Economic comparison of multiple techniques for recovering leaf protein in
biomass processing
Bryan Bals and Bruce E Dale*
Supplementary Online Material
Model Description
As stated in the main text, the protein recovery model uses a material balance of the
process to size capital costs and determine the primary operating costs. Components
in the material streams are water, insoluble non-protein biomass, soluble non-protein
biomass, protein, and any additional chemicals such as ammonia. Both capital and
operating costs are combined and compared to the revenue obtained from the protein
product to determine the overall profit. All economic values are presented in $/Mg initial
feedstock. It is assumed that all other costs and revenues of the refinery, such as
ethanol production or electricity production from lignin, are identical for all possible
protein processes, and thus are not considered.
For capital costs, only major pieces of equipment are considered. These costs are
multiplied by a Lang factor of 5 to determine total capital investment. The final capital
cost per Mg feedstock is determined using straight line depreciation over fifteen years.
Operating costs consist of both direct and indirect costs. Direct costs are material
inputs (which do not include the biomass, as the cost is taken into consideration by the
rest of the biorefinery) and energy inputs (steam and electricity). Indirect costs include
maintenance, insurance, and taxes, which were assumed to be 3.5% of the total
installed cost annually (Aden et al. 2002), and salaries, which were assumed to be
approximately $213,000 annually for both models.
1
It is assumed that no changes are made to the ethanol production facility. This
assumption is based on the major costs associated with ethanol production, which
include the pretreatment operations, the capital cost, enzyme requirements, and
feedstock costs. The revenue streams are ethanol and electricity (Aden et al. 2002).
Protein extraction addition will not impact the enzyme requirement or feedstock costs,
as we are not adding any additional biomass. Capital cost is dominated by the utility
boiler/generator as well as distillation, which should not be adversely affected by protein
extraction (total biomass solids loading during fermentation will remain constant). For
pretreatment, the primary difference between protein extracted biomass and untreated
biomass is the moisture content entering pretreatment. A recent study suggests that
the water content in the biomass is not a large contributor to the cost of AFEX
pretreatment (Bals et al. 2010). While a previous study suggests that protein extraction
may lower the sugar yield during enzymatic hydrolysis (Bals et al., 2007), this is due in
part to the loss of free sugars and oligosaccharides to the extract, which we expect can
be recovered by using the whey as the hydrolysate media. Thus, while some changes
may be present in the ethanol production section of the biorefinery, these changes were
not thought to be significant and thus were not modeled.
The switchgrass used in this model was 15% protein by dry matter, higher than the
protein content in the switchgrass used in the experimental work. Multiple studies have
confirmed that early harvests of switchgrass can contain more than 10% protein. Vogel
et al. (2002) obtained switchgrass with nearly 10% nitrogen that was harvested in late
June. Madakadze et al. (1999) found 12% nitrogen in switchgrass, also in late June.
2
Given that switchgrass can be harvested as early as May, switchgrass with 15% protein
is not unreasonable.
Mechanical Pressing
The process design and economics of the mechanical pressing model was created
based on the pilot plant data provided in Enochian et al. and McDonald et al. (Enochian
et al. 1980; McDonald et al. 1981). Two mills were used for the base case scenario: a
hammer mill and a disk mill. These mills are used to break open the cells, allowing
intracellular protein to be separating from the remaining fiber. It was assumed that 50%
of cells were ruptured using the hammer mill, while 80% of the remaining cells were
ruptured with the disk mill. Electricity requirements for these mills were determined as
an average of multiple runs obtained from McDonald et al. (McDonald et al. 1981). The
capital cost of the hammer mill was based on the value given by Sokhansaj and
Turhollow and scaled appropriately (Sokhansanj and Turhollow 2004). As a first
approximation, it was assumed that the disk mill cost twice as much for a similar
biomass throughput.
Two screw presses were used to remove the juice from the biomass. Vincent
Corporation (Tampa, FL) constructs presses that have been used for LPC production
(Enochian et al. 1980). An estimate of the capital cost was obtained from Fedler et al.,
using the insoluble solid fraction to size the equipment (Fedler et al. 2008). Power
requirements were obtained from Vincent Corporation. Both presses were assumed to
3
have the same power requirements despite different degrees of cell disruption, and both
were assumed to have the same exiting moisture content.
The steam injection system was sized and priced as a simple reaction vessel. It was
estimated that 75 g low pressure steam were required per kg of water to obtain the
desired temperature based on the enthalpy and heat capacity of the protein solution.
Protein was separated using a centrifuge, which was sized and priced based on design
equations obtained from Peters et al. (Peters et al. 2003). Power requirements were
obtained from Enochian et al. (Enochian et al. 1980). It was assumed that the solid
protein product exited at 20% solids.
Much of the deproteinated juice (“whey”) is recycled and added to fresh biomass. This
is done to increase the initial moisture content of the biomass. The model predicts that
protein from ruptured cells freely mixes with all water present prior to the screw press,
and so protein removed from the fiber is a function of both cell disruption and the
proportion of water that is pressed out. By increasing the water present and maintaining
the same final moisture content, a higher proportion of protein can be extracted with the
juice. The initial moisture content of each press was set at 6.7 g water per g insoluble
solids. To prevent buildup of solubles, a portion of the whey is used as the cellulosic
hydrolysate medium, allowing fresh water to also be used to increase the moisture
content of the biomass. Cellulosic hydrolysis is assumed to be performed at 4.56 g
water per g biomass entering the refinery (Lau and Dale 2009), and no water is added
to or removed from the fiber after it leaves the protein extraction module and prior to
4
hydrolysis. The cost of makeup water is not included in this model, as it is offset by the
reduction of water required for cellulosic hydrolysis.
The protein product was then dried using a rotary dryer. Sizing was based on the
amount of water to evaporate, assuming the final product contained 5% moisture.
Capital cost was estimated from Peters et al. (Peters et al. 2003) Energy consumption
of the dryer was set at 3.7 MJ/kg water evaporated (Enochian et al., 1980).
The final capital costs are shown in Table S.1, while operating costs are shown in Table
S.2. Total capital investment is approximately $4.6 million, of which the disk mill and
the presses contribute nearly 75% of the total cost. The disk mill also consumes a large
amount of electricity, which costs $7.25 per Mg feedstock. Electricity costs for the
centrifuge are also high at $3.90/Mg.
Aqueous Extraction
The aqueous extraction model was based primarily on the model proposed by Laser et
al. (Laser et al. 2009) and adapted to the experimental results presented in this paper
and previous work (Bals et al. 2007). As with the mechanical pressing model, a
hammer mill is first used to grind the biomass. The same economic and energy
assumptions are used here as with the mechanical pressing model. A crossflow
extraction column is then used to solubilize the protein. The capital cost and capacity of
the column was obtained from the model used by Laser et al. (Laser et al. 2009).
Extraction conditions were obtained from previous experimental work with switchgrass
5
(Bals et al. 2007). In particular, a residence time of 30 minutes was used for the
extraction, and the alkaline loading was 1% ammonia. The liquid/solid ratio was set at
8:1, despite 10:1 being used in experiments. Experimental work was performed in a
batch setup, rather than a continuous crossflow extraction as simulated, and thus a
higher liquid/solid ratio was required.
After extraction, a pneumapress is used to reduce the moisture content in the insoluble
solids. This press uses air to decrease the moisture content in the insoluble solids to
45%. Capital and electricity requirements for this operation were obtained from Aden et
al., which uses a pneumapress to separate lignin residue from the fermentation broth in
a cellulosic biorefinery (Aden et al. 2002). Because of the high capital and operating
costs of the pneumapress, a simple plate and frame filter press was considered as an
alternative. Equipment sizing details were obtained from design equations taken from
Peters et al. (Peters et al. 2003).
Ultrafiltration is then used to concentrate the protein, with all sizing and operating costs
obtained from Laser et al. (Laser et al. 2009). Operating costs include electricity and
filter replacements, cleaning, etc. All costs are determined as a function of the amount
of water removed from the process. A concentration factor of 30 was used for this
operation, unless the amount of water present was less than 30 times the amount of
recovered protein product. In this case, the protein product is concentrated to 30% of
the water present. The protein product was then dried in the same manner as in the
6
mechanical pressing model, and all costs and assumptions on protein drying are the
same in both models.
In the initial model by Laser et al., the liquid is recycled indefinitely (Laser et al. 2009).
However, as with the mechanical pressing model, only a partial recycle was used so as
to not build up the presence of solubles. Again, enough water is removed for cellulose
hydrolysis, with the remaining solvent recycled and fresh makeup water added as
needed. Because cellulosic hydrolysis occurs in an acidic medium, the ammonia must
be removed or neutralized. An ammonia stripper was designed using Aspen Plus, and
it was estimated that 0.29 MJ/kg water was required to remove 95% of the ammonia.
Alternatively, sulfuric acid was used to neutralize the filtrate. The cost of the ammonia
stripper was determined based on the ammonia stripper used in Laser et al., which was
an operation during ammonia recovery following ammonia fiber expansion (AFEX)
pretreatment.
Although recovered ammonia would need to be compressed and/or condensed in order
to reuse it, this operation is not modeled. Because the protein extraction module is part
of an integrated biorefinery, it is assumed that the ammonia recovery for AFEX
pretreatment can be used here as well. Because the amount of ammonia needed for
protein extraction is small relative to AFEX (approximately 0.06 g NH3 / g biomass vs 1
g/g for AFEX), it is assumed that ammonia recovery costs do not significantly increase.
Alternatively, a scrubber could be used to capture the ammonia if AFEX pretreatment is
not used in the refinery. Again, this is not modeled in the process.
7
Post hydrolysis ultrafiltration and drying are modeled in the same method as the extract.
Yields of protein solubilized and recovery via ultrafiltration were determined from Bals et
al. (2007) and the Figure 3 in the main text. However, the quality and thus selling price
of the product are assumed to be half that of the protein recovered in the extract. This
is due to excess lignin in the final product as well as potential protein degradation after
AFEX pretreatment.
Overall capital costs, and operating costs are shown in Tables S.3 and S.4 respectively.
Total capital investment is similar to mechanical pressing, although operating costs are
lower. The pneumapress dominates capital costs, accounting for 60% of the total
equipment costs in the aqueous extraction model. It also accounts for 15% of the
operating costs. Ammonia stripping is the second most costly process, accounting for
over 20% of the capital cost and 25% of the operating cost. Because of the high costs
associated with these two processes, several alternatives were considered for these
two steps, as shown in the main text.
8
Table S.1: Capital costs for the base case mechanical pressing model
Scaling Factor
kg/Mg biomass
Hammer Mill
1000
Disk Mill
932
Press 1
850
Press 2
730
Steam Injection
148
Centrifuge
9683
Dryer
741
Total Equipment Cost
Total Capital Investment
Depreciation
Capacity
Required
kg
81019
75514
68866
59144
3.33a
392232
33.35b
Equipment
Size
kg
28000
28000
12000
13750
4.72a
450000
40b
a
Capacity and equipment size measured in kg/s
Capacity and equipment size measured in m2
c Cost is measured as $/Mg feedstock
b
9
Number
Required
3
3
6
5
1
1
1
Cost per Item
2008 $
231,022
461,972
175,043
183,267
145,450
247,223
171,367
Total Cost
2008 $
693,066
1,385,916
1,050,260
916,335
145,450
247,223
171,367
4,609,616
23,048,081
8.78c
Table S.2: Operating costs for the base case mechanical pressing model.
Hammer mill electricity
Disk mill electricity
Press 1 electricity
Press 2 electricity
Steam for coagulation
Centrifuge electricity
Dryer Energy
Maintenance
Insurance and Taxes
Salaries
Total Operating Costs
Value
45
145
1.174
1.174
75
8.33
3.7
2
1.5
212,600
Unit
kW*h/Mg dry biomass
kW*h/Mg dry biomass
kW*h/Mg wet biomass
kW*h/Mg wet biomass
g/kg water
kW*h/Mg water
GJ/Mg water
% installed cost
% installed cost
$/yr
10
Cost ($/Mg Feedstock)
2.25
7.25
0.46
0.39
1.96
3.89
2.82
0.85
0.64
1.21
20.57
Table S.3: Capital costs for the base case aqueous extraction model
Scaling Factor
kg/Mg biomass
Hammer mill
1000
Extraction column
9000
Pneumapress
700
Ammonia recovery
6883
Ultrafiltration 1
319
Protein dryer 1
225
Ultrafiltration 2
235
Protein dryer 2
146
Total Equipment Cost
Total Capital Investment
Depreciation
Capacity
Required
kg
20833
93750
14583
143389
6653
1.30a
4897
0.84a
Equipment
Size
kg
30000
100000
15000
150000
6400
0.7a
5400
0.9a
11
Number
Required
1
1
1
1
1
2
1
1
Cost per Item
2008 $
293,789
215,148
3,421,181
1,305,605
83,462
87,564
75,373
101,815
Total Cost
2008 $
293,789
215,148
3,421,181
1,305,605
83,462
175,128
75,373
101,815
5,671,500
28,357,502
10.80b
Table S.4: Operating costs for the base case aqueous extraction model.
Mill electricity
Pneumapress electricity
Steam for NH3 stripper
Filtration costs
Filtration electricity
Drying electricity
Makeup Ammonia
Maintenance
Insurance and Taxes
Salaries
Total Operating Costs
Value
45
31.4
0.29
1.49
35.2
79.4
0.35
2
1.5
212,600
Unit
kW*h/Mg dry biomass
kW*h/Mg dry biomass
GJ/Mg water
$/Mg permeate
kW*h/Mg permeate
kW*h/Mg water
$/kg ammonia
% installed cost
% installed cost
$/yr
12
Cost ($/Mg Feedstock)
2.25
1.57
2.70
0.58
0.68
1.54
0.71
0.81
0.61
1.21
12.67
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