I’m operating a chemical plant and we recently started producing about 3 tons per batch of potassium silicate solution. (90% KOH + H2O +  precipitated silica99% )

After production, we immediately rinse the tank thoroughly with water (right after discharge, before any visible drying). When the tank surface is still wet, it looks completely clean.

However, once the tank dries, we consistently find a significant amount of white powder residue on the inner walls. This residue is quite difficult to remove and requires additional cleaning effort.

We are planning to use this tank interchangeably for both potassium silicate production and cleaning agent manufacturing, so this issue is becoming a serious operational concern.

My questions:

1. Is this a common issue in potassium silicate production?
2. Is this likely due to silicate film formation or silica precipitation during drying?
3. Do most plants use acid rinse or other methods to prevent this?
4. Is it realistic to use the same tank for silicate and detergent production?

Any advice or shared experience would be greatly appreciated.

Thanks in advance!

• Tank Capacity : 2500 m³
• Tank Internal Diameter : 17.2 m
• Height : 12 m
• Liquid_in : 286.1 m³/hr
• Flash point : -43 ^oC

What I've done is calculate the capacity of N2 using API 2000 for inbreathing,

• Q_{inbreathing liquid} : 268.9 Nm³/hr
• Q_{inbreathing thermal} : 422.5 Nm³/hr
• Q_total : 691.4 Nm³/hr = 0.19 Nm³/s

After that, using "Rules of Thumb for Chemical Engineers By Carl Branan"

• Suggested Fluid Velocity N2 (Air = 78% N2) : 4000 fpm = 20.32 m/s
• Typical Design Vapor Velocity (Gas or Superheated Vapor 0 to 15 psig) : 50 - 145 ft/s (pick 65 ft/s = 19.8 m/s so < 20.32 m/s)

After that, 

• A = Q / v = 0.19 / 19.8 = 0.0097 m²
• D = SQRT(4 x 0.0097 / PI) = 111.13 mm

Then with SCH 40 that satisfied the D and common in the market is 6" (154.05 mm)

Does my calculation is correct? 
If wrong, can someone please explain the other way to calculate it because after read the API 2000 that is what I thought.
Thanks in advance !!

On discharge line PSV is provided for a set pressure of 1.3 barg whereas the Separator design pressure is 14 barg.We have strainght guideline in API for single phase check valve failure case but not for Two Phase.Can you please guide how to estimate relief load.

1. The desalted pump suction pressure dropped below the design value, hence low discharge pressure

2. High HFO temperature to the cooling water box

3. Increased pressure at the fractionator overhead.

We have shut down the plant to run on cold recycle and system depressurised  but challenge resurfaced 8 hrs after heating up at a nominal value. 

The product yield is fine but the inability to ramp up rate has been a concern. What could be the challenge? I will appreciate a good insight into this.

We suspect that the exchanger is over-surfaced, which results in very low steam chest pressure inside the exchanger. Because of the low pressure, the condensate pressure is not sufficient to push the condensate to the LP condensate header. Due to this, the exchanger starts getting flooded with condensate until enough pressure builds up to push the condensate out. In our case the condensate needs about 1.5 kg/cm²(g) pressure to move to the header, while the condensate temperature is around 70°C. This intermittent discharge of condensate is creating severe hammering issues in the system.

There was no equalization line between the exchanger and the condensate pot in the original system. We tried installing a temporary equalization connection, but this did not resolve the problem. As a trial, we injected nitrogen (a non-condensable gas) into the exchanger and pressurized the condensate pot to around 1.5 kg/cm²(g). After doing this the system started working smoothly, condensate was flowing continuously to the header, the exchanger did not flood, and the hammering problem disappeared.

Based on this observation there is now a proposal to install a permanent nitrogen connection with a PCV to maintain around 1.5 kg/cm²(g) pressure in the condensate system. However, I am not very confident about implementing this as a long-term solution because I could not find much literature or industry references supporting the practice of injecting nitrogen into a steam condensate system. I am concerned about potential issues such as accumulation of non-condensable gases, possible impact on heat transfer in the exchanger, steam trap performance, two-phase flow instability, or other operational risks over the long term.

Has anyone experienced a similar situation or used nitrogen injection to maintain pressure in a condensate return system? Is this something that is practiced in industry, or would it generally be considered only a temporary workaround? Any insights, experience, or references would be greatly appreciated.

I need your help about line sizing for PSV's, i'm sizing PSV's for a pipeline transporting a gas with the following data:

MW=                        20.53 (mostly methane 71%)

Flowrate=                 320 000 kg/hr

Density upstream = 184 kg/m³

Temperature=          100 °C

Pressure=                240 bar

Z factor =                 0.89

Cp/Cv=                    1.55

i used HYSYS safety analysis for sizing, i choosed blocked outlet scenario as the worst case scenario, so i discharged all the flowrate, HYSYS suggested 3 pilot PSV's, the simulation selected 2 J 3 (2500 x 300)  for each PSV for a flowrate of 106 666 kg/hr, the problem i'm facing now is sizing the downstream lines, i took a mach number of 0.7, density of 2 kg/m3 at relieving.

hand calculation gave me a 10" pipe and the flare header around 20", these results seem  weird to me, i would like to know if these are reasonable or my calculation are missing something.

thank you.  

I would appreciate some input regarding best practice for setting MDMT in gas systems.

In a dry high-pressure nitrogen storage system (around 60 barg), gas is let down through a control valve into a downstream header lower pressure. One scenario considered is a control valve fail-open case (not a formal emergency blowdown).

Should a control valve fail-open case normally be treated as a governing basis for MDMT selection of downstream piping and vessels?

Or is it more common practice to:

• Treat low temperatures as local to the restriction/PSV tie-in region

Assume:

• Dry gas (no liquid flashing)

• No cryogenic service

• Downstream overpressure protected by PSV

• Normal restart controls in place.

I’m interested in typical industry practice rather than specific simulation results.

Thanks in advance.

Do you know any method allowing neutralization of off-spec ethylene oxide in other way than reprocessing in EO plant?
Maybe there are any related regulations or standards.

Regards

Background:

- We distill an organic peroxide using 2 columns under vacuum.  

- The first column separates the peroxide from alcohols and is running fine.

- The organic stream is then mixed with 50%NaOH & DI Water before entering the second column  (7 trays)

- The top is condensed and sent to a plain horizontal decanter.  The top layer splits into majority organic material in water (70/30) while the bottom is around 82/18 organic/water.  This bottom layer is refluxed. while the top goes further into the process.

Before shutting down, we were able to run maximum rates without issue.

When trying to restart, our vacuum pump seized and was replaced.  This has been the only major change to the process.

When restarting, we're unable to get past very minimum rates.  When we increase raw material flow, we end up with a small, white layer at the interface in the decanter and the top product is hazy when it is typically crystal clear.

The only major difference I can see is the column dP.  While it was ~30 mbar with the old pump, it's now ~18-22 mbar.  This column is typically at 140 mbar as well (decanter is under same vacuum) 

I can only assume it's potentially weeping and we need to increase vapor flow (more steam to reboiler), but any experience with this type of probably that could be shared would be welcome.  If there's other information that could be vital, please ask.  We increased the raw material feed to try to load the trays more, giving the vapor resistance, but this caused the entire decanter to turn white, due to mixing I assume.

I am currently revising our existing evaporator design, as the current configuration has been observed to cause oil entrainment. We initially considered two possible causes:

1. Excess heating

2. Insufficient vapor disengagement space

After evaluation, insufficient vapor disengagement has been identified as the primary cause.

In this regard, I am reviewing the existing evaporator design prepared by my former colleague (please see the attached file: Existing Design). I am currently evaluating two revised options:

1. My proposed revision (File name: Own Design)

2. An alternative design suggestion from an external colleague (File Name: Additional Design)

While both options appear technically feasible, I would greatly appreciate your expertise and insights to determine the most suitable approach.

For reference, the process fluid consists of palm olein with excess methanol and water of reaction from the esterification process. Oil entrainment into the vapor stream has been observed, which affects downstream separation.

Thank you in advance for your guidance and recommendations.

I know for Pool fire case scenario, based on API 521 we only have to calculate wetted surface area for bare tubes of air coolers. But for Jet fire scenario considering jet fire occuring in nearby separator, do we need to calculate the surface area of finned air cooler or just bare tubes area?

I guess we need to calculate total surface area of bare tubes (including wetted + non wetted areas) for jet fire case.

Kindly if anyone can provide input on whether we need to include the total finned area of air coolers for blowdown area estimation?

In a district heating network, pressure boosting stations are installed. Each station has a pump installed on the supply pipeline and a control valve installed on the return pipeline.
The boosting station is controlled in such a way that the pump regulates the available pressure so that, at the most distant customer, the available pressure does not drop below 80 kPa.

The control valve on the return line is used to regulate the upstream pressure.

My understanding is that when customers draw too much water, the control valve protects the pump from operating in inefficient regions of its characteristic curve (low head and high flow). By partially closing the valve, the pump is forced to operate at a higher pressure and lower flow, while still maintaining the minimum required available pressure at the end customer.

Can someone confirm this or point out any errors in my reasoning?

To mitigate this, I am considering increasing the pipe size in the horizontal section (approximately 7 m long) from 3" to 5", which would reduce the Froude number to about 0.22 (<0.3). My proposal is to keep the vertical section and dip pipe (about 5 m downward) at 3" diameter.

My rationale is that:

• The larger horizontal section will help vent any entrained or dissolved air.

• The vertical downflow section should remain liquid-full, so a 3" diameter should be sufficient.

Does this approach make sense? Are there any potential issues with transitioning from 5" back to 3" before the vertical section. The location of the transition is it better to do it at the end of the horizontal line, start of the vertical line or bend? I have attached a sketch for more info.

Thank you!