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-------
A minimum installed cost of $70,000 (in December 1979) was assumed for
waste gas streams with flows below 500 scfm. The costs of piping or
ducting from the process sources to the 150-ft. duct noted above were
estimated as for flares. Installed costs were put on a June 1980 basis
using the following Chemical Engineering Plant Cost Indices: the
overall index for thermal incinerators; the pipes, valves, and fittings
index for piping; and the fabricated equipment index for ducts, fans,
and stacks. Annualized costs were calculated using the factors in
Table 8-3. The electricity required was calculated assuming a l.SkPa
(6-inch H20) pressure drop across the system and a blower efficiency of
60 percent. The cost calculation procedure is given in Table E-7.
E.4 CATALYTIC INCINERATOR DESIGN AND COST ESTIMATION PROCEDURE
Catalytic incinerators are generally cost effective VOC control
devices for low concentration streams. The catalyst increases the
chemical rate of oxidation allowing the reaction to proceed at a lower
energy level (temperature) and thus requiring a smaller oxidation
chamber, less expensive materials, and much less auxiliary fuel
(especially for low concentration streams) than required by a thermal
incinerator. The primary determinant of catalytic incinerator capital
cost is volumetric flow rate. Annual operating costs are dependent on
emission rates, molecular weights, VOC concentration, and temperature.
Catalytic incineration in conjunction with a recuperative heat exchanger
can reduce overall fuel requirements.
E.4.1 Catalytic Incinerator Design Procedure
The basic equipment components of a catalytic incinerator include
a blower, burner, mixing chamber, catalyst bed, an optional heat
exchanger, stack, controls, instrumentation, and control panels. The
burner is used to preheat the gas to catalyst temperature. There is
essentially no fume retention requirement. The preheat temperature is
determined by the VOC content of gas, the VOC destruction efficiency,
and the type and amount of catalyst required. A sufficient amount of
air must be available in the gas or be supplied to the preheater for
VOC combustion. (All the gas streams for which catalytic incinerator
control system costs were developed are dilute enough in air and
therefore require no additional combustion air.) The VOC components
contained in the gas streams include ethylene, n-hexane, and other
:-25
-------
TABLE E-7. CAPITAL AND ANNUAL OPERATING COST ESTIMATION
PROCEDURE FOR THERMAL INCINERATORS WITHOUT HEAT RECOVERY
ITEM
VALUE
Capital Costs
Combustion Chamber
Purchase cost
Installed cost
Installed cost, June 1980*
Piping & Ducting (from sources
to main incinerator duct)
Installed cost
Installed cost, June 1980b
Ducts, Fans & Stacks (from
main duct to incinerator
and from incinerator to
atmosphere)
Installed cost0
Installed cost, June 1980d
Total Installed Cost, June 1980
Annualized Costs6
Operating labor
Maintenance material & labor
Utilities
natural gas^
electricity 9
Capital recovery"
Taxes, administration & insurance
Total Annualized Cost
from Figure E-4 for Vcc
purchase cost x 4.0
installed cost x 1.047
see Section E.7 for Qw.g. (scfm)
installed cost x 1.206 for piping
installed cost x 1.288 for ducting
from Figure E-5 for Qw>q.;
use $70,000 minimum
installed cost x T.064
sum of combustion chamber,
piping & ducting, and ducts,
fans, & stacks
1200 hr/yr x $18/hr = $21,600
0.05 x total installed cost
(5.639 x ID'4) (% aux) x LHVW _
x Qw.g. db/hr) W'9'
(0.4955) x Qf>g> (scfm)
0.1627 x total installed cost
0.04 x total installed cost
operating labor + maintenance
+ utilities + capital recovery
+ taxes, administration &
insurance
E-26
-------
Footnotes for Table E-7
allpdated using Chemical Engineering Plant Cost Index from December 1979
(247.6) to June 1980 (259.2).
bPiping updated using .Chemical Engineering Plant Cost pipes, valves,
and fittings index from August 1978 (273.1) to June 1980 (329.3).
Ducting updated using Chemical Engineering Plant Cost fabricated
equipment index from December 1977 (226.2) to June 1980 (291.3).
cFrom Figure E-5 for no heat recovery from Enviroscience (Reference 28),
which assumed 150-ft of round steel inlet ductwork with four ells,
one expansion joint, and one damper with actuator; and costed according
to the 6ARD Manual (Reference 29). Fans were assumed for both waste
gas and combustion air using the ratios developed for a "typical
hydrocarbon" and various estimated pressure drops and were costed
using the Richardson Rapid System (Reference 30). Stack costs were
estimated by Enviroscience based on cost data received from one
thermal oxidizer vendor.
Although these Enviroscience estimates were developed for lower
heating value waste gases using a "typical hydrocarbon" and no dilution
to limit combustion temperature, the costs were used directly because
Enviroscience found variations in duct, etc., design to cause only
small variations in total system cost. Also, since the duct, fan,
and stack costs are based on different flow rates (waste gas, combustion
air and waste gas, and flue gas, respectively) the costs can not be
separated to be adjusted individually.
^Updated using Chemical Engineering Plant Cost fabricated equipment
index from December 1979 (273.7) to June 1980 (291.3).
eCost factors presented in Table 8-3.
f[(% aux) x LHVwgj x Qw>g. (Ib/hr) x
(100 lbw.g.)/100(lbw.g.) x (8600 hr/yr) x (lb-mole/17.4 lbn.gj x
(379 scf at 60°F/lb-mole) x (1040 Btu(HHV)/scf at 60°F) x $5.98/106
Btu (HHV) x (106 Btu)/K)6 (Btu).
SElectricity = (6 in. H20 pressure drop) x Qf>g< (scfm) x (8600 hrs/yr)
x (0.7457 kW/hp) x (5.204 Ib/ft2/in. H20) -h [(60 sec/min) x (550 ft-lb/
sec/hp) x (0.6 kW blower/1 kW electric) x $0.049/kWh].
h!0 percent interest (before taxes) and 10 yr. life.
E-27
-------
easily oxidizable components. These VOC components have catalytic
Ignition temperatures below 315°C (600°F). The catalyst bed outlet
temperature is determined by gas VOC content. Catalysts can be operated
up to a temperature of 700°C (1,300°F). However, continuous use of
the catalyst at this high temperature may cause accelerated thermal
aging due to recrystallization.
The catalyst bed size required depends upon the type of catalyst
used and the VOC destruction efficiency desired. About 1.5 ft3 of
catalyst for 1,000 scfm is required for 90 percent control efficiency
and 2.25 ft3 is required for 98 percent control efficiency.31 As
discussed earlier many factors influence the catalyst life. Typically
the catalyst may loose its effectiveness gradually over a period of
2 to 10 years. In this report the catalyst is assumed to be replaced
every 3 years.
Heat exchanger requirements are determined by gas inlet temperature
and preheater temperature. A minimum practical heat exchanger efficiency
is about 30 percent. Gas temperature, preheater temperature, gas dew
point temperature and gas VOC content determine the maximum feasible
heat exchanger efficiency. A maximum heat exchanger efficiency of 65
percent was assumed for this analysis. The procedure used to calculate
fuel requirements is presented in Table E-8. Estimated fuel requirements
and costs are based on using natural gas, although either oil (No. 1
or 2) or gas can be used. Fuel requirements are drastically reduced
when a heat exchanger is used. Total heat requirements are based on a
preheat temperature of 600°F. A stack is used to vent flue gas to the
atmosphere.
E.4.2 Catalytic Incinerator Cost Estimation Procedure
The capital cost of a catalytic incinerator system is usually
based on gas volume flow rate at standard conditions. For'catalytic
incineration, 70°F and 1 afm (0 psig) were taken as standard conditions. '
The operating costs are determined from the gas flow rate and other
conditions such as gas VOC content and temperature. Table E-9 presents
the basic gas parameters required for estimating system c.osts.
As noted earlier, equipment components of a catalytic incineration
system include blower, preheater with a burner, mixing chamber, catalyst
bed, an optional heat exchanger, stack, controls, and internal ducting
E-28
-------
Table E-8. OPERATING PARAMETERS AND FUEL REQUIREMENTS
OF CATALYTIC INCINERATOR SYSTEMS
Item
Source of information or calculation
Waste Gas Parameters
(1) Flow rate (Q2)» scfm
(2) Amount of air present in
the gas, scfm
(3) Amount of air required
for combustion at 20%
excess, scfm
(4) Net amount of additional
air required (0.3), scfm
(5) Total amount of gas to be
treated (0.4), scfm
(6) Waste gas Temperature at
the inlet of PHR&, °F
(7) Waste gas temperature at
preheater outlet or
catalyst bed inlet, °F
(8) Temperature rise in the
catalyst bed, °F
(9) Flue gas temperature at
catalyst bed outlet, °F
(10) Minimum possible temperature
of flue gas at PHR outlet, °F
(11) PHR efficiency at maximum
possible heat recovery01, %
(12) PHR design efficiency, %
From Table E-9
0, if the waste gas contains VOC and
nitrogen or other inert gas; and
C(l - volume percent VOC) * (volume
percent VOC)] x VOC volume flow (Qj),
scfm, if the waste gas contains VOC
and air
See footnote a.
Item (3) - Item (2); and 0 if
[Item (3) - Item (2)] is negative
Item (1) + Item (4)
From Table E-9
600°F
(25°F/1% LEL) x (%LEL from Table E-9)
Item (7) + Item (8)
See footnote C.
[Item (1) x (Item (7) - 25°F -
Item (6))] * [Item (5) x (Item (9) -
Item (6))]e
See footnote f
E-29
-------
Table E-8. OPERATING PARAMETERS AND FUEL REQUIREMENTS
OF CATALYTIC INCINERATOR SYSTEM (concluded)
Item
Source of information or calculation
(13) Waste gas temperature at
PHR outlet,°F
(14) Amount of heat required by
preheater at additional 10%
for auxiliary, Btu/min
(15) Amount of heat required
for preheater and auxiliary
fuel, 106 Btu/h
(16) Amount of natural gas
required per year, 10^ cfm
0.65 [Item (9) - Item (6)] + Item (6)
Item (5) x [Item (7) - Item (13)] x
[Gas specific heat9, Btu/scf, °F] x
[Item (14) x 60 minutes/hour] x (10%)"
x (106 Btu)/lQ6 Btu
[Item (14) x (8,600 x 60) minutes/year]
x 10-3 r (1,040 Btu/cfm)
aOn volume basis (scfm/scfm): 11.45 for methane, 20.02 for ethane, 28.58 for
propane, 54.31 for hexane, 17.15 for ethylene, and 45.73 for pentane.
Values taken from p. 6-2 in Steam (Reference 18) for 100% total.air and
multiplied by 1.2 for 120% total air or 20% excess air.
^Primary heat recovery unit.
cHeat exchanger should be designed for at least 50°F above the gas dew point.
dThe heat exchanger will be designed for 25°F lower than the preheater
temperature so as to not cause changes in catalyst bed outlet temperature.
eThough the heat recovery to the temperature level of inlet gas is the
maximum heat efficiency possible, in some cases this may not be possible
due to gas dew point condition.
fCost estimates are based on calculated maximum possible heat recovery
up to an upper limit of 65 percent heat recovery.
9Gas specific heat varies with composition and temperature. Used 0.019 Btu/ft3°F
based on average specific heat of air for calculation purpose.
"Auxiliary fuel requirement is assumed to be 10 percent of total.
E-30
-------
TABLE E-9. GAS PARAMETERS USED FOR ESTIMATING CAPITAL AND
OPERATING COSTS OF CATALYTIC INCINERATORSa
ITEM
VALUE
Stream identification
Stream conditions
Temperature,°F
Pressure, psig
VOC content:
Emission factor, kg/Mg
of product
Weight % of total gas
Mass flow rate, kg/h
Ib/h
Organic constituents, wt %
Average mol. wt. (MI), Ibs
Volume flow (Qi), scfm
Heat content (HI),
Btu/scf
Total gas:
Constituents
Mass flow rate, Ib/h
Molecular weight (M2)
Volume flow (Qg), scfm
Air volume flow rate, scfm
VOC concentration (A),
of LEL
Heat content
Btu/total scf
Identify the vent and the polymer
industry from Chapters 3 and 6
(Emission factor, E, kg/Mg) x 1000 Mg/Gg
(Plant production rate, P, Gg/yr) *
(8,600 h/yr)
(kg/h) x (2.205 Ib/kg)
(VOC mass rate, Ib/h) * (60 min./h) }
(Molecular weight (MI), Ibs/lb mole) x
(385 scf/lb-mole at 68°F)= 1.645 (EP/MX)
(174.273)(2.521NC + NH)C
VOC, air and others
(VOC rate, Ib/h) * (wt% of VOC in
gas, Wi/lOOX)
(Gas mass rate, Ib/h) * (60 min/h) •=•
(Gas molecular weight (M2 ), Ib/lb mole)
x (385 ft3 /I b mole) = 1.645
(Total gas flow (Q2), scfm) - (VOC volume
flow (Qi), scfm)
(100) [(Volume flow of VOC, scfm) *
(Volume flow of air, scfm] * LELd
From Chapter
E-31
-------
Footnotes for Table E-9
^Obtain gas parameters from Chapter 3 of the BID, except those to be
calculated.
bCalculate using weight percent values of VOC components.
cif the VOC heating value is not available, calculate it using on heat of
combustion values of 14,093 Btu/lb from carbon converted to C02 and
51,623 Btu/lb from hydrogen converted to water. Nc and NH denote number
carbon and hydrogen atoms in VOC.
dLower explosion levels of ethylene, hexane, methanol, propane, butane,
and pentane are 3.1, 1.32, 7.3, and 2.5, 1.9, and 1.4, respectively.
eTotal gas heat content averages 50 Btu/scf at 100 percent LEL.
E-32
-------
including bypass. Calculations for capital cost estimates are based
on equipment purchase costs obtained from vendors 31,32,33 and application
of direct and indirect cost factors. Table E-10 presents third quarter
1982 purchase costs of catalyst incinerator systems with and without
heat exchangers for sizes from 1,000 scfm to 50,000 scfm. The cost
data are based on carbon steel for incinerator systems and stainless
steel for heat exchangers. The heat exchanger costs are based on
65 percent heat recovery. Catalytic incinerator systems of gas volumes
higher than 50,000 scfm can be estimated by considering two equal
volume units in the system. A minimum available unit size of 500 scfm
was assumed.34'35 The installed cost of this minimum size unit (which can
be used without addition of gas or air for stream flows greater than
about 150 scfm35) was estimated to be $53,000 (June 1980). The heat
exchangers for small size systems would be costly and may not be practical.
Table 8-2 presents the direct and indirect installation cost component
factors used for estimating capital costs of catalytic incinerator
systems. The geometric mean of the two vendor estimates for each flow
rate was multiplied by the ratio of total installed costs to equipment
purchase costs of 1.82 developed for a skid-mounted catalytic incinerator.
Actual direct and indirect cost factors depend upon the plant specific
conditions and may vary with system sizes.
Since the equipment purchase cost presented in Table E-10
represents the third quarter of 1982, the cost data was adjusted to
represent June 1980 by using a cost index multiplying factor of
82.3 percent (based on Chemical Engineering plant cost indices of
259.2 for June 1980 and 315.1 for August 1982). The direct and
indirect capital cost factors were applied to the adjusted purchase
costs and the resultant estimates of catalytic incinerator installed
capital costs as of June 1980 are presented in Figure E-6.
Installed costs of piping, ducts, fans, and stacks were estimated
by the same procedure as for thermal incinerators. Installed costs
were put on a June 1980 basis using the following Chemical Engineering
Plant Cost indices: the overall index for catalytic incinerators; the
pipes, valves, and fittings index for piping; and the fabricated
equipment index for ducts, fans, and stacks.
E-33
-------
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Figure E-6. Installed Capital Costs for Catalytic Incinerators
.-.' With and Without Heat Recovery
E-35
-------
Table 8-3 presents cost bases used for annualized cost estimates.
The operating labor requirement value is based on conversations with
vendors. The capital recovery factor is based on capital recovery
period of 10 years and an interest rate (before taxes) of 10 percent.
(Actually the current tax regulations allow the control system owners
to depreciate the total capital expenditure in the first 5 years.)
Fuel cost is the major direct cost item.
The total annual operating costs are calculated using the cost
bases shown in Table 8-3 and the fuel requirements calculated in
Table E-8. Table E-ll presents a procedure for calculating total
annualized cost estimates of catalytic incinerators.
The amount of catalyst required usually depends upon the control
efficiency. According to a vendor,33 typical catalyst costs are about
$3,000 per ft3. Indirect additional costs involved in replacing the
catalyst every 3 years are assumed to be 20 percent. Therefore, for
98 percent efficient systems, the annual catalyst replacement costs
amount to $2.70/scfm.
Electricity cost calculations are based on pressure drops of
4 in. water for systems with no heat recovery and 10 in. water for
systems with heat recovery, and at 10 percent additional electricity
required for instrumentation, controls, and miscellaneous. Therefore,
at the conversion rate of 0.0001575 hp per inch of water pressure
drop per cubic foot per minute, 65 percent motor efficiency, and $0.049/kWh
electricity unit cost, the total annual electricity costs amount to
$0.335/scfrn for units with no heat recovery (i.e., for 4 in. H20 pressure
drop) and $0.838/scfm for units with heat recovery (i.e., for 10 in. ^0
pressure drop).
E.5 SURFACE CONDENSER DESIGN AND COST ESTIMATION PROCEDURE
This section presents the details of the procedure used for
sizing and estimating the costs of condenser systems applied to the
combined material recovery streams from the continuous process polystyrene
model plant and from the DMT-process polyethylene (terephthalate) model
plant. An outline of the design and costing of the condenser system
for the polystyrene model plant is presented in this section as an
example of the procedure used. Details of either application are given
in the docket.36
E-36
-------
Table E-ll. CAPITAL AND OPERATIONS COST ESTIMATION FOR
CATALYTIC INCINERATOR SYSTEMS
Item
Value
Capital Costs
Incineration system
Installed cost, June 1980
Piping & ducting (from sources
to main incinerator duct)
Installed cost
Installed cost, June 1980a
Ducts, fans & stacks (from main duct
to incinerator and from incinerator
to atmosphere)
Installed costb
Installed cost, June 1980C
Total Installed Cost, June 1980
Annualized Costs
Direct costs
Operating labor
Maintenance material and
labor
Catalyst requirement
Utilities:
Fuel (natural gas)
Electricity
From Figure E-6
See Section E-7 for source flow
rates, scfm.
Installed cost x 1.206 for piping
Installed cost x 1.288 for ducting
From Figure E-5 for waste gas flow
(0,2), scfm; use $70,000 minimum
Installed cost x 1.064
Sum of incineration systems,
piping & ducting, and ducts,
fans, & stacks
$11,200 for systems with no heat
recovery; and $16,700 for systems
with heat recovery
(0.05) x (Total installed capital
cost, $ from Figure E-6)
$2.7 x (Total gas volume flow(Qd)a,
scfm, item 5 from Table E-8) =
($2.7 x Q4)
($6.22/lQ3ft3) x (Amount of natural
gas required, I03ft3, item 16 of
Table E-8)e
($0.335/scfm) x (Total gas volume
flow rate (Q4), scfm, Item 5 from
Table E-8) for units with no heat
recovery; and ($0.838/scfm) x (Total
gas volume flow rate (0.4), scfm,
Item 5 from Table E-8) for units with
heat recovery
E-37
-------
Table E-ll. CAPITAL AND OPERATIONG COST ESTIMATION FOR
CATALYTIC INCINERATOR SYSTEMS (Concluded)
Item
Value
Indirect Costs
Capital recovery
Taxes, insurance and
administrative charges
Total Annualized Costs
(0.1627) x (Total installed capital
cost, $ from Figure E-6)
(0.04) x (Total installed capital
cost, administrative charges
$ from Figure E-6)
Sum of total di rect costs and
total indirect costs
aUpdated using Chemical Engineering Plant Cost Index from December 1979
(247.6) to June 1980 (259.2).
^Piping updated using Chemical Engineering Plant Cost pipes, valves, and
fittings index from August 1978 (273.1) to June 1980 (329.3). Ducting
updated using Chemical Engineering Plant Cost fabricated equipment index
from December 1977 (226.2) to June 1980 (291.3).
GSee footnote c, Table E-7 for discussion on application of these costs
developed by Enviroscience (Reference 28).
^Updated using Chemical Engineering Plant Cost fabricated equipment index
from December 1979 (273.7) to June 1980 (291.3).
eTotal gas flow including waste gas and additional combustion air.
E-38
-------
Two types of condensers are In use in the industry: surface
condensers in which the coolant does riot contact the gas or condensate;
and contact condensers in which coolant, gas, and condensate are intimately
mixed.
Surface condensers were evaluated for the combination of the
following two streams from the polystyrene model plant: the styrene
condenser vent and the styrene recovery unit condenser vent. These
streams may consist of styrene and steam, which are immiscible, or of
styrene in air, a non-condensable. The nature of components present in
the gas stream determines the method of condensation: isothermal or
non-isothermal. The condensation method for streams containing either
a pure component or a mixture of two immiscible components is isothermal.
In the isothermal condensation of two immiscible components, such as
styrene and steam, the components condense at the saturation temperature
and yield two immiscible liquid condensates. The saturation temperature
is reached when the vapor pressure of the components equals the total
pressure of the system. The entire amount of vapors can be condensed
by isothermal condensation. Once the condensation temperature is
determined, the total heat load is calculated and the corresponding
heat exchanger system size is estimated.
The condensation of styrene mixed with a non-condensable, such as
air, can not be considered to be isothermal. Therefore, systems to
condense in the presence of a non-condensable are usually designed
by considering the heat exchanger to be divided into a number of sections,
each with a portion of the total temperature drop. In general, the
condensation of styrene in air is accomplished less readily, and thus
more expensively, than the condensation of styrene in steam.
As new plants were assumed to use vacuum pumps, which result in
styrene-in-air emissions, the costs of condensing styrene in air were
estimated for regulatory alternatives of 90 percent through 98.5 percent
styrene emission reduction. The condenser inlet stream was assumed to
contain saturated styrene in air at 27°C (80°F) since most industry
data showed the emission stream to be near that temperature, and since
the stream was the outlet from previous process or emission control
condensers that would emit at or near saturation. The following general
procedure and assumptions were used in evaluating the condensation
E-39
-------
systems for the combined streams containing styrene in air from the
continuous polystyrene model plant. The presented procedure, however,
is for the entire heat exchanger while the actual procedure36 utilized
an iterative, multiple section analysis.
E.5.1 Surface Condenser Design
The condenser system evaluated consists of a shell and tube heat
exchanger with the hot fluid in the shell side and the cold fluid in
the tube side. The system condensation temperature is determined from
the total pressure of the gas and vapor pressure data for styrene 'in air.
As complete vapor pressure data are not readily available, the conden-
sation temperature is estimated by a regression equation of available
data points^? using the Clausius Clapeyron equation, which relates the
stream pressures to the temperatures. The total pressure of the stream
is equal to the vapor pressures of individual components at the conden-
sation temperature. Once the condensation temperature is known, the
total heat load of the condenser is determined from the latent and
sensible heat contents of styrene and air (see Table £-12). The design
requirements of the condensation system are then determined based on
the heat load and stream characteristics. The coolant is selected
based on the condensation temperature. The condenser system is sized
based on the total heat load and the overall heat transfer coefficient
which is established from individual heat transfer coefficients of the
gas stream and the coolant. An accurate estimate of individual coef-
ficients can be made using such data as viscosity and thermal conductivity
of the gas and coolant and the standard sizes of shell and tube systems
to be used.
The styrene-in-air refrigerated condenser systems were designed
according to procedures for calculating shell-side,40 tube-side,41
and condensation^ heat transfer coefficients, mass transfer coef-
ficients,^ and, finally, an overall heat transfer coefficient for
condensation in the presence of a noncondensable using a multiple
section analysis.44 Heat exchanger45»46 an(j refrigeration unit4?
characteristics were developed from vendor information in conjunction
with information given primarily in the Chemical Engineers'
Handbook.40,48 Refrigerant characteristics were taken primarily from
the Chemical Engineers' Handbook49'50 and publications of the American
E-40
-------
Table E-12. PROCEDURE TO CALCULATE HEAT LOAD
OF A CONDENSATION SYSTEM FOR STYRENE IN AIR
Item
Value
Heat exchanger type
Source identification
Source production capacity
(CAP), Gg/yr
Source emission factor (E),
kg VOC/Mg product
Desired mass emission reduction,
(% Red'n), %
Gas stream condition
Partial pressure of styrene
at inlet (Pin)
Composition of gas stream
at inlet
Styrene mass fl owrate
(Ws), lb/hrd
Gas stream volumetric
flowrate (V), acfrn6
Gas stream mass flowrate
(W), lb/hrf
Partial pressure of styrene
at outlet (PoutK mm M9
Shell and tube heat exchanger
with hot fluid in the shell
side and the cold fluid in the
tube side
Identify the polymer industry and
the vent from Chapters 3 and 6
From model plant in Chapter 6
From model plant in Chapter 6
From regulatory alternative in
Chapter 6
Assumed saturated styrene in
air at 80°F, 1 atm.
7.959 mm Hga
0.01047 ft3 styrene/ft3 gasb;
0.002767 Ib styrene/ft3
0.2564 x CAP x E
6.022 x W,
4.400 x V
100%-.(I Red'n) x 7.959 „„„ „ 9
TDD 3
Temperature required for reduction
fri \ oi/h
(' out'5 *•
Temperature required for reduction
(Tout). °P
Latent beat change of styrene (Qstv)
Btu/hr1 y
4847.95 * [18.2440 - In (Pout)3
(1.8 x T'out) - 459.67
166.36 x Ws x (% Red'n)
E-41
-------
Table E-12. PROCEDURE TO CALCULATE HEAT LOAD
OF A CONDENSATION SYSTEM FOR STYRENE IN AIR (Concluded)
Item
Value
Average (bulk) gas temperature
i k
Density of air (Pair), lb/ft3
Specific heat of air((cD)al>),
Btu/lb-°F
Sensible heat change of air (Qair)»
Btu/hr
Specific heat of styrene ((c0) ),
Btu/lb-°F , sty
(80 + Tout) * 2
1 * [(0.002517 x Tb) + 1.157]
From API Report 441
V x pal> x (cp) . x (80-Tout) x
K air
60 min/hr
From API Report 44"1
Sensible heat change of styrene (Q'sty)
Btu/hr
Ws x (cp) x (80-Tout)
sty
Total design heat load (Qtot)» Btu/hr
(Qsty + Qair + Q'sty)
^Calculated from Clausius Clapeyron curve fit
(In p = -4847.95 ( 1 \ + 18.2440) of styrene vapor pressure versus
~
temperature data given on p. 3-59 of the Chemi cal Engi neers ' Handbook
(Reference 37) for 80°F (see temperature required for reduction).
DVolume fraction of styrene = 7.959 mm Hg = 0.01047 ft3 styrene/ft3 gas.
760 mm Hg
cAssuming ideal gas:
v = RT = 1545 ft lbf/lbm - °R x §40 JR , 394.13ft3/lb-mole;
F p~ 14.7 lbf/in/ x 144 in^/ft^
styrene content (Ib/ft3 gas) =
0 .010 47 ft3 styrene x Ib-mole x 104.14 Ib styrene
ft3 gas 394.13 ft3 Ib-mole ~
E-42
-------
Footnotes for Table E-12 (concluded)
dCAP.Gg product/yr x 1000 Mg/Gg x E kg VOC/Mg product
86UO nr/yr x 0.4536 kg/lb
eCAP,Gg product/yr x 1000 Mg/Gg x E kg VOC/Mg product = i 544 x CAP x E
8t>uu nr/yr x u.453b kg/lb x u.uoz/67 ib/ft3 x bu min/hr
fV.acfm @ 80°F x «28.9 Ib/lb-mole x 60 mln/hr
394.13 acf/lb-mole @ 80"F
9Using mass emission reduction as initial approximation of volume
(partial pressure) reduction. After outlet temperature is approximated,
the outlet gas molar volume and partial pressure can be calculated and
iterated upon until an acceptable outlet temperature is calculated.
"Solving Clausius Clapeyron curve fit of styrene vapor pressure data
(r2 = 0.99995) referred to in footnote a for temperature.
isiope, m, of Clausius Clapeyron curve fit = - \/R
latent heat of styrene, \= -m x R =
Btu-lb-mole
4847.95 (°K) x 1.9853 cal/g-mole-°K x 1.8 cal/g-mole
104.14 lb/lb-mole
JT,°K = T,°C + 273.15 = 5 (T,°F-32) + 273.15 = 0.5556 T,°F + 255.37 -
?
kFor an ideal gas (pV = mRT/MW), P± = mi/Vj_= T£ ; pa1r = 0.08081b/ft3
P2
at 0°C (ChE Hndbk, p. 3-72) (Reference 38)
Pair @ T,°K = 0.0808 x 273.15°K = 0.808 x 237.15
TT'KU.5556 X T,"F + 255.37
1(cP}
'°'769
air
°-231
where
and
are
specific heats of nitrogen and air, respectively, available by
interpolation from API Report 44, p. 652 (Reference 39).
m(cp) vs T,°F, values are available for interpolation on p. 682
sty
of API Report 44 (Reference 37).
E-43
-------
Society of Heating, Refrigerating, and Air-Conditioning Engineers
(ASHRAE).5l The total required heat transfer area and refrigeration
capacity then were calculated from the total heat load, temperature
difference, and overall heat transfer coefficient, and commercially
available sizes were selected. A tabular procedure for calculating
heat exchanger and refrigeration system size for a single section heat
exchanger and a refrigeration unit is presented in Table E-13.
E.5.2 Surface Condenser Cost Estimation Procedure
Since the gas volumes of the two streams are low, the calculated
required (selected) heat transfer areas are also low (about 1.7 (2.3)
and 4.5 (7.2) ft2, respectively, for 90 and 98 percent reduction of
styrene emissions from a single process line). The purchase costs of the
heat exchanger43,44 anc[ refrigeration systems 45 were estimated from data
provided by vendors. An installation factor of 1.39 (see Table 8-2) was
used to estimate installed condenser costs.
The corresponding required refrigeration capacities for 90 and
98 percent styrene reduction were only 0.056 and 0.080 tons (0.20 and
0.28 kW or 670 and 960 Btu/hr), respectively. For 90 percent reduction,
the required capacity is much smaller than the 0.117 tons (0.412 kW or
1405 Btu/hr) available capacity for the minimum available size
refrigeration unit of 1/4 compressor horsepower (0.186 kW compressor)
and a coolant temperature of 0°F (-17.8°C). For 98 percent reduction,
however, 1/2 compressor horsepower (0.373 kW compressor) was required
to provide sufficient available refrigeration capacity for a coolant
temperature of -30°F (-34.4°C).
Installed costs were put on a June 1980 basis using Chemical
Engineering Plant cost indexes. No additional piping was costed since
the condenser unit is so small (_< 5 in.) that it should be able to be
installed adjacent to the source. Table E-14 presents the procedure
for estimating capital and annual operating costs for condensation
systems.
E.6 ETHYLENE GLYCOL RECOVERY SYSTEMS DESIGN AND COST ESTIMATION PROCEDURE
This section outlines the basis and procedures used to design and
estimate costs of the baseline and regulatory alternative ethylene glycol
recovery systems. The resulting costs for the two systems are presented.
E-44
-------
Table E-13. PROCEDURES TO CALCULATE HEAT TRANSFER
AREA OF A CONDENSATION SYSTEM OF STYRENE IN AIR
Heat exchanger configuration
Source Identification
Coolant temperature, TC,°F
Shell-side heat transfer
Coefficient (h0), Btu
Assume appropriate size unit.3
Identify the polymer industry and
the vent from Chapters 3 and 6
Tout-10, rounded to nearest
multiple of 10°F.
Calculate using procedure in
Chemical Engineers' Handbook,
pp. 10-25 thru 10-28
(Reference 38)b
Coolant
Select chilled water at Tc > 60°F;
50% ethylene glycol - 50% water
solutions at Tc = 40,50°F;
Freon-12 at -40°F < Tc < 30°F;
and Freon-502, at Tr<-50°F.c
Tube-side Reynold's Number (NRe)'
(12 x rH x p ) * JJL
Tube-side heat transfer
Coefficient (), Btu
hr-ft^-T
Calculate using appropriate equations
for forced convection in pipes.6
Coolant flow (Wc), Ib/hr
Temperature change of coolant
Coolant flow (Vc), gpm
Across-tube heat transfer
coefficient (ht), Btu
Calculate for assumed heat exchanger
and cool ant.f
Qtot * (Cp x Wc)
Calculate for assumed heat exchanger.9
Calculate for assumed tube
size."
hr-ft^-T
E-45
-------
Table £-13. PROCEDURES TO CALCULATE HEAT TRANSFER
AREA OF A CONDENSATION SYSTEM OF STYRENE IN AIR (Concluded)
Condensation heat transfer
Coefficient (hc), Btu
hr-ft2-"F
Calculate using procedure in
Applied Process Design for
Chemical and Petrochemical Plants,
vol. 3, p. so (Reference 42).'
Overall tube-side heat
transfer coefficient (hj)
C(l/hc)
Mass transfer coefficient (Kg).
1 b- mole
hr-ft2-mra Hg
Calculate using procedure in
Applied Design for Chemical and
Petrochemical Plants. Vol. 3,
pp. 100, 101, 104 (Reference 43).
Clean overall heat transfer
coefficient (Uc), Btu/ft^hr-T
Dirty overall heat transfer
coefficient (Ud), Btu/f^-hr-TJ
Log mean temperature difference (LMTD), °F
Required heat transfer area (A), ft2
Required heat transfer area with 10% safety
margin (A1), ft*
Required refrigeration capacity (RC1), tons1
Selected refrigeration capacity (RC),
compressor horsepower
Horsepower per ton of refrigeration (Hp/ton),
Hp/ton
Calculate from h-j, Kg, and hj
using procedure in Applied Process
Design for Chemical and Petro-
chemical plants, Vol. 3, pp. 100-106
(Reference 44), which iterates on the
condensate film temperature to balance
the tube-side and shell-side heat
transfer within ±5% and then uses
the average to calculate the clean
overall heat transfer coefficient.
[(1/UC) + 0.001 + 0.001]
- AT2) - In
"1
Qtot
x LMTD>k
A x 1.1
if A1 > Aj (available area for
assumed size), try a larger size
heat exchanger (see footnote a)
Qtot r 12,000
Select from vendor information
(such as Reference 47) for design
heat load and coolant temperature
Based on vendor information for
coolant temperature (see Reference 36)
E-46
-------
Footnotes for Table E-13
w»
If no information is available on approximate condenser size required,
assume an overall heat transfer coefficient between 3 and 15 Btu/ft2-hr-°F,
calculate required heat transfer area as noted near end of this table, and
select an appropriate size condenser. Condenser and tube characteristics
for large units (> 34 ft2) can be found in pp. 11-1 thru 11-18 of the
Chemical Engineers' Handbook (Reference 48). Characteristics of smaller
units can be obtained from vendor information.45,46 Characteristics of
units developed for this analysis can be found in the docket.36
Characteristics needed for design calculations:
Tube: outer diameter, D0 (in.); inner diameter, D-j (in.);
thickness, Xw (in.); cross sectional area, Ax (ft'/tube);
specific external surface area, Ae (ft2/ft of tube);
tube side hydraulic radius, r^ (ft) = D-j/(4 x 12 in./ft).
Condenser: shell diameter, Ds (in.);
tube count, NT (no. of tubes);
tube pitch, p (in.);
length, L (ft);
effective length, Leff (ft) = L - (Nts x Lred/12 in/ft)
where Nts = no. of tube sheets (can assume 2)
Lrec| (in.) = reduction in effective length per
tube sheet (can assume = 0.25 in.,
for Ds < 8 in.; = 1.5 in.,
for Ds >. 8 in.);
total tube area, AT (ft') = Ae x Leff x NT.
b
Assumed baffle cut, lc = 0.25 Ds, Ds < 10 in.; = 0.33 Ds, 10 < Ds < 18 in.;
and = 0.45 Ds, Ds 2 18 in.; radial clearance between shell and outer
tube limit = 1/4 in., Ds < 8 in.; = 7/16 in., for 8 < Ds < 25 in.;
= 1/2 in., for Ds 2 25 in.; baffle spacing, ls = 0.6 Ds; diametrical shell-
to-baffle clearance = 0.05 in., for Ds < 4 in.; = 0.075 in., for 4 < Ds
< 7.5 in.; = 0.10 in., for 7.5 < Ds < 14 in.; = 0.125 in., for 14 $ Ds
< 18 in.; = 0.150 in., for 18 < Ds < 25 in.; = 0.30 in., for 25 < Ds
< 42 in.; = 0.35 in. for Ds 2 42 in.; number of sealing strips ^ 0.5 x Nc
GCoolant characteristics can be interpolated or extrapolated for the
coolant temperature, Tc, from the Handbook of Chemistry and Physics,
p. F-36 (Reference 50) and from The Chemical Engineers' Handbook:
pp. 3-71, 126, & 214 (Reference WF Tor water; and pp. 12-46 thru 12-48
(Reference 49) for ethylene'glycol-water solutions; and from Thermophysical
Properties of Refrigerants pp. 9 thru 11 for Freon-12 and pp. 105 and 106
for Freon-502 (Reference 51). Characteristics required are dynamic viscosity
( H-C), Ib/ft-hr; density ( PC), lb/ftj; specific heat C(cD)c] Btu/lb-°F;
thermal conductivity (kc), Btu/hr-°F-ft; and specific gravity ( Ych
dimensionless = Pc/62.42 lb/ftj.
E-47
-------
Footnotes for Table E-13 (Continued)
dFor coolant velocity, V = 3 fps (3-10 fps recommended by Kern in Process
Heat Transfer (Reference 52)).
eFrom The Chemical Engineers' Handbook, =pp. 10-12 thru 10-15 (Reference 41)
, / \ 0-14
Assumed / P-b \ =1 since little change in coolant temperature.
\^w/
(1) For turbulent flow (NRe > 10,000) (from Eq. 10-51):
h = 0.023 x V, ft/hr x P lb/ft3 x cp, Btu-1b-°F x (ib °-14
(NRe)0'2
if 0.7 < NPr < 700 & L/D > 60;
(2) For transition flow (2000 < NRe < 10,000) (from Eq. 10-49):
h
0.029 k
-125)Npr1/3 [l WDi)2/3]/!^0-
L vu/ J\RW/
14
(3) For laminar flow (NRe < 2100) (from Eq. 10-40):
h. = 0.465 k Np.1/3 /fibX0-14 x 0.87 (1 + 0.015
rH
where
(NRe x NPr x 4 x rH) f L.
coolant velocity of 3 fps,
Wc, Ib/hr = Ax x NT x 10,800 ft/hr x Pc
9For coolant velocity of 3 fps.
Vc, gpm = Ax x NT x 180 ft/mi n x 7.48 gal /ft3.
where: K^ = thermal conductivity of tube = 64.2 Btu/ft-hr-°F for
Admiralty brass given on pp. 23-49 of The Chemical
Engineers' Handbook (see Reference 53)1
DL = (Do - °i) * In (
""Must assume temperature of condensate film on tubes, which is the
variable that is iterated upon to balance overall shell -side and
overall tube-side heat transfer.
E-48
-------
Footnotes for Table E-13 (Concluded)
JAssuming both shell-side and tube-side fouling coefficients equal
to 0.001 as given for cooling tower water or refrigerants and for
industrially clean gases, clean hydrocarbon vapors, or atmospheric
air in Applied Process Design for Chemical and Petrochemical Plants,
Vol. 3, Tables 10-10 and 10-11 (Reference 54). :
kMore complicated relationships, considering the average and end point
values for individual sections and the entire exchanger, were used for
the multiple section analysis {see References 36 and 44).
E-49
-------
Table E-14. CAPITAL AND ANNUAL OPERATING COST ESTIMATION
PROCEDURE FOR CONDENSERS WITH REFRIGERATION
Item
Value
Capital Costs
Purchase cost, July 1984
Purchase cost, June 1980b
Total Installed Cost, June 1980C
Annual 1 zed Costsd
Operating labor
Maintenance materials & labor
Utilities
Electricity, pumping'
Electricity, refrigeration
Coolant, make-up
Capital recovery1
Taxes, administration
& insurance
Total annualized cost
without recovery credit
Styrene recovery credit-3
Net Annualized Cost
after recovery credit
[(16.816 x AT) + 125.151]
if Freon coolant, x >1.3a
Purchase cost, Sept. 1982 x 0.871
Total Purchase Cost x 1.39
See footnote e
0.05 x total installed cost
6.104 x Yc x Vc
See footnote g
See footnote h
0.1627 x total installed cost
0.04 x total installed cost
Operating labor & maintenance
+ utilities + capital recovery
+ taxes, administration &
insurance
2767 x Ws x (% Red'n. * 100)
Total annualized cost - styrene
recovery credit
E-50
-------
Footnotes for Table E-14
aBased on vendor information (References 45 and 46).
bDeflated using Chemi cal Engi neeri ng Fabricated Equipment Cost Index from
July 1984 (estimated) (334.6) to June 1980 (291.3).
cBreakdown of installed cost factor given in Table 8-2.
dCost factors presented in Table 8-3.
20 to 34 ft2 units: Operating labor cost = 1 hr/wk x 52 wk/yr
x 1.15 (with supervision/without supervision) x $18/hr (including
overtime). For < 20 ft2 units: Operating labor cost = 2 hr/mon
x 12 mon/yr x 1.15 (with supervision/without supervision) x $18/lir
(including overtime).
fUsing Equation 6-2, p. 6-3 in The Chemical Engineers' Handbook
(Reference 55) for V = 3 fps , assuming a pumping height of
of 50 ft. and a pump efficiency of 65%:
50 ft x Vc x Vc x 0.7457 kW
gpm Tt/hp h"p
8600 hr/yr
0.65 pump efficiency
x $0.049/kwh
where Yc = specific gravity of coolant = P~ * 62.42 lb/ft3
(the density of water)
9 RC' x (hp/ton of refrigeration for Tr)
u.85 compressor efficiency x 0.85 motor efficiency
x 0.7457 KW x 8600 hr/yr x $0.049/kwh
np
"For coolant velocity of 3 fps:
For ethylene glycol-water solutions, Freon-12, and Freon-502, assume one
replacement per year of coolant in condenser and refrigeration system
and coolant volume in condenser and refrigeration unit twice that of
condenser alone.
coolant volume, gal = Ax x NT x L x (2 vol. in
cond & refrig/vol. in cond)
= 2 x Ax x NT x L.
For ethylene glycol-water solutions:
cost of coolant = j[Xw ($0.30/1000 gal x 7.48 gal/ft3)]
+ [XEG ($0.27/lb x PEG Ib/ft3)]} x coolant
volume,ft3 x 1 replacement per year
= ($0.00224 Xw + $0.27 X EG pEG)
x annual coolant replacement, ft3/yr.
E-51
-------
Footnotes for Table E-14 (Concluded)
where,
Xyj = volume fraction of water in solution
= volume fraction of ethylene glycol in solution
For Freon-12:
Cost of coolant = (coolant volume,ft3 x 1 replacement/yr
x $1.31/lb) * PC, lb/ftj
For Freon-502:
Cost of coolant = (coolant volume,ft3 x 1 replacement/year
x $3.11/lb) * PC, lb/ft3
For chilled water, assume 0.1% make-up (99.9% recycle):
Cost of coolant = $0.30/1000 gal x Vc,gpm x 60 min/hr
x 8600 hr/yr x 0.001 make-up/total
= 0.1548 x Vc
ilO percent interest (before taxes) and 10 yr. life.
JVL.lb styrene emitted/hr x 8600 hr/yr x (% Red'n in condenser * 100)
x 0.90,fraction of reduction recovered x $0.3575/lb styrene.
E-52
-------
E.6.1 Ethylene Glycol Recovery System Design
The baseline system for a plant producing a low viscosity product
or high viscosity product with a single end finisher was represented
by a system that recovers ethylene glycol (EG) emitted from the
polymerization reactors through use of EG spray condensers and from
the esterifiers through use of reflux condensers. The baseline
system for a plant producing a high viscosity product was represented
by a system that recovers ethylene glycol emitted from the polymerizers
through the use of EG spray condensers on the initial end finishers
and a distillation column on the cooling water tower and from the
estrifiers through the use of reflux condensers.
The regulatory alternative control systems utilized distillation
columns. For those PET plants producing a low viscosity product or a
high viscosity product with a single end finisher, further control of
EG emissions was obtained by the installation of a distillation column
that reduces the EG concentration in the cooling water tower. For
those PET plants producing a high viscosity product with multiple end
finishers, further control of EG emissions was obtained by increasing
the flow rate of cooling water to the existing distillation column,
which results in a lower EG concentration in the cooling tower.
The equipment selected to comprise the ethylene glycol recovery
systems, as well as the design and operating parameters, was based on
information provided by industry sources. The industry information
(much of which was considered confidential) was used in conjunction
with standard engineering references such as the Chemical Engineers'
Handbook17, and McCabe and Smith's Unit Operations of Chemical
Engineering56, and engineering judgment.57 These design procedures are
summarized in the footnotes to Tables E-15 and E-16 for the baseline
systems.
E.6.2 Ethylene Glycol Recovery System Cost Estimation Procedure
The cost estimates and their bases are presented in Table E-15
for the baseline ethylene glycol recovery system for PET plants producing
a low viscosity product or a high viscosity product with a single end
finisher and in Table E-16 for PET plants producing a high viscosity
product with multiple end finishers. The costs of the baseline systems
E-53
-------
Table E-15. EG RECOVERY COSTS FOR BASELINE SYSTEM FOR PLANTS PRODUCING
A LOW VISCOSITY PRODUCT OR A HIGH VISCOSITY PRODUCT WITH A SINGLE END FINISHER
(June 1980 dollars)
Item (Number of Item)
Model
Plant
Process
Line
Capital Costs
Spray Condensers (28 @ $17,384
each)
Reflux Condensers (14)
Pumps (28)
Heat Exchangers (42)
EG Recovery System (1)
Refrigeration System (14)
Installed Capital Cost Factor
Total Installed Capital Cost
$486,750*
$243,400b
$23,30QC
$64,300d
$386,3006
$128,700f
4.249
$5,234,000
$1,628,400^
Annual i zed Costs
Operating Labor
Operating Materials
Maintenance Materials and Labor
Electricity
Steam
Water
Taxes, Insurance
and Administration
Capital Recovery
Recovery Credit
Total (Annual Costs -
Recovery Credit)
$173, SOO1
0
$21,70Qk
$86,6501
$917,930"!
$4,650n
$209,400°
$853,100P
($1,218,300)3
$1,048,700
$22, 100 J
0
$3,100J
$12,4000
$13 1,100 J
$66 4J
$65,100
$265,400
( $174,000 )J
$325,900
E-54
-------
Footnotes for Table E-15
aBased on size estimate from Tennessee Eastman and cost estimate from
Missouri Boiler. Two spray condensers plus two spares per line in
model plant.
bGoing to this system replaces feed lines from estifiers to distillation
column (CL-2) with reflux condensers, one per line plus one spare per line
The size and cost were assumed to be the same as for EG spray condensers
on the reactors.
A 1 hp pump is required per
Reference 57, p. 14, for pump sizing.
line at a cost of $831 per pump.
dSee Reference 57, pp. 7-9, for exchanger sizing. Only sizing done
for reactors in which industrial resins are produced. Assume size and
costs for heat exchangers associated with reactors producing textile
resins would be the same. Seven scraped surface exchangers (plus 7
spares) for end finishers at $1,130 each; 7 scraped surface exchangers
for prefinishers (plus 7 spares) at $1,930 each; 7 exchangers (plus
7 spares) for esterifiers at $1,530 each (average of $1,130 plus $1,930).
eCosts obtained from a industry source were considered confidential. A
correction factor of 0.764 was obtained from this source to scale the
costs down to our model plant capacity. Using the total equipment and
steam jet ejector system cost yielded the EGRS cost given.
Reference 57. Seven systems for prefinisher at $9,283 each, and
7 for finishers at $9,100 each.
9See Table 8-2, factors based on sum of piping, insulation, painting,
instruments, and electrical factors equaling 1.12 A for larger capacity
system given in confidential industry information. Does not apply to
refrigeration systems.
"Model plant is comprised of seven process lines. Estimate of total
installed capital cost of equipment for just one process line was
obtained with the following equation:
,0.6
x $5,234,000 = $1,628,400
1 Based on 9,640 man-hours per year at $18 per year. The number of
man-hours has been scaled down from the number of man-hours provided by
a confidential industry source.
JEstimate of cost for a single process line was obtained by the following
equati on:
(I) * ACi
where:
3-j = operating labor, operating materials, maintenance
materials and labor, electricity, steam, and water.
E-55
-------
Footnotes for Table E-15 (Concluded)
^Based on maintenance labor requirements and maintenance materials cost
given in confidential industry information for a larger system.
cost comes from 4 sources: pumps, recovery system, heat exchanger
to chill water, and for the reflux condensers. The 14 pumps at 1 hp use
91,500 kWh per year. Recovery system was estimated to use 412,800 kWh
per year based on confidential industry information for a larger system.
Chilled water is necessary for chilling the spent EG used in the EG
spray condensers and is necessary for reactors producing high tenacity
(high viscosity) industrial resins. The electricity requirement to
maintain the chilled water is 660,100 kWh per year for a plant producing
a high viscosity product. Electricity usage by the exchangers for the
reflux condensers was assumed to be the same as that required by the
spray condensers {i.e., 603,900 kWh per year). Assuming all the lines
in a model plant produce high tenacity (high viscosity) industrial
resins, total electricity usage in the model plant- is 91,500 plus
412,800 plus 660,100 plus 603,900, which is equal to 1,768,300 kWh per year.
Cost of electricity is $0.049/kWh.
mBased on confidential industry information and scaling steam usage in
the EG recovery system and its vacuum system proportionately according
to plant capacity, steam usage was estimated to be 1.34 x 10° Ibs/year.
An estimated additional 14.532 x 106 Ibs of steam per year would be
required over the baseline system by the vacuum system servicing the
polymerization reactors. A cost of $6.18/1,000 Ib of steam was used.
"Based on water consumption of 15.5 x 106 gallons per year (from confidential
. industry information scaled down by proportioning relative capacities).
°Based on 0.04 x Installed Capital Cost.
PBased on a capital recovery factor of 0.163.
QBased on a total EG recovery of 19.5 kg of EG/Mg of product and a
recovery credit of $0.595/kg ($0.27/lbs of EG from Chemical Marketing
Reporter). The 19.5 kg of EG/Mg of product is an increase in the
recovery of ethylene glycol from the polymerization reactors plus
ethyl ene glycol recycled from the estifiers that would otherwise have
to be replaced with fresh feed if the baseline system was used.
E-56
-------
TABLE E-16. EG RECOVERY COSTS FOR BASELINE SYSTEM FOR PET
PLANTS PRODUCING A HIGH VISCOSITY PRODUCT WITH MULTIPLE END FINISHERS
(June 1980 dollars)
Item (Number of Item)
Capital Costs
Spray Condensers (4)
Pumps (4)
Heat Exchangers (4)
Reflux Condensers (4)
Distillation Column, cooling tower
Refrigeration system (2)
Installed Capital Cost Factor
Total Installed Capital Cost
Annuali zed Costs
Operating Labor
Operating Materials
Maintenance Materials and Labor
Electricity
Steam
Water
Taxes, Insurance
and Administration
Capital Recovery
Recovery Credit
Total (Annual Costs -
Model
Plant
$82,640*
3,945&
8,220?
82,640d
661,8006
22,065^
3.249
1,258,800
29,500"
-
3,940i
21, 640 J
5,170k
-
50, 350 1
204,800"!
(233,490)"
81,910
Process
Line
$41,320
1,970
4,110
41,320
568,000
11,030
3.24
866,480
14,750
-
1,970
10,820
2,635
-
34,660
140,980
(116,745)
89,070
E-57
-------
Footnotes for Table E-16
number of spray condensers is based on one per line with one spare
per line and two lines per plant. The cost per spray condenser was
estimated by multiplying the spray condenser cost in Table E-15 by the
relative process line sizes (20 Gg vs. 15 Gg) raised to the 0.6 power
as follows:
$17,384 x /20\°-6 = $20,660
\T57
bTwo pumps plus two spares per process line,
from Table E-15 as follows:
Cost per pump adjusted
$830 x/20\°-6 =
YI57
$986
cTwo heat exchangers plus two spares per process line.
exchanger adjusted from Table E-15 as follows:
Cost per heat
[2 scraped surface exchanges for prefinishers at $1,930 each plus 2
exchangers at $1,530 each for esterifiers] x/20\0-° = $8,220
\T5V
dOne reflux condenser plus one spare per process line. Cost per reflux
condenser assumed the same as the spray condenser in footnote a.
^Based on adjusting size and cost information claimed confidential from
an industry source. In general, the size was adjusted based upon
proportioning the flow to the distillation column directly on plant
capacity and then sizing the distillation column on the basis of the
new flow divided by the flow to the original distillation column and
the result raised to the 0.6 power. The cost was adjusted to June 1980
dollars using a cost index factor of 0.8208 based upon the Chemical
Engineering Plant Cost Index of 320.3 for January 1984 and 262.9 for
June 1980.
per refrigeration system adjusted from Table E-15 as follows:
0.6
$9,283 x
= $11,030
9See Table 8-2. Applies to spray condensers, pumps, heat exchangers,
and reflux condensers. Does not apply to distillation column and
refrigeration system.
hconfidential industry information on operating labor scaled down to
assume 1,640 hours per year at $18 per hour.
''Based on maintenance labor requirements and maintenance materials cost
given in confidential industry information for a larger system.
Assumes $1,330 for materials and 145 hours for maintenance labor.
JElectricity usage per source was assumed to be proportional to plant
capacity. The 4 pumps at 1.33 hp each would use 17,385 kWh/year. The
electricity requirement to maintain the chilled water (refrigeration
E-58
-------
Footnotes for Table E-16 (Concluded)
system) is 212,070 kWh/year. Electricity used by the exchangers for
the reflux condensers was assumed to be the same as that required by
the spray condensers (i.e., 212,070 kWh/year). Total electrical usage
in the model plant for these controls is 17,385 plus 212,070 plus
212,070, which is equal to 441,525 kWh/year. The cost of electricity
is $0.049/kWh.
k Includes only steam used in distillation column. Distillation steam
total equals 836,625 Ib/yr (97,575 Ib/yr to heat the water; 5,853 Ib/yr
to heat the ethylene glycol; and 733,197 Ib/yr to vaporize the
water). Steam cost of $6.18 per 1000 Ib of steam.
1 Total installed capital cost x 0.04.
•"Total installed capital cost x 0.1627.
"Based on total EG recovery of 17.65 Ibs of EG/1,000 Ibs of product
and a recovery credit of $0.15 per Ib. Of the total recovered EG,
about 16.68 Ibs EG/1,000 Ibs comes from the initial end finishers and
0.97 Ib EG/1,000 Ibs product from the distillation column.
E-59
-------
were estimated based on the design estimate developed and standard
engineering procedures.
The costs of the regulatory alternative distillation columns were
derived from confidential cost data provided by an industry source for
a similar system on a larger capacity plant. In general, the distil-
lation columns were sized and costed on the basis of the flow rate
from the cooling tower to the column using the ratio of the target flow
rate divided by the base flow rate and the result raised to the 0.6 power.
Utility costs, which are primarily steam costs, were calculated based
on the flow rate from the cooling tower to the distillation column.
E.7 PIPING AND DUCTING DESIGN AND COST ESTIMATION PROCEDURE
Control costs for flare and incinerator systems included costs of
piping or ducting to convey the waste gases (vent streams) from the
source to a pipeline via a source leg and through a pipeline to the
control device. All vent streams were assumed to have sufficient
pressure to reach the control device. (A fan is included on the duct,
fan, and stack system of the incinerators.)
E.7.1 Piping and Ducting Design Procedure
The pipe or duct diameter for each waste gas stream (individual
or combined) was determined by the procedure given in Table E-17. For
flows less than 700 scfm, an economic pipe diameter was calculated
based on an equation in the Chemical Engineer's Handbook59 and simplified
as suggested by Chontos.60»6l>62 The next larger size (inner diameter)
of schedule 40 pipe was selected unless the calculated size was within
10 percent of the difference between the next smaller and next larger
standard size. For flows of 700 scfrn and greater, duct sizes were
calculated assuming a velocity of 2,000 fpm for flows of 60,000 acfm
or less and 5,000 fpm for flows greater than 60,000 acfm. Duct sizes
that were multiples of 3-inches were used.
E.7.2 Piping and Ducting Cost Estimation Procedure
Piping costs were based on those given in the Richardson Engineering
Services Rapid Construction Estimating Cost System30 as combined for
70 ft. source legs and 500 ft. and 2,000 ft. pipelines for the cost
analysis of the Distillation NSPS.63 (See Tables E-18 and E-19).
Ducting costs were calculated based on the installed cost equations
given in the GARD Manual.64 (See Table E-20).
E-60
-------
Table E-17. PIPING AND DUCTING DESIGN PROCEDURE
Item
Value
(1) Pipe diameter, D
(a) Piping^
For Source Legs:
D (in.) = 0.042 x Q (scfm) + 0.472, for Q < 40 scfm
D (in.) = 0.009 x Q (scfm) + 2.85, for 40 12 in. or Q>700scfir
and Q<60,000 acfm ^______
D (in.) = (0.1915) VQ(acfm), for Q > 60,000 acfm
Select size that is a multiple of 3 inches.
Assumed 70-ft. source leg from each source to the
pipeline. Assumed separate pipelines for large
{ > 35,000 scfm) intermittent streams and for all
continuous streams together. Selected pipeline ; ,
length of 70, 500 or 2,000 ft. if calculated safe
pipeline length within 10 percent of standard length;
if not, selected calculated length between standard
values.
Assumed 70-ft. source legs from each source to the
pipeline. Used duct, fan, and stack cost from
Enviroscience (Reference 28) which assumes a 150-ft.
duct cost based on the GARD Manual (Reference 64).
aEconomic pipe diameter equations from Reference 62 (which is based upon References 5S
and 60 ) .
bFrom continuity equation Q _ IT D2y ; assumed velocity, V, of 2,000 fpm for lower
~~~~
flows and 5,000 fpm for higher flows.
E-61
-------
Table E-18. PIPING COMPONENTS3
Equi pment
Type
Check Valves
Gate Valves
Control Valves
Strainers
El bows
Tees
Flanges
Drip Leg Valves
Expansion Fittings
Bolt and Gasket Sets
Hangers
Field Welds
Pipe Length,
(Schedule 40) (ft)
Number of Equipment Type in Pipe Leg Type
Source
1
4
1
1
8
6
15
1
2
. 15
9
18
70 .
Compressor
1
2
-
1
6
2
10
1
1
12
4
12
20
Pipeline (500 ft)
1
3
1
1
6
2
20
1
5
21
50
40
500
Pipeline (2,000 ft)
1
3
1
1
6
3
35
1
20
38
200
120
2,000
From Reference 62.
E-62
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Costs of source legs were taken or calculated directly from the
tables. Costs of pipelines for flares were interpolated for the safe
pipeline lengths differing by more than 10 percent from the standard
lengths of 70, 500, and 2,000 ft.
E-65
-------
E.8 REFERENCES
1. Kalcevic, V. Control Device Evaluation: Flares and the Use of
Emissions as Fuels. In: Organic Chemical Manufacturing Volume 4:
Combustion Control Devices. U.S. Environmental Protection Agency.
Research Triangle Park, N.C. Publication No. EPA-450/3-80-026.
December 1980. Docket Reference Number II-A-18.*
*
2. Reference 1, p. IV-4.
3. Memo from Sarausa, A.I., Energy and Environmental Analysis, Inc.
(EEA), to Polymers and Resins File. May 12, 1982. Flare costing
program (FLACOS). Docket Reference Number II-B-39.*
4. Telecon. Siebert, Paul, Pacific Environmental Services, Inc. (PES)
with Straitz, John III, National AirOil Burner Company, Inc.
(NAO). November 4, 1982. Availability of Flare Cost Data and
Flaring of High Air-Content Docket Reference Number II-E-59.*
5. Straitz, J.F. III. Make the Flare Protect the Environment.
Hydrocarbon Processing. 56. October 1977. Docket Reference
Number II-I-32.*
6.
7.
8.
10.
11.
12.
Oenbring, P.R. and T.R. Sifferman. Flare Design... Are Current
Methods Too Conservative? Hydrocarbon Processing. 59:124-129.
May 1980. Docket Reference Number II-I-58.*
Telecon. Siebert, Paul, PES, with Keller, Mike, John Zink Co.
August 13, 1982. Clarification of comments on draft polymers and
resins CTG document. Docket Reference Number II-E-18.*
Telecon. Siebert, Paul, PES with Fowler, Ed,
1982. Flare Design and Operating Parameters.
Number II-E-60.*
NAO. November 12,
Docket Reference
Telecon. Siebert, Paul, PES with Fowler, Ed, NAO. November 5,
1982. Purchase costs and Design and Operating Criteria for Steam-
assisted, Elevated Flares. Docket Reference Number II-E-58.*
Telecon. Siebert, Paul, PES with Fowler, Ed, NAO. November 15,
1982. Additional Flare Cost Estimates and Flare Design Criteria and
Procedures. Docket Reference Number II-E-61.*
Telecon. Siebert, Paul, PES, with Knock, Cor, NAO. May 2, 1983.
Steam requirements for intermittent flares. Docket Reference
Number II-E- 68.*
Telecon. Siebert, Paul, PES, with Keller, Mike, John Zink, Co.
May 12, 1983. Steam- and air-assisted intermittent flare guidelines.
Docket Reference Number II-E-69.*
E-66
-------
13. Memo from Senyk, David, EEA, to EB/S Files. September 17, 1981.
Piping and compressor cost and annualized cost parameters used in
the determination of compliance costs for the EB/S industry.
Docket Reference Number II-B-33.*
14. Memo from Mascone, D.C., EPA, to Farmer, J.R., EPA. June 11,
1980. Thermal incinerator performance for NSPS. Docket Reference
Number II-B-4.*
15. Air Oxidation Processes in Synthetic Organic Chemical Manufacturing
Industry - Background Information for Proposed Standards. U.S.
Environmental Protection Agency, Research Triangle Park, N.C.
Draft EIS. August 1981. p. 8-4. Docket Reference Number II-A-26.*
16. Blackburn, J.W. Control Device Evaluation: Thermal Oxidation.
In: Chemical Manufacturing Volume 4: Combustion Control Devices.
U.S. Environmental Protection Agency, Research Triangle Park, N.C.
EPA-450/3-80-026. December 1980. p. 1-1. Docket Reference Number
II-A-18.*
17. Perry, R.H. and C.H. Chilton, eds. Chemical Engineers' Handbook,
fifth edition. New York, McGraw-Hill Book Company. 1973.
p. 8-9. Docket Reference Number II-I-16.*
18. Steam: Its Generation and Use. New York, Babcock & Wilcox Company,
1975. p. 6-10. Docket Reference Number II-I-20.*
19. Memo from P. Siebert, PES to Polymers and Resins File. March 16,
1983. Distillation NSPS Thermal Incinerator Costing Computer
Program (DSINCIN). May 1981. p. 2. Docket Reference Number
II-B-67.*
20. Reference 15, p. 8-13.
21. Reference 16, pp. V-3, V-15.
22. Reference 16, p. III-8.
23. Reference 16, Fig. A-l, p. A-3.
24. Reference 15, p. 8-9.
25. Reference 19, p. 4.
26. Reference 16, p. 1-2.
27. Reference 15, p. G-3 and 6-4.
28. Reference 16, Fig. V-15, curve 3, p. V-18.
29. Neveril, R.B. Capital and Operating Costs of Selected Air
Pollution Control Systems. U.S. Environmental Protection Agency,
Research Triangle Park, N.C. Publication No. EPA-450/5-80-002.
December 1978. Docket Reference Number II-A-7.*
E-67
-------
30.
31.
32.
33.
34.
Richardson Engineering Services.
Estimating Standards, 1980-1981.
II-I-52.*
Process Plant Construction Cost
1980. Docket Reference Number
35.
36.
37.
38.
39.
40.
41.
42.
43.
44.
Telecon. Katari, Vishnu, Pacific Environmental Services, Inc.
with Tucker, Larry, Met-Pro Systems Division. October 19, 1982.
Catalytic incinerator system cost estimates. Docket Reference
Number II-E-41.*
Telecon. Katari, Vishnu, Pacific Environmental Services, Inc.,
with Kroehling, John, DuPont, Torvex Catalytic Reactor Company.
October 19, 1982. Catalytic incinerator system cost estimates.
Docket Reference Number II-E-40.*
Letter from Kroehling, John, DuPont, Torvex Catalytic Reactor
Company, to Katari, V., PES. October 19, 1982. Catalytic incinerator
system cost estimates. Docket Reference Number II-D-66.*
Key, J.A. Control Device Evaluation: Catalytic Oxidation. In:
Chemical Manufacturing Volume 4: Combustion Control Devices.
U.S. Environmental Protection Agency™ Research Triangle Park,
N.C. Publication No. EPA-450/3-80-026. December 1980. Docket
Reference Number II-A-18.*
Telecon. Siebert, Paul, Pacific Environmental Services, Inc.,
with Kenson, Robert, Met-Pro Corporation, Systems Division. July 22,
1983. Minimum size catalytic incinerator units. Docket Reference
Number II-E-73.*
Memo from Paul Siebert, PES to Polymers Manufacturing Industry NSPS
File. October 5, 1984. Condensation System Design and Cost Computer
Program for Polystyrene and Poly(ethylene terephthalate). Docket
Reference Number II-B-93.*
Reference 17, p. 3-59.
Reference 17, pp. 3-72.
Rossini, F.D. et al. Selected Values of Physical and Thermodynamic
Properties of Hydrocarbons and Related Compounds, Comprising the
Tables of API Research Project 44. Pittsburgh, Carnegie Press,
1953. pp. 652 and 682. Docket Reference Number II-I-5.*
Reference 17, pp. 10-25 through 10-28.
Reference 17, pp. 10-12 through 10-15.
Ludwig, E.E. Applied Process Design for Chemical and Petrochemical
Plants, Volume 3. Houston, Gulf Publishing Company. 1965. p. 80.
Docket Reference Number II-B-93, Attachment D.*
Reference 42, pp. 100, 101, and 104.
Reference 42, pp. 100 through 106.
E-68
-------
45. Telecon. Meardon, Ken, Pacific Environmental Services, Inc., with
Mahan, Randy, Brown Finntube Company. July 30, 1984. Cost estimates
for various size condensers (4.6 ft2 up to 34 ft2). Docket Reference
Number II-E-90.*
46. Telecon. Meardon, Ken, Pacific Environmental Services, Inc. with
Kurtz, Ned, American Standard Heat Transfer Division. July 30, 1984.
Cost estimates for various size condensers. Docket Reference
Number II-E-91.*
47. Memo from K. Meardon, PES to Polymer Manufacturing NSPS File.
August 17, 1984. Refrigeration Units for Condensers.
Docket Reference Number II-B-88.*
48. Reference 17, pp. 11-1 through 11-18.
49. Reference 17, pp. 3-71, 3-126, 3-214, and 12-46 through
12-48.
50. Weast, R.C., ed. Handbook of Chemistry and Physics, fifty-third
edition. Cleveland, The Chemical Rubber Company, 1972. p. F-36.
Docket Reference Number II-I-122.*
51. Thermophysical Properties of Refrigerants. New York, American
Society of Heating, Refrigerating and Air-conditioning Engineers, Inc.
1976. pp. 9 through 11, 105, and 106. Docket Reference Number II-B-93
Attachments AF and AG.* '
52. Kern, .D.Q. Process Heat Transfer. New York, McGraw-Hill Book
Company, 1950. p. 306. Docket Reference Number II-I-3.*
53. Reference 17, pp. 23-49.
54. Reference 42, pp. 57 and 58.
55. Reference 17, pp. 6-3.
56. McCabe, W.L. and J.C. Smith. Unit Operations of Chemical Engineering,
second edition. New York, McGraw-Hill Book Company. 1967.
1007 p. Docket Reference Number II-I-12.*
57. Meardon, Ken, Pacific Environmental Services, Inc. to E.J. Vincent,
EPArCAS. January 1983. Designs and cost estimates for ethylene
glycol recovery systems. Docket Reference Number II-B-62.*
58. Telecons. Meardon, Ken, Pacific Environmental Services, Inc.,
with R. Smith, Allied Fibers. December 1982 and January 1983.
Design and operating parameters of ethylene glycol recovery systems
Docket Reference Number II-E-54.*
E-69
-------
59. Reference 17, p. 5-31.
60. Chontos, L.W. Find Economic Pipe Diameter via Improved Formula.
Chemical Engineering. £7_( 12): 139-142. June 16, 1980. Docket
Reference Number II-I-59.*
61. Memo from Desai, Tarun, EEA, to EB/S Files. March 16, 1982.
Procedure to estimate piping costs. Docket Reference Number
II-B-37.*
62. Memo from Kawecki, Tom, EEA, to SOCMI Distillation File. November 13,
1981. Distillation pipeline costing model documentation. Docket
Reference Number II-B-36.*
63. Memo from Paul Sie,bert, PES, to Polymers and Resin File. March 16, 1983.
SOCMI Distillation NSPS Pipeline Costing Computer Program (DMPIPE),
1981. Docket Reference Number II-B-66.*
64. Reference 29, Section 4.2, pp. 4-15 through 4-28.
* References can be located in Docket Number A-82-19 at the
U.S. Environmental Protection Agency Library, Waterside Mall,
Washington, D.C.
E-70
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-450/3-83-019a
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
5. REPORT DATE
Polymer Manufacturing Industry
tion for Proposed Standards
- Background Informa-
September 1985
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Pacific Environmental Services, Inc.
1905 Chapel Hill Road
Durham, N.C. 27707
10. PROGRAM ELEMENT NO.
11. CONTRACT/GRANT NO.
68-02-3060
12. SPONSORING AGENCY NAME AND ADDRESS
13—T.YPE.OF REPORT AND PERIOD COVERED
U.S. Environmental Protection Agency
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina 27711
14. SPONSORING AGENCY CODE
EPA/200/04
15. SUPPLEMENTARY NOTES
16. ABSTRACT
Standards of performance for the control of volatile organic compound emissions from
the polymer manufacturing industry are being proposed under the authority of
Section 111 of the Clean Air Act. These standards would apply to new, modified,
and reconstructed facilities that manufacture polypropylene, polyethylene, polystyrene,
or poly(ethylene terephthalate). This document contains background information
and environmental and economic impact assessments of the regulatory alternatives
considered in developing the proposed standards.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b. IDENTIFIERS/OPEN ENDED TERMS
c. COS AT I Field/Group
Air Pollution
Volatile Organic Compounds
Polymers
Resin
Polyethylene *
Polypropylene
Polystyrene
Poly(ethylene terephthalate)
Air Pollution Control
13B
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS /This Report)'
Unclassified
21. NO. OF PAGES
512
20. SECUR! TY CLASS /This page)
Unclassified
22. PRICE
EPA. Form 2220-1 (Rev. 4-77)
PREVIOUS EDITION !S OBSOLETE
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