1

To: EPA

Title: Techno-economic analysis of microalgae-derived biofuel production

Author: Ryan Davis

Date: November 10, 2009

Abstract

The purpose of this study is to evaluate the economics for a feasible method of producing microalgalderived oil in the near-future. Specifically, production via both open-pond and closed photobioreactor 

(PBR) systems was evaluated for a 10 MMgal/yr oil facility based on assumptions for what can feasibly 

be achieved in the near-term, represented by “2022” target production. The scope of this study is based 

on analysis for the process beginning at CO2 delivery to the facility and ending at lipid production

(refining into biodiesel is fairly well-established and is not included here). Upon completing the base 

case scenarios, the cost of lipid production to achieve a 10% return was determined to be $11.25/gal for 

open pond production and $19.50/gal for PBR’s. This cost has potential for significant improvement in 

the future if better microalgal strains can be identified that would be capable of sustaining high growth 

rates at a high lipid content. Given that algal biofuel technology is still in a relatively early stage of 

development, there is intrinsically a higher degree of uncertainty associated with potential process 

performance and cost relative to more established technologies. As such, many assumptions had to be 

made in this study; the assumptions with the highest economic impact were found to be the optimum 

amount of nutrients required, the CO2 delivery cost, and the flocculant requirement for harvesting, as 

well as the material costs for the PBR production system.

Keywords: Algae, techno-economic, pond, PBR, biodiesel, oil

Introduction

It is well-established that microalgal-derived biofuels have the potential to make a significant 

contribution to the US fuel market, due to unique advantages inherent to algae. While these 

advantages are outlined elsewhere [1, 2], they include the fact that microalgae have a higher 

photosynthetic efficiency than terrestrial plants at converting solar energy into biomass and also can 

possess a very high oil content, resulting in drastically lower land requirements that must be dedicated 

to biomass growth per unit of oil produced, relative to other oil-based crops. Furthermore the land used 

to grow microalgae can be low-value or otherwise not suitable for growing traditional agricultural crops. 

Additionally, CO2 from a nearby emissions source can be used to grow the algae and the only source of 

CO2 from the algae facility itself can be recycled, resulting in a “net zero” emissions process.

Technical Memorandum
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In this study, a microalgae facility producing lipids on a basis of 10 MMgal/yr was modeled using Aspen 

Plus software, in order to get more accurate mass balance information than what is typically assumed in 

other studies. Additionally the model was used to estimate the overall amount of power used vs 

generated throughout the plant. Using the resulting flow rates, the cost of each process unit was 

estimated based on vendor information, cost equations, or previous studies to determine the cost of 

lipid production. It is important to note up front that algal biofuel technology is still in a very early stage 

of development, and there are numerous opinions on the optimal configuration and conditions at each 

stage of the production process; as such, the estimated cost of production varies in other studies from 

$1 to >$40/gal [3]. There are also various opinions on when the technology will even become feasible 

on a large scale. Thus the “2022” startup date should be treated as being representative of what can 

reasonably be expected in the “current to near-term” time frame.

The open pond and PBR cases will be discussed together, as most of the technology employed in either 

case is the same downstream of the production step. To summarize the basic flow scheme chosen for 

this study, the open pond configuration is as follows (see Figure 1 for a block flow diagram): pure CO2 is 

concentrated out of flue gas from a nearby power plant and delivered to the facility. The CO2 material 

cost used in this study accounts for all upstream operations required to concentrate CO2 out of the flue 

gas stream (by means such as amine scrubbing). The CO2 is transferred to the ponds via sumps. The 

ponds are simple unlined, paddle wheel-mixed raceway ponds 30 cm deep, providing constant 

circulation to promote algal growth. The microalgae grow to a concentration of 0.5 g/L, and are 

continuously harvested at a rate equal to the growth rate. Harvesting is accomplished first in a simple 

settling tank which concentrates the algae to 1% (10 g/L), and then in a secondary step by flocculation 

using an organic polymer flocculating agent and subsequent collection via dissolved air flotation (DAF), 

which thickens the material to 10% (100 g/L). Extraction is accomplished primarily by mechanical 

methods using homogenizers to lyse the cell, followed by hot oil extraction to purify the oil phase. The 

oil, water, and biomass phases separate by gravity in a clarifier and the oil is drawn off, a fraction of 

which is recycled and heated to provide the “hot oil” solvent. The water and residual biomass are sent 

to anaerobic digestion, which produces biogas burned in a gas turbine to aid in supporting power 

requirements throughout the plant. The turbine flue gas is recycled to the ponds, while the treated 

water is discharged. Makeup water is added to replace this discharge, as well as evaporation from the 

ponds and a small amount of water entrained in the lipid product. Nutrients from the anaerobic 

digestion step are recycled to minimize the amount of fresh nutrients required.

In the PBR case, pure CO2 is delivered to the facility by being injected directly into the feed header, 

eliminating the need for sumps as in the open pond case. The PBR system is made up of clear plastic 

tubes in a parallel configuration, with a degassing/pumping section (achieved by airlift columns) every 

80m to strip out accumulated oxygen and provide circulation through the tubular system. Due to the 

concentrated nature of PBR culture, the algae are grown to a much higher density of 4 g/L. The algae 

are harvested at a rate equal to the growth rate, and the material is again concentrated to 10 g/L in the 

primary settling step then subsequently to 100 g/L in the DAF unit, similar to the open pond case. The 

remaining steps (extraction, purification and spent biomass utilization) are identical to the open pond 

case, as shown in Figure 1.
Figure 1: Block flow diagram of the production scheme evaluated here

Three cases were run for each scenario to evaluate the impact that research advancements will have on 

the production cost. These are summarized below and results are presented at the end of this report.

Table 

Open pond

Algae productivity1

25

[g/m2/day]

Cell density [kg/m3] 0.5

Lipid yield [dry wt%] 25%

CO2 consumed [lb/lb 

algae] 2

Operating days/yr 330

Solar exposure 

hours/day 12

1. Productivity is on an areal basis (g/m

The base case shown above represents a reasonable (though still challenging) 

productivity and lipid content, while the “aggressive” case assumes identification of a strain with near

optimal characteristics, and the “max

near-theoretical maximum growth and lipid content 

be described in more detail throughout the report.

Process Area Descriptions and Major Assumptions for Model/Cost

As noted, there are numerous possibilities at each stage of the 

no clear winner has yet been established

only way to achieve the end result, but is one of the more “likely” scenarios to be feasible at this scale, 

and was chosen by conducting an extensive literature review and building off of previous work 

conducted within NREL. With this in mind, an overview of each production stage as well as a summary 

of the chosen process and underlying assumptions at each stage is de

: Block flow diagram of the production scheme evaluated here

Three cases were run for each scenario to evaluate the impact that research advancements will have on 

cost. These are summarized below and results are presented at the end of this report.

Table 1: Algae production analysis scenarios

Base case Aggressive case

Open pond PBR Open pond PBR

/day]

1.25 

[kg/m3/day]

40

[g/m2/day]

2.0 

[kg/m3/day]

4 0.5 4

25% 50% 50%

2 2 2

330 330 330

12 12 12

Productivity is on an areal basis (g/m2/day) for open ponds and a volumetric basis (kg/m3/day) for PBR’s

The base case shown above represents a reasonable (though still challenging) near-term target for algae 

productivity and lipid content, while the “aggressive” case assumes identification of a strain with near

optimal characteristics, and the “max growth” case is provided as a comparison to demonstrate the 

rowth and lipid content based on photosynthetic efficiencies [

be described in more detail throughout the report.

Process Area Descriptions and Major Assumptions for Model/Cost

As noted, there are numerous possibilities at each stage of the production process, and in many cases 

yet been established. Thus the proposed process evaluated here is by no means the 

only way to achieve the end result, but is one of the more “likely” scenarios to be feasible at this scale, 

chosen by conducting an extensive literature review and building off of previous work 

With this in mind, an overview of each production stage as well as a summary 

of the chosen process and underlying assumptions at each stage is described below.

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Three cases were run for each scenario to evaluate the impact that research advancements will have on 

cost. These are summarized below and results are presented at the end of this report.

Max growth case

Open pond PBR

60

[g/m2/day]

3.0 

[kg/m3/day]

0.5 4

60% 60%

2 2

330 330

12 12

term target for algae 

productivity and lipid content, while the “aggressive” case assumes identification of a strain with near-

” case is provided as a comparison to demonstrate the 

based on photosynthetic efficiencies [1]. This will 

Process Area Descriptions and Major Assumptions for Model/Cost

production process, and in many cases 

. Thus the proposed process evaluated here is by no means the 

only way to achieve the end result, but is one of the more “likely” scenarios to be feasible at this scale, 

chosen by conducting an extensive literature review and building off of previous work 

With this in mind, an overview of each production stage as well as a summary 
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Algae Production: Open Pond

Open pond production is commonly employed today at an industrial level for producing algal-based 

health food supplements and other higher-value products [1]. The main types of open pond systems are 

summarized below.

 Unstirred ponds: Unstirred ponds are the simplest in design, as they are essentially comprised of 

basic uncovered beds or natural lakes. As there is no circulation or continuous harvesting, they 

are limited to a small number of microalgae capable of out-competing other environmental 

contaminants. These ponds are generally used for wastewater treatment and β-carotene 

production [4], but would not be amenable to a continuous large-scale fuel process.

 Circular ponds: Similar to unstirred ponds, circular ponds have been used for wastewater 

treatment and algae production destined for various health food applications. Although they 

are somewhat more robust than unstirred ponds in terms of sustaining a culture, they are 

limited in size due to a centrally-pivoted rotating arm which provides mixing. Additionally, due 

to their circular shape they are slightly less space-efficient which is an important consideration 

on this large scale.

 Raceway ponds: Raceway ponds are generally considered the most efficient and low-cost option 

for large-scale microalgae production [4]. They consist of two or more “passes” separated by a 

baffle, with circulation usually provided with a paddle wheel to prevent settling. Raceway ponds 

are considered “high-rate” due to a much higher achievable productivity than unstirred ponds, 

and are more space efficient on large scales due to their physical shape and shallower depth. 

These three pond types are shown below in Figure 2.

Figure 2: Schematic of open pond types [4]

Given the brief comparison described above, raceway ponds were chosen for this system. The basic 

design of the pond system is based on a compilation of multiple detailed analyses by Benemann and 

Oswald [5], and essentially targets the lowest-cost design possible due to the relatively low value of the 

final product (fuel). Feed (makeup water, nutrients and CO2) would be added just after the paddle 
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wheel, and a fraction of the material would be harvested just before the circulation returns back to the 

paddle wheel. The pond characteristics (which feed into the cost estimate) are summarized below:

 Unlined ponds: even a simple plastic liner could double the capital cost of the ponds [6], thus 

only pond bottom compaction is included, to control percolation and silt suspension

 Site is on low-value, flat, relatively bare land: minimize cost of leveling and site preparation 

 Earthen levees with a geotextile liner to prevent erosion

 30 cm depth: This is an important parameter as it affects system volume and water usage. 

Typical depth is 15-30 cm, with 30 cm optimal to limit CO2 outgassing [7]

 Mixing by paddle wheels requiring 18 kwh/ha/day of power [B&O]

These characteristics formed the basis of the Benemann and Oswald cost estimate for the ponds. The 

B&O study estimated capital and operating costs for every stage of production, and related the resulting 

cost to “$/hectare” of growth ponds for simplicity; the majority of other subsequent economic analyses 

used this basis also at every stage, but with different parameters such a method of cost estimating loses 

its accuracy for downstream units. Thus this basis was avoided for most costs not related to the growth 

ponds, and original cost estimates were used instead for this study. However, for the actual growth 

ponds and associated equipment, a “$/ha of pond” basis is directly applicable. Given this basis, the 

specific cost estimate from B&O for site preparation and pond materials was $6,000/ha, with another 

$5,000/ha for paddle wheels (note these prices are in 1996 dollars).

Equally important as the physical design of the pond system is the resulting algae yield and lipid content 

that could reasonably be expected of such a system. The “base case” representative of near-term 

potential between now and 2022 was set at a growth rate of 25 g/m2/day based on areal pond size, with 

25 wt% lipid content. This was based on NREL’s recent Report to Congress summarizing the current 

state of the technology, which states that the same lipid content and 20 g/m2/day is a “reasonable 

(though still challenging) near-term target” [1]. Other studies have achieved 15-25 g/m2/day, but not on 

an ongoing/sustained basis [8]. Thus it is reasonable to assume that 25 g/m2/day at a sustained rate of 

330 operating days/year is feasible within a 10-year timeframe. The assumption of 330 operating 

days/year also implies that a site location is chosen which receives high year-round solar exposure; two 

such sites which could sustain this high (90%) operating factor are Brownsville, TX or Tarpon Springs, FL 

[9], as an example that such an aggressive target is plausible. In addition, a solar exposure of 12 

hours/day was assumed. While all downstream operations are assumed to operate 24 hr/day, algae 

undergo growth only during daylight hours and CO2 would be fed only during this time (although the 

ponds or PBR’s would be continuously mixed 24 hr/day to prevent settling). 

Two other cases were also studied as shown in Table 1, representing feasible longer-term research 

advancements in strain improvement and the theoretical maximum that could possibly be achieved 

based on photosynthetic efficiency limitations [1]. The open pond case was assumed to reach a cell 

density of 0.5 g/L, held constant across each scenario [10]. Additionally, in all cases a CO2 requirement 

of 2.0 g/g algal biomass was assumed [1], a stoichiometric ratio independent of production method. 

This value corresponds well to another study which uses a ratio of 1.83 assuming 100% CO2 uptake 

efficiency [11], thus a ratio of 2.0 allows for 90% uptake efficiency which is typical.
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Algae Production: Closed PBR

The other means to produce microalgae is by way of a closed photobioreactor (PBR) system. While this 

option offers many distinct advantages over an open pond configuration, it is generally stated that the 

cost can be an order of magnitude larger than for open-pond systems [5] and as such there have been 

very few rigorous techno-economic studies done for PBR production. A brief comparison between open

pond and PBR production is given below [4, 12]:

Open Pond

 + Lower capital and maintenance cost

 + Easy to scale up (add more ponds)

 + Easy to integrate with wastewater 

treatment

 + Readily available technology

 - High water evaporation losses

 - Low cell density = high land 

requirement and harvesting cost

 - Susceptible to invasive species (limits 

the number of viable strains and can 

derail economics if too many re-starts 

are required)

PBR

 + Higher cell density = lower land use 

and downstream processing cost

 + Higher volumetric productivity (higher 

surface area/volume ratio)

 + Reduced contamination risk (not open 

to environment)

 + Low evaporative losses (but 

temperature control can be a problem)

 - Higher capital and maintenance cost

 - Requires degassing to limit O2

accumulation (which inhibits growth)

 - Not demonstrated on large-scale

Although a closed PBR system mitigates contamination risk and reduces harvesting cost through a much 

higher cell concentration, the capital cost is usually expected to outweigh these benefits. While there 

are many possible PBR configurations, they can broadly be classified into two groups:

 Flat plate: Flat plate PBR’s are made up of a clear plastic containment system, such as a 

rectangular box or series of plastic bags housing the microalgal culture. They usually are 

designed to allow for a short light path to maximize light exposure. Flat plate PBR’s can have 

lower capital cost than tubular reactors due to their more simple design, but may not be as 

readily scalable [13].

 Tubular: Tubular PBR’s have historically been used the most frequently for pilot scale PBR 

studies [4]. They consist of clear (plastic or glass) tubes that carry a circulation of culture 

between degassing and harvesting. They can be arranged in a number of different 

configurations (vertical, horizontal or inclined) to maximize efficient land use or sun exposure. 

Although productivities are comparable, some studies have found tubular PBR’s to have higher 

productivity than flat plate reactors [4]. Additionally, they are generally considered more 

feasible for large scale use as they are modular and accommodate higher flows simply by adding 

more parallel passes [2, 13].

These PBR types are depicted in Figure 3.
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Figure 3: Depiction of flat plate (left) and tubular (right) PBR [14]

Unlike raceway ponds for the open pond case, no clear winner has been universally established yet for 

the PBR type. Based on the above comparison (notably the fact that a tubular system may be more 

amenable to large-scale production), a horizontal tubular PBR was chosen here. In this system, the 

culture would be circulated through the tubes, with degassing columns located throughout the system 

to strip out oxygen that accumulates as a product of photosynthesis. CO2 would be injected 

immediately following the degassing columns. The system forms a closed-loop recycle scheme with a 

fraction of the circulation drawn off for harvesting. This flow configuration is shown below in Figure 4.

Figure 4: Basic flow scheme for airlift-driven PBR [15]

The specific tubular PBR characteristics for the design assumed here are summarized as follows:

 Tube ID = 8 cm: The diameter is restricted to <10 cm to ensure adequate light penetration, and 

various other studies assume 6 cm for a smaller-scale plant [2]

 Tube section length = 80 m: The maximum tube length in a single section between degassing 

zones is 80 m, as too much oxygen accumulates at longer lengths which inhibits cell growth [11]. 

Given these dimensions for a single tube, the total number of tubes required can be calculated 

based on the total system volume (which is set by the volumetric productivity).

 1 hectare land use = 500 m3culture volume: This is the translation between culture volume and 

areal footprint of the tubular system, specific to one study [16]. This also corresponds closely to 

a second study [11], thus the same footprint was again assumed here.
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 Circulation via airlift columns: Airlift columns work by pumping air into the bottom of a riser 

column, which imparts a lower bulk density to the fluid and causes the fluid to rise. At the top 

of the column the liquid spills over into a downcomer, which in turn provides head to circulate 

the fluid. This eliminates any shear force associated with mechanical pumps that could damage 

or kill the cells. In addition, this doubles as a degassing column (which is required anyway), by 

allowing the oxygen entrained in the liquid to disengage and exit out of the top with the 

bubbled air [2]. As mentioned previously, an airlift column would be required every 80 m to 

degas the accumulated oxygen.

 Temperature controlled via sprinkler system: The temperature can rise over the length of the 

tube and must be controlled to 20-30 °C [2]. Various methods have been proposed to control 

the temperature; the more popular methods are either to submerse the tubes in a pool of water 

or use a conventional sprinkler system to spray the tube surface (evaporative cooling). While 

the latter may only be feasible in dry climates, it would be far cheaper and is the option chosen 

by various other PBR studies [16, 17].

Given this basis for the physical configuration of the system, the PBR costs could then be estimated. As 

noted, there have been very few thorough economic analyses done for PBR systems, thus costs were 

either scaled off of a basis if available or estimated using engineering judgment. Consequently there is a 

higher degree of uncertainty around capital costs for this system, and this is captured in the sensitivity 

analysis later in this report. 

The largest cost item was the material cost for the tubes themselves. While there are various materials 

that could be feasible for the tubes, one option commonly assumed is clear plastic LDPE with a UVprotectant coating [16]. Given that LDPE is not frequently used for rigid piping (especially in bulk), this 

cost was approximated using bulk HDPE pipe at $1.50/ft [18]. This assumption can be justified in that 

HDPE and LDPE resins have a very similar price [19], and any marginal extra cost associated with UV 

coating would likely be offset by the fact that the quoted price was for a slightly larger pipe size. A 1.5x 

installation factor was applied to this base estimate to account for modular installation as well as 

manifolds and valves [16].

The next cost assumption made was the cost for airlift columns. These columns are typically used only 

for special applications and are custom-built for the purpose; consequently a bulk cost was not 

available. Other studies have assumed a cost number for this item, ranging from 20-200% of tube 

material cost [20, 16]. Given the fact that there would be far fewer airlift columns relative to the 

number of 80m tube sections (roughly one million tubes) and that the airlift columns should also be 

fairly simple in design, the low-end estimate seems much more reasonable here, 20% of tube material 

cost. In the Aspen model, a power requirement assumption for air compression was also required; the 

assumptions here were to introduce air into 20 ft tall columns at a 1:1 volume ratio of air:liquid.

The cost and water use associated with sprinkler cooling was estimated based on previous TE studies 

[16, 17]. The capital cost used in the prior study was for agricultural sprinklers at $2,500/ha of PBR area; 

while this may not fully cover every inch of pipe, it is a reasonable first-pass assumption to provide 

enough cooling to ensure proper performance. The water requirement used in Aspen to estimate 
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circulation pump power was 200 L/min/ha. An evaporative water loss at 1% of this circulation was 

assumed, which requires fresh water makeup at the same rate.

The final cost item is a simple plastic liner on the ground. To maximize light utilization, the ground 

beneath the PBR system is usually painted white or covered with a white liner [11]. A liner was chosen 

here to also provide simple drainage, and the cost was estimated at $3,200/ha [16].

Again, a critical element impacting the overall production cost is the specific microalgal growth that can 

be expected of the given system described above. Productivity has been shown to range from 1-1.5 

g/L/day [21], with a common value for maximum productivity around 1.25 g/L/day [22]. As these 

numbers were again a short-term maximum rather than a sustained result, it was assumed here that a 

reasonable near-term “base case” target would be 1.25 g/L/day sustained over 330 days/year. The 

same “base case” algae strain is assumed as for open ponds, with 25% lipid content. Similar to the open 

pond scenario, an “aggressive” and “max growth” case was also evaluated in a sensitivity study to 

evaluate potential improvements in PBR production cost, with parameters shown in Table 1. The algal

productivity for these additional cases was originally only given for the open pond scenario, thus various 

assumptions were made to estimate reasonable corresponding values for the PBR scenario. The 

“aggressive” case was estimated by assuming the same percentage improvement in productivity as was 

given for open ponds, while the “max growth” case was calculated by scaling the “max” open pond 

productivity by the ratio of (surface area/volume) for PBR tubes vs ponds given the specific design of 

each. All PBR cases were assumed to achieve a cell density of 4 g/L [11]; this is an 8x improvement over 

open ponds, and results in drastically lower water use and consequently downstream harvesting cost.

Water Delivery

Water usage is a critical sustainability metric in microalgal biomass production. Given that the algae 

grow in a very dilute medium, a large-scale production facility would process vast amounts of water. 

Even if water losses in the system were kept to a relative minimum, a substantial amount of makeup 

water could still be required; thus water source and means of delivery is key, especially in a dry climate. 

To that end, a decision was made for the base case to assume low-value brackish or saline water is 

pumped from an underground source rather than drawing from fresh water resources (the latter case is 

considered in a sensitivity study). Although relatively little is known about saline/brackish groundwater 

resources, there are many areas in the US with relatively easy access to such water, and the ability of 

certain microalgal strains to grow in brackish or saline water means that there is no need to further 

strain freshwater resources [1]. This option of pumping water from underground was analyzed by 

Benemann & Oswald [5], and the same capital cost is assumed here; namely $4,200/ha (1996$) for open 

ponds, which translates roughly to $72/(m3/day) on a flow rate basis for the PBR case. This includes 

pumps and all infrastructure required to deliver the water from underground, although a specific depth 

was not specified. The associated pump power requirement used in Aspen was back-calculated to be

37.5 kw/(MM m

3

/yr) of water [5]. Being pumped from an underground aquifer, the brackish water has 

no additional cost beyond capital and operating cost of delivery.

Overall makeup water demand was determined for the open pond case by assuming a 0.3 cm/day rate 

of evaporation [1] and 95% process water recycle. 95% water recycle implies 5% water lost in addition 
10

to evaporative losses; of the 5% loss, a small fraction is water entrained in the lipid product and the 

remainder is assumed for mass balance purposes to be blow down to treatment (anaerobic digestion), 

to control contaminant buildup in the system. For the PBR case, being a closed system there is no 

evaporative loss of process water, thus brackish water makeup is only required to replace the 5% drawoff. If the water has a high salinity, using the same saline water rather than fresh water to make up for 

blowdown would result in salt accumulation in the system; however, it is assumed here that the 

brackish water has sufficiently low salinity in keeping with the original source study [5] (the cost 

difference is small either way). Finally, as noted above there is a separate sprinkler water circulation for 

PBR cooling; this water must be fresh (non-saline) water to prevent scaling on the outside of the tubes. 

The total sprinkler water circulation was set to 200 L/min/ha [17], of which 1% was assumed to be lost 

to evaporation and must be replaced with fresh water (priced at $0.0013/gal [23]).

CO2 Delivery

Another important consideration is CO2 delivery. Again for this item, the analysis conducted by 

Benemann & Oswald is still the most detailed to date, and the same cost basis is used in this analysis [5]. 

There were two options considered in the B&O study for CO2 delivery, one scenario assumed using the 

total flue gas material while the other assumed that pure CO2 is concentrated out of the flue gas. In the 

“flue gas” case, the flue gas from a nearby power plant or other emissions source is piped 2.5 km 

(roughly 1.5 mi) to the facility at low pressure. A cost of $5,000/ha of ponds is given which includes 

pipes and blowers. Additionally, to ensure efficient CO2 uptake and minimize outgassing in the case of 

open ponds, the gas is transferred to the liquid by way of a 1.5 m deep sump with a baffle and CO2

sparger, costing an additional $5,000/ha (again note all prices from B&O are in 1996$). The associated 

power requirement to transport the flue gas this distance is given as 15,400 kwh/ha/yr.

The second option considered in the B&O study is a “pure CO2” case, where instead of transporting the 

entire flue gas material, the CO2 is scrubbed out and transported under pressure to the facility. As the 

piping and valving system for this case would be more simple than for the flue gas case, a cost of 

$300/ha is quoted for supply and distribution. Additionally, for open ponds the gas again must be 

efficiently transferred by way of a sump, costing slightly less than the flue gas case at $4,000/ha. The 

B&O study did not include a power estimate for the “pure CO2” case in a similar manner as the “flue 

gas” case; instead, all operating costs associated with concentrating CO2 out of flue gas (i.e. by a method 

such as amine scrubbing) as well as transporting the CO2 are captured in an overall CO2 material cost of 

$40/tonne, which although a rough value corresponds well to other estimates [5, 49]. 

For both the open pond and PBR cases, the pure CO2 scenario was chosen as the “base case” for several 

reasons. For the PBR case, the choice was obvious given the system cost; the tube cost was already 

found to be extremely high with just liquid + CO2, so to alternatively use bulk flue gas at a typical 15% 

CO2 content the amount of tubes required and thus the cost would be prohibitive. Additionally, utilizing 

flue gas in general implies that the microalgal culture can tolerate the nitrogen compounds (NOx), sulfur 

compounds (SOx), and other minor contaminants. While there is some speculation that the NOx in flue 

gas could be used to meet the nitrogen requirements for the microalgal culture, this has not been 

proven and additionally the use of bulk flue gas could result in pH imbalances [12, 24]. For these

reasons as well as to provide a consistent basis for comparison between the cases, the pure CO2 option 
11

was used for the “base case” in both scenarios; the flue gas option is also considered for the open pond 

case in the sensitivity analysis.

As noted, the total capital cost for CO2 delivery and distribution for the open pond case is $4,300/ha [5]. 

For the PBR case, the delivery portion was retained ($300/ha), with no cost included for sumps as the 

gas could be directly injected into a pipe header since it is a closed system. While the cost basis is 

price/ha of growth ponds for an open pond scenario, as noted above the CO2 requirement per unit algae

growth is constant for PBR and open ponds at 2.0 g/g biomass, and the biomass growth rate is the same 

between the two scenarios for the base case (~158,500 tonne microalgae/yr to produce 10 MM gal/yr 

lipids); thus the CO2 rate is equal and the open pond land area can be applied to this cost for PBR CO2

delivery. 

The operating cost for the pure CO2 case is based on the $40/tonne estimate described above; however, 

this basis was slightly adjusted to account for power. Although the $40/tonne cost implicitly includes all 

power required to capture and transport the CO2 from the flue gas [5], the actual quantity of power

could be substantial. As the plant power balance is an important consideration in the context of this 

study, it was decided that the amount of power associated with the CO2 capture step should be explicitly 

included in the overall power balance in the Aspen model to provide a more representative estimate. A 

basis of 250 kwh/ton CO2 was assumed, based on a preliminary estimate of flue gas carbon capture 

efficiency provided by the EPA [25]. This translates to $16.14/tonne CO2 as an electricity cost, which is 

reasonable at roughly 40% of the total material cost basis. Thus to ensure that this item was not 

double-counted, the resulting power was added to the total plant power balance, and $16.14/tonne was 

subtracted from the cost basis, resulting in an adjusted operating cost of $23.86/tonne CO2.

Harvesting

The next step after production is to harvest and concentrate the biomass to bring the throughput down 

to a manageable range for extraction. While there are many options to accomplish this task, it is critical 

that the selected option be low cost and scalable to process large volumes of water. The most feasible 

options for harvesting are summarized below.

 Bioflocculation: Bioflocculation is a phenomenon where small microalgal colonies spontaneously 

flocculate into larger colonies, which promotes some degree of settling and concentration (up to 

4% solids concentration has been observed [1]). While this is not sufficient for final 

concentration, the process would be the simplest and lowest cost option, as the required 

equipment would essentially consist of settling tanks. For this reason, most open pond studies 

assume bioflocculation as a primary separation method [1, 5].

 Filtration: Filtration has frequently been studied as a means to concentrate microalgal cells, 

either under pressure or vacuum. The process consists of particles either flowing through a 

membrane or being retained on the feed side depending on size. While filtration of larger 

microalgal species has been proven on the lab scale to achieve concentration factors up to 180x 

[22], large-scale continuous filtration is still inefficient, due to the spherical and gelatinous 

nature of cells and their propensity to clog filter membranes [1, 12, 26].
12

 Centrifugation: Centrifugation separates solid and liquid phases using centrifugal force. It is an 

established technology and is an effective method to concentrate most microalgal species, 

achieving up to 150x concentration factor [22]. However, it is very energy-intensive and has 

fairly high capital cost, so its use would be limited to secondary recovery to limit the flow rate 

processed [5, 12].

 Flocculation: Flocculation consists of neutralizing repulsive negative charges on the microalgal

cells to promote cell agglomeration, using a cationic agent as a coagulant. The coagulant can be

a multivalent salt such as alum, although large doses are required and the agent is not 

biodegradable which limits downstream processing options. Alternatively, organic polymers can 

be used, which in addition to neutralizing the cells act as a bridge to physically link flocs together 

[12]. After the coagulant is added, the mixture is gently agitated using a paddle wheel or similar 

device, and the solution is transferred to a settling basin, where the agglomerated flocs either 

settle by gravity or are forced to the surface using flotation methods. While the equipment is 

relatively low-cost, the largest cost can be the material cost of coagulant which can become 

significant depending on the amount of material processed.

Based on this information, it was decided that at least the open pond case would require primary 

concentration using bioflocculation. This is a simple and low-cost process which is important given the 

large amount of water processed at this stage (roughly 150,000 GPM). It is a somewhat speculative 

assumption that this process will work as intended, as the “natural settling” phenomenon is not wellunderstood and has only been observed on a small scale [1]. However, the general conclusion is that 

this is the only economically viable option for primary settling on a large scale [1, 5]. Therefore an 

important aspect of future research is to better understand bioflocculation and how it can be achieved 

on a large scale using different microalgal species. For costing purposes, an equation was used for 

clarifier settling basins by Wang et al. [27], which is based on a compilation done by the EPA of various 

wastewater clarifiers in the US, using data regression to establish a cost equation based on throughput. 

In the Aspen model a conservative assumption was made that the solids are only concentrated from 0.5 

g/L to 10 g/L (1%). For the PBR case, primary concentration is not as critical as for open ponds given that 

the culture is more concentrated at 4 g/L; however, it was decided that primary settling would also be 

used for the PBR case given that the simple settling tank has lower capital and operating cost than 

secondary concentration. The same outlet concentration of 10 g/L was assumed as the open pond case.

For secondary harvesting, flocculation followed by flotation was chosen for both open pond and PBR 

cases. Although a centrifuge is another viable option here, some experts in industry conclude that it will 

not be the most practical option at this scale due to cost and power use [28, 29]. Still, for comparison a

centrifuge is also examined in the sensitivity study. Chitosan (an organic polymer) was chosen as the 

coagulant as it is fairly well-studied, relatively low cost, and biodegrades downstream in anaerobic 

digestion, thus does not build up in the system or cause toxicity concerns [12, 30]. One note to be 

aware of is that chitosan has a relatively high C:N ratio which could exacerbate a nitrogen imbalance in 

the digester, in which case the feed to the digester would have to be adjusted for proper operation [31]. 

Upon formation of flocs, the agglomerated cells would be collected by skimming the top layer of a 

dissolved air flotation (DAF) tank. DAF was chosen over sedimentation because it can achieve higher 
13

solid concentrations and also requires residence times on the order of minutes rather than hours, 

greatly reducing equipment size and cost [32]. DAF essentially works by pressurizing the liquid feed 

along with air, sending the mixture to a retention tank to allow the air to dissolve, then reducing the 

pressure across a valve which releases the dissolved air as very small bubbles. These bubbles attach to 

the flocs as they rise and bring them to the surface where they are skimmed off the top. For an algae 

process, a recycle scheme is optimum as the pressurizing pump is located on the clarified water line 

rather than the feed, thus the cells are not subjected to the pump’s shearing action. This configuration 

is shown below in Figure 5.

Figure 5: DAF process with recycle [33]

This process shown in Figure 5 was modeled in Aspen using the following assumptions:

 10% solids concentration: This represents a 10x concentration factor and is based on one source 

stating 8-10% could be achievable [5]. Although published data is scarce on actual results for a 

DAF system concentrating algae using chitosan, this appears to be a reasonable target.

 Pressurizing to 5 atm: Typical value given by a vendor [34]

 Recycle ratio = 50% of incoming flow: Typical recycle value [33]

 Air/solid ratio = 0.01 mL/mg: Typical value used in wastewater [33]

 Chitosan required = 40 mg/L solution: This value is specific to the species being cultivated, but is 

an average of several studies’ conclusions of 15-80 mg/L required, with 40 mg/L fairly common 

[30, 35]; this range is examined in the sensitivity analysis.

 Power requirement for agitation is the same on a per-flowrate-basis as the paddle wheel power 

use given for open ponds by Benemann & Oswald [5]: This is assumed, but should be fairly 

reasonable as a paddle wheel is also used in this application to agitate the coagulant mixture.

Finally, the cost for the DAF unit was estimated in a similar manner to the clarifier cost equation, where 

a cost based on flowrate was established based on a compilation of DAF units in wastewater 

applications throughout the US [27]. The material cost for the chitosan was set at $4.84/lb [36].
14

Extraction

After the cells are concentrated, the lipid stored inside the cell is extracted and purified to produce the 

oil product. The extraction step is commonly regarded as the most speculative in terms of large-scale 

feasibility, but also among the most costly [5]; thus extraction will be a critical area of research going 

forward to achieve practical algal lipid production. Existing technology commonly employed to extract 

seed oils (e.g. from soy beans, canola, etc.) does not lend itself well to algae oil extraction [1]. Plantbased oils are extracted by crushing the seeds then using mechanical presses and/or hexane extraction

to recover the oil. The process is not as simple for microalgae, given that a cell is very small and 

gelatinous with a thick cell wall protecting the lipids, and even after concentration the solids content is 

still only 10-20%. Some of the more common methods discussed for algal oil extraction are summarized 

below.

 Solvent extraction: As noted, solvent extraction is a common process used to extract seed oils. 

The process involves using an organic solvent such as hexane to exploit the miscibility of the oil 

phase into the solvent. The oil and solvent are then separated downstream and the solvent is 

recycled. While this method is plausible for use in an algae application, it would only be 

effective if the suspension was dried to a low water content [1]. This would introduce significant 

energy costs, both to dry the material as well as to separate the solvent off downstream using 

distillation. This has led some in industry to conclude that this method will not be economical 

on a large scale [29].

 Supercritical Fluid Extraction: Supercritical CO2 could eliminate the problem of downstream 

distillation to separate the solvent from oil. Supercritical CO2 has properties of liquid and gas, 

allowing it to penetrate the cell wall and act as an organic solvent without forming a single liquid 

phase requiring subsequent separation [1]. However, its use has largely been limited to 

removing caffeine from coffee, and is currently too expensive and power-intensive to be 

feasible for a lower-value product such as fuel. However, this would be an ideal application for 

the technology if further advancements are made in the future.

 Mechanical Extraction: Mechanical extraction employs some type of device which disrupts the 

cell wall. Even if not used as the only extraction technique, some form of mechanical extraction 

would likely be required to weaken the cell wall for a subsequent step such as solvent extraction 

[22]. Mechanical devices could include crushers and presses similar to what is used in seed oil 

extraction (but only after dewatering) or alternatively could employ sonication, homogenization,

or bead milling to treat the entire wet slurry [1]. Sonication consists of using high-frequency 

sound waves to create cavitation bubbles which rupture the cell; however, this is a new 

technology which needs further development before it becomes feasible. Pressure 

homogenization uses rapid pressure swings to lyse the cell, and is commercially available on a 

fairly large scale [37]. Bead milling involves mixing the slurry with small beads to lyse the cell 

using shear forces; this could create problems separating the beads back out, and also is only 

available on a relatively small scale [26].

Given these options, it was decided that mechanical extraction using pressure homogenizers would be 

the primary method to extract the oil from the biomass. Solvent extraction, although proven, would 
15

create unnecessary additional drying and separation steps, and is not seen as the appropriate solution at 

large scale as described above. Homogenizers are commonly used in the dairy industry and are also 

being introduced in the wastewater industry to lyse bacteria [38], thus the technology is established and 

has been shown to work for a similar application. Furthermore, of the wet processing devices they are 

available at the closest scale to what is required here.

As described above, the goal of homogenization would be to lyse the cell and release the contents. 

Once the oil is released, it would be purified and removed using a simpler form of solvent extraction 

described by Benemann & Oswald [5]. The process as envisioned would utilize hot oil as the waterimmiscible solvent to extract the lipid phase, similar to existing technology used for extracting betacarotene. However, this would take it a step further and use the lipid product itself as the oil solvent, 

recycling and heating a portion of the purified product and sending it back to the mixture. It is 

important to note that although this is a novel technology as applied to lipid extraction and hasn’t been 

proven on a large-scale, the underlying fundamentals are well established for beta-carotene purification. 

One difference from the Benemann study is that after the recycled oil is introduced to the mixture, the 

Benemann study goes on to separate the phases (oil, water, and spent biomass) using a 3-phase 

centrifuge, which would have a high capital cost and power requirement. Since the water phase will 

now contain high amounts of dissolved proteins and carbohydrates that were released from the cells, it 

should not be recycled back to the process anyway, but should be treated as wastewater. Thus it is not 

important to separate the water and solid phases, and a simple clarifier settling tank would suffice [39]. 

This is somewhat similar to the concept presented by OriginOil, where the cells are ruptured using 

electromagnetism then the various phases simply settle out without mechanical aid [40].

The underlying assumptions behind the extraction step are summarized below:

 90% cell rupturing efficiency: While a maximum cell rupture of 85% is suggested to be currently 

possible [37], it is assumed here that with advancements in technology as well as homogenizers 

designed specifically for algae processing, this can be increased to 90%

 Temperature rise across homogenizer = 25 °C: Per vendor [41]. This is advantageous in that the 

extractability of the oil phase increases with temperature [5]

 Homogenizer power requirement = 200 kwH/tonne solids processed: Per vendor [42]

 Lipid recycle ratio for oil extraction = 5:1 – Conservative estimate based on envisioned recycle 

ratio of up to 6:1 [26]

 Recycle stream heated to 65 °C: Hot oil solvent is envisioned to be heated to 60-70 °C [5]. This 

would be accomplished by cross-exchange with the turbine flue gas stream

 5% lipid loss into the water phase: The recovery of beta-carotene using hot oil extraction is 

>98% [5], so it is reasonable to assume 5% of the lipid is left entrained in the water/solid phase 

sent to anaerobic digestion. Additionally, the same quantity of water is assumed to be 

entrained in the oil phase as a conservative estimate for the water balance.

Given these assumptions, the unit was modeled in Aspen and resulting costs were calculated. The 

maximum capacity for the homogenizers was quoted by a vendor at 17 GPM/unit; at a flow rate of 

roughly 800 GPM through the homogenizers (for both open pond and PBR cases), this equates to 47 
16

homogenizer units. The total installed cost for each unit was quoted at $1MM/unit, thus the 

homogenizer system would cost close to $47MM as economies of scale cannot be assumed here (each 

unit is already operating at maximum capacity). The clarifier settling tank was costed in the same 

manner as the open pond primary settling tank, using a flowrate-based cost equation [27]. The resulting 

streams leaving the extraction clarifier are the oil product and the contaminated water + spent biomass 

solids. Although the oil product may require an upgrading step such as degumming to remove 

phospholipids prior to biodiesel conversion (similar to traditional seed oil biodiesel processing), this step 

was not included as there is no information in the literature regarding the process logistics specific to 

algal-derived oil, and the cost would be marginal relative to the other units.

Spent Biomass Utilization

After the cell is ruptured and the oil product is drawn off, the remaining spent biomass material can be 

utilized in several ways to provide a byproduct credit. As noted previously, the water will contain high 

levels of dissolved protein and carbohydrates after the cell is ruptured, and as such it should not be 

returned to the system. This is the blowdown stream described above in the “Water Delivery” section. 

Thus both the water and solid phases are drawn off and sent to a “byproduct utilization” stage, of which 

there are two main options commonly discussed:

 Animal feed: The protein and nutrient content of the microalgae makes the remaining biomass 

amenable to formulation into specialty animal feeds. Although bulk animal feed may have too 

low of a value for sufficient co-product credit, certain nutrients such as carotenoids have a 

higher value and could be used in specialty feeds [1]. However, some experts in industry 

caution against this option, as extensive testing for toxicity and other properties would be 

required and there is no guarantee that the product could be sold into this market [43].

 Anaerobic digestion: The spent biomass can also be sent to anaerobic digestion to produce 

biogas, which can in turn be burned in a gas turbine to offset the power requirements for the 

plant. This also improves sustainability as it enables the plant to be more self-sufficient and 

either requires less power from the grid or none at all, depending on plant power requirements. 

This has an added benefit of providing partial treatment to the blowdown water as well.

Given these options, anaerobic digestion and power generation is utilized in this study as power usage is 

another important sustainability metric. The anaerobic digestion block in the biochemical ethanol 

Aspen model was assumed to be applicable here [23], so the stream data from the algae model was 

exported to the ethanol model digester inlet stream, and the digester block was modified to include 

properties of the new algae components. For the purposes of anaerobic digestion calculations, the

microalgal cell was assumed to be composed of 25% lipid with the remainder split evenly between 

protein and carbohydrate as a simple first-pass approximation [1]. Of the carbohydrate, the 

composition was split evenly between glucose and cellulose (for the wall structure); although in reality 

the cellulose content may be lower [44], this assumption is conservative in that cellulose does not 

degrade in the digester model and thus does not contribute to the biogas yield. The output from the 

digester model was exported back to the algae model, and the stream sent to a flash drum to separate 

the gas (burned in the turbine) from the liquid (containing nutrients to be recycled and purified water to 

be discharged).
17

Process assumptions behind the digestion and gas turbine units are summarized below.

 Nutrients are required for anaerobic digestion, and the required amount calculated by the 

ethanol model is still valid for this process. The nutrients for digestion are composed of 

“primarily caustic with some phosphoric acid, urea and micronutrients” [23].

 The remaining nutrients after anaerobic digestion are recycled to the microalgal production 

stage, to mitigate the amount of fresh nutrients required

 The water exiting digestion is treated further (described later in “other” capital costs), then 

discharged from the system

 Thermal efficiency = 33%: The generated biogas is combusted in a gas turbine with a 

stoichiometric amount of air. The resulting thermal efficiency to generate power (work 

extracted/higher heating value) was assumed to be 33% [5]

 The turbine flue gas is recycled to the production stage 

After modeling the units in Aspen, capital cost was calculated. All costs for equipment in anaerobic 

digestion were calculated in the same manner as is done in the biochemical ethanol study [23], namely 

by scaling a previous cost estimate by flow rate according to the equation:

  = ( ) 

 

 



The gas turbine capital cost was taken from a vendor quote in Benemann & Oswald, stated as $1,000/kw 

in 1995$ [5].

Other Capital Costs

In addition to the main process units described above, there would be additional systems and 

miscellaneous equipment required to support the process which carry additional capital cost, as 

described below.

Inoculum Production

A culture inoculum would be required for startup as well as to restart the growth process after the 

system is shut down for cleaning, culture crashes etc. This is a cost which isn’t usually included in most 

studies; some studies justify this by assuming that the strain is robust enough to not require inoculation 

[12]. However, this is a step which would very likely be required and it is important to at least estimate 

what the cost might be. For this study, a system described by Benemann & Oswald is assumed [5]. The 

configuration as envisioned would consist of a series of steps cascading from a small lab-scale 

photobioreactor up to a medium-size covered pond before the culture is released to the full scale open 

ponds. The analysis concluded that it would be feasible to build such a system at $9/m2and requiring 

no more space than 10% of the open pond area. Another conclusion was that at such a small scale, 

capital cost would dominate and associated operating costs could be neglected. For the PBR case, it was 

assumed that only half of the resulting cost for the open pond case would be required, as the final 

stages using covered ponds would not be necessary for PBR’s.
18

Land Use

After calculating the land use required for the growth ponds themselves, the total plant land use was 

estimated assuming open ponds require 2/3 of the total land for the plant [5]. This was decreased to 

50% for the PBR case, given that PBR’s are more space-efficient. The land cost used was $3,000/acre 

assuming low-value land not intended for crop production.

Nutrient Delivery

The capital cost for nutrient delivery equipment for the open pond case was set at $1,000/ha in 1996$ 

[5]. For the base case scenario, this translates to $96/(kg/hr) of algae produced, a basis which can then 

be applied to the PBR case.

Secondary Waste Treatment

As mentioned previously, after the blowdown water containing dissolved proteins and carbohydrates is 

treated in anaerobic digestion, a second line item for waste treatment was also included in the 

Benemann & Oswald analysis [5] and is included here. The cost stated for this process was $1,000/ha 

(1996$), which for the PBR case translates to $0.33/m3of total production system volume (which is the 

basis for the amount of blowdown ultimately leaving the system). 

Electrical Supply

A separate line item was included in the Benemann & Oswald study for electrical supply infrastructure

[5]. This was originally set at $2,000/ha (1996$) but is decreased here to $1,000/ha given that much of 

the new equipment costs used here already include all electrical elements. This is an appropriate basis 

for both the open pond and PBR case.

General Machinery

The final “miscellaneous” capital cost item included in Benemann & Oswald was for general machinery

[5]. While it was not stated what this would include, it is assumed here that this includes pumps to 

circulate the material throughout the plant; relative to the large units in each processing step, the 

circulation pumps would carry a marginal cost. This cost was set at $500/ha (1996$).

Operating Costs

Some of the operating costs have been described previously, but all costs are summarized as follows:

Power

All power requirements throughout the plant are met using electricity. This is supplied from the gas 

turbine, and any additional electricity if required would be purchased from the grid. Power cost was 

updated to 2007 at $0.06/kwh per the thermochemical ethanol design report [45], and the same price 

was used for a power credit if the power generated exceeded the total power required. As described 

previously, all units downstream of the growth stage were assumed to operate 24 hours/day and 

electricity use was calculated as such. For the growth step (ponds or PBR’s), nutrients and CO2 would be 

fed only 12 hr/day corresponding to the daylight hours when algae are grown (thus the energy use 

associated with CO2 capture described above was based on 12 hr/day operation). Therefore, given that 

power (an instantaneous rate of energy) is used at a different rate during the day than at night, it is 

more appropriate to report the total amount of energy used annually, as shown in the “Results” section.
19

Nutrients

Nutrients for microalgal cultivation constitute a significant portion of operating cost. Many other 

studies assume that for a first-pass analysis, nutrients can be approximated as “fertilizer” [46]. 

However, the optimal amount of fertilizer (or other nutrient medium) varies widely across different 

studies, even for the same strain [12]. The minimum (stoichiometric) ratio of nutrients to microalgal 

cells was calculated assuming cell N content of 5.5% and P content of 1.1% of dry weight [12], met by 

using urea (46% N) [50] and diammonium phosphate (18% N, 20% P) [51]. This results in a total nutrient

requirement of 1:6.5 ton/ton biomass growth, at a composition of 64% urea and 36% DAP. This ratio 

was used as the “base case” here, and the variation is studied in the sensitivity analysis. As noted 

previously, nutrients from digestion are recycled, thus the amount of fresh nutrients required is lower. 

The associated price for each fertilizer type was extrapolated to the year 2022 according to EIA 

projected oil prices [47] as these prices are correlated. The resulting blended fertilizer cost is $819/ton 

[48], to which an additional 20% was added to allow for additional micronutrients such as iron and 

vitamins [12]. It is important to note that there are other options for nutrient sources. For example, it 

has been suggested that the NOx in the flue gas could potentially provide the required nitrogen, but this 

has not been proven and could present additional pH problems [12], thus this was not examined further. 

Additionally, the plant could be set up to treat wastewater, where some nutrients would already be 

present and thus fresh nutrient requirements would be either mitigated or eliminated.

CO2

As described previously, both cases assume CO2 purified from the flue gas at an adjusted cost of 

$23.86/tonne, which accounts for upstream operating costs (excluding power) to purify the CO2 from 

flue gas. The power associated with this step was estimated at 250 kwh/ton CO2 [25]. The sensitivity of 

the oil production cost on this CO2 price is evaluated in the sensitivity analysis. Additionally, an alternate 

option of piping the total flue gas stream to the facility is considered in the sensitivity analysis, and in 

this case the flue gas is assumed to be free beyond the power required to get it there.

Coagulant

Chitosan is used as the harvesting coagulant in DAF for reasons described above, and is required at 40 

mg/L and $4.84/lb. As both PBR and open pond production utilizes primary settling to concentrate the 

biomass to 1% prior to the DAF unit, the same amount of chitosan is required for either case.

Waste Disposal

A marginal operating cost associated with miscellaneous waste disposal was included in the Benemann 

& Oswald study [5] at a rate of $1,000/ha (1996$). This is assumed here for open ponds, and translated 

to PBR’s at $0.48/(m3system volume)/yr.

Water

As described previously, for the base case brackish water is pumped from underground and assumed to 

be free beyond the associated power requirement. Thus the only water cost is for the fresh makeup 

water to replace evaporative losses in the cooling system for the PBR. Fresh water price was based on 
20

the biochemical ethanol design report at $0.0013/gal [23]. An alternate option of purchasing all 

required water as a utility at this rate is considered in the sensitivity analysis.

Labor and Overhead

The number of operators required to run the plant was estimated as follows:

 Open ponds: 50 ponds/operator (results in 26 operators)

 PBR: 25 hectare/operator (results in 31 operators)

 Primary settling: 1,000 man-hours/10,000 ft2of settler tank area [27] (15 operators for open 

pond, 3 for PBR)

 DAF: 2,800 man-hours/1,000 ft2of DAF area [27] (5 operators for open pond or PBR)

 Homogenizers: 4 operators required (assumed)

 Product clarifier tank: 1,000 man-hours/10,000 ft2of settler tank area [27] (1 operator)

 Digester/turbine: 4 operators required (assumed)

All other employees were set at the same number and pay rate as the biochemical and thermochemical 

ethanol reports [23, 45]. The overhead rate was set at 60% of labor cost, per the biochemical model.

Maintenance, Taxes, Insurance

For the open pond case, maintenance was set at 2% of the installed equipment cost per the biochemical 

model. Insurance and taxes were set at 1.5% of the total installed cost per the same basis. For the PBR 

case, given that the capital cost is much larger, using this same basis would result in unrealistically high 

rates. Thus twice the resulting cost as open ponds was assumed, given that maintenance for PBR’s 

would be higher than for open pond.

Financial Parameters

The remaining financial parameters assumed for this study are as follows:

 Indirect capital cost: rates based on the thermochemical and biochemical models [23, 45]:

 Open pond

 Site development: 9% of installed cost of process equipment

 Warehouse: 1.5% of total installed equipment cost

 Prorateable costs: 10% of total installed cost

 Field expenses: 10% of total installed cost

 Home office and construction: 25% of total installed cost

 Contingency: 10% (given the unknowns associated with algae production)

 Other costs: 10% of total capital investment (see Aden et al report for a 

description of these items)

 Working capital: 25% of operating costs [5]

 PBR: Set all indirect capital expenses to same value as open pond case, as these items 

would all be fairly similar and using the same basis for PBR’s would be unrealistic

 10% return on investment

 20-year plant life

 35% tax rate

 MACRS depreciation
21

Results

Base Cases 

Given the assumptions described above, the resulting oil production cost is shown in Figure 6 for all 

three scenarios presented in Table 1 (note TAG= triacylglyceride, the oil constituent). For consistency 

with NREL’s other current studies, all costs are in 2007 dollars.

Figure 6: Cost to produce 10 MM gal/yr oil at each growth scenario for open pond and PBR production

The oil production cost for the open pond case ranges from $11.25/gal in the current case to $3.99/gal

for the max growth case; this corresponds well with other studies [3]. The cost for PBR production is 

roughly twice the cost for the open pond case. To put these costs into a larger context, it is important to 

note that “max” in this case merely means the maximum algae growth and oil content (based on 

efficiency limits) applied to this specific configuration and associated assumptions. Thus while it results 

in the lowest costs for this study, it does not imply that these are the absolute lowest costs that can ever 

be achieved as there are numerous other process options and even unforeseen additional 

improvements that could be realized as the technology develops. 

The results from Figure 6 show that cost of land is marginal in all cases, while for the open pond case the 

majority of the cost is split fairly evenly between capital and operating costs. The PBR case is dominated 

by capital cost of the tubular production system, although it also has a higher operating cost than open 

ponds; this is mainly due to higher maintenance, tax, and insurance cost (a reflection of higher capital 

cost). The specific cost breakdown, as well as a summary of land and resource requirements, is shown 

further below in Table 2.

Table 2 demonstrates the advantage of PBR’s over open ponds with respect to land and water use, with 

PBR’s requiring <10% of the water demand for the open pond scenario (more than half of which is lost 

to evaporation in the base case), and roughly 50% of the open pond total land requirement. While the 

$0

$2

$4

$6

$8

$10

$12

$14

$16

$18

$20

OP 

base

PBR 

base

OP 

aggressive

PBR 

aggressive

OP 

max growth

PBR 

max growth

Cost of TAG ($/gal)

Cost of Algal Lipid Production

Operating ($/gal of lipid)

Capital ($/gal of lipid)

Land ($/gal of lipid)
22

CO2 requirement per unit weight of microalgae is constant, the required rate of algae production 

decreases in the more aggressive cases due to higher lipid content, thus the CO2 demand also decreases. 

The overall flue gas rate required to meet the pure CO2 demand can be calculated for a given flue gas 

composition. At a bulk vol composition of 15% CO2, 6% O2, and 79% N2 representative of one particular 

flue gas [24], the flue gas rate required to meet the CO2 demand shown here would be approximately 

1.3 – 0.7 – 0.6 MM ton/yr, for each growth case respectively. Note that the CO2 demand shown in Table 

2 accounts for recycling the turbine flue gas to the production stage; in this respect, it could be 

considered that the plant itself has “zero” air emissions, and the only other stream leaving the plant 

(beyond the lipid product) is treated water at the “blowdown” rate shown in Table 2.

As noted previously, the fertilizer nutrient requirement was calculated at a minimum stoichiometric 

ratio of 1 ton/6.5 tons biomass grown, which is further decreased by recycling the nutrients from 

anaerobic digestion. This implies that the nutrients from anaerobic digestion can provide the same cell 

growth rate as the actual fertilizer nutrients. Thus it is important to note that the fertilizer requirement 

shown in Table 2 is an optimistic assumption due to these two factors (although this can be justified 

given that the nutrient cost calculated on the basis presented by Benemann & Oswald [5] would be even 

lower). Even under these optimistic assumptions, the nutrient cost is still a significant portion of the 

total operating costs (roughly 40%). The cost impact of adjusting the nutrient requirements is presented 

in the sensitivity analysis.

In all cases evaluated, there was a net electricity import required after considering all power demands in 

the plant (including CO2 capture) beyond what can be met with the gas turbine, thus there was an 

associated cost to purchase additional electricity from the grid. This stands in contrast to other studies 

which typically show a net power export [5], however these studies do not consider the CO2 capture 

step in the power balance, which in this study was found to be the single largest power draw in the 

plant; if CO2 capture were not explicitly included here, all cases would show a net power export. It is 

important to note that additional assumptions had to be made in the Aspen models (which otherwise 

had no effect on the mass balances) in order to estimate power usage as outlined in the process area 

descriptions above. Thus the electricity requirement presented in Table 2 is more accurate as a 

qualitative comparison (e.g., the open pond case requires more power due to the higher amount of 

water being processed) rather than on an absolute quantitative basis. 
23

Table 2: Summary of results for microalgal biomass production at each growth scenario

Basis = 10 MM gal/yr lipid production Base case Aggressive case Max growth case

Open 

pond PBR Open pond PBR Open pond PBR

Yield

Lipid yield [gal/ton algae produced] 57 57 114 114 138 138

Land Use:

Pond/PBR land size [acre] 4,743 1,897 1,482 593 823 329

Total plant land required [acre] 7,079 3,795 2,212 1,186 1,229 659

Resource Assessment:

Net water demand [MM gal/yr] 9,740 720 3,830 320 2,710 250

Net water demand [gal/gal lipid] 974 72 383 32 271 25

 -Water evaporated [gal/gal lipid] 556 19 174 6 96 3

 -Water blowdown to treatment/discharge [gal/gal lipid] 418 52 209 26 174 22

CO2 consumed [lb/lb biomass produced] 2 2 2 2 2 2

Net CO2 used from offsite flue gas [ton/yr]1290,000 290,000 150,000 150,000 120,000 120,000

Nutrients (fertilizer) required for algae [ton/yr]223,920 23,880 12,000 11,980 10,010 10,000

Nutrients required for anaerobic digester [ton/yr]32,960 3,000 1,440 1,460 1,190 1,200

Chitosan required for flocculation [ton/yr] 690 690 350 350 290 290

Net annual electricity imported [MM kwh/yr]

4

60.2 35.7 27.0 18.8 19.8 16.0

Production Cost:

Overall lipid production cost [$/gal] $11.25 $19.49 $5.11 $10.07 $3.99 $6.62

 -Land cost [$/gal] $0.38 $0.21 $0.12 $0.06 $0.07 $0.04

 -Capital cost [$/gal] $5.10 $13.21 $2.16 $6.75 $1.63 $4.07

 -Operating cost [$/gal] $5.76 $6.07 $2.83 $3.26 $2.29 $2.51

1. After recycling flue gas from gas turbine

2. Net requirement after recycling digester nutrients. Fertilizer composition approx. 64% urea, 36% DAP [12]

3. Primarily caustic with some phosphoric acid, urea, and micronutrients [23]

4. After supplementing with power from gas turbine; includes CO2 capture step. Imported power is purchased from grid
24

Sensitivity Analysis

Given that there are still many uncertainties in developing a microalgal oil production technology and 

thus a number of assumptions had to be made, the potential cost impact of such assumptions can be 

captured in a sensitivity analysis. Individual parameters were varied from the established base case over 

a reasonable range, and the results are shown in Figures 7-8.

Figure 7: Open pond sensitivity analysis

Figure 7 shows that for open ponds, the amount of nutrients required has the highest impact on 

production cost of the variables evaluated. As noted previously, the total nutrient requirement was set 

at a minimum ratio of 1 ton/6.5 tons biomass produced, with fertilizer being used as the primary 

nutrient source after recycling the nutrients from anaerobic digestion (which was a minimal amount at 

roughly 10% of the total nutrient requirement). However, to avoid nutrient limitation, nutrients may be 

added in excess and the optimum amount varies widely [12]. For example, another study states that it 

is reasonable to assume that an “average” microalgal strain will require fertilizer at a ratio closer to 1:3 

[46]. At such a high ratio, the study acknowledged that the associated nutrient cost alone could cause a 

project to become infeasible. This high cost sensitivity is reflected in the figure above. Even though the 

1/6.5 ratio assumed here appears to be the minimum stoichiometric amount and could thus be 

optimistic, it is a reasonable first-pass estimate as further justified by the fact that the nutrient cost 

calculated on the basis of the Benemann & Oswald study [5] would actually be on the opposite end of 

the scale, roughly 5x lower than even this cost calculated here (shown as the “low” limit above). This 

wide lack of agreement in the literature highlights that this is an important aspect of future research.

The rest of the parameters evaluated are fairly self-explanatory, but can be summarized briefly. The 

option of delivering flue gas rather than concentrated CO2 was evaluated with the process differences 

highlighted previously, and the resulting cost impact is shown to be fairly low, roughly $0.50/gal 

improvement. The water source was also examined, with a sensitivity around purchasing water as a 

-$3 -$2 -$1 $0 $1 $2 $3 $4

Lipid extraction (clarifier : 3-phase centrifuge)

Inoculum system (not required : required)

Homogenizers (process 1/2 flow : full flow)

Flocculant required (15 : 40 : 80 mg/L)

Water supply (underground : utility delivery)

CO2 delivery (flue gas : pure CO2)

Nutrients (ref [5] basis : ref [46] basis)

Change to production cost ($/gal)

Open Pond Sensitivity Costs
25

utility at the rate outlined in Aden et al [23] and eliminating the somewhat substantial power 

requirement associated with pumping water from deep underground. This increased the production 

cost which further demonstrates that this is not the optimum choice for water source. Next, the 

amount of chitosan coagulant required was studied, as the optimum amount is not well-published. The 

few studies that were done have found that the amount varies by strain but falls mostly between 15-80 

mg/L , thus this was the range evaluated here. 

The last three parameters shown in Figure 7 are related to capital cost. As mentioned, the optimum 

technology for the extraction step is still fairly speculative, and homogenizers are only one feasible 

option. The base cost assumes that the entire material must be homogenized to achieve 90% cell 

disruption and subsequent lipid release. However, if it were possible to achieve the same results by only 

processing half of the material (analogous to only “weakening” the cell wall), the high capital and power 

cost of homogenization could be reduced, resulting in nearly $1/gal improvement. Next, the Benemann 

& Oswald study [5] assumed the same extraction scheme with the exception that a 3-phase centrifuge 

was used for phase separation instead of the simple clarifier assumed here. The additional capital and 

power use results in a marginally higher production cost. Finally, the cost impact of the inoculum 

system was investigated. As noted, many studies do not include inoculation cost, and additionally the 

cost assumed here was based on very rough assumptions. However, the cost impact is not extremely 

large, as eliminating this unit would result in roughly $0.75/gal improvement.

Figure 8 shows the results of the sensitivity analysis for the PBR system:

Figure 8: PBR sensitivity analysis

As shown in the figure above, the single largest cost item in the PBR system is the cost of the tubes 

themselves, as this cost has a very high economic impact. As described previously, the base cost of 

$1.50/ft is for a similar-size HDPE pipe, as an approximation to the specialized UV-coated LDPE material. 

While this is a fairly low price per foot, the system would be composed of nearly one million 8cm ID x 

-$6 -$4 -$2 $0 $2 $4 $6

Water supply (underground : utility delivery)

Flocculant required (15 : 40 : 80 mg/L)

CO2 delivery cost ($5 : $40 : $75/tonne)

Nutrients (ref [5] basis : ref [46] basis)

Inoculum system (not required : required)

Harvesting method (DAF : centrifuge)

Airlift column cost (5% : 20% : 50% of tube cost)

Tube cost (-50% : $1.50/ft : +50%)

Change to production cost ($/gal)

PBR Sensitivity Costs
26

80m pipes to accommodate the design specifications listed above in the “PBR” section, thus this cost 

becomes very large. The +50% range results in a +25% overall production cost impact, implying that the 

cost of the tubes alone accounts for nearly half of the production cost. Further compounding the large 

uncertainty here is that the price quoted in other studies varies even more than this 50% range. For

example, one study used a cost equivalent to $1.45/m2of tube surface area which would result in a 

$10/gal cost reduction relative to the baseline here [16], while another study assumed $150/m2on an 

areal basis, which would increase the cost by $7.50/gal [12]; furthermore the latter study stated that 

such a cost basis was even quite optimistic. Such wide variations highlight the challenge in determining 

a credible production cost for PBR’s at this stage.

The next cost item evaluated was the cost of the airlift columns. As these are generally specialized and 

custom to the application, a cost estimate was not available for these columns. Thus the cost was 

assumed to be 20% of the tube cost given that these would be fairly simple in design and far fewer 

would be required than the number of 80-meter tubes. It would seem feasible for this cost to range 

anywhere between 5-50% of the tube cost, and over this range the cost impact is not overwhelming.

Many PBR studies assume the harvesting would be done in one step using centrifuges since the biomass 

is more concentrated for PBR’s than open ponds [22]. However, for consistency the DAF option was 

assumed in both cases for this study. As shown in the sensitivity, the use of a centrifuge instead of DAF 

appears on a first pass to result in higher production cost, although only marginally.

Another important parameter is the CO2 delivery cost. As described previously, the base cost assumed 

was $40/tonne [5] (adjusted internally to $23.86/tonne after removing CO2 capture energy), and varying 

this parameter over a reasonable range from $5-$75/tonne has a cost impact of roughly +$1/gal (the 

same degree of cost impact is equally applicable to open ponds). Thus the production cost is somewhat

sensitive to CO2 price, although not as strongly as some of the other parameters evaluated. The 

remaining operating cost items evaluated in the PBR sensitivity were discussed above for open ponds.

A final note should be mentioned regarding economies of scale. Although 10 MM gal/yr oil production 

is not excessively large, the total system throughput is considerably higher due to the dilute nature of 

the process. Thus even at the scale evaluated here, the associated flow rates are much higher than the 

largest wastewater treatment plants in the US [42]. Consequently, all of the equipment would already 

be modular and multiple units would be required to accommodate the flow rate in excess of any given 

unit’s maximum capacity. Therefore, it is not expected that economies of scale would play a significant 

role in reducing the production cost at larger plant scales (for example doubling capacity to 20 MM 

gal/yr). However, as noted above if the equipment itself could be improved and tailored specifically to 

microalgae production, significant cost improvements could result (i.e. developing a low-cost material 

that could be produced in bulk for PBR production, modifying homogenizers to handle more flow etc).
27

Concluding Remarks

The overall lipid cost shows that in the near-term, the economics of a microalgal biofuel facility may only 

be feasible if production also targets higher-value coproducts which are produced commercially today. 

However, the production cost could improve substantially with identification of algae species capable of 

maintaining high sustained growth rates with a high lipid content; thus this is one major area important 

for future research. Additional critical areas as outlined previously include reaching a consensus on the 

optimum nutrient requirements, understanding the mechanisms behind bioflocculation to prove its 

large-scale feasibility, and developing an effective low-cost method for lipid extraction. Additionally, 

there is a large potential for cost reduction associated with development of a low-cost material suitable 

for use in a PBR system. 

The intent of this report was to provide a first-pass analysis of the near-term feasibility of microalgal 

lipid production, by performing a ground-up estimate of material balances using Aspen modeling as well 

as equipment costing by individual unit. This is a more detailed approach than has been taken in most 

prior studies, but still carries a relatively high degree of uncertainty in both the process steps as well as 

costing methods, intrinsic to the early state of the technology (most of the processes have only been 

demonstrated for algae applications on a lab or small scale). The results presented here are specific to 

one set of processing options and assumptions, among many other feasible solutions. However, both 

the assumptions and results correspond well to other studies [3, 7].
28

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[18] Personal communication with Secor (pipe vendor), October 2009.
29

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419-422.

[31] Personal communication with Phil Pienkos (NREL), November 2009.

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report prepared for the Department of Energy (1984), contract no. DE-AC02-83CH10093

[33] Armenante, Piero. “Flotation.” PowerPoint notes, http://cpe.njit.edu/dlnotes/CHE685/Cls05-1.pdf

[34] DAF vendor website, http://www.envirowise.gov.uk/uk/Dissolved-air-flotation-DAF-applicationand-design-.html

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of Coagulated By-Products.” Proceedings of the First International Conference on Chitin/Chitosan

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EPA 832-R-06-005 (2006), publicly available online at: 

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[40] From OriginOil website, http://www.originoil.com/technology/low-cost-oil-extraction.html

[41] From homogenizer vendor website, http://www.microsludge.com/microsludge/in_detail.php

[42] Personal communication with Microsludge (homogenizer vendor), September 2009.

[43] Willett, Elizabeth. “Mars Symbioscience.” Presented at the 2009 Algae Biomass Summit, San 

Diego, CA (October 2009).

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online: http://www.biofuelreview.com/content/view/1897/1

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31

Appendix

Capital and operating cost breakdowns for each growth case are summarized below.

Table 3: Open pond capital cost breakdown

Base Aggressive Max growth

Ponds $15,865,290 $4,957,903 $2,754,391

Mixing (paddle wheels) $13,221,075 $4,131,586 $2,295,325

CO2 delivery + sumps $10,611,925 $3,961,923 $2,566,174

Primary harvesting (settling) $30,265,124 $14,119,617 $11,547,902

Secondary harvesting (DAF) $1,352,781 $621,301 $501,157

Cell rupturing $47,000,000 $24,000,000 $20,000,000

Clarifier tank $222,920 $50,953 $42,007

Water/Nutrient/Waste/Electrical Supply $19,038,348 $5,949,484 $3,305,269

Land Costs $21,238,326 $6,636,977 $3,687,209

Anaerobic Digestion $6,370,177 $4,445,692 $4,700,790

Gen-Set $13,171,225 $6,399,650 $5,267,786

Inoculum production system $23,797,935 $7,436,855 $4,131,586

General machinery $1,322,107 $413,159 $229,533

Initial Water Charge $0 $0 $0

TOTAL INSTALLED DEPRECIABLE 

CAPITAL $182,238,906 $76,488,122 $57,341,920

TOTAL INSTALLED NONDEPRECIABLE CAPITAL $21,238,326 $6,636,977 $3,687,209

TOTAL INSTALLED CAPITAL $203,477,232 $83,125,099 $61,029,129

Table 4: Open pond operating cost breakdown

Base Aggressive Max growth

Power $3,524,525 $1,583,830 $1,158,392

Power Flue Gas Supply $0 $0 $0

Nutrients (fertilizer) + wastewater nutrients $24,763,242 $12,405,622 $10,345,256

CO2 $6,344,268 $3,190,159 $2,559,819

Flocculant $6,678,039 $3,381,286 $2,789,561

Waste Disposal $2,748,351 $858,860 $477,144

Water $0 $0 $0

Labor and Overhead $6,966,277 $4,092,688 $3,517,970

Maintenance, tax, ins $6,618,268 $2,782,828 $2,089,100

Total Algal Pond Op Costs ($/yr) $57,642,969 $28,295,272 $22,937,243
32

Table 5: PBR capital cost breakdown

Base Aggressive Max growth

Photobioreactor equipment $642,527,429 $200,715,037 $111,558,210

Mixing (paddle wheels) $0 $0 $0

CO2 delivery + sumps $740,367 $276,413 $179,035

Primary harvesting (settling) $3,703,417 $1,727,705 $1,413,742

Secondary harvesting (DAF) $1,359,727 $616,989 $501,200

Cell rupturing $47,000,000 $24,000,000 $20,000,000

Clarifier tank $222,920 $51,062 $42,132

Water/Nutrient/Waste/Electrical Supply $4,474,680 $2,005,971 $1,564,527

Land Costs $11,383,743 $3,557,420 $1,976,344

Anaerobic Digestion $3,989,533 $3,558,729 $3,237,034

Gen-Set $13,350,757 $6,488,982 $5,342,252

Inoculum production system $11,898,967 $3,718,428 $2,065,793

General machinery $528,843 $165,263 $91,813

Initial Water Charge $0 $0 $0

TOTAL INSTALLED DEPRECIABLE 

CAPITAL $729,796,640 $243,324,580 $145,995,739

TOTAL INSTALLED NONDEPRECIABLE CAPITAL $11,383,743 $3,557,420 $1,976,344

TOTAL INSTALLED CAPITAL $741,180,383 $246,881,999 $147,972,083

Table 6: PBR operating cost breakdown

Base Aggressive Max growth

Power $2,088,003 $1,102,799 $937,955

Power Flue Gas Supply $0 $0 $0

Nutrients (fertilizer) + wastewater nutrients $24,740,772 $12,392,695 $10,335,425

CO2 $6,268,379 $3,152,476 $2,633,460

Flocculant $6,762,571 $3,381,286 $2,789,561

Waste Disposal $183,223 $57,257 $31,810

Water $125,122 $24,111 $13,395

Labor and Overhead $6,295,773 $3,613,756 $3,065,162

Maintenance, tax, ins $14,257,046 $8,896,888 $5,337,682

Total PBR Op Costs ($/yr) $60,720,890 $32,621,268 $25,144,449