2

0.22 12 18[inch] W

D Lt

SD+

= (28)

where D is outside shell diameter (ft), L is vessel height (tangent to tangent length, in ft), and the factor of 18 allows for the column cage ladders, which adds additional effective diameter to the column. When there is wind load, the girth seam must withstand the combined load of the wind and the internal pressure, the latter computed using:

[inch]2 0.4

d iG

d

P DtSE P

=+

(29)

The thickness of the bottom of a vertical vessel is then given by: B W Gt t t= + (30)

To estimate the vessel thickness (assumed constant), use the average of the top and bottom thicknesses, plus the corrosion allowance, tC, usually 0.125". Thus the values of wall thickness are computed as follows:

( )0.5SC B S Ct t t= + + t (31)

Flash Vessels (in the current design: V-100, V-101 and V-102).

Flash vessels are generally designed to be nominally half-full of liquid and to allow for a liquid residence time of 10 minutes. Thus, the required vessel capacity is:

10 min0.5 60 min/hr 3Q QV = × = (32)

where Q (m3/hr) is the volumetric flow of liquid in the flash liquid effluent. Assuming a vertical cylindrical vessel of diameter D

L (m) and height L (m), and assuming the standard geometry with a

vessel height three times the diameter, the vessel height is determined as:

( )2 1 32 3 364 4

L LD L VV Lππ ⎛ ⎞= = ⇒ = ⎜ ⎟π⎝ ⎠

(33)

In special cases, different geometries can be used. Eq.(19) is used to estimate the vessel wall thickness, and Eq. (18) is used to compute the vessel weight. The cost of vertical pressure vessels, such as flash drums, is estimated using Eq. (23), with the cost of platforms and ladders computed using Eq. (24). The cost of a demister, , is estimated by assuming a mesh of diameter equal to the vessel, with a height of 1/2 ft, and a cost of $500/ft

DMC3. For complete details for the EPL design, see the spreadsheets

V-100 Costing, V-101 Costing, and V-102 Costing.


© D. R. Lewin 2006

Heat Exchangers (in the current design: E-101 to E-107).

These, and any additional heat exchangers required by your redesign, are tube-in-shell models, with purchase costs estimated using Seider et al (2004), pg. 523:

( ) ( ) 2exp 11.0545 0.9228 ln 0.09861 lnHEX M PC A A F F⎡ ⎤= − ⋅ + ⋅ ⎡ ⎤⎣ ⎦⎣ ⎦

(34)

In the above, HEXC is the cost of the heat exchanger, A is the heat exchanger surface area, in ft2, FM is

a factor accounting for materials of construction and FP is a factor accounting for operating pressure. Following Seider et al (2004), the FM is computed for stainless steel shells and tubes using the expression:

( )0.072.7 100MF A= + (35)

For all exchangers in the synthesis gas section, FP = 1, while for the high pressure synthesis loop, FP = 2. The heat transfer area is computed (in m2) using:

2 Pay attention to m units here!lm

QAU T∆

←= (36)

In the above, Q is the heat exchanger heat transfer duty in kcal/hr, U is the heat transfer coefficient, which can be taken at 250 kcal/m

2 oC for all applications, and lmT∆ (in oC) is the log mean temperature

difference, which for counter-current heat transfer, is computed as:

( ) ( ), , , ,

, ,

, ,

ln

h in c out h out c inlm

h in c out

h out c in

T T T TT

T TT T

− − −∆ =

⎛ ⎞−⎜ ⎟⎜ ⎟−⎝ ⎠

(37)

Note: In the case of a heater fed by a condensing heat utility stream (e.g., steam) or an evaporating cold utility stream (refrigerants), the utility stream inlet and outlet temperatures are taken as equal to the utility supply temperature. In the case of a cooler fed by non-condensing cold utility streams (cooling water, chilled water, or refrigerated brine), a 20

oC temperature rise is assumed for the utility

stream. Thus, for example, CW enters at 20 oC and exits at 40

oC.

Compressor (in the current design: K-100 and K-101).

These are centrifugal compressors, designed to compress the air fed to the oxidation reactor (K-100), tand to compress the synthesis gas before its injection into the synthesis loop (K-101). The purchase cost of centrifugal compressors can be estimated using:

( )exp 7.2223 0.8 ln [$ f.o.b.]COMP C MC P F= ⎡ + ⋅ ⎤⎣ ⎦ (38)

where PC is the power consumption of the compressor in hp and FM is a factor accounting for materials of construction. For the air compressor, FM = 1, while for all other applications, FM = 2.5 for stainless steel. Gas Turbine (not installed in the current design).

As an option, you may consider the installation of a gas turbine to recover electrical power from a process stream, through reduction in the stream pressure. The adiabatic efficiency of this device is


© D. R. Lewin 2006

assumed to be only 35%. The purchase cost of this equipment item, following Seider et al (2003), pg. 554, is:

0.81420 [$ f.o.b.]TURB T MC P F= ⋅ (39)

where PT is the recovered duty in Hp, and FM = 2.5 for stainless steel. If not used directly in the flowsheet, electricity generated is credited as $185/kW/year.

Special separation units (in the current design X-100 and X-101).

These component splitter units are used to represent separation systems not modelled by UNISIM.

Adsorbing beds (for water and CO2): X-100 and X-101 each model separation systems of two identical adsorbing beds that are intended to remove the small amount of water left in the synthesis gas that leaves the flash unit V-100 as overhead vapor (X-100) and the small amount of CO2 left in the synthesis gas that leaves the flash unit V-101 as overhead vapor (X-101). For the case of X-100, it is assumed that all of the remaining water as well as any remaining oxygen are removed. For the case of X-101, all of the remaining CO2 as well as 99.99% of the remaining CO are removed. These assumptions define the split fractions in the two units. Also, the temperatures of the two effluent streams are assumed to be 40 oC, and the two effluent pressures are set to be equal to the feed pressure. Energy streams is attached to close the material and energy balances (but their small values are ignored). The vessel capacities are computed assuming that one unit will be on-line for an entire shift (8 hours) with a second unit being regenerated at the same time:

8V Q= × (40)

where Q (m3/hr) is the volumetric flow of liquid that needs to be removed from the gas stream. Given V, the dimensions of the two vessels are computed as with a flash drum (see Eq. 33), and the cost of the vessels, platforms and ladders in the same way. Note that we need two such vessels. The one-time cost of adsorbant is computed assuming a cost of $10,000/m3 adsorbant (recall that we need two beds in each case), but the cost of energy for regeneration is neglected (this would be carried out by using hot nitrogen). For complete details for the EPL design, see the spreadsheets X-100 Costing and X-101 Costing.

Membrane Separation of Hydrogen from Synthesis Gas.

The MEDALTM membrane technology3, commercialized by Air Liquide, enables hydrogen to be separated from a mixture containing the other species as permeate, with the remaining gases removed as residue. A schematic of a typical setup, taken from the Air Lique website, is shown in Figure 4. The usage of highly selective polyvinylchloride membranes enables almost perfect separation of the hydrogen, with a recovery of as much as 95%. This technology was not adopted by the EPL team, but you may wish to consider it to recover the hydrogen from the purge gas

3 See http://www.medal.airliquide.com/en/membranes/hydrogen/ammonia.asp

Figure 4. Typical membrane set-up.


stream. Here are guidelines for modeling using UNISIM should you wish to consider implementing this technology:

Modeling in UNISIM: A membrane separation unit is approximated in UNISIM using a component splitter. You may assume 95% of the H2 and 10% of the NH3 is recovered as permeate. Both the permeate and residue stream pressures should be taken as one half of the feed pressure, and you may assume all gases exit the membrane at 0 oC. The component splitter is set up with an energy stream, whose duty is ignored.

Approximate costing: A single unit is required. The membrane is installed in a vessel, assumed to have a void fraction of 0.5, and a vapor residence time of 5 min. Thus, the vessel capacity is:

5 min0.5 60 min/hr 6Q QV = × = (41)

where Q (m3/hr) is the volumetric flow of vapor fed to the unit. The vessel dimensions are determined assuming an appropriate ratio of height to diameter. The calculation of vessel wall thickness, weight and purchase cost of vessel and platforms and la ers is as shown previously for flash vessels. The cost of the membrane is taken as $200,000.

Heat-Mass-Exchange (HME) Technology applie

HME technology4 involves a specially heat-insulated pair of adsorbing vessels that perform their normal role of adsorption while also transferring heat from the hot regenerant stream to the cold process stream from which one or more component is removed by adsorption, as depicted schematically in Figure 5.

In NH3 synthesis we need to preheat the synthesicool the stream coming out of the converter to blargely done by means of a heat exchanger concentration of ammonia in the stream coming therefore essentially independent of the ammoniainstall a HME unit to completely or partially reconverter effluent at the expense of its feed. This cby a few percents that translate into an increase loop. The HOT RICH stream fed to the HME unconsider this technology in their design eitheshould you wish to consider implementing this tec

4 See Lavie, R. “Ammonia Synthesis Enhancement Through Heat-Ma

dd

© D. R. Lewin 2006

Ma

ss

He

at

H M EH M E U nitH o t & R ic hH o t & R ic h

C o ld & L ea nC o ld & L ea n

C o o le d &C o o le d &E n r ic h e dE n r ic h e d

H e a te d & H e a te d & D e p le tedD e p le ted

C

d to Ammonia Synthesis.

s gas directed to the converter and we also need to e able to condense the ammonia. This is normally between the two streams. Also, the maximum out of the converter is limited by equilibrium and concentration in the feed to the converter. If we

place the heat exchanger, then we also enrich the an increase the conversion-per-pass in the converter of 10-20% in ammonia production from the same it should be at least 150 oC. While EPL did not r, here are guidelines for modeling using UNISIM hnology:

ss-Exchange,” Plant/Operations Progress, 6(2),122-126, April 1987

Figure 5. HME schematic.


© D. R. Lewin 2006

Modeling in UNISIM: A HME unit is approximated in UNISIM using the PFD shown in Figure 6, involving a comp-onent splitter, a mixing TEE and a heat exchanger. The adiabatic component splitter X-100 transfers a fraction ηY of the NH3 in COLD LEAN to stream 1, assuming the temperatures of streams 1 and 2 are the same. The energy balance on E-100 is closed assuming a temp-erature of TP for the COOLED ENRICHED stream. Values for ηY and TP are computed assisted by the chart in Figure 7, which shows the relationship between ηY , the separation extent, and ηT, the heat exchange extent achieved in the HME:

F PT

F CW

T TT T

−η =

− (42)

where TF and TP are the temperatures of the HOT RICH and COOLED ENRICHED streams, and TCW is the supply temperature of cooling water. Evidently, from Figure 7, one can obtain a value of ηY of about 0.5 for values 0.35 ≤ ηT ≤ 0.45.

Figure 7. Design chart for HME (from Lavie, 1987).

Approximate costing: Two adsorbing beds are required, with one unit adsorbing ammonia from the COLD LEAN stream stream and releasing the HEATED DEPLEATED stream, and the second being fed the HOT RICH stream and releasing the COOLED ENRICHED stream. The adsobant is packed into a vessel, assumed to have a void fraction of 0.5, and an NH3 residence time 10 min. Thus, the vessel capacity is:

Figure 6. Approximate UNISIM model for HME.


© D. R. Lewin 2006

10 min0.5 60 min/hr 3Q QV = × = (43)

where Q (m3/hr) is the volumetric flow of NH3 transferred from the COLD LEAN stream to the COOLED ENRICHED stream. Given V, the dimensions of the two vessels are computed as with a flash drum (see Eq. 33), and the cost of the vessels, platforms and ladders in the same way. Note that we need two such vessels. The one-time cost of adsorbant is computed assuming a cost of $10,000/m3 adsorbant (recall that we need two beds)

h) Utility Costs.

The table below indicates the annual cost of a unit of enegy consuption rate in $/year, as well as the temperature of the utility, if relevant. For example, the usage of high pressure stem (HPS) implies a hot utility temperature of 350 oC, costing $250,000 per year for each 106 kcal/hr consumed. In contrast, using BFW as a coolant at 350 oC to raise HPS, makes a profit of $250,000 per year for each 106 kcal/hr of coolant duty.

Table C. Data for Utility Streams. Utility Type C/H Supply Temperature Unit of energy consumption rate Annual cost ($/year)

Pump power n/a n/a 1 kW 0

Compressor power n/a n/a 1 kW 185

Air cooling, AC C 30 oC 106 kcal/hr 2×103 Cooling water, CW C 20 oC 106 kcal/hr 6×103 Chilled water, CH C 10 oC 106 kcal/hr 1.8×104

Refrigerated rrine, RB C 0 oC 106 kcal/hr 4×104 Ammonia refrig, AR C -30 oC 106 kcal/hr 1.5×105

Ethane refrig, AR C -90 oC 106 kcal/hr 2.5×105 Methane refrig, MR C -160 oC 106 kcal/hr 5×105 Nitrogen refrig, NR C -190 oC 106 kcal/hr 106

Boiler feed water, BFW C 200 oC 106 kcal/hr -1.5×105 Boiler feed water, BFW C 350 oC 106 kcal/hr -2.5×105 Fuel gas (furnace), FG H n/a 106 kcal/hr 1.15×105

Low press. steam, LPS H 130 oC 106 kcal/hr 105 Inter. press. steam, IPS H 200 oC 106 kcal/hr 1.5×105 High press. steam, HPS H 350 oC 106 kcal/hr 2.5×105

i) References. Lavie, R., Ammonia Synthesis Enhancement Through Heat-Mass-Exchange,” Plant/Operations Progress, 6(2),

122-126, April 1987 (1987). Lewin, D. R., and R. Lavie, “Optimal Operation of a Tube Cooled Ammonia Converter in the Face of Catalyst

Bed Deactivation,” I.Chem. Eng. Symp. Ser., 87, 393-368 (1984). Parisi, D. R. and M. A. Laborde, “Modeling Steady-state Heterogeneous Gas-solid Reactors using Feedforward

Neural Networks,” Comp. and Chem. Engng., 25, 1241-1250 (2001). Seider, W. D., J. D. Seader and D. R. Lewin, Product and Process Design Principles, 2nd Ed., John Wiley and

Sons, New Jersey (2004) Wolf, D., M. Höhenberger and M. Baerns, “External Mass and Heat Transfer Limitations for the Partial

Oxidation of Methane over a Pt/MgO Catalyst – Consequences for Adiabatic Reactor Operation,” Ind. Eng. Chem. Res., 36, 3345-3353 (1997).


© D. R. Lewin 2006

j) Deliverables. 1) Each project should be packed as a zip file with the following initial da and the I.D number of the

student who submits the project. i.e. student with I.D. number 123456789 should submit a file by the name da123456789.zip

2) The zip file should contain: a. Typed engineering report file in MS WORD format. Use the template provided on the website. b. UNISIM files supporting your solutions. c. A simple ASCII/ANSI file “readme.txt” should be provided with the short purpose description

of all relevant files in the packaged (can be prepared by NOTEPAD). For example:

case1.hsc - the solution of the MER for DT=7C case2.hsc -the final (best) solution …and so on.

3) The cover of the report should indicate the names and I.D numbers of the students who submit the project and the group number.

4) No late submission will be allowed. 5) You are allowed to work in groups of up to five students. 6) Each group should work independently. The sharing of files and data between groups is not

allowed.

Submission deadline is Thursday 8th February 2007 at 12:00 Grading: (a) Technical Presentation. Correct report structure (including executive summary, description of work done, results, summary and conclusions, appendices). Care in presentation (neatness, clarity, use of graphics as appropriate). Be concise in your project write-up! [Max. grade = 10]. (b) Part 1. Improvements to the EPL Design. The grade given here will depend on the number of new features implemented successfully and the degree to which they improve the profitability of the design. We are also looking for demonstration of correct engineering practice. This is by far the most important part of this project [Max. grade = 40]. (c) Part 2.a-c. Data Extraction and targeting for HEN Design. Correct interpretation of the effect of phase changes is critical. MER Targets should be estimated both using the problem table by hand, using the value of ∆Tmin assigned to your group. You can use the LNG module, but no extra credit will be given if you do so. [Max. grade = 15]. (d) Part 2.d. MER Design and Refinement. Care should be taken to ensure targets and constraints are satisfied. Examine the existing trade-off between equipment and utility costing to minimize annual cost. This may require you to change the configuration designed up to now and to change the value of ∆Tmin [Max. grade = 10]. (e) Solution Performance. Performance ≡ VP [Max. grade = 25]. The “performance” grade in part (e) given to each group will be computed on the basis of the veture profit attained relative to those of competing solutions of other groups (i.e., the best solution will get 100% of 25, and the worst will get 0% of 25). If the all of the groups provide similar solutions, the grade will be calculated from a comparison of the student solutions to the one obtained by the course staff.

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