Attachment 4 - Process Design Basis Documentation.pdf

For the best experience, open this PDF portfolio in

Acrobat X or Adobe Reader X, or later.

Get Adobe Reader Now!

http://www.adobe.com/go/reader

To: NREL Procurement and Subcontract Vendors From:

Date: 4/4/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Isomerization Stripper

Unit Operation: Isomerization Product Stripper

Feed Rate(s): 1.00 BBL/DAY of isomerized distillate

Assumptions:

• Standard conditions are 60 deg F and 12 psia on South Table Mountain Campus

• 10 theoretical stages in the column

• Cooling water supply is available at 24 deg C (approx. 75 deg F) at approximately 50 psig

• BBL/DAY = 42 gal/day

Heat and Material Balance Table:

INLET(S)

Stream Description Mass Flow, lb/hr Volume Flow, BBL/DAY

Molar Flow, lbmole/hr

Isomerization Reactor Effluent

11.45 0.9978 0.05

OUTLET(S)

Stripper Overhead Vapor

0.02 4.86 SCFD 0.0005

Aqueous Phase

0.0 0.0 0.0

Organic Phase

0.25 0.022 0.0012

Stripper Bottoms 11.19 0.97 0.044

Process Design Basis:

Isomerized distillate is fed to the top stage of a 10 theoretical stage column. The overhead vapor is condensed and can be reflux back into the tower for improved mass transfer, or drawn off as a light product consisting of hydrogen through pentane material. The material may be pressured to a storage tank, the operating pressure is 13 psig on the overhead of the stripper. A flanged connection for a pump upstream of the control valve should be placed if a pump is needed. The overhead receiver should be sized for five minutes hold up and surge time. The vessel could be sized larger if more space is necessary for level and pressure control instrumentation.

The column is reboiled by a clamshell reboiler. The reboiler should provide approximately 4020 BTU/hr. The bottoms stream should be cooled with an air cooler and a trim cooler, providing approximately 4000 BTU/hr of cooling. A pump for level control and flow measurement was added and delivers 150 ft of head. The bottoms temperature is 650 deg F prior to air cooling and cooling water trim cooling.

If the bottoms product is temperature sensitive and cannot be as high as 650 deg F, the distillation column will be rated for full vacuum and a vacuum pump or ejector could be added to the overhead to accomplish this mode of operation. The ejector or vacuum pump is a separate equipment module but would be appropriate to use in tandem with this module. Another option would be to inject live steam into the bottom of the tower to strip the light end material from the distillate product. This tower design does not currently allow for that capability.

The tower internals should be designed to remove easily between research experiments. Packing or cartridge trays are feasible as internals for this tower. The tower itself was designed with a height equivalent theoretical plate (HETP) of 1 ft per stage. This tower was modeled with 10 theoretical stages. The packed section of the tower will be 10 ft in height and additional elevation is necessary for liquid distribution and collection as found in a commercial tower. The sump should be sized to hold five minutes of hold up capacity.

Metallurgy for these vessels should be constructed of 316 SS at a minimum.

This process will need to be controlled via a programmable logic controller (PLC) using a human machine interface (HMI). Currently the facility uses Opto 22 but is open to other hardware options.

The facility HMIs are on Dell Computers located within the central control room of the pilot plant.

Equipment List

Distillation Column

Internal Diameter, in

Packing Height, ft

Vessel Height, ft MAWP

Stripping Section

1.75 10 12 50 psig

Pumps

Total Developed Head, ft Brake Horsepower, hp Overhead Light Organic Product Pump

150 0.00002

Bottoms SAF Product Pump 150 0.0009

Heat Exchangers

Duty, BTU/HR MAWP Type of Exchanger Overhead Condenser

103 50 psig Shell and Tube

Stripper Reboiler 4020 N/A Clamshell Heater Bottoms Product Air Cooler

3100 100 psig Air Cooler

Bottoms Product Trim Cooler

905 100 psig Shell and Tube

Date: 6/11/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Kerosene/Distillate Isomerization

Unit Operation: Kerosene or Distillate Isomerization Fixed Bed Reactor Unit Operation

Feed Rate(s): 1.0 BBL/DAY of distillate or kerosene range material and approximately 1900 SCFD

• BBL/DAY = 42 gallons/day

INLETS

Stream Description Mass Flow, lb/hr Molar Flow, Volume Flow, Stabilized Distillate 11.47 0.045 1.002 Hydrogen 0.417 0.207 1882 SCFD

OUTLETS

Stream Description Mass Flow, lb/hr Molar Flow, Volume Flow, 1st Stage Compressor Drain

0 0 0

2nd Stage Compressor Drain

3rd Stage Compressor

High Pressure Flash

0.43 0.203 1845 SCFD

0.0 0.0 0.0

Low Pressure Flash

0.014 0.005 43 SCFD

Isomerized Product 11.45 0.05 0.9978

Stabilized distillate is transferred to the Isomerization unit from either the Hydrotreater Stripper Bottoms pumps or from a feed tank. The hydrogen feed is preheated up to 400 deg F, which requires a duty of 283 BTU/hr. The lipid feed must be cooled when pumped straight to the isomerization unit from the hydrotreater. A duty of -2203 BTU/hr is necessary to cool the feed to 400 deg F. The hydrogen is delivered via a four stage compressor. The outlet temperature of each stage was limited to 250 deg F to minimize damage to the seals and increase the reliability of the machine.

The hydrogen to hydrocarbon ratio is 4.5 at the reactor inlet. Hydrogen is not consumed unless distillate is cracked into light ends, which does occur with the acid function found in isomerization catalyst. The reactions were modeled in an equilibrium reactor because isomerization reactions stop when they reach equilibrium. A pressure drop of 5 psig was assumed across the catalyst bed.

A small amount of cracking was assumed in order to model the stripper downstream of the isomerization unit.

The designed flux is 3000 lbs/hr/ft2 for this fixed bed reactor. The flux matches the Hydrotreater.

Oftentimes a bed of isomerization catalyst is in the same reactor as the hydrodeoxygenation catalyst, so the same flux can be used. This led to an inner diameter of 0.85 inches. The bed height of the reactor was sized using a liquid hourly space velocity (LHSV) of 4 hr-1. The bed height was calculated at 17 feet tall, with an additional 1 – 2 feet of height for feed distribution and collection.

The total reactor height is 19 ft. Due to height limitations, the reactor needs to be split in half to accommodate the LHSV of 4hr-1. Each reactor should be 10.5 feet tall with 8.5 foot catalyst bed height. The metallurgy listed below is a minimum requirement.

Vapor, aqueous phase and organic phases are separated by gravity downstream of an air cooler and cooling water trim cooler. The air cooler duty is 1300 BTU/hr and the trim cooler duty is 1150 BTU/hr.

The reactor effluent is flashed and separated in a high pressure flash drum operating at 1615 psig.

The vapor will be metered, sampled, and routed to the flammable vent header. The aqueous phase will be sampled, metered, and collected with the other aqueous waste from the downstream low pressure flash drum.

The organic phase is then routed to a downstream low pressure flash drum operating at 115 psig.

Again, the vapor, aqueous and organic phases are separated via gravity. The vapor and aqueous phases are metered, sampled, and routed appropriately. The organic phase is pressured to a downstream stripper where any remaining light ends can be removed from the final fuel product.

Both the high pressure and low pressure separators should be sized for a surge and holdup time of five minutes. The vessels could be larger to accommodate the necessary instrumentation for adequate level and pressure control.

Reactors

Catalyst Bed Height, ft

Vessel

MAWP Metallurgy

Reactor 1 0.85 8.5 10.5 2500 psig Alloy 347 Reactor 2 0.85 8.5 10.5 2500 psig Alloy 347

Hydrogen Heater 285 2500 psig Undetermined Hydrotreated Feed

-2205 2500 psig Air Cooler

Reactor Effluent Air

-1130 2500 psig Air Cooler

Reactor Effluent

-1150 2500 psig Shell and Tube

H2 Compressor – 1st Stage Cooler

-347.5 500 psig – Shell 300 psig – Tube

Shell and tube

H2 Compressor – 2nd

-178 700 psig – Shell 420 psig – Tube

H2 Compressor – 3rd

-178 1250 psig – Shell 750 psig – Tube

Compressors

Adiabatic Head, ft Brake Horsepower, hp H2 Compressor 1st Stage 245,100 0.074 H2 Compressor 2nd Stage 253,600 0.068 H2 Compressor 3rd Stage 258,000 0.068 H2 Compressor 4th Stage 263,300 0.0695

Date: 6/4/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Liquid Liquid Extraction

Unit Operation: Liquid Liquid Extraction

Feed Rate(s): 1 BBL/DAY of Mixed Esters and 10% Water Wash for Methanol Removal

• Liquid liquid extraction has 5 theoretical stages

• Height Equivalent for a Theoretical Plate (HETP) for each theoretical stage is 3 ft in LLE column

• Solvent Recovery Column has 4 theoretical stages with HETP of 3 ft per stage.

Ester Feed 12.59 1 0.05 Water Wash 1.259 0.09 0.07

Glycerin Column Feed 1.525 0.11 0.075 Solvent Recovery

0.097 26.76 SCFD 0.0029

Overhead Liquid

0.4175 0.03 0.0015

Bottoms Product

11.81 0.94 0.04

Process Design Basis Narrative:

Liquid Liquid Extraction (LLE)

The unit is designed with a liquid liquid extraction column to water wash a mixed ester feed and a drying stripper to remove the water and methanol from the mixed ester stream. This is based on biodiesel production facility operations commonly found in bio-refineries. The liquid liquid extraction column water to move the methanol from the ester stream into the water phase in order to more easily separate the methanol, glycerin and ester. This method is less energy intensive than a distillation column separating all three materials.

The water wash was designed to be 8 – 10 wt% of the mixed ester feed. In many cases for liquid liquid extraction, the solvent to raffinate ratio is 3 to 1 or up to 6 to 1. If this were the case, a LLE column would need a diamter of 2 inches. The current design of this LLE column is a 0.8 inch diameter. The LLE column was designed with five theoretical stages and an HETP for packing of 3 ft was used based on data from a similar system operating commercially under very similar pressures and temperatures.

The LLE column pressure is controlled with a nitrogen bubble at the top of the column. There is a give/take line that will allow for pressure control on these columns. If the columns are glass, they will not be rated for anything above atmospheric pressure, however if the columns are constructed of steel, they will need this pressure control scheme. There is also a liquid level control on the top and bottom of the LLE column to control the interfaces within the column. The streams exiting the column will have flow indication, including the vent stream to the thermal oxidizer vent line.

The construction of both columns should allow for changing internals easily to accommodate changes in research.

Solvent Recovery Column

The solvent recovery column is designed to separate methanol and water from the esters. Methanol is used as a solvent in biodiesel production. Recovering as much methanol as possible is economically desirable.

The column is a water and methanol stripper. The distillation column operates under a vacuum of up 15 mm Hg. The feed is introduced at the top of the tower on tray 1. The tower should be constructed with multiple feed nozzles for future research needs. The reboiler is a clamshell heater.

The tower is designed as a packed column with four stages. The HETP of random packing is assumed to be 3 feet tall based on other systems similar to this. That requires the column height needs to be 12 feet plus an additional 2 – 3 feet of liquid distribution hardware and vapor/liquid disengagement. The tower should be designed to be 15 feet tall. The column diameter is 3 inches.

The overhead receiver should be sized for a minimum of five minutes of surge and five minutes of hold up time. The vessel will probably need to be larger than necessary in order accurately measure a level. A pump will be necessary to recirculate the solvent. Another pump is necessary to move product from the tower sump to a product tank.

The construction of the solvent recovery column should allow for changing internals easily to accommodate changes in research.

Packing

Solvent Recovery Column

3 12 15 50 psig & Full Vacuum

Hastelloy

Liquid-Liquid Extractor

Packing

Ester Wash Column

0.8 15 17 50 316 SS

Separators

Height (Tangent to Tangent), in

Diameter, in MAWP

Overhead Receiver Size to have 5 minute surge and holdup at a minimum.

50 psig

Solvent Recovery Bottoms Product

160 0.001

Solvent Recovery Solvent Recycle

130 0.00003

Reboiler in Solvent Recovery Column

2136 Not applicable Clamshell Heater with six zones

Date: 4/1/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Solvent Extraction

Unit Operation: Solvent Extraction

Feed Rate(s): Based on feeding 5000 kg/hr of algae slurry to pretreatment reactor with 15% solids.

Solids contain 20 wt% lipids.

• The algae slurry contains 15 wt% solids.

• The solids are 20 wt% lipids

• 55% of the lipids are triolein (C18 triglyceride), the remainder are tripalmitin (C16 triglyceride)

Stream Description Mass Flow, lb/hr Volumetric Flow

Flashed Preatreatment

14.92 1.75 0.045

Hexane Solvent Makeup

29.98 3.1 0.35

Ethanol Solvent Makeup

12.85 1.1 0.28

Stream Description Mass Flow, lb/hr Volumetric Flow, 1st Stage Decanter Water Vapor Phase

0.214 108 SCFD 0.012

2nd Stage Centrifuge Water Phase

0.013 0.001 0.0007

Solvent Purge 47.03 5.02 0.64 Lipid Product 10.5 1.06 0.015

The solvent extraction unit is designed to be able to extract valuable products from a mixture containing 5 – 20 wt% solids and over 50 wt% water. Streams similar to this composition are generally created in processes using terrestrial biomass as well as other biological feeds such as algae slurry. The basis for this heat and material balance is an algae slurry containing 15 wt% solids. The solids contain 20 wt% lipids, which need to be extracted with hexane and ethanol.

The slurry is treated in the existing pretreatment reactor in the south high bay using sulfuric acid and steam. From there it is flashed to drive off noncondensibles and some of the water vapor. This is not a part of the scope of this project, but is included for reference.

The slurry is then fed from a tank to a decanter, where the water is separated from the slurry as much as possible. The feed pump requires a total developed head (TDH) of 195 ft. The decanter will be designed so that the solids are moving countercurrent to the water being removed at a slower pace in order to keep the system continuous, as opposed to batch. This does not always work well for solids that have a “slimy” consistency because they are not able to be moved efficiently.

Following the decanter there are two centrifuges used in a counter-current process flow. Fresh solvent is a volumetric mixture of 1 part ethanol, 3 parts hexane. Because this system tends to foam, ethanol is added to help reduce the foaming and decrease the emulsion formed. The solvent mixes with the 2nd stage centrifuge organic phase outlet and feeds the 3rd stage centrifuge. The 2nd stage centrifuge operates with an inlet pressure of 10 psig and it was assumed that the outlets are atmospheric pressure. A water byproduct stream is pumped to aqueous phase storage for waste water testing and possible stripping. A pump is required to deliver 70 ft of TDH. The lipid rich stream is combined with solvent and the pump needed to move those to the 3rd stage centrifuge needs 126 ft of TDH.

The 3rd stage centrifuge separates the hexane, ethanol, and any remaining water from the lipid stream. The solvent rich stream is recycled to the 2nd stage centrifuge. This requires a pump with 145 ft TDH. The products, lipids, contain a small amount of hexane which can be recovered in a stripper downstream if necessary. A pump requiring 155 ft of TDH is required to get the stream to a stripper or a downstream storage tank.

Lipid Feed Pump 195 0.0016 1st Stage Centrifuge Water Pump

70 4.9x10-7

Solvent Makeup Pump 125 0.04 Solvent Recycle Pump 145 0.03 Lipid Product Pump 154 0.003

Date: 3/12/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Thin Film Evaporator

Unit Operation: Thin Film Evaporator or Wiped Film Evaporator

Feed Rate(s): 1.5 gpm – 2 gpm of still bottoms/algae slurry/enzymatic hydrolysate

Did not create one because this unit operation is generally an “off the shelf” purchase. Will need power requirements for the unit to determine building power limitations when this is operating simultaneously with other unit operations.

Process Design Narrative:

A thin film evaporator is needed for the IBRF high bay, preferably one that can process flammable materials. The feed rate is expected to be 1.5 gpm up to 2 gpm of fermentation broth, algae slurries, and/or enzymatic hydrolysate with between 10 – 15 wt% solids. The evaporator needs to evaporate between 500 lb/hr – 750 lb/hr of water. If there is a boiler reservoir, the reservoir needs to hold a minimum volume (less than 200 L) so that the WFE can be used for smaller batch systems with adequate level control. If there is no boiler, the heating fluid can be controlled via a temperature control valve on the inlet or outlet depending on the fluid, with appropriate temperature indication on the evaporator body.

There are many fermentations, both anaerobic and aerobic, that research requires to produce renewable fuel precursors. These precursors are oftentimes acidic, and therefore the materials of construction of the evaporator needs to be rated for a pH of approximately 3.5 up to 7.5.

In addition to acidic fermentation broth, the evaporator will also process slurries containing fatty acids that are sensitive to heat and will degrade with high residence time or too much heat input.

Therefore, a vacuum operation is necessary for the evaporator. The evaporator may be steamed for cleaning and needs to be rated for low pressure saturated steam of approximately 15 psig. Clean in place (CIP) connections are also required for decontamination efforts after research campaigns are completed.

A vertical wiped film evaporator is preferred because of plot space limitations. The high bay ceiling height is approximately 25 ft tall, and that will limit the total height of the equipment.

A feed tank, pump, flow meter and control valve are necessary along with the remainder of the instrumentation needed to monitor and control temperature. These items will need to be modular so that they can possibly be used on other process units. The heating jacket will need its own portable heating system, glycol based or possibly hot oil based. There is also an opportunity to use the house saturated steam available at approximately 350 psig or 35 psig.

Cooling water supply is available at approximately 50 psig and 75 deg C. The temperature does fluctuate depending on the time of year.

Product tanks also need to be engineered as a part of this system as well. These should be modular in design like the feed tank system. All tanks need level control, flow meters, scales, pressure relief valves, and pressure instrumentation.

Date: 6/12/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Anaerobic Digester

Unit Operation: Anaerobic Digestor

Update: Based on new feedback from researchers the new desired vessel volume is 100 L. This represents a 40x reduction in scale relative to the process design basis listed below. Upstream solids feeding needs to be incorporated into the design to accommodate 5 kg/hr max throughput with turndown to 0.1 kg/hr.

Feed Rate(s):

Liquid phase:

12 lb/hr 10-25 wt% solids slurry, anaerobic digester is operated fed-batch with an average residence time of 25 days, resulting in a steady-state volume of 4000 L

• 1 bpd SAF from biogas is not possible, targeting 1 gpd possibility o Desired products will require extraction from liquid phase (e.g., volatile fatty acids)

• Operations are heavily feedstock dependent o Conventional CSTR can handle slurry containing 10-25 wt% solids

• Feedstock will likely be municipal solid waste (MSW), so this unit will be operated in a biosafety level 2 (BSL-2) facility

• Mesophilic operation from 35-60°C

• This unit is nominated to replace hollow-fiber membrane bioreactor

Inlet Biogas VFA (LLE) Waste (biosolids) Flow (lb/hr) 12 0.5 1-5 6-11 T (°F) 77 95 95 95 P (psia) 17 12 12 12

Process Flow Diagram:

Two feedstocks showing promise in SAF pathways are CO2 and volatile fatty acids (VFA), organic carboxylic acids with carbon numbers ranging from two to seven. Anaerobic digestion can produce both CO2 and VFA’s from waste carbon streams in addition to methane in the biogas. Biosolids and wastewater are two other carbon-containing streams which could be utilized in conversion pathways. Wild-type colonies inside anaerobic digesters can be pathogenic, which is compatible with the future BSL2 process area on the east side of the revitalized IBRF North High Bay.

Anaerobic digestors tend to be very large volumes, in the hundreds and thousands of cubic meters.

This facility cannot tolerate equipment at that scale, so the possibility of producing 42 gallons of SAF per day is not possible. However, the amount pathways utilizing MSW as a feedstock are only increasing and this is a unique opportunity to provide research on actual anaerobic digester product streams.

Biochemical processes involved in anaerobic digestion are varied and slow. This requires an average residence time of 25 days to fully convert incoming feedstocks. Homogenization upstream of the digester is required to achieve a slurry containing 10-25 wt% solids. This can be pumped directly into the anaerobic digester vessel.

Operating temperatures range from ambient to mesophilic (35-60°C). This is achieved primarily with the steam jacket on the primary vessel. A recirculating loop will aid with mixing and internal heat transfer. A mechanical agitator and internal baffles will further contribute to good mixing.

Slipstreams off the recirculating loop provide a means for wastewater and biosolids removal to maintain a constant volume. Optional VFA removal via solvent extraction is also possible at this location. Finally, biogas is removed through a biofilter and H2S scrubber overhead. Additionally, iron hydroxide may be fed into the digestion to suppress H2S production, instead precipitating FeS.

Steam sterilization is the primary kill method, with internal manifolds available in all three vessels.

This is preferable to waste generated from chemical kill methods.

NATIONAL RENEWABLE ENERGY LABORATORY

National Renewable Energy Laboratory

15013 Denver West Parkway

Golden, CO 80401

Operated for the US Department of Energy by

The Alliance for Sustainable Energy, LLC

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Absorber

Unit Operation: Pressure- and/or Temperature-Swing Absorption

Gas phase feed: 29 lb/hr syngas (1/6, 1/6, 2/3 of CO, CO2, H2)

• This carbon flow rate is 350 mol C/hr which is roughly equivalent to 1 bpd (42 gpd) SAF

Liquid phase adsorbent: 106 lb/hr (Methanol)

• Feed pressure comes from upstream process

• Water is either removed before absorber (cf. adsorber) or is removed downstream (cf.

distillation)

• Throughput is low enough to allow random packing mass transfer material

• HETP varies between 3-12 inches

• Minimum absorption temperature limited to -20°C to avoid exotic refrigeration systems

• Key performance indicators are:

o Less than 5 mol% CO2 in methanol recycle o Greater than 75 mol% CO2 in Gas B product stream o Less than 2 mol% methanol in Gas B product stream

Rectisol separation of CO2 & H2:

A B C E G I GasA GasB Flow (lb/hr)

29 123 123 111 111 111 13 13

T (°C) 20 2 0 75 40 -20 5 2 P (bar) 21 20 4 4 4 21 1 1

CO2 0.167 0.0929 0.0929 0.0370 0.0370 0 0.0127 0.774

CO 0.167 0.0163 0.0163 0 0 0 0.170 0.215

H2 0.667 2.24e-8 2.24e-8 0 0 0 0.816 2.95e-7 CH3OH 0 0.891 0.891 0.963 0.963 1 8.07e-4 0.0108

Energy requirements

HX1 HX2 HX3

Q (W) 1370 1500 130 T1 (°C) 22.4 74.9 20 T2 (°C) -20 40 5

The model process for the absorber unit operation is the Rectisol process. Developed by Lurgi AG, currently a subsidiary of Air Liquide, the process uses cold methanol to remove acid gases from gas streams. The process is primarily used in synthesis gas (syngas) purification applications.

Biomass-to-fuels processing famously suffers from heteroatom impurities and their associated compounds. The Rectisol process removes most of these including: CO2, H2S, COS, NH3, HCN, CS2, mercaptans, BTX, and carbonyls. Being a binary separation process, one can focus either on the absorbate (Gas B) or effluent (Gas A) stream. For instance, one can achieve high-purity CO2 or NH3 capture and recycle with the Rectisol process. Methane purification from biogas streams is another viable application.

Notably, this design does not proactively address contaminant building in the solvent phase (e.g., water). This can be addressed by pairing with the modular distillation column to provide contaminant removal as needed.

A variety of gas-liquid and gas-gas heat exchangers are employed to conserve energy within the process. The refrigeration duty required to achieve continuous absorption operations at 40°C results in high capital and operating costs, so energy conservation matters.

The liquid-phase absorbent for this design basis is methanol. It is understood methanol has a low flash point and poses human toxicity hazards. It is also a low-cost, high-capacity absorbent for acid gases. Solvent selection varies with the application. Examples include ionic liquids, monoethanolamine (MEA), methyldiethanolamine (MDEA), and aqueous K2CO3 (UOP Benfield process).

The absorption process is based on the concept of relative solubilities. Relative solubilities for syngas compounds in methanol at -25°C are shown in Table 1 (from (David A Bell 2010)).

Table 1: Relative solubilities in cold methanol

Compound Relative solubility Benzene 59.5

CH4 0.05

C2H4 0.46 C2H6 0.42 C3H8 2.35

CO 0.02

CO2 1*

COS 3.92

H2 0.005 H2S 7.06 N2 0.01

* Basis for relative solubility

One can see how light hydrocarbons and diatomic gases have low affinity for the cold methanol absorbent, whereas CO2, acid gases, and heavier hydrocarbons partition into the liquid phase.

This unit operation consists of two packed columns operating in tandem. The absorber captures CO2 gas into the cold methanol liquid. The desorber liberates the CO2 from the methanol phase by heating in a stripping column. A circulating chiller provides ~4 kW of cooling duty to achieve a minimum absorption temperature of -20°C. Various heat exchangers, pumps, and valves make up the balance of plant.

The absorber requires 12 stages to achieve performance at -20°C and 20 bar. The MAWP should be 500 psig to permit future permutations. The stripping column operates at 4 bar and requires 8 stages for this process design basis.

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Adsorber

Unit Operation: Pressure- and/or Temperature-Swing Adsorption

Gas phase:

Nominal: 150 sLPM inlet (12 lb/hr)

Regeneration: minimum 40 sLPM

• Solid adsorbent particle size is in range of 1-5mm (4-18 mesh, 0.04-0.2 in), both spherical and cylindrical

• This application is primarily for gas-phase separations

• This fits a classic design paradigm of an alternating dual-bed system with regeneration to achieve continuous operations

• MAWP of vessels is 750 psig to accommodate up to 500 psig for typical operations

• MAOT is 750°F for thermal regeneration

• This unit will be used to process both hydrogen (embrittlement) and acid gases (humid CO2, H2S) in the future. Materials of construction should be selected accordingly.

Methane dehydration serves as the process design basis:

Inlet Outlet Regeneration F (lb/hr) 12 12 6.6 (N2) T (°F) 77 77 600 P (psia) 130 120 20

The purpose of the adsorption unit operation is to achieve continuous separations of gas-phase components. Various approaches are available to achieve this goal:

• Immobilize target compound, recover during regeneration (e.g. enriched air)

• Immobilize undesired compounds, send to waste during regeneration (e.g. H2S scrubbing)

This dual-bed adsorption unit is configured to operate in downflow during adsorption and upflow during regeneration. Additional ports are indicated to allow for future features such as bypass, upflow adsorption, and downflow regeneration. The beds can operate in the following modes:

• Parallel o Continuous (online, offline/regen) o Batch (operate until exhaustion)

• In series o Lead-lag configuration with simultaneous regeneration o Mixed media for custom applications

Regeneration can be accomplished either in situ or via physical media removal. Inert gas flow is available to facilitate convective mass transfer. An economizing feed-exhaust heat exchanger is included for heated regeneration steps. Available regeneration modes are:

• Temperature-swing o Temperature cycling achieved via heat exchange of working fluid with house chilled water and house hot oil systems

• Pressure-swing o Pressure cycling is achieved between feed gas pressure via blower or compressor and ambient or house vacuum pressures. Optional pressure accumulators can mimimize depressurization/repressurization duration for upstream and downstream process stability.

Each individual vessel can

• Single adsorbent

• Mixed/layered bed

• Short packed (with appropriate back-flow spring-support screens or similar media retention strategies)

Packed adsorbent media will tend to be micro-/mesoporous, resulting in weak adsorption (Type IV & V) isotherms. This results in longer mass transfer zones (MTZ). These parameters favor higher-pressure operations (~10-40 bar), and high aspect ratio vessels to delay breakthrough. This requires adsorbent media to be mechanically robust enough to withstand a taller (>2m) bed height. Note this approach is not compatible with polymeric ion exchange media.

Materials of construction should be minimum SS316 and Kalrez elastomers (or similar) due to possibility of acid-gas processing (H2S, CO2, etc.).

Flanged access to vessel interior required to modify internals and load/empty solid adsorbent media. Solid retention schemes include graded ceramic balls (Al2O3 or similar) or internal screens and supports.

Other applications include:

• Hydrogen recovery via pressure-swing adsorption

• Direct-air or point-source carbon dioxide capture/recovery/removal

• High-purity methane (RNG) production from bio-/thermochemical processes

• H2S & siloxane removal from biogas

This unit operation is dependent upon upstream processes for a pressurized feed. An optional compressor-accumulator is shown in the PFD to illustrate the point. A mechanical forward-pressure regulator determines the operating point.

Each vessel is jacketed to permit cool temperatures for adsorbing and hot temperatures for desorbing. The feed gas flows downward through the online bed where water is immobilized on a 3A molecular sieve. The exiting methane stream is dried to a dew point less than 100°F. The product stream leaves a back pressure valve and is immediately available for use.

Regenerating inert gas is introduced through an economizing heat exchanger to preheat against hot gas leaving the offline bed. Regeneration occurs in the upflow direction where water is thermally

Date: 7/8/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Decanter

Unit Operation: Decanter

Feed Rate(s): 100 – 150 gal/hr

From Perry’s Chemical Engineers’ Handbook, 9th Edition

The decanter needs to be able to separate three phases, two liquid phases of differing densities and a solids phase. The feed rate should be approximately 100 to 150 gallons per hour, or 500 liters per hour, and the feed contains 10 – 20 wt% solids. The feed could contain solvents and possibly vapors, and therefore needs to be electrically classified. At this point, the particle distribution of the solids content is not known, and a middle-of-the-road particle size should be chosen for a conceptual phase bid.

The decanter needs to be able to be steamed for decontamination and sterilization as well as cleaned in place (CIP). The operating pressures could approach 200 psig and the operating temperatures could be up to 400 deg F.

In addition to solvent and solids, sea water or briny water could be processed in the decanter. The high chloride content of the sea water could lead to stress chloride cracking issues in stainless steel alloys at high temperatures, and therefore metallurgy needs to be considered in estimating cost and design.

Typically the decanter would be downstream of a biomass pretreatment reactor or a bioreactor, but upstream of any solvent extraction, or possibly the first step of solvent extraction. Removing as much of the water as feasible first prior to further processing, reduces the sizing of the downstream equipment.

The decanter requires a feed tank, feed pump, two liquid tanks for the two liquid phases and a third tank for solids product.

Date: 7/3/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Microfiltration/Ultrafiltration Unit

Unit Operation: Microfiltration/Ultrafiltration Unit

Feed Rate(s): Approximately 50 L/min

• Fermentation broth consist of mostly water

A microfiltration unit capable of filtering 50 L/min and removing 1 – 100 µm bacteria/yeast cells from fermentation broth with up to 10 wt% solids is necessary. This same filtration unit should be able to be operated with a separate set of filter units so that ultrafiltration of 100 – 500 Dalton for acid separation is also possible with the same filtration unit. In this case it would separate acids from a fermentation broth.

The filtration unit needs to contain a backwash as well as a recirculation pump. A tank and feed pump module will be necessary to feed this filtration unit.

The filter needs to be able to operate up to 50 psig or more. The filter must also be able to be steamed for sterilization with saturated low pressure steam at 40 psig or a clean in place skid.

The filtration unit will need metallurgy capable of handling high chlorides content water as well as low pH, acidic fermentation broth.

This process will be controlled via a programmable logic controller (PLC) using a human machine interface (HMI). Currently the faciltiy uses Opto 22 control cabinets but is open to other options.

The HMIs are on Dell Computers located throughout the pilot plant and within the control room.

Date: 6/11/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Hydrotreater Stripper

Unit Operation: Distillation Column – Hydrotreater Stripper

Feed Rate(s): 1.0 BBL/DAY from Hydrotreater of oxygen free C18 Alkanes

• The column is packed with 0.5” random packing

• The column is operated at the lowest pressure possible for the separation to conserve energy

Drum Organic Phase

11.96 1.07 0.057

Stripper Overhead 0.49 108.9 (SCFD) 0.012 Stripper Bottoms 11.475 1.002 0.045

Unstabilized distillate from the hydrotreater low pressure flash drum is fed to the top of the of the distillation column. The column has a reboiler on the bottom to supply heat for vaporization. The top of the column has a partial condenser to condense light end hydrocarbons formed in the process. The condenser will also cool the molecules too light to condense at the operating pressure. The overhead receiver can separate vapor, light organic phase and a water phase. The light organic phase can be refluxed back into the tower if necessary. The bottoms product is stabilized distillate for fuel consumption.

The reboiler design currently is a clamshell heater with zone control to control the heat profile within the column. The supplied duty is approximately 4200 BTU/hr.

The internals of the column will be designed so that they can be removed and changed frequently.

The design for the stripper of the hydrotreater is based off of 0.5” Pall ring random packing and a height equivalent theoretical plate (HETP) of 1 ft per stage. The HETP of the packing will change with each system using this distillation column and will need to be included in the engineering study before each different experiment. The bottoms sump should be sized for a five minute hold up time.

The pressure drop with the current hydrotreater model is approximately 0.2” psig, which is normal for this height of a column. The diameter of the column for the vapor rates in the column is 0.18 ft, which is just over 2.15” internal diameter.

The overhead condenser is sized for a duty of 1.6 BTU/hr. The overhead receiver should be sized for five minutes of holdup and surge time at a minimum. The vessel may be larger to accommodate the instrumentation needed for level and pressure control.

The stripper bottoms stream is stabilized distillate that can be used as a finished product, or proceed downstream to an isomerization unit if necessary. The bottoms stream will need to be cooled with an air cooler and trim cooler if it is being sent to storage. The pump will need to provide 4525 ft of total developed head if the stream is being routed to the downstream isomerization unit, which operates at approximately 1500 psig.

The distillation column operating pressure is 15 psig. The maximum allowable pressure of this equipment will be 100 psig for the tower, condenser, and overhead receiver. The bottoms temperature does reach 660 deg F to vaporize the distillate for mass transfer. If the bottoms product is temperature sensitive and cannot be as high as 675 deg F, the distillation column will be rated for full vacuum and a vacuum pump or ejector could be added to the overhead to accomplish this mode of operation. The ejector or vacuum pump is a separate equipment module but would be appropriate to use in tandem with this module. Another option would be to inject live steam into the bottom of the tower to strip the light end material from the distillate product. This tower design does not currently allow for that capability.

Packing Height, ft

Vessel Height, ft MAWP

Stripping Section 2.15 10 12 100 psig

Overhead organic phase pump

150 1.3x10-10

Bottoms Product Pump 4525 (if fed directly to Isomerization Unit)

0.028

Duty, BTU/HR Type of Heat Exchanger Reboiler 4,200 Clamshell Heater Overhead Condenser 1.6 Cooling Water Condenser

Date: 5/29/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Hydrotreater

Unit Operation: Hydrotreater

Feed Rate(s): 1 BBL/DAY of triolein and 2200 SCFD of 100% H2.

• The feed is 100% triolein, which is a vegetable oil type feed requiring a liquid and hydrogen recycle stream to quench the exotherm.

• The hydrodeoxygenation reaction occurs in two reactors in series. The reactor is modeled as one with no interbed hydrogen injection.

• The reactor mass flux is 4000 lb/hr/sq ft

• Cooling water supply is available at 24 deg C (approx. 75 deg F) at approximately 50 psig

• BBL/DAY = 42 gal/day

• Control valve dP = 25 psi, when it is NOT a pressure control valve

• Hydrogen supply is 200 psig.

Triolein – oil feed 13.375 1 Hydrogen 0.49 2200 (SCFD)

1st Stage Compressor

0 0

2nd Stage Compressor Drain

0.07 0.0048

3rd Stage Compressor

0.0048 0.0003

4th Stage Compressor

0.002 0.0002

Hydrogen Purge on Recycle

0.18 294 (SCFD)

1.61 0.1105

0.144 88.6(SCFD)

0.0 0.0

Organic Phase

11.96 1.07

Water Stripper Vapor Phase

0.00009 0.156 (SCFD)

A hydrotreater is designed to process approximately 1 bbl/day, or 42 gal/day, of a triolein feed.

Hydrogen is compressed from the house hydrogen system. The compressor stages are minimized to an outlet temperature of 250 deg F to limit seal damage and prolong equipment life and reliability. Because of the limit, there are four stages of compression necessary to achieve an outlet pressure of 1900 psig. Each stage of compression requires a brake horsepower of 0.07 hp.

The combined feed is preheated from approximately 100 deg F up to 650 deg F. The reactor exotherm is approximately 75 deg F. To maintain this exotherm, a liquid recycle volumetric ratio of 10 to 1 versus fresh triolein feed is necessary. The combined feed also includes a hydrogen recycle.

The hydrogen is recycled as much as possible to minimize the fresh hydrogen needed from the house utility. A small purge of approximately 50 SCFD of the recycle hydrogen is necessary to purge any buildup of non-condensables.

The reactors are operated at 1800 psig and an outlet temperature of 590 deg F. The pressure is controlled via a back pressure control valve on the high pressure flash drum. The exotherm is minimized to ensure that there is enough hydrogen in the reactors to prevent coking the catalyst quickly.

The reactor is designed to maintain plug flow and adiabatic operation. A mass flux of 4000 lbs/hr/sq ft was chosen as the basis based on other commercial HDO hydrotreater reactors. To maintain that flux, the reactor inner diameter needs to be 2.5 inches. The other important reactor sizing criteria is space velocity. For this unit operation, a liquid hourly space velocity (LHSV) of 1 hr-1 leads to a catalyst height of 4 ft tall. LHSV calculations were completed with fresh feed, not the total hydrocarbon feed rate. The recycle was assumed to be 100% converted at the reactor outlet.

The reactor effluent coolers are comprised of an air cooler and a cooling water trim cooler. The air cooler is directly downstream of the second reactor outlet and is designed to cool to 250 deg F and the trim cooler is designed for an outlet temperature of 100 deg F.

Downstream of the coolers is the high pressure flash drum controlled by a back pressure control valve on the vapor outlet. The vapor, which consists of mostly hydrogen, is recycled back to the feed to decrease the overall hydrogen need. The recycle compressor requires 0.009 hp. And the recycle flow is approximately 5600 SCFD, and is 87 mol% hydrogen. The high pressure flash drum has an aqueous phase draw and the organic phase is split into a low pressure flash drum feed stream and a large liquid organic phase recycle to the reactor feed. The recycle stream is approximately 10 to 1 volumetric ratio versus fresh feed. The recycle can be adjusted to limit the exotherm needed for the research. The pump, flow meter and control valve may need to be modified, but that scope is possible.

The low pressure flash drum pressure is approximate 200 psig. A control valve between the high pressure and low pressure flash drums controls the pressure drop. This valve will need to be sized so that choked flow is not a concern.

The low pressure flash drum pressure is controlled with a back pressure control valve on the vapor outlet. The vapor outlet is routed to the thermal oxidizer, where it is later burned. There is a flow meter on that stream for mass balance closure. A sample station on that stream will be added in the future. The aqueous phase from the low pressure flash drum is metered and routed to a holding tank. The organic phase is level controlled from the drum to the downstream stripper for fuel finishing.

As an alternative design case, a hydrocracker should be considered. The hydrocracker reactors, three beds instead of two as shown in the current design basis, would operate at 2000 psig. The temperature profile would be nearly the same, as far as exotherm expected, and there would need to be a liquid and hydrogen recycle loop as there is in the current hydrotreater design.

Reactors

Catalyst Bed

Reactor 1 2.5 4 6 2750 psig Alloy 347 at a minimum Reactor 2 2.5 4 6 2750 psig Alloy 347 at a mimimum

Separators

Height (Tangent to Tangent), in

Diameter, in

High Pressure Separator

36 12 2750 psig Alloy 347 at a minimum

Low Pressure Separator

18 6 650 psig Alloy 347 at a minimum

1st Stage Makeup H2 Compressor

KO

9 0.25 2750 psig 304SS at a minimum

2nd Stage Makeup

3rd Stage Makeup

4th Stage Makeup

Oil Feed Pump 5600 0.05 Liquid Quench Pump 215 0.02

H2 Compressor – 1st

-347.5 500 psig – Shell 300 psig – Tube

H2 Compressor – 2nd

-178 700 psig – Shell 420 psig – Tube

H2 Compressor – 3rd

-178 1250 psig – Shell 750 psig – Tube

Reactor Preheater 52,790 2750 psig

Not Specified

Reactor Effluent Air

-54,870 2750 psig Fin Fan

-5,929 2750 psig – Shell 1650 psig – Tube

Compressors

Adiabatic Head, ft Brake Horsepower, hp H2 Compressor 1st Stage 245,100 0.074 H2 Compressor 2nd Stage 253,600 0.068 H2 Compressor 3rd Stage 258,000 0.068 H2 Compressor 4th Stage 263,300 0.0695 Recycle Compressor 4,965 0.0089

Date: 6/23/2024

Subject: Process Design Basis – FEL 1 – IBRF Pilot Plant Revamp – Ion exchange

Unit Operation: Ion Exchange Columns

Liquid phase:

Nominal: 500 lb/hr (1500 gpd = 36 bpd)

Maximum: 5000 lb/hr (15,000 gpd = 360 bpd)

• Solid adsorbent media is crosslinked polymeric resins, less rigid than ceramic adsorbents

• Very small particle size, 0.1-1.5 mm (100-1500 µm) diameter beads

• Not attempting to duplicate simulated moving bed (SMB) due to cost, complexity o Incorporation of spare ports accommodate future beds, valves, and pumps

• Ion exchange vessels tend to have low aspect rations (L/D) to mitigate pressure drop over the small particles. Typical heights range from 1-2 meters, but the bed should be sized for a maximum 150 kPa (1.5 bar = 22 psig) pressure drop o Will need column internals to:

Ensure liquid distribution for entire cross section Permit segmentation for mixed beds or for short-packing o Johnson screens and other internal technology required to manage packing at scale

• Liquid phase processing is way more common than gas phase

• Liquid feed can contain chloride ions, nominally 35 ppt (seawater) o This has unacceptable corrosion rates for SS316 so materials of construction need to be more robust, likely similar to Hastelloy

• Regeneration media will include 1-molar solutions of sodium hydroxide and sulfuric acid

• Maximum operating temperature = 100°C (typical media doesn’t tolerate > 60°C)

• Vessel MAWP = 500 psig o Typical MAOP will be 200 psig, with most operations significantly below that

• Beds are jacketed to accommodate cooling or heating fluid for isothermal operations

Application: Organic sugar and organic acids recovery from lignocellulosic biomass hydrolysate

Species Inlet (lb/hr) Outlet (lb/hr) Product (lb/hr)

Total 499 493 6 H2O 481.639 481.639 0

GLUCOSE 3.9828 0.796560 3.18623

GALACTOSE 0.160548 0.0321095 0.128438

MANNOSE 0.0673626 0.0134725 0.0538901

XYLOSE 2.71816 0.543632 2.17453

ARABINOSE 0.327933 0.0655866 0.262346

CELLOBIOSE 0.0236371 0.00472742 0.0189097

SUCROSE 0.0459644 0.00919287 0.0367715

GLUCOLIG 0.17083 0.170830 0

GALAOLIG 0.00432451 0.00432451 0

MANOLIG 0.00181448 0.00181448 0

XYLOLIG 0.0724843 0.0724843 0

ARABOLIG 0.00883321 0.00883321 0

EXTRACTIVES 1.52606 1.52606 0

SOLUBILIZED LIGNIN 0.105798 0.105798 0

HMF 0.0289652 0.00289652 0.0260687

FURFURAL 0.11901 0.119010 0

LACID 0.00552237 0.000552237 0.00497013

NH4SO4 0.520864 0.520864 0

OIL 0.00057783 0.000577830 0

CELLULOSE 0.189097 0.189097 0

GALACTAN 0.0118924 0.0118924 0

MANNAN 0.00498982 0.00498982 0

XYLAN 0.00302018 0.00302018 0

ARABINAN 0.000368051 0.000368051 0

LIGNIN 2.01046 2.01046 0

ACETATE 0.615791 0.0615791 0.554212

PROTEIN 2.79505 2.79505 0

ASH 1.79616 1.79616 0

ENZYME 0.264978 0.264978 0

DENATURED ENZYME 0.0294365 0.0294365 0

TRICHO (ENZ BIOMASS) 0.0552979 0.0552980 0

TAR 0.006075 0.00607500 0

Ion exchange is a special type of adsorption. Whereas conventional adsorbents depend largely on porosity and microstructure for performance, ion exchange media is surface-functionalized and nonporous. Form factors can be membranes (PEM), walls (chromatography), or beads. This application focuses on packed beds of ion exchange beads. Typical approaches involve immobilization of desired components (sugars, organic acids) or removal of undesired components (salts). Depending on affinity between solutes and ionic media, a variety of techniques exist for product separation.

An ion exchange bed experiences multiple stages to during an operating cycle: product separation, elution, regeneration, and rinsing. Product elution is achieved with an aqueous or organic process fluid. Regeneration is achieved by washing the bed with stronger-affinity ions, typically acids, bases, or brines. A rinse step can follow regeneration to avoid contamination between feedstock and products. Dual beds allow for simultaneous operation and regeneration if operated in parallel, or in a lead-lag configuration or mixed media if operated in series. The complex sequencing relies on recipe logic and breakthrough detection with product and effluent monitoring.

The representative composition is lignocellulosic biomass hydrolysate from dilute acid pretreatment (no deacetylation) of lignocellulosic biomass. This is a common pretreatment step performed in our IBRF pilot plant. Desired components are sugars and organic acids which are immobilized in the beds. Product concentration is ~1 wt%, with typical product concentrations ranging from 1-10 wt%.

A key design element is the amount of variability in flows and compositions encountered in ion exchange applications. The separation process itself is highly empirical and finely tuned. The adsorption isotherm determines the length of the mass transfer zone. Bed height is determined by keeping pressure drop below 1.5 bar at maximum liquid flow rate and minimum ion exchange particle size. Vessel diameter scales with flow rate. Internal liquid distribution should be considered for smaller aspect ratios.

Hydrolysate enters the unit through a degasser. This suppresses bubble formation in the column which causes channeling and poor utilization of ion exchange media. The hydrolysate is pumped into vessel A through node C. Pressure is moderately elevated at 75 psia and temperature is close to ambient. Sugars and organic acids are slowed or immobilized in the bed while the balance of aqueous hydrolysate passes through node E. Valving is adjusted for effluent to leave through node H. A back pressure valve is available if operating at elevated pressures. Simultaneous monitoring of pH, conductivity, and UV/Vis is a common approach to detect all phases of operation and inform valve sequencing. Depending on the phase of operations, valves are adjusted so effluent can be sent to waste, solvent recovery for recycle, fraction collection for characterization, or a downstream unit operation.

In the offline bed vessel B, various…